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CO₂ removal in power systems using calcium-based sorbents Sun, Ping 2007

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C 0 2 R E M O V A L IN P O W E R SYSTEMS USING C A L C I U M - B A S E D SORBENTS by PING SUN B. Sc., Northeast Institute for Electric Power Engineering (Jilin), 1993 M. Sc., Northeast Institute for Electric Power Engineering (Jilin), 1996 D. Eng., Zhejiang University (Hangzhou), 1999 A THESIS SUBMITTED IN PARTIAL FULFILMENT OF THE REQUIREMENTS FOR THE DEGREE OF DOCTOR OF PHILOSOPHY in THE F A C U L T Y OF GRADUATE STUDIES (CHEMICAL AND BIOLOGICAL ENGINEERING) THE UNIVERSITY OF BRITISH COLUMBIA April 2007 © Ping Sun, 2007 ABSTRACT Bench scale studies were carried out, focusing on the application of calcium-based sorbents in fossil-fuel-fired combustion and gasification systems, with conditions ranging from atmospheric to elevated pressures and at practical combustion and gasification temperatures. In the kinetic study of CaO carbonation, the reaction order changed abruptly from first- to zero-order when the CO2 partial pressure exceeded the equilibrium value by more than -10 kPa. A Langmuir mechanism successfully explained the experimental information, with the intermediate complex CaO»C02 postulated to saturate CaO sites immediately at high CO2 partial pressure. The activation energies for rate constants were found to be 29±4kJ /mol and 24±6kJ/mol for Strassburg limestone and Arctic dolomite, respectively. A discrete-pore-size-distribution-based model was formulated, with the aid of which the kinetic study was extended to obtain diffusivities through the solid product layer formed during carbonation, with activation energies of 215 and 187 kJ/mol for the limestone and dolomite, respectively. Sorbent cyclic C02, removal ability was investigated based on pore size distribution measurements. Several important features observed from measurements could be predicted by a mechanistic model which included simultaneous sintering and calcination in the fixed bed. It was found that the decay in the reversibility of limestone capture/regeneration was insensitive to operating conditions; the achievable carbonation extent of each cycle depends on the <~220 nm. pore volume that decreases monotonically during cycling. Co-capture of SO2 and CO2 was attempted at fluidized bed combustion temperatures. Parametric studies with Strassburg limestone and Arctic dolomite found that the presence of SO2 impeded carbonation even at low concentrations of SO2 relative to CO2 concentrations. This finding is significant for the application of calcium-based sorbents in fluidized-bed combustors ii (FBCs), given the initial assumption that sulphur should not be problematic given the low sulfur/carbon ratio in fuels. The mechanism of the impeding effect of SC^was investigated forJ_ seven sorbents at both atmospheric and elevated pressures. It was found that direct sulphation becomes dominant after completion of an initial fast stage of carbonation, enveloping the sorbents and inhibiting further carbonation. Among the techniques tested, increasing the C O 2 partial pressure was found to be the most helpful way to improve sorbent reversibility. It was also shown that often-cycled sorbents can be reactivated and achieve improved reversibility by the use of low-temperature steam or liquid water. CO is not appropriate as an agent to cyclically regenerate CaO from CaS0 4 because of slow regeneration of CaS04. Among the inert dopants tested, only A I 2 O 3 mixed with CaO at a 1:1 molar ratio, was able to achieve satisfactory C O 2 capture reversibility. As extensions of the applications of calcium-based sorbents, sequential S O 2 and C O 2 capture -were investigated for fluidized bed combustion. Among the four options examined, the best was found to be to apply spent sorbent after cyclic C O 2 capture to remove S O 2 from atmospheric -pressure combustors. A novel concept of co-capture of C 0 2 and H 2 S in a gasifier-based process was also investigated. Unlike the findings for co-capture of S0 2 and C0 2 , no obvious impeding effect of H 2S was observed on cyclic C 0 2 capture. Parametric studies indicated that it should be feasible to co-capture H 2S and C0 2 . With C 0 2 sorbents in a gasifier, one-step hydrogen production via gasification should be achievable. T A B L E O F C O N T E N T S ABSTRACT '....ii T A B L E OF CONTENTS ' :. i v LIST OF TABLES viii LIST OF FIGURES ix ACKNOWLEDGEMENTS xx CO-AUTHORSHIP STATEMENT. xxi CHAPTER 1 INTRODUCTION, THERMODYNAMICS AND EQUIPMENT 1 1.1 Background and structure of this thesis: using calcium-based sorbent to removal C0 2 . , 1 1.2 Thermodynamics underlying this thesis 5 1.2.1 Sorbent properties, calcination and carbonation 5 1.2.2 Sorbent calcination temperature 6 1.2.3 C O 2 capture in FBC: Effect of S O 2 on CaO carbonation 11 1.2.4 C O 2 capture during gasification 12 1.2.5 Sorption-enhanced hydrogen production using a calcium-based C O 2 sorbent 14 1.2.6 Effect of H 2 0 , low-temperature eutectic between CaO-Ca(OH)2-CaC03 16 1.3 Experimental equipment and materials 17 1.3.1 Apparatus 17 1.3.2 Sorbents 20 1.4 Nomenclature 21 CHAPTER 2 DETERMINATION OF INTRINSIC RATE CONSTANTS OF THE CaO-C O 2 REACTION.... 22 2.1 Introduction 22 2.2 Experimental details 23 2.2.1 Elimination of physical limitations in ATGA tests 24 2.2.2 PTGA tests : 26 2.3 Results and discussion 27 2.3.1 Direct measurement of carbonation kinetics: use of a gas-solid reaction model 27 2.3.2 High-pressure measurement 31 2.3.3 Carbonation data from equilibrium analyses 32 2.3.4 Comparison with earlier results 33 2.3.5 Mechanistic explanations for the variable order reaction 34 2.3.6 Other issues 36 iv 2.4 Conclusions ^ ' 2.5 Nomenclature 38 CHAPTER 3 A DISCRETE-PORE-SIZE-DISTRIBUTION BASED GAS-SOLID MODEL AND ITS APPLICATION ON THE CaO+C02 REACTION 48 3.1 Introduction 4 8 3.2 Experimental details 51 3.3 Model development 52 3.3.1 Pore overlap 52 3.3.2 Model description 53 3.3.3 Rate expressions 56 3.3.4 Reaction front, pore evolution and conversions by volume balance 58 3.3.5 Algorithm -59 3.4 Results and discussion v-60 3.5 Conclusions 66 3.6 Nomenclature 68 CHAPTER 4 INVESTIGATION OF EFFECT OF SINTERING ON CYCLIC C02 CAPTURE UNDER FLUTDIZED BED COMBUSTION CONDITIONS 81 4.1 Introduction 81 4.2 Experimental studies 82 4.3 Pore size distribution 84 4.4 Model development 86 4.4.1 Pore evolution during cyclic calcination/carbonation 86 4.4.2 Reactor model for calcination 90 4.4.3 Macroscopic sintering during cyclic calcination/carbonation cycles 93 4.4.4 Results and discussion 95 4.5 Conclusions 98 4.6 Nomenclature 99 CHAPTER 5 SIMULTANEOUS C02 AND S02 CAPTURE AT FLUD3IZED BED COM BUSH ON TEMPERATURES 114 5.1 Introduction 114 5.2 Experimental details 116 5.3 Results and Discussion: 117 5.3.1 Baseline runs: carbonation test with no S0 2 in gas stream 117 5.3.2 Simultaneous sulphation and carbonation: Effect of SO2 on C O 2 capture 118 V 5.3.3 Simultaneous sulphation and carbonation: Effect of CO2 on SO2capture 123 5.4 Conclusions 124 CHAPTER 6 R E M O V A L OF C02 BY CALCIUM-BASED SORBENTS IN THE PRESENCE O F S O 2 135 6.1 Introduction 135 6.2 Experimental details 136 6.3 Results and Discussion 138 6.3.1 Tests with no S0 2 among the gaseous reactants, referred to hereafter as "no SO2" tests ; ; 138 6.3.2 Simultaneous S0 2 and CO2 capture .141 6.3.3 Change of operating conditions in an effort to improve co-capture 147 6.4 Conclusions.... 150 CHAPTER 7 AN INVESTIGATION OF ATTEMPTS TO IMPROVE CYCLIC CO z CAPTURE BY SORBENT HYDRATION AND MODIFICATION 163 7.1 Introduction 163 7.2 Experimental details: 164 7.3 Results and Discussion..... 167 7.3.1 Tests involving steam or liquid water 167 7.3.2 Regeneration of CaO from CaS0 4 using CO during calcination 173 7.3.3 Modification of sorbents (no S0 2 present) 175 7.4 Conclusions 179 CHAPTER 8 SEQUENTIAL CAPTURE OF C02 AND S02 UNDER FBC CONDITIONS ; ; 192 8.1 Introduction 192 8.2 Experimental Details 195 8.3 Results and discussion. ..196 8.4 Conclusions 205 CHAPTER 9 CO-CAPTURE OF H 2S AND C02 IN A PRESSURIZED-GASIFIER-BASED PROCESS 219 9.1 Introduction....' 219 9.2. Experimental Details 220 9.3 Results and Discussion..... 223 9.3.1 "Once-through" tests 223 9.3.2 Parametric tests 225 vi 9.3.3 Effect of CO, and H 2S Partial Pressures .227 9.3.4 Effect of temperature 229 9.3.5 Effect of particle size 230 9.3.6 Effect of residence time 231 9.3.7 Effect of total pressure 231 9.3.8 Effect of sorbent type 232 9.3.9 Cycled sorbents in sulfidation 233 9.4 Conclusions 234 C H A P T E R 10 CONCLUSIONS AND RECOMMENDATIONS F O R F U T U R E WORK246 10.1 Conclusions 246 10.2 Recommendations for future work 250 L I T E R A T U R E CITED .....253 APPENDIX I ABOUT KINETIC DATA FITTING 265 APPENDIX II M O L A R V O L U M E RATIO FOR A DOLOMITE EMPLOYED IN THE MODEL OF CHAPTER 3 273 APPENDIX HI INITIAL PORE SIZE DISTRIBUTION EMPLOYED EST THE MODEL OF CHAPTER 3 : 275 APPENDIX IV A FORTRAN PROGRAM EMPLOYED IN CHAPTER 3 280 APPENDIX V LINEARIZATION OF EQUATION (4.5) 285 APPENDIX VI A M A T L A B PROGRAM EMPLOYED IN CHAPTER 4 287 APPENDIX VII SUPPLEMENTARY RESULTS WITHIN THE SCOPE OF CHAPTER 6... 294 APPENDIX VUI SUPPLEMENTARY RESULTS WITHIN THE SCOPE OF CHAPTER 7.299 APPENDIX LX A CASE STUDY BASED ON OPTION B PROPOSED IN CHAPTER 8 304 APPENDIX X PRELIMINARY TEST RESULTS FOR L i 4 S 1 0 4 CARBONATION AND SULPHATION 315 Vl l LIST OF T A B L E S Table 1.1 Eutectic formation conditions for the CaO-Ca(OH)2 and CaC0 3-Ca(OH) 2 system 17 Table 1.2 Chemical analyses (% by wt in each case) 20 Table 1.3 Summary of sorbents and reactors in each chapter 21 Table 3.1 Comparison of activation energies for effective diffusivity in reactions between CaO or CaC0 3 and gases of interest in this work 63 Table 4.1 Fitted results for Ubed (in equation 4.5) 95 Table 7.1 Regeneration conditions and CaS/CaO test 173 Table 7.2 Test conditions of CaS04 regeneration by CO 175 Table 7.3 Summary of experiments to test possible additives to improve sorbent reversibility. 181 Table 8.1 Gas compositions for the four options examined. All reactions were conducted at 850°C 195 Table 8.2 Sulphation extents after direct sulphation 201 Table Al . 1 Fitting results with nonlinear equations. Values in [] are values with 95% confidence level. Values in bold are utilized in Chapter 2 267 Table AI.2 Fitting results with linear fit to logarithmic equations. Values in [] are values with 95% confidence level; Values in bold are adopted in Chapter 3 268 Table ATX. 1 Proximate and ultimate analysis for the fuels simulated 307 Table ATX.2 Summary of simulation results for overall mass balance 308 Table ATX.3 Summary of gas emissions 308 Table ATX.4 Calculation results for all streams in Figure VTH.1 (with CaO conversion to CaC03 is assumed as 0.4 in the CARBONER block) 314 Vlll LIST O F FIGURES Figure 1.1 Schematic diagram of the process studied in this thesis work to apply calcium- based sorbents 4 Figure 1.2 Phase diagram for calcite and aragonite. Adapted from the work of Redfern et al., (1989) 6 Figure 1.3 Ellingham plot for the three carbonate decomposition reactions with isobaric lines for partial pressure of C O 2 8 Figure 1.4 Maximum C O 2 removal efficiency as a function of total pressure and molar fraction of C 0 2 in the flue gas at the inlet of reactor for a temperature of 850°C 10 Figure 1.5 Ellingham plot for CaO and MgO hydration reactions isobaric lines for partial pressure of H 2 O 11 Figure 1.6 Equilibrium S O 2 concentrations for reaction given in equation (1.8) as the function of reactor temperature and total pressure in the presence of 3%v O2 12 Figure 1.7 Thermodynamic predictions for gaseous products as a function of temperature for 0.76 MPa total pressure. Gas feed: l%v F£2S, 20%v C0 2 , 12.6%v H 2 and 1.5%v CO with N 2 balance 13 Figure 1.8 Effect of C O 2 capture on H 2 production in a simulated gasification system. Feed: 1 kmol carbon, 1.5 kmol steam, 0.01 kmol H 2 S , with and without lkmol CaO 15 Figure 1.9 Schematic of SHTMADZU-based atmospheric thermogravimetric analyzer system (ATGA) 18 Figure 1.10 Schematic of pressurized thermogravimetric analyzer (PTGA) system 19 Figure 1.11 Schematic of atmospheric thermogravimetric reactor (ATGR) system 20 Figure 2.1 High-resolution SEM picture, 38-45 pm Strassburg limestone, calcined under isothermal heating, final temperature 850°C 41 Figure 2.2 Slope extraction with the aid of the grain model during early stages of carbonation for 38-45 pm Strassburg limestone particles at 700°C with 15% C 0 2 and 85% He 41 Figure 2.3 Typical grain model plots. Early stage of carbonation for 38-45 pm Strassburg limestone particles 42 Figure 2.4 Reaction order plot (Squares: 850°C; Triangles: 600°C) for fully calcined 38-45 pm Strassburg limestone with varying C O 2 partial pressure, helium making up the balance of the gas stream. RMSE=Root mean square error 42 Figure 2.5 Reaction order plot (Squares: at 850°C; Triangles at 600°C) for fully calcined 38-45 pm Arctic dolomite with varying C O 2 partial pressure, helium making up the balance of the gas stream. RMSE= Root mean square error 43 Figure 2.6 Arrhenius plot for carbonation reaction with 38-45 pm Strassburg limestone particles. 43 Figure 2.7 Arrhenius plot for carbonation reaction with 38-45 pm Arctic dolomite particles 44 I X Figure 2.8 Conversion vs time for typical kinetic run showing induction period and how the initial rates were obtained based on the maximum slopes. Strassburg limestone, 38-45 urn, 800 kPa, carbonation with 100% C0 2 , 690°C 44 Figure 2.9 Conversion vs time for typical kinetic run showing induction period and how the initial rates were obtained based on the maximum slopes. Arctic dolomite, 38-45 urn, 0.8 MPa, carbonation with 100% C 0 2 at 764°C 45 Figure 2.10 Arrhenius plot comparing PTGA runs (at a total pressure of 0.8 MPa, with 100% C0 2 ) with A T G A runs for 38-45 um Strassburg limestone 45 Figure 2.11 Arrhenius plot. Comparison of PTGA runs (at total pressure of 0.8 MPa, with 100% CO.) with A T G A runs for 38-45 urn Arctic dolomite 46 Figure 2.12 Arrhenius plot for CaO-C0 2 reaction based on calcination data of Borgwardt (1985). .•: 46 Figure 2.13 Illustration of energy levels for CaO+C02<=> CaC0 3 . 47 Figure 3.1 Schematic of a two-pore system with overlap 70 Figure 3.2 Schematic of a two-pore system after evolution 70 Figure 3.3 Pore size distribution results for Strassburg limestone calcines. Calcination conditions: isothermal calcination at 850°C in 100% N 2 71 Figure 3.4 Pore size distribution results for Arctic dolomite calcines. Calcination conditions: isothermal calcination at 850°C in 100% N 2 71 Figure 3.5 Fitting results. Experiments were in ATGA, with 80%v C 0 2 , 20% N 2 balance and 35-45 urn Strassburg limestone particle. Experiment points: 72 Figure 3.6 Fitting results. Experiments were in ATGA with 100%v C 0 2 and 35-45 um Arctic dolomite particle 73 Figure 3.7 Arrhenius plot for diffusivity, D p , Strassburg limestone. E=215 kJ/mol 74 Figure 3.8 Arrhenius plot for diffusivity, D p , Arctic dolomite. E=187 kJ/mol 74 Figure 3.9 Experimental carbonation data showing effect of varying C 0 2 partial pressure for 38-45 urn Strassburg limestone. Pco2.,eq is calculated from equation (3.19) (a) ATGA at 600°C (b) ATGA at 850°C (c) PTGA at 800 kPa compared with ATGA tests 75 Figure 3.10 Experimental carbonation data showing effect of varying C 0 2 partial pressure for 35-45 um Arctic dolomite. Pco2.,eq is calculated from equation (3.19) (a) ATGA at 600°C (b) ATGA at 850°C (c) PTGA at 800 kPa compared with ATGA tests 76 Figure 3.11 Predicted effect of varying C 0 2 partial pressure for 35-45 um Strassburg limestone vs experimental data. PCo2, e q is 48 kPa for 850°C and 0.36 kPa for 600°C, as calculated from equation (3.19) 77 Figure 3.12 Predicted effect of varying C 0 2 partial pressure for 3 5-45 um Arctic dolomite vs experimental data. Pco2,eq is 48 kPa for 850°C and 0.36 kPa for 600°C, as calculated from equation (3.19) 78 Figure 3.13 Predicted C 0 2 concentration at reaction front for Strassburg limestone at 600°C, with varying PCo2- Pco2,eq is 0.36 kPa for 600°C, calculated from equation (3.19) 79 Figure 3.14 Predicted C 0 2 concentration at reaction front for Arctic dolomite at 600 °C, with varying P c o 2 . Pco2, e q is 0.36 kPa for 600°C, calculated from equation (3.19) 79 Figure 3.15 Pore size distribution evolution for Strassburg calcine carbonation at 600°C and 80 kPa C O 2 partial pressure 80 Figure 3.16 Pore size distribution evolution for Arctic dolomite calcine carbonation at 600°C and 100 kPa CO-partial pressure 80 Figure 4.1 Pore size distribution: effect of carbonation time. Experiments: 850°C for calcination and carbonation in the TGR. 212-250 pm Strassburg particles. Calcination with 100% N 2 ; carbonation with 100% C 0 2 , fast stage completed 102 Figure 4.2 Pore size distribution: carbonate before and after mild grinding. Same test conditions as in Figure 4.1 103 Figure 4.4 Pore size distribution: effect of calcination time or mode. Same test conditions as in Figure 4.1 104 Figure 4.5 Pore size distribution: calcines after various number of calcination/carbonation cycle. Test conditions as in Figure 4.1 104 Figure 4.6 SEM pictures of cycled Strassburg calcine samples after 15 cycles of 850°C calcination/carbonation cycles. Test conditions as in Figure 4.1 105 Figure 4.7 Conversion of CaO to CaC03: Experiments vs Predictions with pore volume. Experimental conditions: 850°C for calcination and carbonation in the TGR. 212-250 pm Strassburg particles. Calcination with 100% N 2 , carbonation with 100% C O 2 , fast stage completed 105 Figure 4.8 Specific surface area after each cycle of calcinations, experimental results vs. predictions. The predictions show the sensitivity to Sg. Strassburg limestone, TGR test, 850°C calcination in 100% N 2 , 850°C carbonation in 100% C O 2 . Fast stage of carbonations completed 106 Figure 4.9 Schematic of sintering progression during cyclic calcination and carbonation 107 Figure 4.10 Reversibility under different test conditions in the TGR or TGA. All with Strassburg limestone, calcination in 100% N 2 , 850°C carbonation in 100% C0 2 . Fast stage of carbonation finished for each cycle of carbonation 108 Figure 4.11 CaO conversion profiles for several calcination cycles: experimental results vs. predictions. 212-250 pm Strassburg limestone, TGR test, 850°C calcination in 100% N 2 , 850°C carbonation in 100% C O 2 . Fast Stage of carbonation finished for each cycle of carbonation 108 Figure 4.12 Reversibility: experimental results vs. predictions for 212-250 im Strassburg limestone in the TGR. Calcination in 100% N 2 , 850°C carbonation in 100% C0 2 . Fast Stage of carbonations is allowed to finish for each carbonation cycle 109 Figure 4.13 Reversibility: experimental results vs. predictions for 38-45 im Strassburg limestone in the TGR or TGA. Calcination in 100% N 2 , 850°C carbonation in 100% C0 2 . Fast Stage of carbonations is finished for each carbonation cycle 110 Figure 4.14 Calcination time: experimental results vs. predication. Calcinations: in 100% N 2 , carbonation: 850°C, in 100% C0 2 , Fast Stage of carbonation is finished for each carbonation step I l l Figure 4.15 Predicted CaO utilizations for 1000 cycles. Same calculation conditions as in Figure 4.12 and 4.13 for each case 112 Figure 4.16 Effect of carbonation time on cyclic C 0 2 capture performance: experimental results. Starting from 850 mg of 212-250 im fresh Strassburg limestone. Calcinations: in 100% N 2 , 850°C; carbonation: 850°C, in 100% C0 2 . Carbonation time at each carbonation stage: FSF-Fast stage finished, comparing with 3-minute and 8 minute for each cycle. 113 Figure 5.1 C 0 2 cyclic capture performance. 212-250 um Arctic dolomite and Strassburg limestone. Calcination: 850°C in 100% N 2 ; Carbonation 850°C with 80 or 100% C 0 2 (no S02). Sorption time is 3, 30 min for each sorption stage or FSF 125 Figure 5.2 Illustration for turning point selection in a 100% C 0 2 capture test with 212-250 um Strassburg limestone 125 Figure 5.3 Test results for 212-250 um limestone, 850°C calcination and 850°C sorption with 2900 ppm SO. and 80% C 0 2 126 Figure 5.4 Comparison of initial calcination rate of each cycle for the limestone, 850°C calcination/850°C carbonation, 212-250 urn particle 127 Figure 5.5 Effect of cumulative reaction time. 850°C calcination/850°C sorption, gas compositions for S0 2 /C0 2 sorption: 80% C0 2 , 2900 ppm S0 2 , 3% 0 2 , balance N 2 . (Baseline, 80% C 0 2 20% N2) 128 Figure 5.6 Total calcium utilization change with cycles for 850°C calcination/850°C sorption, gas compositions for sorption: 80% C0 2 , 2900 ppm S0 2, 3% 0 2, balance N 2 129 Figure 5.7 Effect of reaction temperature C 0 2 concentration for each cycle, 212-250 urn, 3-minutes sorption time, 850°C calcination, gas compositions for S0 2 /C0 2 sorption: 80% C 0 2 , 2900 ppm S0 2, 3% 0 2, balance N2.(Base lines, 80% C0 2 , 20% N 2) 130 Figure 5.8 Effect of S0 2 concentration for successive cycles for 212-250 um particle 850°C calcination/850°C sorption, gas compositions for S0 2 /C0 2 sorption: 80% C 0 2 , 3% 0 2, balance N 2 . (Baseline, 80% C 0 2 20% N2) 131 Figure 5.9 Effect of C 0 2 concentration for successive cycles with 212-250 um limestone. 850°C calcination/850°C sorption. Gas compositions for S0 2 /C0 2 sorption: 80% C 0 2 , 2900 ppm SO.?. 3% 0 2 , balance N2.(Base lines, 80% CO 220% N 2) 132 Figure 5.10 Effect of C 0 2 on S0 2 capture. 850°C calcination/850°C sorption. Gas compositions or S0 2 /C0 2 sorption: 80% C0 2 , 2900 ppm S0 2, 3% 0 2 (Base lines: 2900 ppm S0 2, 3% ();) : 134 Figure 6.1. Calcium utilization over 15 cycles for all seven sorbents with 212-250 urn particles and no S0 2 present. ATGR tests. Carbonation: 80% C0 2 , 20% N 2 , 850°C and 101 kPa, Fast Stage Finished. Calcination: 100% N 2 , 850°C and 101 kPa 152 xii Figure 6 . 2 High-resolution SEM pictures of calcines, with no S O 2 present. Same test conditions as in Figure 6 . 1 . (a) Strassburg limestone, after 1 5 cycles (b). Kelly Rock limestone, after 1 5 cycles; (c) Arctic dolomite, after 2 0 cycles 1 5 3 Figure 6 . 3 Evolution of pore size distribution with calcination/carbonation cycling at 8 5 0 ° C . Test conditions: same as in Figure 6 .1 for no S O 2 test and in Figure 6 . 6 for co-capture.(a) Strassburg limestone; (b) Kelly Rock limestone; (c) Arctic dolomite 1 5 4 Figure 6 . 4 Cyclic performance with no S0 2 present: effect of calcination type. Calcination/carbonation cycling at 8 5 0 ° C with 2 1 2 - 2 5 0 pm Strassburg particles 1 5 5 Figure 6 . 5 Cyclic performance with no S O 2 present: effect of total pressure. Calcination/carbonation cycling at 8 5 0 ° C PTGA test: 2 1 2 - 2 5 0 pm Arctic dolomite 1 5 5 Figure 6 . 6 Performance of all seven sorbents for co-capture. ATGR tests, calcination/carbonation cycling at 8 5 0 ° C and 1 0 1 kPa with 2 1 2 - 2 5 0 pm particles. Sorption: 8 0 % C0 2 , 2 9 0 0 ppmv S0 2 , 3%v 0 2 and balance N 2 , 8 minutes for each cycle. Calcination: 1 0 0 % N 2 . (Lines show corresponding results with no S O 2 present for two of the sorbents.) 1 5 6 Figure 6 . 7 ATGR once-through tests, at 8 5 0 ° C and 1 0 1 kPa with 2 1 2 - 2 5 0 pm Strassburg limestone. Top and bottom curves and for limiting case where there was no S O 2 or C O 2 respectively. Points are for co-capture case showing total (squares) calcium utilization, utilization for C O 2 capture (triangles) and utilization for S O 2 capture (circles) 1 5 7 Figure 6 . 8 Relation between sulfate content and calcination rate for three sorbents. ATGR co-capture tests, same conditions as in Figure 6 1 5 8 Figure 6 . 9 Cyclic C O 2 retention performances in 8 5 0 ° C PTGA tests, effect of total pressure. Co-capture with 2 1 2 - 2 5 0 pm Strassburg limestone. Sorption: 8 %v C0 2 , 1 1 2 5 ppmv S0 2 , 3 % O 2 , and balance N 2 , 4 minute for each cycle. Calcination: 1 0 1 kPa, 1 0 0 % N 2 1 5 8 Figure 6 . 1 0 Cyclic C 0 2 retention performances in 8 5 0 ° C PTGA tests: effect of total pressure. Co-capture with 2 1 2 - 2 5 0 pm Arctic dolomite. Sorption: 8 %v C 0 2 , 1 1 2 5 ppmv S0 2 , 3 % 0 2 , and balance N 2 , 4 minute for each cycle. Calcination: 1 0 1 kPa, 1 0 0 % N 2 . 8%v C0 2 , 850T 1 5 9 Figure 6 . 1 1 Sulfur mapping for non-calcined ATGR samples from varied co-capture cycles, same test conditions as in Figure 6 . 6 . (Light points mark sulfur) (a) Strassburg limestone after 1 2 cycles (b) Kelly Rock limestone after 1 5 cycles (c) Arctic dolomite, after 7 cycles '. : '. • '. 1 5 9 Figure 6 . 1 2 High-resolution SEM pictures for calcines after ATGR co-capture tests, (a) Strassburg limestone, after 1 2 t h cycle (b) Kelly Rock limestone, after 1 5 t h cycle (c) Arctic dolomite, after 7 t h cycle 1 6 0 Figure 6 . 1 3 Cyclic performances during PTGA co-capture tests: effect of PCo2- 2 1 2 - 2 5 0 pm Straussburg limestone. Sorption at 8 5 0 ° C and 1 .8 MPa, with 1 1 2 5 ppmv S0 2 , 3%v O2 and balance N 2 , 4 minutes for each cycle and calcination at 8 5 0 ° C and 1 0 1 kPa, 1 0 0 % N 2 . (a) For CaO conversion to CaC0 3; (b) For CaO conversion to CaS04 1 6 1 Figure 6 . 1 4 Cyclic performance during PTGA co-capture tests: effect of PCo2- 2 1 2 - 2 5 0 pm Arctic dolomite particles. Sorption at 8 5 0 ° C , and 1 .8 MPa, with 1 1 2 5 ppmv S0 2 , 3%v 0 2 Xlll and balance N 2 , 4 minutes for each cycle and calcination at 850°C and 101 kPa, 100% N 2 . (a) For CaO conversion to CaC03; (b) For CaO conversion to CaS0 4 162 Figure 7.1 High-resolution SEM photos of calcines derived from initial calcination of 212-250 urn Strassburg limestone particles; (a) h-CaO (b) c-CaO 182 Figure 7.2 Cyclic performance: comparison of c-CaO and h-CaO (no S0 2 present, sorbent derived from 212-250 um Strassburg limestone). Test conditions: 850°C calcination and carbonation, carbonation in 100% C 0 2 , calcination in 100% N 2 . Fast stage of carbonation finished 182 Figure 7.3 Pore size distribution: comparison of c-CaO and h-CaO (no S0 2 present, sorbent derived from Strassburg limestone). Samples are the same as in Figure 7.2........ 183 Figure 7.4 Effect of steam calcination on cyclic capture (No S0 2 present, 212-250 um Strassburg limestone Strassburg limestone). Test conditions. 850°C calcination and carbonation, Carbonation in 100% C0 2 . Calcination in 95% steam, balance N 2 . Fast stage of carbonation finished. 183 Figure 7.5 Pore size distribution: effect of steam calcination (No S0 2 present, Strassburg limestone). Same samples as in Figure 7.4 184 Figure 7.6 Calcium utilization for C 0 2 capture: effect of varying operating conditions, 212-250 um particles Test conditions: 850 °C calcination and sorption. Sorption in 80% C0 2 . 3% 0 2, 2900 ppm S0 2 and balance N 2 or with steam. Calcination in 100% N 2 . 8 minutes for each sorption , 185 Figure 7.7 Effect of hydration on co-capture (C02+S02). 212-250 um Strassburg limestone. Procedure: 1-h co-capture or S0 2 sorption followed by 30 min hydration and 8 min cycles of re-capture: (a) Conversion of CaO to CaC03; (b) Conversion of CaO to CaS04 186 Figure 7.8 Effect of intermediate hydration of carbonates on further carbonation (no S0 2 present) A: 10 min carbonation plus 30 min steam hydration at 300°C plus 30 min carbonation; B: 60 min carbonation plus 30 min steam hydration at 300°C plus 30 min carbonation; C: 60 min carbonation plus 10 min liquid water steam hydration at room temperature plus 30 min carbonation. 850°C for both calcination (100% N2) and carbonation (100% C0 2 ) 187 Figure 7.9 Effect of intermediate steam (95%v) or water (100%) hydration at various temperatures of sintered calcines on cycling (no S0 2 present). 850°C for both calcination (100% N2) and carbonation (100% C0 2 ) 188 Figure 7.10 Variation of mass increase due to sorption divided by total initial mass of CaO with cycling. Test conditions are provided in Table 7.3 189 Figure 7.11 Effect of various dopants on CaO reversibility in cyclic calcination/carbonation. Cycling conditions: 850°C for calcination (in 100% N2) and carbonation (in 100% C0 2): (a) y-Al203; (b) Si0 2 and Kaolinite; (c) Zr0 2 ; (d) MgO; (e) dolomite, Ti0 2 , precipitated calcium carbonate; (f) other dopants as identified in Table 7.4 190 Figure 8.1 Candidate processes for C 0 2 and S0 2 removal with calcium-based sorbents; Options A and C. sulphation before calcination/carbonation; Options B and D: calcination/carbonation before sulphation; Options C and D involve pressurized fluidized-bed combustion (PFBC) 207 XIV Figure 8.2 Sorbent performance: conversion history based on option A. 212-250 pm Strassburg limestone. Sulphated first for 10 minutes at 850°C and 0.1 MPa. 850°C calcination/carbonation cycles. Baseline conditions for fresh limestone: see Table 8.1 for C O 2 sorption stage : 208 Figure 8.3. SEM photos for calcines from different runs; for (b), (c) and (d), see Table 8.2, for test conditions, (a) CaO sulphation, followed by 4 calcination/carbonation (c/c) cycles, 12% calcium utilization for sulphation (on molar base). Test conditions: as for solid points in Figure 8.2. Option A. (b) Co-capture for 600 s followed by 4 c/c cycles, Option C. (c) Direct sulphation, 5000 ppm S O 2 , for 600 s followed by 4 c/c cycles, Option C. (d) Direct sulphation, 5000 ppm S O 2 , for 2400 s followed by 4 c/c cycles, Option C 209 Figure 8.4 Sulphation of cycled sorbents after different number of calcination/carbonation cycles. 212-250 pm Strassburg limestone. (Option B) 210 Figure 8.5 EDX sulfur-mapping of calcines for 212-250 pm Strassburg limestone. Test conditions as in Figure 8.4. (a) Sulphation of Strassburg limestone for 2400 s (Fresh sorbent) (b) Sulphation of Strassburg sorbent after 15 calcination/carbonation cycles 210 Figure 8.6 Sulphation of cycled sorbents after different number of calcination/carbonation cycles. 212-250 pm Arctic dolomite (Option B). For conditions, see Table 8.1 211 Figure 8.7 Mass change profiles during typical runs for Option C with 212-250 pm Strassburg limestone, (a). Direct sulphation for 600 s. (b). Simultaneous sulphation and carbonation for 10 minutes. Test conditions: Both direct sulphation and simultaneous sulphation and carbonation: 5000 ppm S0 2 , 15% C O 2 , 3% 0 2, N 2 balance before calcination/carbonation cycling, 1.82 MPa, 850°C 212 Figure 8.8 Cyclic calcination/carbonation performance of 212-250 pm Strassburg limestone after direct sulphation (Option C). See Table 8.1 for calcination/carbonation cycling conditions. 213 Figure 8.9 Cyclic performance of 212-250 pm Strassburg limestone after simultaneous sulphation and carbonation, and a complete calcination. Option C with 212-250 pm Strassburg limestone. See Table 8.1 for calcination/carbonation cycling conditions 213 Figure 8.10 Cyclic performance of 212-250 pm Arctic dolomite after simultaneous sulphation and carbonation, and calcination. Option C. Baseline conditions for fresh sorbent: See Table 8.1 for calcination/carbonation cycling conditions 214 Figure 8.11 Cyclic performance of 212-250 pm Arctic dolomite after direct sulphation and calcination. Option C. Baseline test conditions for fresh sorbent: See Table 8.1 for calcination/carbonation cycling conditions 214 Figure 8.12 Surface texture of cycled 212-250 pm Arctic particles different sulphation history (a) Same test conditions as in Figure 8.10 (b) Same test conditions as in Figure 8.11 215 Figure 8.13 Mass change profiles with simultaneous carbonation and sulphation after 7 cycles of carbonation and calcination. Option D with 212-250 pm Strassburg particles 216 Figure 8.14 Sulphation and carbonation extents during co-capture with cycled 212-250 pm Strassburg limestone particles. Option D vs. Option B. See Table 8.1 for conditions 216 xv Figure 8.15 Sulphation and carbonation extents during co-capture with cycled 212-250 pm Arctic dolomite particles. Option D vs. Option B. See Table 8.1 for conditions... 217 Figure 8.16 Effect of small residual S02 on calcination/carbonation cycling. 212-250 pm Strassburg limestone particles. Points: Sorption with 14% C02, 100 ppmv S02,3% O2 and balance N 2 at 850°C and 1.82 MPa. Calcination in 100% N 2at 850°C and atmospheric pressure. Cycling conditions for fresh sorbent with no SO2 present: see Table 8.1 for carbonation conditions 218 Figure 9.1 Illustration of operating procedure and masses used to calculate of calcium utilization for C 0 2 and H 2S capture. Co-capture at 850°C and 0.76 MPa. Calcination: in 100% N 2 . Sorption: l%v H 2S, 20%v C0 2 , 12..6%v H 2 and 1.5 %v CO with the balance N 2...,. 235 Figure 9.2 Once-through tests with 212-250 pm Strassburg limestone. Initial calcination at 850°C and 101 kPa with 100% N 2 . Co-capture at 850°C and 0.76 MPa with l%yH 2S, 20%v CO2, 12.6%v H2 and 1.5 %v CO, balance N 2 . Carbonation with no H 2S: same as for co-capture tests except with no H2S. Sulfidation with no CO2: same as co-capture tests except with no CO.. 236 Figure 9.3 Cyclic capture performance for H2S and CO2 with 212-250 pm Strassburg limestone at850°C and 0.76 MPa for sorption, 850°C and 101 kPa for calcination. 3-minute sorption for each cycle. Sorption gas compositions except specified H2S and CO2 concentrations in legends: 12.6%v H 2 , l%v CO, balance N 2 . Calcination: 100% N 2 237 Figure 9.4 High-resolution SEM photos of co-capture 212-250 pm Strassburg limestone calcines: (a) H 2S and C02 capture in l%v H 2S, 20%v C02, 12.6%v H 2 , l%v CO, balance N 2 at 0.76 MPa, 3 min sorption for each cycle, (b) S02 and C02 capture in 1125 ppmv S02, 8%v CO2, 3%v O2 and balance N2 at 1.8 MPa, 4 min of sorption for each cycle. Sorption and calcination temperatures for both cases were 850°C. All calcinations in 100% N2 238 Figure 9.5 EDX sulfur mapping for 212-250 pm reacted Strassburg limestone, (a) Calcine from 1 h continuous sulfidation at 850°C, 0.76 MPa with l%v H 2S, I2.6%v H 2 , 1.5 %v CO, balance N 2 . (no CO2). (b) Calcine from cyclic H2S/CO2 sorption for 15 cycles at 850°C, 0.76 MPa with l%v H 2S, 20%v C02) 12.6%v H 2 , 1.5 %v CO, balance N 2 ; Calcination at 850°C and 101 kPa with 100% N 2 ; 3 min for each sorption stage 238 Figure 9.6 Mass changes during first and fifth cycles of H2S/C02 capture tests. 212-250 pm Strassburg limestone. Tests conditions except where specified: sorption at 850°C and 0.76 MPa with 20% C02, 1%H2S, 12.6%v H 2 , l%v CO, balance N 2 , 850°C; Calcination at 850°C and 10.1 kPa with 100% N 2 . 3 minutes for each sorption stage 239 Figure 9.7 Effect of changing temperature on cyclic capture of H 2S and C02. 212-250 pm Strassburg limestone. Sorption at 0.76 MPa, all with 20% C02, 1% H 2S, 12.6%v H 2 , l%v CO, balance N 2 . Calcination at 850oC and 101 kPa with 100% N 2 . 3 min for each sorption stage :., 240 Figure 9.8 Effect of particle size on cyclic capture of H 2S and C02. Strassburg limestone. Sorption at 850°C and 0.76 MPa, with 20% C02, 1% H 2S, 12.6%v H 2 , l%v CO, balance N 2 . Calcination at 850°C and 101 kPa with 100% N 2 . 3 min for each sorption stage 241 Figure 9.9 Effect of cyclic sorption time on cyclic capture of H 2S and C02. 212-250 pm Strassburg limestone. Sorption at 850°C and 0.76 MPa, with 20% C02, 1% H 2S, 12.6%v X V I H 2 , l%v CO, balance N 2 . Calcination at 850°C and 101 kPa with 100% N 2 . 3 min for each sorption stage 242 Figure 9.10 Effect of total pressure on cyclic capture of H 2S and C0 2 . 212-250 um Strassburg limestone. Sorption at 850°C, with 20% C 0 2 , 1% H 2S, 12.6%v H 2 , l%v CO, balance N 2 . Calcination at 850°C and 101 kPa with 100% N 2 . 3 min for each sorption stage .243 Figure 9.11 Effect of sorbent type on cyclic capture of H 2S and C0 2 . 212-250 urn. Sorption at 850°C and 0.76 MPa, with 20% C0 2 , 1% H 2S, 12.6%v H 2 , l%v CO, balance N 2 . Calcination at 850°C and 101 kPa with 100% N 2 . 3 min for each sorption stage 244 Figure 9.12 SEM photos for different sorbents after 15 co-capture cycles. Same test conditions as for Figure 9.10. (a) Arctic dolomite (b) Kelly Rock limestone 245 Figure 9.13 Sulfur mapping for calcines of different sorbents after co-capture. Same test conditions as for Figure 9.10. (a) Arctic dolomite after 15 cycles, (b) Kelly Rock after 15 cycles 245 Figure A l l Comparison of experimental and fitting results (Dataset 1) 269 Figure AI.2 Comparison of experimental and fitting results (Dataset 1). Same as above but in Arrhenius plot 269 Figure AI.3 Comparison of experimental and fitting results (Dataset 2).. 270 Figure AI.4 Comparison of experimental and fitting equation (Dataset 2) Same as above but in Arrhenius plot 270 Figure Al.5 Comparison of experimental and fitting equation,(Dataset 3) in Arrhenius plot (Data predicted by the nonlinear fitted equation is off this chart) 271 Figure A1.6 Comparison of experimental and fitting equation (Dataset 4) in Arrhenius plot (Data predicted by the nonlinear fitted equation is off this chart) 271 Figure AJJI.l Illustration of the relation between initial pore radii based on measured pore volume data 277 Figure AVTJ. 1 (Comparing with other runs in Figure 6.2) High-resolution SEM pictures of calcines, with no S 0 2 present. Same test conditions as in Figure 6.1. Danyang limestone. : : 294 Figure AVT1.2 (Comparing with other runs in Figure 6.11) Sulfur mapping for non-calcined ATGR samples. 850C for both tests. S0 2 only test: 2900 ppm S0 2 , 3% 0 2 and balance N 2 for 2 hours; Co-capture test conditions as in Figure 6.6, 15 cycles. (Light points mark sulfur). Danyang limestone 294 Figure AVTJ.3 (Comparing with other runs in Figure 6.3) Evolution of pore size distribution with calcination/carbonation cycling at 850°C. Test conditions: same as in Figure 6.1 for S0 2 test and in Figure 6.6 for co-captures. Danyang limestone 295 Figure A V E 4 (Comparing with other runs in Figure 6.3) Evolution of pore size distribution with calcination/carbonation cycling at 850°C. Test conditions: same as in Figure 6.1 for no S0 2 test and in Figure 6.6 for co-captures. Havelock limestone 295 Figure AVTJ.5 (This is for conversion of CaO to CaS04, comparing that for CaO to CaC03 shown in Figure 6.9) Cyclic C 0 2 retention performances in 850°C PTGA tests: effect of XV11 total pressure. Co-capture with 212-250 urn Strassburg limestone. Sorption: 8 %v C O 2 , 1125 ppmv S O 2 , 3% 0 2 , and balance N 2 , 4 minute for each cycle. Calcination: 101 kPa, 100% N 2 296 Figure AVTJ.6 (This is for conversion of CaO to CaS04, comparing that for CaO to C a C 0 3 shown in Figure 6.10) Cyclic C 0 2 retention performances in 850°C PTGA tests: effect total pressure. Co-capture with 212-250 um Arctic dolomite. Sorption: 8 %v C O 2 , 1125 ppmv S O 2 , 3% 0 2 , and balance N 2 , 4 minute for each cycle. Calcination: 101 kPa, 100% 8%v C O 2 , 850°C 296 Figure AVH.7 Cyclic performances in no S O 2 only tests: effect of total pressure. Atmospheric calcination. 850qC, 8% C O 2 balanced by N 2 , 212-250 um Straussburg limestone 297 Figure AVTJ.8 Cyclic performances in no S O 2 only tests: effect of temperature. 8%v C O 2 balanced by N 2 , 212-250 um, and Straussburg limestone 297 Figure AV1L9 Cyclic performances in no S O 2 only tests: effect of Pco2- Pt=18.2 bars C O 2 only, Straussburg, 212-250 um 298 Figure AVTH. 1 Performance of c-CaO and h-CaO (no S O 2 present, sorbent derived from 212-250 um Arctic dolomite). Same test conditions as specified in Figure 7. 2 299 Figure AVT1L2 Pore size distribution: comparison of c-CaO and h-CaO. (no S0 2 present, sorbent derived from Arctic dolomite) 299 Figure AVTH.3 Effect of steam on cyclic capture (No S O 2 present, 212-250 um Strassburg limestone). Test conditions: 850°C calcination and carbonation; Carbonation in 86% C02.and 14% steam; Calcination in 100% N 2 300 Figure AVffl.4 Effect of steam on cyclic capture (No S O 2 present, 212-250 um Arctic dolomite). Test conditions: 850°C calcination and carbonation; Carbonation in 86% C02 .and 14% steam; Calcination in 100% N 2 300 Figure AVTJI.5 Calcium utilization for S0 2 capture: effect of varying operating conditions Comparing with Figure 7.6 Calcium utilization for S O 2 capture: effect of varying operating conditions, 212-250 um particles Test conditions: 850 °C calcination and sorption, Sorption in 80% C0 2 . 3%> 0 2 , 2900 ppm S0 2 and balance N 2 or with steam. Calcination in 100% N 2 . 8 minutes for each sorption 301 Figure AVTH6 Close-up view of one cycle of S 0 2 / C 0 2 sorption followed by a C a C 0 3 calcination in N 2 and then a slow reduction of CaS04 at cycle 9. test conditions are described in Table 7.3 .< 302 Figure AVTJX7 EDX element mapping for modified sorbents with the light color points marking the element of the dopants 303 Figure ATX. 1 Case study flowsheet for FBC-based S02/C02 sequential capture process (based on Option B in Chapter 8)...'. 306 Figure ATX.2 Sensitivity of CaC03 feed to average CaO utilization in carbonator for a 262 MW fossil-fuel-fired conventional FBC unit 310 Figure AEX.3 Sensitivity of energy production to average CaO conversion in the carbonator for of a 262 M W fossil-fuel-fired conventional FBC unit 311 XVl l l Figure ATX.4 Sensitivity of total CaCCh needed in FBC on sulfur content in raw fuel fed to the FBC for a 262 M W fossil-fuel -fired conventional FBC unit 312 Figure A X . l Ellingham plot for the reaction (AJX.2) with isobaric lines for partial pressure of CO., .316 Figure AX.2 Cyclic C O 2 capture performance and subsequent S O 2 capture. Test conditions: carbonation-650°C, 100% CO z ; Calcination- 800°C, 100% N 2 ; Sulphation-850°C, 2900 ppm S0 2 , 3% 0 2 , balance N 2 317 Figure AX.3 Comparison of cyclic C O 2 capture ability with literature test 318 Figure AX.4 Test history for cyclic C O 2 capture and S0 2 capture. Test conditions: carbonation-650°C, 100% C0 2 ; Calcination- 800°C, 100% N 2 ; Sulphation-850°C, 2900 ppm S0 2 , 3% 0 2 , balance N 2 319 xix A C K N O W L E D G E M E N T S I would like to express my gratitude to my supervisors, Drs John Grace, C. Jim Lim and Edward J. Anthony, who provided me the opportunity to work in this excellent group, encouraged me to think freely and critically, guided me with great patience and encouragement on the road to complete this thesis work and always supported me whenever I have questions and difficulties to deal with in my research and life. Many thanks to Drs Naoko Ellis and Peter V. Barr for being in my supervising committee. Advices and encouragement from Herbert Andrus and John Chiu are highly appreciated. Interesting discussions with Professor Said Elnashaire about gas-solid modeling are greatly appreciated. I do not think I could thank enough for the constant mental support from my parents, who are always behind me and encouraging me on my career journey. Their way of loving and having faith in me sustained me in building my strength and persistence in the face of difficulties. My wife, Mengzhu, deserves my most sincere gratitude from the bottom of my heart for her love, encouragement and her wonderful gifts to make our life so enjoyable. I would like to thank everyone of my colleagues, friends, who helped and cared for me, like roses, spreading their fragrance when I was walking in their presence. xx CO-AUTHORSHIP S T A T E M E N T This is a manuscript-based thesis. Chapters 2-9 are independent manuscripts with no overlaps. They are either published papers in peer-reviewed journals/international conference proceedings or manuscripts prepared for publication in peer-reviewed journals. I declare that my contributions to these manuscripts lie in the following areas: • Proposing research work and designing research plan. • Performing experiments, modeling and data analyses. • Preparing manuscript drafts and working with co-authors in revising the drafts. The use of these manuscripts in this thesis is under the permission of all the co-authors. x x i Chapter 1 Introduction, thermodynamics and equipment CHAPTER 1 INTRODUCTION, THERMODYNAMICS AND EQUIPMENT 1.1 Background and structure of this thesis: using calcium-based sorbent to removal C 0 2 CO2 emissions, as the major source of greenhouse gases, are receiving increasing attention due to the impact of greenhouse gases on global climate change. Affordable CO2 control technology has become the focus of worldwide research. Among all the point sources of CO2 emissions, power/thermal plants firing fossil fuels account for the principal anthropogenic source (about 6 gigatonnes carbon/year) Power plants implementing CO2 capture from power system and storage could become effective net sinks of CO2 from the atmosphere (Obersteiner et al. 2001). To capture and dispose of CO2 safely before it is released into the atmosphere, three major steps are involved (FETC, 1999): capture, storage/transportation and permanent disposal. Among these, capture is the most costly step (Wallance, 2000). This thesis work is dedicated to CO2 capture. C 0 2 capture technologies under development can be classified into post-combustion capture, pre-combustion technology and oxy-fired technologies (Davison and Thambimuthu, 2004). Among the post-combustion technologies, amine-based CO2 wet scrubbing is believed to be the most technologically mature (Rao and Rubin, 2002), but it requires high capital and operational costs. Another technology, similar to the approach studied in this work, is to use a sodium-based sorbent to cyclically capture CO2, with a calciner to drive off a high-purity CO2 stream for storage (Liang et al., 2004). 1 Chapter 1 Introduction, thermodynamics and equipment Pre-combustion technology removes CO2 during syngas reforming. Syngas comes from natural gas or from fossil fuels gasification. C O 2 removal can be realized by chemical absorption, e.g. based on CaO (Ortiz and Harrison, 2001), chemical adsorption, e.g. using hydrotalcite (Hufton and Sircar, 2001) or membrane separation. The C02-free syngas after further treatment can then be burnt for power generation, or converted to other products such as methanol. The Zero-Emission-Coal concept (Ziock et al., 2004a; 2004b) is an example of a pre-combustion process with limestone as C 0 2 sorbent. In the pre-combustion processes, capture of CO2 can also enhance the water-gas-shift (WGS) reaction to produce more hydrogen according to the Le Chatelie' principle. Furthermore, the exothermic carbonation reaction reduces the energy requirement i n the reformer or gasifier. Oxy-fired combustion has been studied intensively, including in a small demonstration unit (e.g. Crosietand Thambimuthu, 2000). Recently a relatively new combustion concept, chemical looping combustion has been developed (Lyngfelt et al., 2001). In the combustor, gaseous hydrocarbon is oxidized by contacting solid oxidizing agents (instead of air), and the reduced solid species is then transferred to a separate chamber where it is oxidized by reacting with air. As a result, the only gases produced from the combustion system are C O 2 and H 2 0 . The latter can then be condensed and separated, leaving a relatively pure C 0 2 stream, suitable for storage by various sequestration methods. Reaction 1.1 has recently received renewed research interest (Shimizu et al. 1999; Abanades et al. 2003) as a means of removing C 0 2 between a C0 2 - lean fluidized bed combustor (FBC) and a calciner where nearly pure C 0 2 is driven off for storage. In the combustion temperature range, CaO is also reactive to another air pollutant, S 0 2 . If co-capture of S 0 2 and C 0 2 could be realized, the CaO looping system could be readily retrofitted to existing F B C 2 Chapter I Introduction, thermodynamics and equipment plants. This concept does not belong to any of the above categories, instead being an in situ process. C a O + C 0 2 ^ C a C 0 3 AH 2 9 8 J t=-178 kJ/mol (1.1) This thesis focuses on the application of reaction (1.1) in a variety of energy systems, from steam reformers of hydrocarbons, to combustors of fossil fuels and gasifiers of fossil fuels. A schematic diagram of process in which such carriers could be applied is shown in Figure 1.1, with calcium-based sorbents (limestones or dolomites) serving as C 0 2 carriers between a carbonator and a calciner. Heat must be supplied to the calciner to regenerate the sorbent. Petroleum coke (Wang et al., 2004), natural gas (Abanades et al., 2005), steam (Abanades et al., 2005), solid heat carriers (Abanades et a l , 2005), waste heat from fuel cell plants (Ziock et al., 2004a; b) and pure C 0 2 (Ortiz and Harrison, 2001) have all been proposed as heating sources or carriers. When relatively pure C 0 2 is generated from the calciner, the task of capturing C 0 2 is completed. The relatively pure C 0 2 stream is then further transported for ultimate safe disposal (sequestration). 3 Chapter I Introduction, thermodynamics and equipment C0 2-free flue gas/syngas Fuel Steam/Air/O, For further gas cleaning or heat exchanger CaC0 3-rich solids Reformer/Combustor/ Gasifier (Carbonator) CaO+C0 2=>CaC0 3-178kJ/mol Spent sorbent| /Fuel ash I Relatively pure C02-stream for safe disposal Heat Calciner CaC0 3=>CaO+C0 2 +178 kJ/mol CaO-rich solids Fresh limestone /dolomite Figure 1.1 Schematic diagram of the process studied in this thesis work to apply calcium-based sorbents The thesis starts from fundamental topics, such as kinetics of the CaO-CCh reaction, and extends to a series of practical topics, including the application of calcium-based sorbents (mainly limestones and dolomites) in FBC-based processes and other energy systems. The cyclic reactivity of the calcium-based sorbent is the key unifying subject throughout this thesis. This thesis is organized as a series of manuscript-based chapters (from Chapter 2 to 9), each dealing with an independent topic and having an independent structure. Content relevant, but not included in the manuscripts, because of its supplemental nature, is relegated to Appendices. This thesis consists of ten chapters. After this introductory chapter, it includes kinetic study of the CaO-CC>2 reaction (Chapter 2), modeling of the gas-solid reaction (Chapter 3), a mechanistic study of sorbent sintering (Chapter 4), application of calcium-based sorbents in a FBC based process for simultaneous capture of C O 2 and S O 2 (Chapter 5, 6, 7), consideration of Chapter 1 Introduction, thermodynamics and equipment sequential capture of C 0 2 and S0 2 (Chapter 8) and work relevant to simultaneous capture of C 0 2 and H 2S in coal-fueled gasification (Chapter 9). Chapter 10 presents overall conclusions and recommendations. 1.2 Thermodynamics underlying this thesis As the basis of the following chapters, several thermodynamic issues need to be first addressed: 1.2.1 Sorbent properties, calcination and carbonation Naturally occurring limestones and dolomites are the major sorbents considered throughout this thesis. Boynton (1979) gave a detailed introduction to the properties of limestones, dolomites and limes. Limestones usually contain calcium carbonate and other impurities. Calcite and aragonite are the two major polymorphs of calcium carbonate, with calcite being the more common. As seen in the phase diagram in Figure 1.2, adapted from the work of Redfern et al., (1989), an extremely high pressure is needed for aragonite to form by carbonation (Note: in the original figure the pressures were kbar): Given the test conditions employed in this work and given that all of the limestones tested were initially calcitic, only calcite is dealt with in this work. Dolomites contain calcium carbonate and magnesium carbonate, as well as some impurities. Usually a dolomite contains a solid solution of calcium carbonate and magnesium carbonate, as well as a small portion of free magnesium carbonate. During calcination of dolomites, the breakdown of the chemical bonds between the two carbonates has been investigated using in situ XRD measurements (Engler et al., 1988). It was also found that after full calcination, the mixture consists only of a mechanical mixture of calcium oxide and magnesium oxide. 5 Chapter 1 Introduction, thermodynamics and equipment 1400 10 16 20 P r o c u r e k b a r 2 3 Figure 1.2 Phase diagram for calcite and aragonite. Adapted from the work of Redfern etal., (1989). 1.2.2 Sorbent calcination temperature The temperature required for calcination is predicted using a thermodynamics database. In this work, HSC Chemistry 4.0 (Ronie, 1997) is used to make the equilibrium predictions. In this database, equilibrium concentrations are calculated based on non-stoichimetric Gibbs free energy minimization. In these methods, no details about reaction mechanism or kinetics are 6 Chapter 1 Introduction, thermodynamics and equipment required in the numerical solution. In this thesis, the ideal gas model was also applied in the calculations. All solid phases involved were assumed to be pure condensed phases with activities of unity. More details about the non-stoichimetric Gibbs free energy minimization method is included in the monograph of Smith and Missen (1982). The Figure 1.3 shows plots of standard free energy change versus temperature, sometimes called Ellingham diagrams (Gaskell, 1995) for the calcination reactions, C a C 0 3 ^ C a O + C 0 2 Atf 2 9 8 i , =178 kJ/mol (1.2) M g C 0 3 ^ M g O + C 0 2 &H29sk =100.9 kJ/mol (1.3) CaMg(C0 3 ) 2 ^MgO+C0 2 +CaC0 3 A ^ 2 9 8 ^ =124.5 kJ/mol (1.4) Standard state refers to 101 kPa pressure and any temperature of interest. The tie-lines (broken lines) show the change in Gibbs free energy when the partial pressure of C 0 2 is varied from the standard state to the C 0 2 partial pressure level of interest. The incipient calcination temperatures correspond to the intersection of the two lines. Figure 1.3 shows that when the temperature increases at a constant C 0 2 partial pressure, the free MgC0 3 part of a dolomite decomposes first, followed by the half-calcination of CaMg(C0 3) 2 into a calcite, and finally by calcite calcination. While calcite calcination is seen to be more favourable at high temperatures than at low temperatures in the absence of C 0 2 , calcination of a limestone in a pure N 2 environment does not occur in practice until temperatures of at least 650°C. This probably occurs because calcination kinetics are very slow at lower temperatures. For calcination in vacuum, an equation cited by Dennis and Hayhurst (1987) gives ^calcination ~ ^calcination C^C02,eq ~ PcOl,true ~^o) 0-5) This shows that even when the true C 0 2 partial pressure is zero, the temperature has to be high enough to overcome the P0 term. This term is, only important when the driving force is very low 7 Chapter I Introduction, thermodynamics and equipment (e.g. less than ~6 kPa depending on temperature level). For decomposition of both limestones and dolomites, it was found that the incipient decomposition temperature is also related to particle size and crystalline properties (Boynton, 1979). 0 200 400 600 800 1000 T(°C) Figure 1.3 Ellingham plot for the three carbonate decomposition reactions with isobaric lines for partial pressure of CO2. It should also be pointed out that Figure 1.3 assumes ideal gas behaviour. The compressibility factors (Perry et al., 1997) for all relevant gases H 2 0, 0 2 , SO2, CO2, CO2, H 2and N 2 are all close to unity for the test conditions of this work (150-400°C for steam and 500-950°C for all other gases, 101-2000 kPa), so that ideality is a good assumption. 8 Chapter I Introduction, thermodynamics and equipment The equilibrium partial pressure of C O 2 for CaO carbonation can also be predicted by the correlation (Baker, 1962), P c o ^ ( i n k P a ) = 1 0 - 8 3 0 ™ 7 9 ( T , n K ; (1.6) Figure 1.3 also indicates that for MgO to carbonate at temperatures of interests for FBC, e.g., 850°C, extremely high P c o 2 (>50 MPa) would have to be maintained, well beyond the pressures of interest in this thesis. Therefore carbonation of MgO is neglected here. Equation (1.6) can also be used to estimate the C O 2 removal efficiency. In a pressurized reactor, assuming that the inlet C O 2 molar fraction is fCQ and the total pressure is PT , the maximum achievable C O 2 removal efficiency is, " C O l = — f (1-7) rTJCOl The removal efficiency is seen to be a function of the overall pressure and the molar fraction of C O 2 at the reactor inlet. Figure 1.4 illustrates how the maximum C 0 2 removal efficiency varies with total pressure and inlet molar fraction of C O 2 at 850°C. Because of possible hydration when using CaO to capture C 0 2 in a reformer or gasifier where there is a high steam partial pressure, the conditions for CaO hydration are shown in Figure 1.5, constructed in the same manner as Figure 1.3. It is seen that, MgO needs a higher H 2 O partial pressure to be hydrated than CaO. 9 Chapter I Introduction, thermodynamics and equipment o 1 0.8 D "3 0.6 o e S 0.4 O u x 0.2 <3 500 1000 1500 Total pressure (kPa) 2000 2500 Figure 1.4 Maximum C 0 2 removal efficiency as a function of total pressure and molar fraction of C 0 2 in the flue gas at the inlet of reactor for a temperature of 850°C. 10 Chapter 1 Introduction, thermodynamics and equipment 3 0 r -30 1 1 — J : ' ' — * . 1 0 0 - 3 0 0 5 0 0 7 0 0 9 0 0 T ( ° C ) Figure 1.5 Ellingham plot for CaO and MgO hydration reactions isobaric lines for partial pressure of H2O. 1.2.3 C0 2 capture in FBC: Effect of S02 on CaO carbonation When dolomites are employed as CO2 sorbents in FBC-based processes, the MgO part, although inert to CO2, might not be inert to SO2. For the reaction MgO+S0 2+l/20 2^MgS0 4 AH29iK =-386.5 kJ/mol (1.8) 11 Chapter 1 Introduction, thermodynamics and equipment HSC 4.0 was employed to the equilibrium SO2 concentration in the presence of 3% O2 at varying temperatures and total pressures. The results appear in Figure 1.6. They show that 2250 ppm SO2 are needed for MgO sulphation at 1 atm and 850°C. The equilibrium SO2 concentration decreases with increasing total pressure in good agreement with experimental data of Dewing and Richardson (1959) summarized by Hartman and Svoboda (1985). 100000 r •3 0.1 ' 1 1 — L 1 — ' ->' c r W 100 500 900 1300 1700 2100 2500 Total pressure (kPa) Figure 1.6 Equilibrium SO2 concentrations for reaction given in equation (1.8) as the function of reactor temperature and total pressure in the presence of 3%v 0 2. 1.2.4 C0 2 capture during gasification Chapter 9 deals with CO2 and SO2 removal for two limestones and a dolomite under simulated gasification conditions. With a simulated syngas, gas phase reactions need to be taken into account. The typical gas feed in the experiment, i.e. l%v H2S, 20%v CO2, 12.6%v H 2 and 1.5%v CO with the balance N2, was analyzed with HSC 4.0. The predicted results are shown in Figure 1.7 as a function of reactor temperature for a total pressure of 0.76 MPa. 12 Chapter I Introduction, thermodynamics and equipment Figure 1.7 Thermodynamic predictions for gaseous products as a function of temperature for 0.76 MPa total pressure. Gas feed: l%v H 2S, 20%v C 0 2 , 12.6%v H 2 and 1 5%v CO with N 2 balance. Figure 1.7 shows that for the specified gas feed, the water gas-shift reaction favours production of H 2 0 and CO, and consumption of C0 2 , causing the actual C 0 2 concentration to be a few percentage points lower than in the inlet gas stream. Gaseous sulfur is retained in the form of H 2S, with negligible H 2S decomposition, as revealed by the low concentration of S2. The concentrations of N H 3 and COS are also very low. The calculation also shows negligible carbon deposition and negligible reaction of CaS with C0 2 . In the experiments described in Chapter 9, relatively high concentrations of hydrogen and CO were maintained, as suggested in the literature (Fenouli and Lynn, 1995a) to avoid H 2S decomposition and CaS oxidization by C0 2 . 13 Chapter 1 Introduction, thermodynamics and equipment Based on the equilibrium CO2 and H 2 0 concentrations with the given feed shown in Figure 1.7, the equilibrium H2S concentration can be predicted for the other two reversible reactions studied in Chapter 9, i.e., CaO sulfidation (reaction 1.6) and CaC03 direct sulfidation C a O + H 2 S ( g ) ^ C a S + H 2 0 (g) A/f 2 9 8 J^ =-59.4 kJ/mol (1.9) C a C 0 3 + H 2 S ( g ) ^ C a S + H 2 0 ( g ) + C 0 2 ( g ) Atf 2 9 8 J C =118.7 kJ/mol (1.10) For a typical operating conditions, such as 850°C and 0.76 M P a , the equilibrium H 2 S concentration was predicted to be 114 ppm and 235 ppm for these two reactions, respectively, using the equilibrium partial pressures Pco2=~14%v and PH 2 O=~8%V , predicted from Figure 1.7. 1.2.5 Sorption-enhanced hydrogen production using a calcium-based C O 2 sorbent Enhancement of hydrogen production with C 0 2 capture with CaO during steam reforming of methane has been investigated by Ortiz and Harrison (2001). They showed that CaO removal could shift the chemical equilibrium to produce relatively pure hydrogen. Sorption-enhanced hydrogen production in a bubbling fluidized-bed reformer was recently demonstrated experimentally by Johnsen et al. (2006). The thermodynamic predictions of Ortiz and Harrison (2001) show that when CaO hydration is favoured, the maximum achievable hydrogen purity is lower than under non-hydration conditions. 14 Chapter 1 Introduction, thermodynamics and equipment 1 r 0 I 1 J- — ' - J 500 700 900 1100 1300 Temperature (°C) (a) Effect of reactor temperature (at 1000 kPa) 0.8 r Si c o o S 0.6 0.4 0.2 •With CaO - Without CaO 0 200 400 600 800 1000 1200 Total pressure (kPa) (b) Effect of reactor pressure (at.850°C) Figure 1.8 Effect of C 0 2 capture on F£2 production in a simulated gasification system. Feed: 1 kmol carbon, 1.5 kmol steam, 0.01 kmol H 2 S , with and without 1 kmol CaO. 15 Chapter I Introduction, thermodynamics and equipment As discussed in Chapter 9, calcium-based sorbents are also applicable to C O 2 and H 2 S removal in fossil-fuel gasifiers. Removal of C 0 2 could also enhance the water-gas shift reaction to produce hydrogen of higher concentration. HSC 4.0 was used to simulate the effect of C O 2 removal by CaO gasification for a feedstock of 1 kmol C, 0.01 kmol H 2S and 1.5 kmol steam. Illustrative results are shown in Figure 1.8 as a function of reactor temperature and pressure. Figure 1.8 clearly shows that when the conditions are favourable for C O 2 capture by CaO, e.g., for T<900°C at P=l MPa or for P>0.3 MPa and T=850°C, H 2 production can be appreciably enhanced. This finding is similar to that for sorption-enhanced steam methane reforming. 1.2.6 E f f e c t o f H2O, l o w - t e m p e r a t u r e e u t e c t i c b e t w e e n C a O - C a ( O H ) 2 - C a C 0 3 The operation of a steam reformer or gasifier using a calcium-based C O 2 sorbent should also consider the possible formation of a low-melting-point eutectic. It is commonly believed that the CaO-Ca(OH)2-CaC03 system likely produces a low-temperature eutectic (Curren et al., 1967; Fuerstenau et al., 1981). Curran et al. (1967) calculated the eutectic temperature for the binary systems. CaO-Ca(OH)2 and CaC03-Ca(OH)2 by equilibrium analyses. The results are summarized in Table 1.1, where the equilibrium steam pressures to maintain solid-phase Ca(OH)2 are also shown for the CaO hydration reaction, CaO+H20 (g)=Ca(OH)2 (s) A H 2 9 i a , =-109.1 kJ/mol (1.11) However, there are some limitations in this prediction because the predicted H 2 O equilibrium partial pressures differ from those predicted by HSC 4.0 database (see Figure 1.5), i.e., 4200 kPa for 790°C and 690 kPa for 638°C. The phase diagram for such a binary system could be more rigorously constructed by the method introduced by Gaskell (1995). However, the specific heat data needed for imaginary materials such as CaO (1), Ca(OH)2 (1) and Ca(OH)2 (1) are mostly 16 Chapter 1 Introduction, thermodynamics and equipment absent from the literature. Other eutectics, as reviewed by Fuerstenau et al. (1981) for the binary system CaC03-Ca(OH) 2 , are also shown in Table 1.1. T a b l e 1.1 E u t e c t i c f o r m a t i o n c o n d i t i o n s for the CaO-Ca(OH) 2 a n d C a C 0 3 - C a ( O H ) 2 s y s t e m System Ca(OH )2 -CaO Ca(OH) 2-CaCOj Ca(OH) 2 -CaC0 3 Ca(OH )2 -CaC03 Min. Eutectic formation temperature 790 °C (Steam pressure 5400 kPa) 638°C (Steam pressure 880 kPa) 675°C 645°C(1 , 100 MPa total pressure) Molar fraction where eutectic forms 69% CaO 41% Ca(OH)2 37%CaC0 3 Reference Curranetal. 1976 Curran et al. 1976 Originally from Wyllie and Turtle (1960) Reviewed by Fuerstenau et al. (1981) Originally from Wyllie and Raynor. (1965); Reviewed by Fuerstenau et al. (1981) Fuerstenau et al. (1981) determined the composition of the ternary eutectic system CaO-Ca(OH) 2-CaC0 3 using DSC techniques. They found eutectics to form in the temperature range 640-789°C and at pressures of 1.7 to 7.0 MPa, with CaO accounting for 2.3%-5.0% and CaC0 3 9.5 to 64.5% molar fraction. 1.3 Experimental equipment and materials 1.3.1 Apparatus The experimental work carried out in this project was all conducted in bench-scale equipment. Three thermogravimetric reactors were utilized, as portrayed in Figures 1.9, 1.10 and 1.11 17 Chapter 1 Introduction, thermodynamics and equipment Figures 1.9 shows the atmospheric pressure thermogravimetric analyzer S1TJMADZU TA60 based system (ATGA). The balance has a precision of 1 ug. The sample mass is usually <30 mg. This apparatus was mainly used for kinetic studies with C O 2 partial pressures from 0 to 101 kPa. Gases flow from top to bottom. All gas flows were controlled by mass flow controllers. More details can be found in the chapters where the A T G A system was employed. Needle Mass flow valve controller Computer Figure 1.9 Schematic of SHTMADZU-based atmospheric thermogravimetric analyzer system (ATGA). Figure 1.10 provides a schematic of the pressurized TGA (PTGA) system, with operating pressures up to 2.4 MPa. It was primarily employed to test kinetics at relatively high pressures. This unit includes a Cahn 100 balance with 1 ug sensitivity. The reactor itself is made of Inconel 600 alloy. A pressure regulator adjusted the pressure to the desired level. The mass detection range is adjustable, usually between 0-100 mg. All gas flows were controlled by mass flow controllers. More details are provided in later chapters where the PTGA system was employed. 18 Chapter 1 Introduction, thermodynamics and equipment Needle valve Figure 1.10 Schematic of pressurized thermogravimetric analyzer (PTGA) system The atmospheric T G reactor (ATGR) is shown in Figures 1.11. This was utilized in a variety of experiments involving sulphation, carbonation, calcination and hydration. It was especially useful for cyclic calcination/carbonation tests. A load cell at the top of the reactor chamber allows much larger sample sizes, i.e., 0-10 grams of total suspension mass (samples plus basket and wire). The advantage is that enough samples can be produced for topology analyses. The TGR allows operation with S O 2 , C O 2 , CO, air, steam etc. The flows of gases except steam were controlled by mass flow controllers, whereas the flow rate of steam was controlled by a water pump (powered by a AC motor). Most sample basket were made from stainless wire mesh. Stainless steel wire experienced some oxidization under high-temperature operation, but the mass change was usually less than a few milligrams, negligible for samples of mass 400 mg. More details can be found in Laursen et al. (2000, 2001). 19 Chapter 1 Introduction, thermodynamics and equipment Needle Mass flow Figure 1.11 Schematic of atmospheric thermogravimetric reactor (ATGR) system 1.3.2 Sorbents Seven commercially available sorbents were tested: Strassburg limestone (US), Danyang limestone (Korea), Havelock limestone (Canada), Kelly Rock limestone (Canada), Cadomin limestone (Canada), Arctic dolomite (Norway) and GS dolomite (US). Table 1.2 gives their chemical analyses. Table 1.2 Chemical analyses (% by wt in each case) Sorbent S i0 2 A1 20 3 Fe 2 0 3 MgO CaO Na 2 0 K 2 0 LOI Strassburg 0.94 0.19 0.94 1.25 53.7 0.02 0.08 42.9 Danyang 1.55 0.9 0.29 1.03 52.94 0.01 0.32 42.9 Havelock 1.5 <0.38 <0.55 0.59 53.99 <0.17 O.08 43.34 Cadomin 1.5 <0.38 <0.55 2.25 55.12 <0.17 0.21 42.77 Kelly Rock 5.31 1.54 0.36 0.58 51.74 0.07 0.36 43.14 Arctic Dolomite 2.12 0.17 1.3 21.25 30.51 0.15 0.04 44.4 GS Dolomite 0.87. 0.16 0.15 18.69 33.41 0.04 0.08 46.5 20 Chapter 1 Introduction, thermodynamics and equipment The apparatus and sorbents used in the major chapters are summarized in Table 1.3. Table 1.3 Summary of sorbents and reactors in each chapter. Applicat ions and chapter Reactor used Sorbent used Kinetic study (Chapters 2, 3) A T G A Pressurized TGA ATGR Strassburg limestone Arctic dolomite Sintering modeling (Chapter 4) ATGR Strassburg limestone Co-capture of SO2 and C02 (Chapters 5, 6) ATGR PTGA Limestones: Strassburg, Danyang, Havelock, Kelly Rock, Cadomin Dolomites: Arctic, GS Effect of steam, sorbent modification etc. (Chapter 7) ATGR Strassburg limestone Arctic dolomite Sequential capture of S02 and C02(Chapter 8) PTGA Strassburg limestone Arctic dolomite Co-capture of H2S and C02(Chapter 9) PTGA Strassburg and Kelly Rock limestone Arctic dolomite 1.4 Nomenclature Symbols calcination calcinatio co2 P p Pt Po Calcination rate Calcination rate constant Inlet C O 2 molar fraction True partial pressure of C O 2 Equilibrium partial pressure of C O 2 Total reactor pressure Constant in Equation (1.5) mol/m Is r mol 1 s'm2 (kPa) kPa kPa kPa kPa Greek letters Vco, Maximum achievable C O 2 removal efficiency 21 Chapter 2 Determination of intrinsic rate constants of the CaO-C02 reaction C H A P T E R 2 D E T E R M I N A T I O N O F INTRINSIC R A T E CONSTANTS O F T H E C a O - C 0 2 R E A C T I O N A version of this chapter was prepared for publication in Chemical Engineering Science. Authors. Sun, P., Grace, J. R., Lim, C. J. and Anthony, E. J. This manuscript will be submitted soon. 2.1 Introduction Extensive research recently has been devoted to C 0 2 capture techniques because of growing concerns over greenhouse gas emissions. Calcium-based materials have attracted particular attention (Silaban and Harrison, 1995; Silaban et al. 1996; Shimizu et al., 1999; Abanades et al. 2003) as potential sorbents for cyclic CCvcapture processes because of their potential for regeneration. Among the possible applications of calcium-based sorbents for C O 2 removal are steam reformers (Han and Harrison, 1994; Balasubramanian et al., 1999; Ortiz and Harrison, 2001; Johnsen et al., 2006), gasifiers of fossil fuels (Lin et al. 2001; Jukkola et al., 2005) to enhance water-gas-shift reaction giving higher hydrogen yields, and fluidized bed combustors with in-situ C O 2 capture (Shimizu et al., 1999; Abanades et al, 2003; Abanades et al, 2004a). These applications all involve sorbent cycling between calcination and carbonation. However, the kinetics of the carbonation reaction have been studied much less than its reverse reaction, calcination. As a typical gas-solid reaction producing a solid product, carbonation is initially fast, followed by a much slower stage. In the absence of pore diffusion, it is controlled jointly by surface reaction and product layer diffusion. However, among the limited number of kinetic studies, only a handful have focused on the intrinsic surface reaction kinetics. Nitsch (1970) 22 Chapter 2 Determination of intrinsic rate constants of the CaO-CC>2 reaction observed a zero-activation energy for an initial growth stage over a narrow temperature range, 800-850°C. Bhatia and Perlmutter (1983) studied the kinetics via a random pore model to fit the early portion of the experimental data. Their experiments, carried out in an atmospheric TGA with 0-10%v C 0 2 , suggested a zero-activation energy. The zero activation energy was further supported by Dennis and Hayhurst (1987) based on an equilibrium analysis of their calcination data. However, a zero-activation-energy is rare. More likely, the activation energy for the carbonation is actually small, i.e. near zero. Based on tests in an atmospheric TGA, Kyaw et al. (1996) reported activation energies of 78 kJ/mol for a limestone and 35 kJ/mol for a dolomite. Another issue with respect to carbonation kinetics is the uncertainty with respect to intrinsic reaction order. Bhatia and Perlmutter (1983) claimed a first order reaction for Pco2 from 0-10 kPa, whereas Kyaw et al. (1996) measured a close-to-zero order for a much higher C 0 2 partial pressure. Work is needed to cover a wider range of Pco2. This chapter focuses on an intrinsic kinetic study for the CaO-C0 2 reaction with two commercial calcium-based sorbents, one a limestone and the other a dolomite. A wide range of C 0 2 partial pressures was tested using both atmospheric and pressurized TGA equipment. 2.2 Experimental details Two fixed-bed thermogravimetric analyzers were used in this study, one operated at atmospheric pressure and the other under pressurized conditions. The atmospheric SHTMADZU TA60 system (ATGA) was used for most tests to measure rate constants, with C 0 2 partial pressures varying from zero to 0.1 MPa. Gases flow from top to bottom. A schematic is provided in Figure 1.9. The inner diameter of the reactor is -0.02 m, whereas the diameter of sorbent pan is 0.004 m. The pressurized TGA (PTGA) was used to test kinetics at much higher C 0 2 partial pressures. This unit includes a Cahn 100 balance of 1 ug sensitivity and a reactor column made 23 Chapter 2 Determination of intrinsic rate constants of the CaO-C02 reaction of Inconel 600 alloy. A pressure regulator adjusted the overall pressure to the desired level. When the required pressure was achieved, the reactant gas stream flowed upwards, passing the suspension sample basket and then leaving through the outlet of the reactor at the middle of column. A schematic of the PTGA system appears in Figure 1.10. Ultra-pure grade C O 2 , N 2 and He were mixed at the inlet of both reactors. Mass flow controllers accurately controlled the gas flow rates to obtain the desired C O 2 concentrations. To minimize the influence of calcination heating rates, pure C 0 2 was fed initially to both reactors, preventing CaC03 decomposition during heating to 850°C at 50°C/min in the ATGA or at 20°C/min in the PTGA. In both cases, the calcination was carried out at atmospheric pressure. When 850°C was reached, the C O 2 was replaced by He or N 2 to initiate calcination. After completion of calcination, a pre-determined carbonation temperature was then approached while maintaining the flow of inert gas. For PTGA tests, the desired pressure was achieved by rapidly injecting inert gas from the top of the vessel. After adjusting the control valves to establish an inert gas flow rate which would provide the desired C O 2 concentration, carbonation was initiated by opening the C O 2 control valve. Strassburg limestone and Arctic dolomite were selected for these tests. Chemical analyses of these two sorbents are provided in Table 1.2 2.2.1 Elimination of physical limitations in A T G A tests Preparatory runs were conducted with Strassburg limestones to eliminate physical constraints in the A T G A measurements. It has been shown that, after a fast initial stage, the CaO-C0 2 reaction becomes an extremely slow product-layer-controlled reaction (Bhatia and Perlmutter, 1983; Abanades et al., 2003). Kinetic control in the initial stage was achieved by careful control of the operating conditions. 24 Chapter 2 Determination of intrinsic rate constants of the CaO-COj reaction In the A T G A tests, C O 2 was introduced after no appreciable further calcination was observed during exposure to helium at a pre-determined flow rate. After initiating the C O 2 flow, before significant reaction-related mass change, there was a small step increase of apparent mass due to drag, as confirmed by both drag force calculation and a blank test. This had to be accounted for. The calcium utilization was calculated on a molar basis, ^ reacted Ca moles inrilf)(t)-mCafl(Qi) ,. • ... . X = = C a 0 y ' C a 0 X ' KmCaCm (0) x purity 1100) (2.1) total Ca moles 4 4 V C a C 0 3 W ' y ' where 4 4 and 1 0 0 are the C O 2 and CaCCb molar weights, in g/mol. Reducing sample size and dispersing particles as well as possible can minimize the "packing" effect causing inter-particle mass transfer resistance. External mass transfer limitations can be minimized by using high gas flow rates. Experiments using varying sample size showed that starting with 2 .9 mg fresh limestone gave no difference in mass breakthrough profile compared with 4 mg. Thus in this study, all runs started with samples smaller than 2 .9 mg. Tests at 600°C with P C o 2 - 5 0 kPa and the balance He and with varying gas flow rates showed no difference in mass breakthrough profiles for flow rates from 4 0 0 up to 8 0 0 ml/min. Therefore, total flow rates around 4 0 0 ml/min were selected for the later runs. All flow rates provided here were measured at room temperature at the inlet of the ATGA. Regarding the mechanism for the elimination of external mass transfer limitations, it Uod, should be noted that although the particle Reynolds number (defined as Re = —) is small v ( ~ 6 x l 0 " 3 ) for a typical case: e.g., for a 4 8 pm particle at 850°C with inlet helium flow rate of 4 0 0 ml/min at room temperature, Sherwood number would be ~2 for a single particle (Szekely et al., 1 9 7 6 ) and almost independent of gas flow rate. However, in a TGA reactor, the fine particles behave like a fixed-bed with the true gas velocity between the particles much lower than the 25 Chapter 2 Determination of intrinsic rate constants of the CaO-C02 reaction superficial velocity and the true Sherwood number (for a fixed bed) is much smaller than for a single particle and still a function of Reynolds number (Bird, 2000). Therefore, an increase of gas flow rate in the T G A would effectively increase the true Sherwood number (Bird, 2000), so that the external mass transfer coefficient would change with flow rate. Since changes in flow rate did not appreciably alter the measured rate, it can be inferred that external mass transfer was not limiting. Similar results with varying particles size showed that using particles smaller than 53-63 um can eliminate the effect of intraparticle mass transfer. In current studies, 38-45 um particles were therefore employed. Finally changing the balance inert gas from N 2 to He did not result in any change in carbonation history, indicating that heat transfer limitations are negligible. In this study, helium was used in all ATGA tests, whereas N 2 was used in the PTGA tests. 2.2.2 P T G A tests The sample pan in the PTGA has a cross-sectional area about 4 times larger than in the ATGA. In each PTGA test, 8-11 mg of sample were used, resulting in a well-dispersed thin-layer sorbent bed. The PTGA tests provided supplemental data for P c o 2 > 101 kPa. Two groups of tests were performed with the Strassburg limestone, one at 101 kPa total pressure to compare with the ATGA run results, and the other at a total pressure of 800 kPa. For Arctic dolomite, all tests were performed at 800 kPa. After calcination, 100% C O 2 replaced 100% N 2 to initiate carbonation. Carbonation was detected 0.5 to 1 minute after fully opening the valve for C O 2 introduction and fully closing the N 2 valve at the same time. Although full openness would give a maximum of 500 ml/min flow rate for the atmospheric reactor vessel, the actual flow rate was only about half this value at 800 26 Chapter 2 Determination of intrinsic rate constants of the CaO-CC>2 reaction kPa total pressure. No appreciable decrease in reaction rate at a gas flowrate of 250 ml/min as seen below, confirmed that the gas film is not rate-limiting. A blank test confirmed that no appreciable buoyancy force was caused by changing the gas from N 2 to C O 2 . 2.3 Results and discussion 2.3.1 Direct measurement of carbonation kinetics: use of a gas-solid reaction model SEM photos of the Strassburg calcines (e.g. Figure 2.1) show that grain shapes are almost spherical, but with some evidence of initial grain-neck growth due to the early stage of sintering. Arctic calcines (not shown here) did not show clear grain boundaries, probably due to much smaller grain size: This indicates that the grain model should be a good descriptive model, although the grains are not perfectly spherical due to earlier sintering for limestone calcines. Therefore a grain model with kinetic control was chosen to determine the kinetic-controlled region. The grain model under kinetic control gives (Szekely et al., 1976) dX = 3>" (2 2) dt(l - X)2n K ' In integral form, [l-(\-X)V3] = rxt (2.3) Equation (2.3) shows that when the reaction is under kinetic control, a plot of (1-(1-X)1/3) vs t should give a straight line of slope r. Figure 2.2 obtained from mass vs time profiles illustrates how the kinetic control region is identified with the help of the grain model. At the start of the reaction, there was usually a short induction period featuring a lower slope. Following this and starting from a point with maximum slope, several seconds of mass breakthrough led to a nearly linear stage (shown in Figure 2.2) 27 Chapter 2 Determination of intrinsic rate constants of the CaO-C02 reaction followed by an obvious slow-down. The short linear stage identifies the kinetic control region with the slope of this stage giving the intrinsic surface reaction rate, defined as r in equation (2.3). The starting points of the linear stage, featuring the maximum slope, usually resides at low conversions, i.e. 5%. Following the kinetic control region, the slow-down stage is controlled by both surface reaction and product-layer diffusion, showing much slower slopes. As the carbonation reaction proceeds further, product-layer diffusion becomes more and more important, resulting in much slower carbonation. In the kinetic control region, the slope (also the maximum slope) is constant and gives the value of r in equation (2.3). It should be noted that the initial induction period right before the linear kinetic-controlled stage should also belong to the kinetic control region, in spite of its slower slope as found for typical gas-solid reactions. Therefore the constant value of r in the kinetic controlled region obtained from the linear plot slope can also be extended to represent the true rate of r at the zero conversion point, i.e., r0=r (2.4) Some other typical plots obtained by applying the grain model are shown in Figure 2.3. The reaction rate for a gas-solid reaction, is usually defined as a specific rate, R-W^] (1/s) (25) When the reaction is under kinetic control, the specific rate can be further expressed in power law form, AY R = dt(l-X)=56kAPc°>-Pco1,qYS = 3r(\-Xr" (2.6) where 56 is the molar weight of CaO, in g/mol. Equations (2.3) and (2.5) lead to, 28 Chapter 2 Determination of intrinsic rate constants of the CaO-CC>2 reaction At the initial time 0, when the surface area is S0, equation (2.7) becomes, 0 dt 3r0=56ks(PCO2-Pco^qyS0 (2.8) In a logarithmic form, equation (2.8) becomes, lnr0=ln(56fc,/3) + nln(P C O j - J P C O 2 ^) + ln5 0 (2.9) Hence, the slope of lnr0 versus \n(PC02 - PCOi eq) plot gives the order of the CaO-CC«2 reaction with respect to the CO2 partial pressure. Such plots are shown in Figures 2.4 and 2.5. They show that at low CO2 partial pressure driving force or Pco2- Pco2,eq, the CaO-C02 reaction is 1st order. However at higher Pco2- Pco2,eq, e.g. for about Pco2- Pco2,eq=6 kPa or higher for 850°C tests and at about 8 kPa or higher for 600°C runs, an obvious transition of the reaction order occurs. The order then quickly approaches zero or close-to-zero order at higher CO2 partial pressure. The transition, if present, as shown in both Figures 2.4 and 2.5, appears to be sharp, suggesting an abrupt shift in control mechanism. For both sorbents at both temperatures, the driving force in the transition region generally lies between 6 and 12 kPa. Driving force (Pco2-Pco2,eq) at -10 kPa is taken as a reasonable approximation of the turning point hereafter. With ks = k0 exp(-E IRIT) at the n=0 region, equation (2.9) can be rewritten as, l n r 0 = l n ( 5 6 * 0 V 3 ) - ( | ) ^ (2.10) The two parameters, activation energy E and pre-exponential factor k0 for both sorbents were obtained by nonlinearly fitting the Arrhenius equation (2.10) with experimental data. Note that a comparison has been made between nonlinear fitting and linear fitting using the logarithmic form of equation (2.10). The results are summarized in Appendix I, showing that 29 Chapter 2 Determination of intrinsic rate constants of the CaO-C02 reaction better results were obtained through nonlinear fitting. The initial specific surface areas S0 for both sorbents were measured by N 2 adsorption, giving 29 m2/g for Strassburg calcines and 49 m2/g for fully calcined Arctic dolomite. The Arctic calcines contain both MgO and CaO, but the surface area for carbonation is only the CaO part. It is reasonable to assume that the surface area ratio of the CaO part to the MgO part is the same as their volumetric ratio. For a fully calcined dolomite, both CaO and MgO crystals should be cubic, with the cube edge length being 0.48 nm for CaO and 0.42 nm for MgO (Boynton, 1979). In the Arctic dolomite, the molar ratio of CaO 0 483 to MgO is 1.1, so the molar volume ratio of CaO to MgO should be *1 1 = 1-64. Thus the CaO surface is of Arctic dolomite is estimated to be 31 m2/g. The data in these figures are extracted from runs with 80%, 50% and 15% inlet C O 2 volumetric fractions for Strassburg limestone and 100% for Arctic dolomite. All correspond to the zero-order regions as discussed above. Arrhenius plots using the fitted parameters E and k0 are shown in Figures 2.6 and 2.7. The resulting parameters are, for limestones, ks = 1.67 x 10"3 exp(——) E = 29 + 4kJ I mol; P Co2-Pco2 ,eq >10 kPa R,T m s • (2.11) for dolomites, ks = 1.04x 10~3 exp( ) — r — ;E = 24± 6kJ I mol; P C02-Pco2 ,eq >10 kPa R,T m s (2.12) where the limits on the activation energy correspond to a 95% confidence level. Since at the turning point from the first-order to zero-order reaction, or at Pco^ -PCOt e q = ^ kPa, the specific reaction rate predicted as the zero-order reaction, i.e. R0 = 56ks(PC0^ -PCOieq)S0 30 Chapter 2 Determination of intrinsic rate constants of the CaO-C02 reaction (Pco2-Pco2,eq - 1 ° ^ a ) should equal the one predicted for the first order reaction, it can be derived that for the first order region, for limestones, ks = 1.67 x 10"4 exp(——) """ \E - 29±4kJ I mol; Pco - Pco eq <10 kPa RtT m skPa 1 2' (2.13) Ti IflOl for dolomites, A: =1.04x10 4 exp( ) — ;E = 24 ±6kJ I mol ;PC0 -Pco , <10kPa RtT m2 skPa 1 °2>" (2.14) 2.3.2 High-pressure measurement It is been shown above that the driving force (Pco2-Pco2,eq) has no further effect on the rate of reaction in the A T G A when it exceeds 10 kPa. In the ATGA tests, the highest Pco2 achievable was 0.1 MPa. To test greater Pco2, higher total pressure levels are required. The PTGA tests were undertaken for this purpose. The same methods as described above were applied to the high-pressure data, except that data were smoothed by moving averages to dampen fluctuations probably associated with more significant drag imparted by the higher-density gas. Typical plots based on the grain model are shown in Figures 2.8 and 2.9, combined with calcium utilization (conversion) curves for Strassburg limestone and Arctic dolomite, respectively. The maximum' slopes are also shown in these Figures, in spite of the sparse data points compared to A T G A tests. The Arrhenius plots in Figures 2.10 and 2.11 compare the high-pressure data with previous ATGA data at various temperatures and 0.1 MPa total pressure. These plots indicate that, despite the greater scatter in the PTGA data, the absolute rates at 800 kPa are very similar to those for 101 kPa ATGA runs. Considering the wide difference of 31 Chapter 2 Determination of intrinsic rate constants of the CaO-C02 reaction absolute PCo2 between the ATGA and PTGA tests, it can be concluded that increasing PCo2 from 0.1 to 800 kPa makes no difference in the intrinsic carbonation kinetics, supporting the above conclusion that for, (Pco2-Pco2,eq) > 10 kPa, the carbonation reaction was of zero order. In view of the PTGA data scatter, the PTGA data were not used for rate constant parameter fitting. 2.3.3 Carbonation data from equilibrium analyses As a check, rate constants can also be estimated by means of the equilibrium constant and kinetic data of the reverse (calcination) reaction, recognizing that K=k c a r bonat ion/koaic inat ion. The kinetics of calcination have received much more attention in the literature than the kinetics of carbonation. Borgwardt (1985) reviewed important results for calcination and made measurements with ultra-fine particles (1-90 pm) in two different types of reactors (a TGA and an entrained flow reactor). The calcination activation energy from his study, 200±13 kJ /mol, has been widely used in calcination modeling and is adopted here. The equilibrium constant has been reported (Baker, 1962) to be, K = 1 0 (-8308/T + 9.079) K has units of kPa when the partial pressures are expressed in kPa. Combining the data utilized by Borgwardt (1985) and equation (2.15), we obtain the rate constants for the carbonation reaction from the slope of Arrhenius plot in Figure 2.12. The activation energy derived from these data is 41.5 kJ /mol, significantly higher than that determined in our ATGA tests for Strassburg limestone. The discrepancy probably arises because at the initial point of carbonation, the assumption of equilibrium is not truly valid, while the lattice structure of the sorbents is also relevant as discussed below. However, the order of magnitude agrees reasonably with the values reported above (24-29 kJ/mol). 32 Chapter 2 Determination of intrinsic rate constants of the CaO-COi reaction 2.3.4 Comparison with earlier results Bhatia and Perlmutter (1983) defined a rate as: dX R= = ks'S'C (2.16) dt(\~X) s v ' and measured ks' = 5.95 ± 0.18 xl0~10.m4/mol/s for a limestone, where C is the driving force in mol/m3. The CaO true density is 3.34xl06.g/m3 (Barin, 1989). As noted above, the initial surface area for Strassburg limestone is 29 m2/g from N 2 adsorption measurement, while the porosity is 3.5xl0"4 m3/kg from mercury intrusion measurements. So at 615°C, 10 kPa Pco2 with total pressure 0.1 MPa and at the initial point (X=l), dX(X = 0) = 3 r = 3 x l 6 7 x l 0 - 3 x56/3x29xexp — =0.054 s"1 forthiswork dt #,(615 + 273) whereas, dX^t~ ^ = k ' ' S ' C = 0 0 4 1 s" ! from Bhatia and Perlmutter (1983). These two values are of similar magnitude. However, both are smaller than the value of -0.5 s*1 obtained in the work of Kyaw et al. (1996). The activation energy derived in this work differs significantly from the results of Bhatia and Perlmutter (1983) who, based on specific rate law analyses, reported that the CaO-C0 2 reaction has a zero activation energy in the kinetic-controlled region. This result was supported by Dennis and Hayhurst (1987) based on equilibrium analyses and calcination data. On the other hand, our activation energy is lower than values reported by Kyaw et al. (1996) who found 78 kJ/mol for limestone and 35 kJ/mol for dolomite. The first order agrees well with results of Bhatia and Perlmutter (1983) who investigated Pco2 only up to 10 kPa at 615°C. The zero order agrees well with the work of Kyaw et al. (1996) 33 Chapter 2 Determination of intrinsic rate constants of the CaO-COj reaction who observed a zero-order reaction for both a limestone and a dolomite based on initial rate analyses with Pco2>20kPa and a total pressure of 0.1. MPa. Our results show that there is a shift in reaction order, with neither of the former studies covering the entire P c o 2 range. 2.3.5 Mechanistic explanations for the variable order reaction In an attempt to explain the variable order based on a Langmuir mechanism, a series of elementary steps is assumed to be possible, with C a O « C 0 2 a s the intermediate complex. Similar to Allen and Hayhurst (1991) in the case of the CaO-H2S reaction, several assumptions are made before applying the Langmuir mechanism. The first is that under the steady-state hypothesis, a fraction 0 of CaO sites is covered by short-lived intermediate C a O » C 0 2 , while a fraction of (1-6 ) is available for C O 2 adsorption. The second assumption is that the intermediate forms C a C 0 3 quickly and that the formation of CaCOs is not the rate-limiting step. Similar to the sketch of Allen and Hayhurst (1991), the intermediate lies between CaO and CaC03. It is assumed that C a C 0 3 does not provide any diffusion resistance to C O 2 within the kinetic control region. The shift in apparent reaction order suggests that the C a O - C 0 2 reaction might follow a series of elementary steps k, CaO+C0 2 <=>CaO«C0 2 step 1 C a O « C 0 2 < = > C a C 0 3 step 2 kd Step 1 involves a reversible process with C O 2 molecules colliding with CaO sites to produce C a O » C 0 2 intermediate complexes. Step 2 involves adsorption to produce solid product. 34 Chapter 2 Determination of intrinsic rate constants of the CaO-C02 reaction Because of the assumption that step 2 is at equilibrium, so that step 1 is rate-controlling and, k3x&-k4=0 (2.17) or 6=k4lk3 (2.18) Then, R=fcx(l-0)xpC O 2-A;2x e (2.19) Re-arrangement of equations (2.18) and (2.19) leads to, i? = ^ [ ( l - ^ ) P C 0 2 - ^ l ] (2.20) Overall equilibrium gives ^ 2 ^ 4 = — = Pc02 , so that ktk3 K k3 K Since step 2 is assumed to be fast, site coverage of CaO«C0 2 complexes is very low, so that 0 = f c / £ ? « l . Equation (2.21) then becomes, R- k] (PC02 - Pc07,eq ) (2.22) This equation indicates the rate dependence through the Langmuir mechanism. However, from measurements at initial points we found that the apparent first order dependence only holds for Pco2 ~Pco2,eq -10 above which a zero-order dependence was observed. Based on this observation, it can be concluded that there must be a shift in reaction mechanism when Pco2 ~ Pcoi.eq exceeds this critical value, such as ~10 kPa. When the driving force (PC 0 2 - PC02 eq) is greater than the critical value, the concentration of transition complex CaO«C02 is high enough that the CaO sites are mostly saturated so that 9 approaches 1. The assumption of low 35 Chapter 2 Determination of intrinsic rate constants of the CaO-CC>2 reaction site-coverage of complex CaO'CCh is no longer valid. At this stage, step 2 becomes rate-limiting instead. Thus, R= k3* 6 -k4 = k3*l-k4= k3-k4. At the critical value of PC02 - PC02<ea (~10 kPa), the rate determined from equation (2.22) reaches its maximum value, numerically equal to the rate determined as R=&r&4, independent of {Pc02 ~ Pc02,eq )• Based on above analyses, the derived rate law is further written as, R = k{(PC02-PC02eq) (*, isin —1—) forP C O 2 - .P C O 2 e ? <10kPa (2.23) 7 s• kPa R = k3-k4= (PC02-PC0Xeq)kx = 10/c, (kx is in — J — ; k3,k4 are in - ) s• kPa s forPCO2-PCO2j„>10kPa (2.24) The equations (2.23) and (2.24) defining rate constants are equivalent to equations (2.11) and (2.13) for a limestone or (2.12) and (2.14) for a dolomite. In equation (2.24), (PC02 —Pco2sq)~^ kPa, so that the units of R are 1/s. 2.3.6 Other issues Our tests indicated a somewhat higher activation energy for limestones than for dolomites, consistent with the results of Kyaw etai. (1996). The difference is probably due to the structure-related difference during nucleation or formation of the solid product. For the CaO-C02 reaction with swelling solid product, the rate of reaction is not only related to chemical energy change, but also to mechanical energy, such as strain energy between different grains. Young (1966) illustrated that during the nucleation stage, the solid structure might influence the overall free energy change, with 36 Chapter 2 Determination of intrinsic rate constants of the CaO-CC>2 reaction AG = b]onn2n + b2nn(AGc +Est) (2.25) Equation (2.25) indicates that to make the reaction occur, the energy barrier involves not only chemical energy barriers, but also a structure-related mechanical barrier. Dolomite-derived calcines, with a mixture of CaO and MgO, are likely to have relatively low strain energies, resulting in lower apparent activation energy than limestones. Other factors, involving surface energy and nucleation properties, may also be significant. Secondly, dolomites differ from limestones in their large proportion of magnesium, causing dislocations at full calcination. These dislocations could play several roles in re-crystallization or nucleation. Young (1966) pointed out that during the incipient nucleation stage, the germ nucleus is "situated at regions of disorders, such as points of emergence of dislocations, vacancies, interstitial or impurity clusters", so that the reactant at these points must re-crystallize more readily than on a normal surface, because fewer bonds per ion need to be broken. This could also result in a lower energy barrier for dolomites. Energy levels between calcination and carbonation are shown schematically in Figure 2.13. Both activation energies obtained from Strassburg limestone and Arctic dolomite, when taking into account the standard deviations of the fitting, are seen to be nearly consistent with thermodynamic analysis, further supporting the experimental results of this study. 2.4 Conclusions Based on a grain model, it appears that kinetic control applies only briefly in the initial stage of carbonation. The rate constant of the CaO-C02 reaction was mainly studied using an atmospheric thermogravimetric analyzer (ATGA). Tests with two calcium-based sorbents gave similar results, showing a variable intrinsic rate dependence on C 0 2 partial pressure. The carbonation reaction was first order only for C 0 2 partial pressures driving force less than 10 kPa, 37 Chapter 2 Determination of intrinsic rate constants of the CaO-CO'2 reaction abruptly changing to 0-order for higher C O 2 partial pressures. Tests in a pressurized TGA showed that further increases in C O 2 partial pressures could not enhance the intrinsic rate, thereby confirming the zero-order dependence above 10 kPa. Based on a Langmuir mechanism, it is plausible that saturation of CaO sites with intermediate complex CaO»C02 above a critical C O 2 partial pressure driving force is responsible for the shift in reaction order. The activation energies are 29±4kJ /mol for limestones and 24±6kJ/mol for dolomites, reasonably consistent with those based on equilibrium analysis. Structural differences between the two sorbents are believed to be responsible for their different activation energies. 2.5 Nomenclature Symbols a Activity bx b2 C Constants in equation (2.25) C O 2 concentration Particle diameter mol/m m E E„ Activation energy Strain energy kJ/mol kJ/mol Pre-exponential factor in equation (2.10) mol Rate constant in equations (2.19) (2.20) etc. 1 s -kPa k2 k3 k4 Rate constant in equations (2.17) (2.18) (2.20) etc. Rate constant in equation (2.6) mol (kPay 38 Chapter 2 Determination of intrinsic rate constants of the CaO-C02 reaction m4/mol/s ks' Rate constant used by Bhatia and Perlmutter (1983) ^ .... . m3/mol K Concentration equilibrium constant mcao (0) Mass of calcine at time zero mcao (0 Mass of calcine at time t ^ mcaco3 (0) Mass of fresh sorbent at start of run ^ n Reaction order defined in equation (2.6) nn Number of molecules in nucleus in equation (2.25) kPa Pco2 Partial pressure of C O 2 p„„ Equilibrium partial pressure of C O 2 kPa Purity Purity of C a C 0 3 in fresh sorbents rQ Reaction rate r at time zero s"1 r Reaction rate in grain model s"1 dX s"1 R Specific reaction rate R = dt(\ - X) Rep Particle Reynolds number R Specific reaction rate at time zero s"1 Rt Gas constant, 8.314xl0"3 kJ/mol/K S Specific surface area m2/g S0 Specific surface area at time zero m2/g S' Specific surface area per unit volume of total bulk space, m 2/ m3 used by Bhatia and Perlmutter (1983) T Reaction time s T Temperature K 39 Chapter 2 Determination of intrinsic rate constants of the CaO-CC>2 reaction Uo Superficial velocity in the ATGA reactor m/s X Conversion of CaO Greek letters 0 Fraction of site occupied by CaO«C02 complexes v Kinematic viscosity m2/s cr Surface free energy kJ/mol AG Overall free energy kJ/mol AGC Chemical free energy kJ/mol 40 Chapter 2 Determination of intrinsic rate constants of the CaO-CC>2 reaction 0.3 r 0.2 >< i 0.1 Maximum slope exists for the linear stage, starting from this point • 1—3 • • o i n o 5 inear stage: y = 0.0199x - 0.0258 0 5 Timet(s) 1 0 15 Figure 2.2 Slope extraction with the aid of the grain model during early stages of carbonation for 38-45 um Strassburg limestone particles at 700°C with 15% C 0 2 and 85% He. 41 Chapter 2 Determination of intrinsic rate constants of the CaO-C02 reaction X 0.4 0.3 0.2 0.1 0 A 8 5 0 ° C 8 0 % v C O 2 • 7 0 0 ° C 25%v CO2 0 6 0 0 ° C 25%v CCh O 5 5 0 ° C 5 0 % v C O 2 , A A A A A A * A A ' ^ ^ o o ^ o o o o o o o o o * 0 0 0 0 0 0 0 0 0 0 0 6 0 5 10 Time t (s ) 15 Figure 2.3 Typical grain model plots. Early stage of carbonation for 38-45 pm Strassburg limestone particles. -5 -4 -3 -2 -1 0 ln[(PCo2-Pco2.eq)/100] Figure 2.4 Reaction order plot (Squares'. 850°C; Triangles: 600°C) for fully calcined 38-45 pm Strassburg limestone with varying C O 2 partial pressure, helium making up the balance of the gas stream. RMSE=Root mean square error. 42 Chapter 2 Determination of intrinsic rate constants of the CaO-CC>2 reaction -2 -r -2 .5 - 3 - 3 .5 -4 -4 .5 - 5 -5 .5 -6 y = 0.0182x - 3.0986 RMSE=0.072 y = 1 0078x - 1.5275 RMSE=0.21 y = 0.0281 x - 3.8331 RMSE=0.066 y = 0.9537X - 2.1882 RSME=0.039 -3 -2 -1 ln[(PCO2-PcO2,eq)/100] Figure 2.5 Reaction order plot (Squares: at 850°C; Triangles at 600°C) for fully calcined 38-45 pm Arctic dolomite with varying C O 2 partial pressure, helium making up the balance of the gas stream. RMSE= Root mean square error. • E x p e r i m e n t ^ F i t t i n g r e s u l t s 3.5 h o -4.5 -5.5 5\ 0.0008 0.001 0.0012 0.0014 1/T (1/K) Figure 2.6 Arrhenius plot for carbonation reaction with 38-45 pm Strassburg limestone particles. 43 Chapter 2 Determination of intrinsic rate constants of the CaO-C02 reaction -2.5 -3 -3.5 -4 -4.5 X Experiment • Fitting results 0.0007 0.0009 0.0011 1/T (1/K) 0.0013 Figure 2.7 Arrhenius plot for carbonation reaction with 38-45 pm Arctic dolomite particles. 0.7 0.6 * 0.4 i 0.3 O X 0.2 0.1 0 o o o o o o o o o o o o o o .CaO conversion or X O O l-(l-X) 1/3 o o ^^^Ivlax. slope: y = 0.0365x- 0.1472 10 20 Time (s) 30 40 Figure 2.8 Conversion vs time for typical kinetic run showing induction period and how the initial rates were obtained based on the maximum slopes. Strassburg limestone, 38-45 pm, 800 kPa, carbonation with 100% C 0 2 , 690°C. 44 Chapter 2 Determination of intrinsic rate constants of the CaO-CQ2 reaction 0.8 X 0.6 O O O O O O O O O ^CaO conversion or X O O O l-(l-X) 1/3 O X 0.4 0.2 o o o o o Max. slope: y = 0.0377x - 0.2546 10 . 20 Time (s) 30 40 Figure 2.9 Conversion vs time for typical kinetic run showing induction period and how the initial rates were obtained based on the maximum slopes. Arctic dolomite, 38-45 um, 0.8 MPa, carbonation with 100% C 0 2 at 764°C. -3.5 -4.5 -5 -5.5 A A A" A A A 101 kPaPTGA 800 kPa PTGA 101 kPa A T G A A A A A 0.0008 0.0009 0.001 0.0011 0.0012 0.0013 0.0014 1/T(1/K) Figure 2.10 Arrhenius plot comparing PTGA runs (at a total pressure of 0.8 MPa, with 100% C0 2 ) with ATGA runs for 38-45 um Strassburg limestone. 45 Chapter 2 Determination of intrinsic rate constants of the CaO-C02 reaction 0.0008 0.0009 0.001 0.0011 1/T (1/K) 0.0012 Figure 2.11 Arrhenius plot. Comparison of PTGA runs (at total pressure of 0.8 MPa, with 100% C0 2 ) with A T G A runs for 38-45 pm Arctic dolomite. -3 r ^-3.5 -o 'fs -4 a o . •e-4 5 o ^ -5 -5.5 -6 -5001.5x +0.9824 0.0005 0.0007 0.0009 0.0011 0.0013 0.0015 1/T (1/K) Figure 2.12 Arrhenius plot for CaO-C0 2 reaction based on calcination data of Borgwardt (1985). 46 Chapter 2 Determination of intrinsic rate constants of the CaO-CC>2 reaction For calcination: 200±13 kJ/mol (Borgwardt (1985) 1 mole CaC0 3 For carbonation 29 ±4 kJ/mol for limestone 24 ± 6 kJ/mol for dolomite 1 mole CaO + 1 mole C 0 2 Reaction heat: 168 kJ/mol by thermodynamic analysis Figure 2.13 Illustration of energy levels for CaO+C02<=> CaC03. 47 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application C H A P T E R 3 A DISCRETE-PORE-SIZE-DISTRIBUTION BASED GAS-SOLID M O D E L AND ITS APPLICATION ON T H E CaO+C0 2 R E A C T I O N A version of this chapter was prepared for publication in Chemical Engineering Science. Authors: Sun, P., Grace, J. R., Lim, C. J. and Anthony, E. J. This manuscript will be submitted soon. 3.1 Introduction The CaO-C02 (carbonation) reaction is of growing interest because of its potential usefulness in C O 2 removal in such industrial systems as steam reformers, gasifiers and fluidized bed combustors (Shimizu et al. 1999; Ortiz, A. L.; Harrison, 2001; Lin et al., 2002; Abanades et al., 2003; Johnsen et al. 2006). In a study of intrinsic kinetics for this reaction, Chapter 2 concluded that the carbonation reaction is first order when the C O 2 partial pressure driving force is less than a critical value, 10 kPa. Above this value, it becomes a zero-order reaction. The activation energies are 29 + 4 kJ/mol and 24 + 6 kJ/mol for surface reaction for the limestone and dolomite investigated. The CaO- C O 2 reaction has been found to consist of two stages, a fast stage followed by an extremely slow stage (Barker, 1973; Bhatia and Perlmutter, 1983; Abanades and Alvarez, 2003; Alvarez and Abanades, 2005a; Alvarez and Abanades, 2005b). As the reaction produces an expanded solid product, the slow stage is believed to be controlled by product-layer diffusion. The sudden change from a fast to a slow stage of reaction is of interest with respect to lime-based sorbents for C O 2 removal. Barker (1973) and Alvarez and Abanades (2005b) attributed the sharp transition to a critical thickness of product layer. Findings in Chapter 4 and Abanades and coworkers (Abanades, 2002; Abanades and Alvarez, 2003) suggest that pore size distribution plays a crucial role for the C a O - C 0 2 reaction. 48 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application When the pore size distribution changes during calcination/carbonation cycles, the reactivity of the sorbent is altered accordingly. In this chapter we consider a mechanistic model based on distributed pore size to describe the CaO-CCh reaction with pore structural evolution. Examples of applications of a descriptive gas-solid model to the CaO-CCh reaction are rare in the literature. Bhatia and Perlmutter (1983) employed a random pore model to obtain kinetic data to fit their experimental data at initial points and during the slow stage of reaction. However the ability of the model to describe the reaction history was not reported. The current work formulates a new gas-solid model based on discrete pore size distribution measurement in order to obtain the effective internal diffusivity, to fit experimental data and to explain some important issues concerning the reaction. We apply intrinsic kinetic data from Chapter 2. Before formulating the model, a brief review of previous gas-solidkmodels and their applicability to the carbonation reaction is in order. The grain model, initiated by Szekely et al. (1976), has been applied widely and successfully, especially in modeling CaO sulphation. Some major subsequent improvements have been made, e.g. the introduction of various grain size distributions by Szekely and Proposter (1975), the expanding grain model proposed by Georgakis et al. (1979) and separation of micropores from macropores by incorporating sub-grains by Dam-Johansen et al. (1991). However, grain models have several disadvantages (Gavalas, 1980; Bhatia and Perlmutter, 1980; Sahimi et al. 1990): For example, the total internal surface area monotonically increases even when the porosity approaches zero. For this study, it is preferable to use a pore model to directly deal with pore evolution. Pore models were initiated by Peterson (1957). His model for the first time considered two dimensional overlapping of pores, but was later superseded by 3-dimensional random pore models. Hoshimoto and Silveston (1976) also proposed a pore model with overlapping and 49 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application distributed pores. However, they introduced a number of fitting parameters, restricting its applicability. Christman and Edgar (1983) proposed a distributed pore size model, which showed a reasonably good ability to predict reactions with expanding product. However, they ignored pore overlapping. Simons and co-workers (Simons and Filson, 1979; Simons and Garman, 1986; Simons et al. 1987) assumed that pores in a particle are distributed such that the widest pores are at the particle surface, while the interior pores decrease in diameter. This model, called a "pore trees model", started with a good concept of distributed pores, but has been criticized as being oversimplified with respect to the pore geometry evolution (Sahimi etal., 1990). In pore models, the issue of overlapping of random pores is important. Detailed analyses have been independently discussed by Bhatia and Perlmutter (1980, 1981a) and Gavalas (1980). Bhatia and Perlmutter (1980) applied the overlap theory originally proposed by Avrami (1940), whereas Gavalas formulated his model statistically using a Poisson pore distribution. Later Bhatia and Perlmutter (1985) pointed out a defect resulting from the thin-product-layer assumption and made appropriate corrections, but at the cost of adding further complexity. Bhatia (1985) further extended the random pore model by incorporating distributed pore sizes. Although the random pore model has been improved based on rigorous derivations, the disadvantages of its complexity and the necessity to obtain information on the structural parameters are obvious when simultaneously considering distributed pore sizes and the build-up of product layer. In this chapter, we develop a simple mechanistic pore model based on distributed-pores with the aid of existing pore overlapping theory. Pore size distribution data based on the mercury intrusion method are applied as input to the model. 50 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application 3.2 Experimental details Strassburg limestone and Arctic dolomite particles of diameter 38-45 pm were used for this study. Their chemical analyses are provided in Table 1.2. Three fixed-bed thermogravimetric analyzers were used in this study, two operating at atmospheric pressures and one at pressurized conditions. An atmospheric SHTMADZU TA60 system (ATGA) was used for kinetic study purposes with the C O 2 partial pressure, Pc.02, varying from zero to 101 kPa. A pressurized thermogravimetric analyser (PTGA) was used to provide additional high-pressure data to test the effect of high P c o 2 on the reaction. Details and test descriptions are provided in Chapter 2. The third apparatus was an atmosphenc pressure thermogravimetric reactor (ATGR), which allows operation with larger sample sizes. It was used to generate samples for pore size distribution measurement. Details of this apparatus are given by Laursen et al. (2001). Heating from room temperature to the pre-set calcination temperature (850°C) in all three reactors was in pure C O 2 to prevent C a C 0 3 decomposition. As all calcinations were conducted isothermally, heating rates of the reactors have no effect of the pore structure of calcines. Thus the thermal history of the calcines from the ATGR and the other reactors were the same so that the pore size distribution of the samples from the ATGR are the same as for samples from the other two TGAs. Pore size distribution of samples were determined by a Micromiritics 9300 Poresizer with maximum pressure -197 MPa, corresponding to a 5.8 nm pore size according to the Washburn equation (Lowell and Shields, 1991). Specific surface area was measured by the multiple point BET method in a Micromeritics ASPS 2010 analyser. 51 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application 3.3 Model development 3.3.1 Pore overlap Figure 3.1 illustrates a simplified two-dimensional two-pore system. The two pores, with radii and R., and finite length /. and lj normal to the plane, overlap as indicated. The difference of the total volume from the summation of pore volume for the two circular pores TER,2/,. + nRj2lj lies in the overlapped part. In a real system, the relation between total pore volume and summation of pore volume for all cylindrical pores needs to be considered through probability analysis. Both reactant and product surfaces are developing at any time, all evolving from a group with the same initial pore radius/?,-0(i =1, N). The volumes enclosed by either the reactant surface or the product surfaces are not necessarily parallel and can spatially overlap. For pore overlappings, both Gavalas (1980) and Bhatia and Perlmutter (1980) used the same correlation, i.e., in 1 m 3 of total space, the increment in the volume enclosed by the overlapped pore system is only the fraction of the growth in the non-overlapped system, dV = (\-V)dVE (3.1) so that, after integration, V = \-exp(-VE) (3.2) where V is the true pore vdlume per unit particle space (m3/m3), whereas VE is the pore volume (m3/m3) obtained by adding the volumes of all the independently developing voids, with no consideration of overlap, e.g. Vj = nR^f for a pore of radius Rt. 52 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application 3.3.2 Model description The basic assumptions for the current model are as follows. 1. The reaction proceeds isothermally. 2. Diffusion in pores and external mass transfer are not limiting factors compared to surface reaction and solid-state diffusion in product layers. 3. The product layer CaC03 is non-porous once formed. Therefore the reaction proceeds through solid-state diffusion crossing the product layer. 4. Pores are circular cylindrical with a total length /. (m/m3). 5. All intraparticle pores are available for carbonation. There is no premature pore-mouth blocking. Figure 3.2 shows the evolution of pore structure for a reaction with swelling solid product for two representative pores of initial Rj0 and Rj0. In a real pore system the assembly of pores is represented by a set of discrete pores with initial pore radii Rl 0 (i=l to N). As the reaction proceeds, the reaction fronts migrate outwards, whereas product boundaries move inwards. Initially, the reaction front is the same as the product surface because there is no product. For cell i, during reaction, the pore of initial radius7?, 0 evolves concentrically. After time t, the pore of initial radius Ri0 has developed into one of reactant radius R. r and product radius Ri , as illustrated in Figure 2. Cylindrical voids enclosed by the product surface form the new pore volume available for further reaction. As the reaction proceeds, the overlap between adjacent surfaces grows. Although the solid product migrates inwards, the reaction fronts may also overlap with each other, especially in a 3-dimensional randomly distributed pore system. As the volume enclosed by product shrinks, VE becomes smaller, and V closely approaches VE." 53 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application Equation (3.2) generally applies both to the volume enclosed by the reaction front, and to that enclosed by the solid product surface, or pore volume, e.g., in Figure 3.2, the volume enclosed by the reaction front is that enclosed by Rt r or Rj r , and pore volume by 7^  p or Rj . The total non-overlapping specific volume enclosed by reactant surface denoted by VirE is obtained by adding all pores enclosed by reactant surface, i.e. ^ = Z ^ . , B = Z ^ V A - (3.3) i=i 1=1 Similarly the total pore volume by product surface is given by, 1=1 1=1 The total length can be determined from the fractional pore volume, i.e. I, =VUQEl{nR\rfi) (3.5) On the other hand, the counterpart variables with consideration of overlap are Vr - ^ V i r =total volume enclosed by the reactant surface, and Vp = ^ Vjp =total volume enclosed by the product 1=1 surface. Applying equation (3.2) leads to, JV iV Vr=l-exp(-VrE) so that £ V i r = 1 - e x p ( - ^ V i r E ) (3.6) ;=i i=i JV N Vp=l-exp(-VpE) so that £ V i p = 1 - e x p ( - £ V i p E ) (3.7) i= i 54 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application It needs to be noted that, equations (3.6) and (3.7) are applicable to both.total volumes enclosed by reactant and product surface. However, this is not exactly true for each discrete pore, i.e., for pore.fi;, ^ * l - e x p ( - ^ r £ ) (3.8) Vip*l-exp(-VipE) (3.9) However, equation (3.1) is applicable to discrete pores, dVr = (1 - Vr )dVrE so that Vir ) = (\-Vr )dj^ VirE (3.10) i=\ i=\ dVp =(l-Vp)dVpE sothrtd&Vip) = (l-Vp)dfiV^E (3.11) i=] Considering that each group of pores randomly overlaps with others, it is reasonable to assume that the ratio of the overlapped volume to the non-overlapped one, as stipulated in equations (3.10) and (3.11), is applicable to each discrete pore, i.e. dVir=(l-Vr)dVirE (3.12) dV,,P=(l-Vp)dVipM (3.13) Note that the discrete variables VirE, Vi. E relate to the radii by equation (3.3) and (3.4), or r,r.B = *R,A (3.14) The specific surface areas with overlap is given by sir=dVir/dRir (3.16) for the reactant front. The area for the product surface, or the true pore surface area, can be estimated in a similar manner, leading to, 55 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application sUp=dV,p/dRip (3.17) The current model is intended to trace the evolution of the radii of both reactant surface and pore surface. Although cylindrical pores overlap with one another according to equations (3.12) and (3.13), the radii are only determined by the non-overlapping volumes r £ .and VipE using equations (3.14) and (3.15). 3.3.3 Rate expressions The material balance for the solid reactant is % ^ ^ = ^ c O T - W " ~ V (3-18) dt M dt From Baker (1962), P c o 2 , e q = 1 0 ^ ™ 079> (kPa) (3.19) Applying equation (3.16) into equation (3.18) leads to ^ ~ ^ - = ks(Pco2-Pco2,q),,; (3.20) The gaseous reactant concentration driving force at the reaction front is determined by applying a psuedo-steady state mass balance across the product layer, i.e. - ^ - [ £ ) ff.^21] = 0 (3.21) The psuedo-steady-state assumption has been critically examined by Bischoff (1963). The significance of this assumption is that the interface remains stationary at any time t, so that equation (3.21) can be integrated as an ordinary differential equation. The boundary conditions are. At product surface R= RUp , CC02 = Cc02 0 (3.22) 56 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application f3C At reactant surface R= Rir, ks(PC02 -PC02eq )n=-Dp (3.23) Att=0 Rip = Rir = R.fi and CC02 -CC02eq = C C O 2 0 -CC02eq (3.24) The partial pressure of CO2 is related to the CO2 concentration by Q 0 2 = — ( 3 - 2 5 ) C02 R T For limestones, kinetic results from Chapter 2 apply to equation (3.20) and (3.23), For n=0, ks = 1.67 x 10"3 exp(—) , E = 29±4kJ with PC02-PCo2,eq > 10 kPa (3.26) RtT m s Forn=l, ks =1.67x10-" exp(^—) ™° E = 29 ± 4kJ with P c o -/> <10 kPa (3.27) Rtl m skPa In the case of dolomites, For n=0, ks = 1.04x 10 - 3 exp(—) E = 24 ± 6kJ with Pco2-Pco2eq >10 kPa (3.28) RT m s tttol F o r n = l , ^ = 1 . 0 4 x l 0 - 4 e x p ( - — ) — — £ = 24 ± 6AJ with Pco - Pco eq<l0 kPa (3.29) RtT m skPa Measured specific surface areas S'0 were 29 m2/g for Strassburg calcine and 48 m2/g for Arctic calcine, or 31 m2/g for the CaO part in the dolomite (See Appendix II for further discussion on dolomites). At reaction front, if Pco2-Pco2,eq <10 kPa, when the reaction order is unity, integration of equations (3.20) and (3.21), together with boundary conditions, leads to C r , - C e = , C ° ~ C e " (3.30) 57 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application When Pco2-Pco2,eq >10 kPa, the reaction order is zero, and integration gives C r , , = C 0 - ^ , > ( ^ - ) (3.31) 3.3.4 Reaction front, pore evolution and conversions by volume balance The relationship between the pore surface and reaction surface can be further related through a volume balance. For the i"1 group of pores, Vip=ZVl0-(Z-l)Vir (3.32) where, Z is the molar volume ratio of CaCC>3 to CaO, about 2.17. For a dolomite, one must also consider the fraction of MgO. In this case, the molar volume ratio Z' is adopted, as discussed i n Appendix II. To trace the evolution of product radius, equation (3.32) is rearranged to give dV,=(\-Z)AVir (3.33) Furtherapplyingd^ = ( l -r p ) dVipE and dVir = (l-Vr) dVirE gives dVipE=^(l-Z)dV,rE (3.34) p Note that, for dolomites, Z needs to be replaced by Z'. Since dVupE^d(7diRip2) = 27diRipdRip and dVirE=d(7dlR,2) = 27d1RirdRir, equation (3.34) leads to l-V Rir dRip=(\-Z) - ^ d R , , (3.35) p l-V R p >,p Combining equation (3.20) with equation (3.35) gives dR. M l-V R ^ = ^ -Z)——~-^-ks(PC02-PC02eqy (3.36) 58 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application Equations (3.20) and (3.36) are integrated to calculate surface positions as a function of reaction time with the initial condition R\. = R,R =RJO- When integrating, it should be noted that the reaction order is variable (1st or zero order) depending on the local driving force as shown in equations (3.26)-(3.29). As the reaction causes a swelling, the radius of product surface monotonically decreases as the reaction proceeds. Once the pore radius becomes zero, there is no space for further reaction, so the reaction stops in that pore, i.e. d ^ R i ^ = 0 when R,. = 0 (3.37) dt '•" V ' The initial pore properties utilized here come directly from experimental measurements. Detailed discussion regarding the relationship between measured values and the initial conditions appears i n Appendix ITJ. Initially, K, = V , , P = V,,r,B = r,,.* = i&iSh = V,O,E (3-38) The initial specific pore volume, Vi0E (m3/m3) relates to measured pore size distribution (see Appendix HI). The overall conversion at any time is expressed as, \-V X = \ r- (3.39) 1 v o JV where Vr is determined by equation (3.6) and V0 = 0 i=i 3.3.5 Algorithm The measured pore size distribution data were divided into 29 discrete pore size groups (i=l to 29) for both sorbents. Detailed discussion about initial pore s i z e distribution input can be 59 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application found in Appendix HI. Other input data include measured kinetic data (equations (3.26-29) for the two sorbents), reaction temperature, bulk C0 2 concentration, molar volume ratio Z or Z ' and time step length. At each time step for each group of pores, the C O 2 concentrations at the reaction front are obtained from equation (3.30). After comparing the converted C O 2 concentrations (in kPa) with critical partial pressure (10 kPa), it is decided whether equation (3.31) is needed to recalculate concentrations. The reaction order and kinetic data may then also need to be revised. Equations (3.20) and (3.36) for the reaction front and pore surface are then solved using 4 t h order Runge-Kutta integration. Once the conditions for equation (3.37) are satisfied for a certain group of pores, the reaction stops inside this group of pores. After finishing a time step for all discrete pores, the volumes enclosed by reactant and product surface, are calculated by means of equations (3.14) and (3.15) for non-overlapped volumes for each discrete pore, equations (3.3) and (3.4) for non-overlapped total volume, and equations (3.6) and (3.7) for total volumes with overlap. Finally the conversion is estimated based on equation (3.39). A Fortran program employed for this calculation is attached in Appendix IV. In this model, the solid-state diffusivity crossing product layer, D p , is the only fitting parameter. The A T G A conversion data obtained at different temperatures with 80%v C0 2 for Strassburg limestone and 100% C0 2 for Arctic dolomite were used to fit this parameter. The best fit is found for each case by least square curve regression. 3.4 Results and discussion The measured pore size distributions for both sorbents are shown in Figures 3.3 and 3.4. Both 38-45 and 212-250 um particles were used in the measurements. For Strassburg limestone, the larger particles show much larger pore volume for sizes > 610 nm; the major difference 60 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application between the samples with different particle size, but with the same thermal history, implies that the pore volume measured for sizes > 610 nm belongs to interparticle voids. Thus only the pore volume for pores <610 nm can be used as input for this model. Similar analysis for Arctic dolomite calcine in Figure 3.4 shows the actual intraparticle pores reside within 360 nm. The smaller division line for Arctic dolomite is probably because some finer particles were generated during calcination due to the weak mechanical strength of dolomitic samples. Comparison of Figures 3.3 and 3.4 shows that the Strassburg calcines had broader pore distributions of a bimodal shape. The volume in 300-610 nm range is probably due to sintering during calcination that shifts the pore volume through a lattice diffusion mechanism (German, 1996). Fitting results are shown in Figure 3.5 and 3.6 for both sorbents. Wider deviation is found for lower temperature runs with Arctic dolomite, e.g. 570-700°C, but reasonably good agreement is achieved for other tests. Compared to Strassburg limestone, the Arctic dolomite generally achieved higher calcium utilizations as shown by others (Dobner et al. 1977; Silaban et al. 1996). The reason for the different utilizations of the two sorbents can be reasonably explained by the current model based on structural difference. As discussed in Appendix II, the different effective molar volume ratio for Z (for limestones) and Z' (for dolomites) indicates that the enhanced occurrence of MgO relative to CaO in a dolomite ensures that the sorbent has enough pore volume to accommodate carbonate product. However, the wider deviation between predictions and experiments for Arctic dolomite at lower temperatures still indicates a deficiency in modeling the mixture structure in the product layer of the dolomite. One possibility is that the product layer zone is not totally non-porous because of the presence of the MgO component. The fitted diffusivity by taking slopes from Arrhenius plots (or linear fitting) in Figure 3.6 and 3.7 is used to find the activation energy, (see Appendix I for discussions on linear fitting and nonlinear fitting) 61 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application — F kJ D = 0.27exp( ) m2/s, £ = 215 for Strassburg limestone, 500°C< T <850°C (3.40) F Rt'T mol — F kJ D =0 .00083exp( ) m2/s, £ = 187 for Arctic dolomite, 570°C< T <850°C (3.41) p RtT mol The activation energies are found to be 215 kJ/mol for Strassburg limestone and 187 kJ/mol for Arctic dolomite. These values are reasonably close to, or slightly higher than, the 175 kJ/mol for temperatures higher than 515°C reported by Bhatia and Perlmutter (1983). A slightly higher diffusivity 238 kJ/mol was reported by Mess et al. (1999) who operated their TGA in the 550-1200°C range with non-porous CaO. As pointed out by Mess et al. (1999) and Bhatia and Perlmutter (1983), the high magnitude of activation energy is due to the solid-state diffusivity. As lattice defects due to the presence of impurities are known to enhance solid-state diffusion, the value is still lower than lattice diffusivities in single crystals of calcite (352 kJ/mol) (Mess et al., 1999). The discrepancies in measured diffusivities may also relate to lattice defects due to different impurity contents in the sorbents. It should be noted that the diffusivities obtained in Equations (3.40) and (3.41) should be referred as to effective diffusivity because it presumably combines the diffusivity in both micropores of the product layer and in the solid product lattice. The temperature could also have a variety of effects on solid-state diffusion, i.e. affecting crystal structure or increasing diffusivity by causing lattice defects. Activation energies obtained for some other typical product-layer-diffusion-controlled gas-solid reactions (sulphation, direct sulphation, sulfidation) are also listed in Table 3.1 for comparisons. Note that all the reactions are similar in that they are controlled by product later diffusion at the later stage of the reaction and that different type of models have been applied to derive the data reported in this table. 62 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application Another feature of the CaO-C0 2 reaction is that increasing PCo2 will not greatly change the conversion-vs-time profile when the bulk C 0 2 concentration is appreciably higher than the equilibrium C 0 2 concentration. This has also been pointed out by Bhatia and Perlmutter (1983), but no analytical explanation has been given for the lack of dependence on P c o 2 Table 3.1 Comparison of activation energies for effective diffusivity in reactions between CaO or CaCQ 3 and gases of interest in this work Reaction Activation energy (kJ/mol) Reference Reaction Activation energy (kJ/mol) Reference CaO carbonation 215 (limestone) 187 (dolomite) This work CaO sulphation 153 Borgwardt and Bruce (1986) CaO carbonation 175 (>515°C) (limestone) Bhatia and Perlmutter (1983) CaO sulphation 137 Borgwardt et al. (1987) CaO carbonation 238 (limestone) Mess et al. (1999) CaO sulphation 120 Bhatia and Perlmutter (1981b) CaC0 3 sulphation 146 Hajaligol et al., (1988) CaO sulfidation 154 Attar and Dupuis, (1979) CaC0 3 sulphation 125-144 Qiu and Lindqvist, (2000) CaO sulfidation 130 Borgwardt et al., (1984) The A T G A and PTGA experimental results are shown in Figure 3.9 and 3.10. Results for two typical temperatures, 600°C and 850°C, with Strassburg limestone are shown in Figures 3.9a and 3.9b. Only runs with C 0 2 partial pressure driving force <10 kPa show either lower conversions or appreciably slower carbonation. Further increase of PCo2 well beyond Pco2,eq cannot enhance carbonation rate and final conversions appreciably. Figure 3.9c compares the PTGA results, where there is much higher driving force, with the atmospheric runs at different temperatures. Comparing the PTGA and ATGA runs at 850°C, similar conclusions can be reached, with a lack of dependence on Pco2. Similar plots for Arctic dolomite appear in Figure 3.10. The trends are similar to those for Strassburg limestone, but with higher conversions. PTGA results in Figure 3.10c do not show appreciable improvement in either the carbonation history or final conversions for the carbonation. Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application The current model is next used to predict the removal history with varying Pco2,eq. The results are shown in Figure 3.11 for the limestone and Figure 3.12 for the dolomite. The dotted points are used to distinguish close prediction results, with experimental data at 600°C and 850°C also plotted for comparison. The predictions confirm that increasing Pco2 beyond the equilibrium partial pressure has little effect on the capture. Only runs at very low driving force, e.g. driving force of 7 kPa at 600°C, show appreciably slower carbonation. In Figure 3.13 and 3.14, the C 0 2 concentrations at the reaction front for pores of initial measured pore size 30 nm are traced in order to explain the lack of dependence of Pco2 Runs with different bulk C O 2 concentrations are compared. The dashed horizontal line marks the converted critical CO2 partial pressure driving force for reaction change (10 kPa). The plots show that there are generally two stages of reactions. During the first, (Pco2-Pco2, eq) at the reaction front is much higher than the critical driving force. At this stage, as shown in Chapter 2 and in equations (3.26)-(3.29) for the two sorbents, the local surface reaction lies in the zero-order reaction zone, so there is no influence of C O 2 concentration. With the quick build-up of the product layers, the slower stage of carbonation quickly becomes dominant. The CO2 concentrations at the reaction fronts drop sharply, so the reaction lies in the first-order reaction zone. At this stage, the reaction is very slow compared to the first stage, presumably because of the controlling product-layer-diffusion step in the overall reaction. As seen in Figure 3.13 and 3.14, at this stage the CO2 concentrations at the reaction.fronts become very close for runs with different bulk concentrations. During the very early stage of carbonation, the CO2 driving force at the reaction front usually exceeds the critical driving force (10 kPa), but the changing order of the CaO-C02 reaction makes the reaction less dependent on Pco2i At the later stage the product layer diffusion becomes rate-controlling, the difference in the bulk C 0 2 concentrations is much reduced causing the C 0 2 driving force at the reaction front to become very low. 64 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application It is also of interest to investigate the reason for the sudden shift of the carbonation reaction from a fast stage of reaction to a slow stage. Barker (1973) and Alvarez and Abanades (2005b) attribute this sharp turn to a critical product layer thickness. In the current work, typical experimental data clearly indicate a sharp turn at the transition point (Figure 3.5 for the 600°C Strassburg limestone run with 80%v C O 2 ) . This abrupt transition occurred at -100 s. However, for similar conditions (100%v C O 2 ) Arctic dolomite shown in Figure 3.6 fails to show a very sharp turn between the initial fast breakthrough and the level-off slow stage. The pore surface (product surface) represented by Rt monotonically decreases as shown by equation (3.36) due to the assumption of unchanging total pore length. The invariable total pore length assumption, adopted in most pore models (Gavalas, 1980; Bhatia and Perlmutter, 1980; Sahimi et al., 1990) is convenient for calculating pore volume. In reality, the pore-filling by swelling product might actually occur in a three-dimensional highly overlapped pore network, and the total pore length may also be reduced. Evidence for this can be found in work by Gullett and Bruce (1987) who measured pore size distribution at different stages of the CaO-SCh reaction. They found reduced pore volume at each pore size, but did not find a shift in pore diameter to smaller values. Based on these considerations, pore size distributions estimated (by calculating VjpE) at different reaction times versus the initial pore sizes, are plotted in Figure 3.15 and 3.16 for both sorbents. Figure 3.15 shows that during carbonation, pore volume shrinks for pores < 250 nm, but no appreciable pore volume change can be observed found for larger pores. At the sharp turn, i.e. -100 s for Strassburg limestone carbonated at 600°C with 80 kPa C 0 2 , nearly all the smaller pores were consumed, marking the turning point from the fast stage of carbonation to the slow stage, as smaller pores contribute more surface area. The remaining carbonation must proceed 65 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application mainly in pores of original size > 250 nm. The reaction proceeds very slowly because of the loss of most surface area. The bimodal pore size distribution of the Strassburg limestone is believed to be responsible for the two-stage pattern of the reaction. The total pore volume for small pores determines what conversions can be achieved for industrial applications where only the fast stage of reaction is of interest. In the case of the Arctic dolomite, the pore volume decreases monotonically as shown in Figure 3.16. After 500 s of reaction, some small pore volume remains, contributing to higher conversions for dolomite than for limestone. The lack of larger pores means that the carbonation of the dolomite does not pass through a sharp turn. Note that bigger particles with appreciable pore diffusion resistance, or even appreciable external diffusion limitations, can be readily integrated into the current model using an overall mass balance on the particle. Usually a partial differential equation is needed to determine local gaseous reactant concentration at the pore surface that can be further used as an input to the current model to replace the bulk gaseous reactant concentration C C O 2 0 . Finally, the current model and mechanism are not limited only to the CaO+C02 reaction, but may be applicable also to other gas-solid reaction. 3.5 Conclusions The CaO-C02 reaction is sensitive to the pore size distribution of calcines. A gas-solid model based on measured discrete-pore-size-distribution is formulated. The measured pore size distribution from mercury intrusion data is applied as input. The pore evolution is traced along with reaction. The method is more straightforward than other distributed-pore based pore models. Previous pore overlap theory is employed to deal with the evolution of reactant and product surface. For dolomite calcines, allowance is made for the smaller effective molar ratio of solid product to solid reactant than for limestones. 66 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application The model is fitted to thermogravimetric reactor data from Chapter 2, with effective diffusivity as the only fitting parameter. The results indicate that the limestone and dolomite examined have activation energies of 215 kJ/mol and dolomites 187 kJ/mol, respectively, due to solid-state lattice diffusion. Experimental data under atmospheric and pressurized conditions demonstrated that showed the reaction is appreciably affected by the CO2 partial pressure when it is significantly higher than the equilibrium CO2 partial pressure. The model predictions with varying Pco2 generally confirm the trend and show a lack of dependence on CO2 partial pressure beyond a critical CO2 partial pressure of 10 kPa. During the very early stage of carbonation, the surface reaction is mostly in the zero-order zone, whereas product layer diffusion becomes rate-controlling later, but concentrations at the reaction front have little influence for different bulk CO2 concentrations. The pore size distribution evolution provides an explanation for the sharp turn of some limestone carbonations. The generally bimodal pore size distribution, probably arising from sintering during calcination, accounts for this partem. When pores smaller then 300 nm are filled, the reaction turns into a slow reaction because of loss of most surface area and increasingly important product-layer diffusion. For the dolomite investigated, the carbonation is generally more gradual due to the difference in pore volumes and size distributions. 67 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application 3.6 Nomenclature Symbols A Measured pore volume m"7 kg C Concentration of C O 2 mol/ m3 £) Diffusivity of gaseous reactant m2/s E Carbonation activation energy kJ/mol ks Rate constant defined by equation (3.20) [ m o 1 (kPa s* m1 M Molecular weight of CaO kg/mol 1 Total pore length m/m3 n Reaction order defined in equation (3.20) nCa0 Number of molecules of CaO per unit particle volume mole/ m 3 N Maximum group size of pores Pco2 Partial pressure of C O 2 kPa Pco2,eq Equilibrium partial pressure of C O 2 kPa P* Critical partial pressure of C O 2 kPa R Pore radius m Pt Gas constant, 8.31 x 10"3 in equation (3.25) kJ/mol/K s Surface area per unit particle volume m2/m3 g' Specific surface area per unit of mass of CaO m2/g t Reaction time s T Temperature K 68 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application v Volume of void m3/ m3 V Volume enclosed by reactant or product surface per unit m3/ m3 particle volume X Conversion of CaO Z Molar volume ratio of product to reactant for limestones Z ' Molar volume ratio of product to reactant for dolomites Greek letters <J) Dimensionless term containing kinetic rate expressions -X Probability density m/ m3 8 Molar volume ratio of CaO to MgO in a dolomite s Volume fraction of voids m3/ m3 t| Molar ratio of Ca to Mg in a dolomite ps Density of CaO kg/m3 Subscripts i, j Index for discrete pore size interval, from 1 to N 0 Initial value m Measured variables, from mercury intrusion data E No consideration of overlap r Enclosed by reaction front p Enclosed by solid product surface CaC03 Calcium carbonate CaO Lime MgO Magnesium oxide 69 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application Figure 3.1 Schematic of a two-pore system with overlap. Figure 3.2 Schematic of a two-pore system after evolution. 70 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application 0.14 n 0.12 0.1 0.08 ft In / I . 212-250 |im 38-45 nm 2 0.06 1/ Intraparticle pores <360 nm 1000 Mean pore diameter (nm) Figure 3.3 Pore size distribution results for Strassburg limestone calcines. Calcination conditions: isothermal calcination at 850°C in 100% N 2 . 0.08 T M e a n pore diameter (nm) Figure 3.4 Pore size distribution results for Arctic dolomite calcines. Calcination conditions: isothermal calcination at 850°C in 100% N 2 . 71 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application 1 -j - - — — ............................ 850°C, prediction Dp=3.29x10"'' ml* ^ ™ Dp= 3 . 3 5 x 1 0 1 3 m 2 / s Reaction time (sec) Figure 3.5 Fitting results. Experiments were in ATGA, with 80%v C 0 2 , 20% N 2 balance and 35-45 um Strassburg limestone particle. 72 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application Dp=1.87xl0"1 4 m 2/s Dp=3.8xl0"'3 m 2/s 100 C, prediction c _o S3 s-< o o o O a U Dp=4.62xl0"'2 m 2/s 850 C, prediction 600 C, prediction Dp=7.4x10'15 m 2 /s 570 C, prediction Dp=3.12x1015 m 2/s • 850 °C, experiment • 800 °C, experiment A 700 °C, experiment O 600 °C, experiment • 570 °C, experiment 500 1000 Reaction time (s) 1500 2000 Figure 3.6 Fitting results. Experiments were in ATGA with 100%v C 0 2 and 35-45 Arctic dolomite particle. 73 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application Figure 3.7 Arrhenius plot for diffusivity, D p , Strassburg limestone. E=215 kJ/mol. .36 -I , —— j 0.0008 0.001 0.0012 1/T (1/K) Figure 3.8 Arrhenius plot for diffusivity, D p , Arctic dolomite. E=187 kJ/mol 74 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application 0.8 a 0.6 O 0.4 0.2 4-0 - i a o o • A 200 :-Pco2.,=' " a i-Pco2.,q=30 kPa rC02~rCO2.eq' Pc02"Pc02.eq: Pc02"Pc02.eq' =60 kPa '80 kPa = 101 kPa 400 600 Reaction time (s) 300 1000 0.8 0.6 > c o o O CO O 0.4 0.2 a a a a a a n a n a a a n n • ^ O O O <*«««»' o e -200 • PCQ2 Pc02.eq - 5 3 k P a • PC02 Pc02,eq = 3 3 k P a o Pc02 Pc02,eq = 2 0 k P a • P C O Z Pc02,eq = 1 2 k P a • Pc02 Pc02,eq = 3 k P a 400 600 Reaction time (s) 800 1000 0.8 o 2 0.6 o a o 0.4 0.2 © 0 ^ -o • A • o<*c<fX^<x*>oc<Zfr O o o <? o o o o <j o o o A A A A A A A A A < V A A A A A A A A A A ^ A A A A A A X A i • 5° D n <Ja n a a a • a i^a • o • 8 5 0 ° C , A T G A , Pco2-Pco2,„=53 k P a 7 0 0 ° C , A T G A , Pco2-Pca2,„=97 k P a 6 0 0 ° C , A T G A , Pco2-Pco2.,"101 k P a 8 4 0 ° C , P T G A , Pco2-Pcc<2.,,=758 k P a 7 6 0 " C , P T G A , Pco2-Pcoi..q=789 k P a 6 0 4 ° C , P T G A , Pco2-Pco2,.,=799 k P a O 6 0 0 ° C , P T G A , P c a 2 - P c o i , . , = 7 9 8 k P a 100 200 300 Reaction time (s) Figure 3.9 Experimental carbonation data showing effect of varying C0 2 partial pressure for 38-45 pm Strassburg limestone. Pco2., e q ts calculated from equation (3.19) (a) ATGA at 600°C (b) A T G A at 850°C (c) PTGA at 800 kPa compared with A T G A tests 75 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application • Pc02 Pc02.eq =7 k P a PC02 Pc02.eq = 16 k P a • PC02 PcQ2.,q =40 k P a  Pc02 Pc02.tq =60 k P a • Pc02 Pc02.tq =80 k P a A Pc02 Pc02.e-5 = 101 k P a 2 0 0 4 0 0 6 0 0 R e a c t i o n t i m e (s) 8 0 0 1 0 0 0 o O CO o 0 . 8 0 . 6 0 . 4 0 . 2 O A O O • • m Pc02 Pc02,eq -4 k P a m Pcoz Pc02,eq = 14 k P a O Pc02 Pc02,eq =24 k P a • P c o z Pc02,«q =33 k P a • Pc02 Pc0 2,eq =53 k P a A a o <3 0.8 H 0.6 0 .4 0 .2 2 0 0 4 0 0 6 0 0 R e a c t i o n t i m e ( s ) 8 0 0 A ft s&OOOO O m 850°C, A T G A , P c 0 ! =53 k P a A. 700'C, A T G A , P C 0 ! Pc02.e =97 k P a • 600°C, A T G A , P c 0 2 Pc02.t =101 k P a O 5 90°C, P T O A , P c 0 2 Pc02.eq -799 k P a A 765°C, P T G A , P c 0 2 =788 k P a O 717"C, P T O A , P c o ! =795 k P a a 840°C, P T G A , P c 0 2 Pc02.to =758 k P a 100 2 0 0 R e a c t i o n t ime (s) 1 0 0 0 A A 3 0 0 Figure 3.10 Experimental carbonation data showing effect of varying C 0 2 partial pressure for 35-45 um Arctic dolomite. Pcoz.eq is calculated from equation (3.19) (a) ATGA at 600°C (b) ATGA at 850°C (c) PTGA at 800 kPa compared with ATGA tests. 76 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application o i o o CO CJ o • A • • • ATGA experiment, 850°C, 101 kPa C 0 2 (upper line) ATGA experiment, 600°C, 101 kPa C 0 2 (lower line) Prediction, 850°C, 101 kPa C 0 2 Prediction, 850°C, 60 lcPa C 0 2 Prediction, 850°C, 800 kPa C 0 2 Prediction, 600°C, 101 kPaC0 2 Prediction, 600°C, 7 k P a C 0 2 Prediction, 600°C, 800 kPaCO 2 Prediction, 600°C, 50 kPa C 0 2 500 1000 Reaction time (s) 1500 2000 Figure 3.11 Predicted effect of varying C 0 2 partial pressure for 35-45 pm Strassburg limestone vs experimental data. PCo2,eq is 48 kPa for 850°C and 0.36 kPa for 600°C, as calculated from equation (3.19). 77 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application f f11**1 o > c o o O o ATGA experiment, 850°C, 101 kPa C 0 2 (upper line) ATGA experiment, 600°C, 101 kPa CO z (lower line) Prediction, 850°C, 101 kPa C0 2 Prediction, 850°C, 60 kPa C 0 2 Prediction, 850°C, 800 kPa C 0 2 Prediction, 600°C, 800 kPa C 0 2 Prediction, 600°C, 7 kPa C 0 2 Prediction, 600°C, 101 kPa C 0 2 Prediction, 600°C, 50 kPa CQ 2 500 1000 Reaction time (s) 1500 2000 Figure 3.12 Predicted effect of varying C 0 2 partial pressure for 35-45 urn Arctic dolomite vs experimental data. P C o 2 , e q is 48 kPa for 850°C and 0.36 kPa for 600°C, as .calculated from equation (3.19). 78 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application Figure 3.13 Predicted C 0 2 concentration at reaction front for Strassburg limestone at 600°C, with varying PC 02 Pco2,eq is 0.36 kPa for 600°C, calculated from equation (3.19). 0 10 20 30 Reaction time (s) Figure 3.14 Predicted C 0 2 concentration at reaction front for Arctic dolomite at 600 °C, with varying PCo2- Pco2,eq is 0.36 kPa for 600°C, calculated from equation (3.19). 79 Chapter 3 A discrete-pore-size-distribution based gas-solid model and its application 1000 Pore size (nm) Figure 3.15 Pore size distribution evolution for Strassburg calcine carbonation at600°C and 80 kPa C O 2 partial pressure. Prediction results. 1000 Pore size (nm) Figure 3.16 Pore size distribution evolution for Arctic dolomite calcine carbonation at 600°C and 100 kPa C 0 2 partial pressure. Prediction results. 80 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions C H A P T E R 4 INVESTIGATION O F E F F E C T O F SINTERING ON C Y C L I C C 0 2 C A P T U R E UNDER FLUIDIZED BED C O M B U S T I O N CONDITIONS A version of this chapter has been submitted for publication in AIChE J. The authors are P. Sun, J. R. Grace, C. J. Lim and E. J. Anthony. 4.1 Introduction Calcium-based sorbents have been receiving increasing attention as possible candidates to remove CO2 in-situ from reactors, such as steam reformers, gasifiers and water-gas shift reactors, where they can also improve hydrogen yields (Han and Harrison, 1994; Ortiz and Harrison 2001; Lin et al. 2001; 2002a; 2002b; Johnsen et al. 2006). Calcium-based sorbents can also provide CO2 removal from fluidized bed combustors (FBC) (Shimizu et al. 1999; Gupta and Fan, 2002; Abanades et al., 2003; Salvador et al. 2003; Abanades et al. 2004a; Abanades et al. 2005). Sorbent lack of reversibility, i.e., the decline of sorbent capability over multiple cycles, is a key factor affecting process economics for calcium-based CO2 sorbents (Abanades et al. 2004b). Among calcium-based sorbents, limestones are most appealing because of their high calcium-content, widespread occurrence and competitive prices. Previous work on cyclic CO2 removal with natural limestones shows that, the reversibility of limestones decays according to a similar trend for a wide variety of test conditions (Barker, 1973; Silaban and Harrison, 1995; Salvador et al. 2003; Abanades et al. 2004a; Abanades 2002; Abanades and Alvarez, 2003). CaO Sintering is believed to be the major cause of deactivation, as evidenced by the change of sorbent surface texture after multiple cycles (Barker, 1973; Abanades and Alvarez, 2003; Abanades etal. 2004a). The surface textures of cycled limestones usually show growth, of macropores, as well as shrinkage of smaller pores (Abanades and 81 Chapter 4 Investigation of effect of sintering on cyclic C02 capture under FBC conditions Alvarez, 2003). These trends are typical of an intermediate sintering stage, described by sintering theory (German, 1996) in which vacancies (or voids) generated by temperature-and-ion-sensitive lattice defects direct void sites from smaller to larger ones, whereas the mass flow is from larger to smaller pores. The importance of sintering of calcined limestones or their hydroxides has been studied because of the importance of limestones in high-temperature S O 2 capture (Borgwardt, 1989a; Borgwardt, 1989b; Silcox et al. 1989; Milne et al. 1990; Fuertes et al. 1993; Ghosh-dastidar et al. 1995; Mahuli et al. 1999). Ionic compounds, such as CaO mostly sinter due to volume diffusion (or lattice diffusion) mechanism. Borgwardt (1989a) introduced extremely mildly calcined CaO to a high-temperature oven to study sintering kinetics, confirming that the controlling sintering mechanism is through lattice diffusion. C O 2 and H 2 O have also been reported to enhance CaO sintering (Ewing et al. 1979; Beruto etai. 1984; Borgwardt, 1989b). Sintering theory has not been previously applied to explain cyclic C O 2 capture. The mechanism behind the very similar sorbent performance under a variety of test conditions also needs to be clarified. It is the aim of this work to formulate and relate the sintering mechanism to sorbent cyclic behaviour and to provide better understanding of cyclic C O 2 capture. 4.2 Experimental studies A fixed-bed thermogravimetric reactor (TGR) was used for the cyclic C O 2 capture tests. For comparison, one cyclic C 0 2 capture run was also performed on a SFDMADZU TA60 thermogravimetric analyzer (TGA). All reactors operated at atmospheric pressure. Further details of the TGR are provided elsewhere (Laursen et al. 2000; Laursen et al. 2001). Mass flow controllers were utilized to achieve desired inlet gas concentrations. A total gas flow of 1600 ml/min was maintained at the reactor inlet for both calcination and carbonation in the 82 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions TGR tests. Calcination temperatures were varied from 720 to 900°C. All carbonation reactions were carried out at 850°C with 100% CO2 for a variety of times in different cycles to allow the fast stage of carbonation to go to completion. The TGA test featured 38-45 um particles with 300 ml/min N2 flow for calcination and 300 ml/min CO2 for carbonation. Physical limitations were explored by varying operating parameters, e.g. particle and sample size, gas flow rate etc., to make sure surface reaction was the rate-controlling step during calcination. In all the TGR and TGA runs, a 100% CO2 atmosphere was maintained during heating from room temperature to the desired calcination temperature to prevent non-isothermal calcination. When suitable temperatures were reached, calcination was initiated by switching to 100% N2. As both heating and cooling rates of the TGR were fast compared to the calcination process itself, the time needed to change between calcination and carbonation operating temperatures, is considered negligible. Each calcination stage continued until no further mass decrease was observed. Previous studies have shown that different natural limestones decay very similarly (Abanades 2002; Abanades et al. 2004a). This leads to an assumption that that the sintering kinetics for natural limestones are similar, probably due to their similar lattice diffusion behaviour. Therefore, in this study, only one sorbent was used as a representative natural sorbent for the experimental and modeling study. Strassburg limestone particles of size ranges 38-45 um and 212-250 um were used for the TGR and T G A tests. Because of different bulk densities, 1 g of 212-250 um or 0.5 g of 38-45 um fresh particles was used at the start of TGR tests. In the TGA test, 4 mg 38-45 um fresh particles were used. After the TGR tests, the sorbents were collected and stored in a desiccator. Samples were transferred to a Micromeritics 9300 Poresizer for pore size distribution (PSD) measurement. The highest pressure achievable in this instrument is around 197 MPa, corresponding to a pore size of 5.8 nm according to the Washburn equation (Lowell and Shields, 1991). 83 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions A separate PSD result based on N 2 adsorption showed that a 212-335 pm Strassburg calcine after initial calcination had a peak pore volume of -30 nm, with very limited pore volume residing at pores <6 nm, beyond the range of the mercury intrusion measurement employed in this work. Under these conditions, the specific surface area obtained by mercury intrusion data should approximate the BET values (Lowell and Shields, 1991). As a check, for initial calcines under 850°C, the two repeated specific surface area measurements by mercury intrusion gave 33 and 31 m2/g, respectively, whereas the BET method gave 36 m2/g. The differences are believed to be due to the pores outside the range of mercury porosimetry, which, while low in pore volume still contribute to surface area. Since the measured specific surface area is neither the main modeling objective nor input data in the current study, the specific surface area provided by mercury porosimetry is adequate for our purposes here. 4.3 Pore size distribution The pore size distributions for calcines after different carbonation times in Figure 4.1 show that changing the carbonation time had negligible effect on the subsequent calcine structure. In these experiments, each carbonation step was allowed to proceed long enough to complete the fast stage. These results imply that lime sintering has no memory of carbonation history for the first cycle, or that carbonation makes no contribution to CaO sintering. It is believed that during calcination the re-crystallization from C a C 0 3 to CaO eliminates all structural differences caused by carbonation. As discussed below, only extended carbonation (24 h or more) was capable of filling of the intraparticle pores of highly cycled sorbents and hence able to rearrange the CaO to give a single-peaked pore structure. Since in this study, all the carbonation steps were only allowed to finish their fast portion, the calcination step is the sole stage considered for sintering. 84 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions To investigate whether the carbonation process experiences early pore blocking as is usually found for CaO sulphation, a calcined sample was carbonated long enough to allow the fast stage carbonation to be completed. The sample was split into two parts, one of which was then mildly ground into finer powder in order to expose blocked pores. Both carbonate samples were analysed using mercury intrusion. Figure 4.2 compares the pore size distributions for the ground and the non-ground carbonates, as well as their calcines prior to being carbonated. No appreciable pore volume was found to be due to the blocked pores for runs under current test conditions. EDX mapping further indicated uniform carbon distribution over the particle cross-section as shown in Figure 4.3. (Note that the more concentrated carbon layer in the background was due to carbon-rich material used as sample bonders.) Figure 4.4 shows the pore size distributions for calcines with different calcination holding times at 850°C. Results show that without cycling, longer holding times reduced the pore volume for pores <220 nm, but the pore distributions were all similar, with one major peak below <220 nm. Cycling the sorbent between carbonation and calcination for a cumulative holding time of 80 minutes resulted in a very different pore size distribution compared to the (single) run with initial calcination prolonged to 82 minutes. For the cyclic run, except for reduced volume of pores <220 nm (denoted V i pores), marked pore volume growth can be observed for pores larger than 220 nm (denoted V 2 pores). This indicates that sintering by holding calcines in N 2 is very different from that after cyclic calcination/carbonation. The cycled samples with shrinkage of the smaller pores and simultaneous growth in the larger pores are typical of intermediate stage solid-state sintering (German, 1996). Given that the presence of C O 2 is able to accelerate sintering (Ewing et al. 1979; Beruto et al. 1984; Borgwardt, 1989b), the cyclically released C 0 2 during each calcination stage must be responsible for the development of the bimodal pore size distribution. 85 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions The pore size distribution results for calcines after different numbers of cycles in Figure 4.5 give further clear evidence of the development of bimodal distributions. Generally with an increase in the number of cycles, the V i pore volume decreased, whereas the V 2 volume increased. Similar observations were reported by Abanades et al. (2004a). However, more tests on different limestones are needed to test whether the trend observed in this work applies more generally. To identify the upper limit of pore sizes in Figure 4.5, a SEM picture of a highly cycled calcine in Figure 4.6 shows that no pores larger than 1 pm can be observed on the surface of the cycled sample. Thus the inflection point (at -610 nm) in the range of 500-1000 nm in Figure 4.5 should mark the limiting size for the largest intraparticle pores. Pores larger than 610 nm are believed to be interparticle voids. Figure 4.5 also shows that there are clear divisions between the two pore size ranges, i.e. pores <220 nm and larger ones of 220-610 nm. 220 nm functions as a division line in this study, but it can also be applied to the work of Alvarez and Abanades (2005) and Fennell et al. (2007). The smaller pores are presumably due to CO2being driving off during calcination, whereas the larger ones are due to sintering shifting vacancies from smaller pores to larger ones, driven by vacancy gradients (German, 1996). 4.4 Model development 4.4.1 Pore evolution during cyclic calcination/carbonation Based on the above observations, a model for pore evolution during the cyclic operation is constructed. The pore evolution during cycling is believed to be responsible for the apparent decrease in CaO utilization. In the carbonation steps, a much slower reaction occurs when pores smaller than a critical pore diameter (around 220 nm in this work) are filled. This phenomenon has been reported elsewhere (Barker, 1973; Bhatia and Perlmutter, 1983; Abanades, 2002; 86 Chapter 4 Investigation of effect of sintering on cyclic CO2 capture under FBC conditions Abanades and Alvarez, 2003). To help clarify this for cyclic operation, the relationship between CaO utilization and available pore volume is plotted in Figure 4.7, where the conversion of CaO to carbonate is based on the measured pore volume given by, X = ^ (4.1) ( l / Z - l ) ( l - * 0 ) Both the total pore volume and the pore volume below 220 nm are used in the calculation. When the predictions are compared to the experimental CaO utilization in the TGR, good agreement is found for the V i pores. This further confirms that, on recarbonation, solid C a C 0 3 product fills pores of small diameters (<220 nm). Once these pores are filled, the carbonation reaction becomes product-layer-controlled and proceeds at a much slower rate, with the carbonation product then slowly filling the V 2 pores. In other words, the pore volume of smaller-diameter pores determines the achievable extent of carbonation during the fast stage of carbonation. The specific surface area of calcines also decreases with increasing cycle number. The specific surface areas after different numbers of cycles are shown in Figure 4.8. In all of these experiments, carbonation and calcination were carried out at 850°C with 212-250 pm Strassburg limestone particles. Given that carbonation depends strongly on the smaller pores (F,) and the monotonic decreasing trend of pore volume of Vx pores and specific surface area with cycling, as well as the fact that pores of larger diameter contribute much less to surface area, the relationship between the volume of <220 nm pores and specific surface area is approximated as dS = AdVx (4.2) with S = Sg when Vl=Vg and S = Sa when V] = Va 87 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions where S and V are the initial specific surface area and specific pore volume. Subscript "g" refers to a "green state" (German, 1996), an assumed state for freshly calcined limestones with zero degree of sintering where the grains are ideally spherical. A value of S =70 m2/g is taken, the lower limit of the range (70-80 m2/g) recommended by Mai and Edgar (1989). As different values are used in different models, sensitivity analyses on this parameter are performed below. Vg is estimated from the theoretical maximum porosity sg : V2 = ^ (4.3) where sg is obtained from equation (4.1) with X=l. Sa and Va are the asymptotic specific surface area and specific pore volume when the sample is sintered for an extremely long time. Various values, from 2 to 4 m2/g, have been suggested for Sa (Mai and Edgar, 1989; Silcox et al. 1989; Milne et al. 1990; Fuertes et al. 1993; Ghosh-dastidar etal. 1995;Mahuli etal. 1999). A value of 1.4 m2/g is taken in this study as our separate BET data, using the five-point BET method for a CaO sample overnight-sintered at 1100°C, gave 1.4 m2/g. Integration of equation (4.2) gives, r,.iizM + rM w l t h A J S ' - s - * l - ^ - ° (44) A s> Based on the above findings for cyclic calcination/carbonation, a pore evolution model is described as follows. It is assumed that after calcination with no sintering, C O 2 released during the calcination produces calcines with voids <220 nm and a single-peaked pore size distribution in the green-state calcine. This is illustrated in Figure 4.9a, where Vx pores are <220 nm within a unit volume occupied by the original nonporous limestone. However, in practice, calcination always involves sintering. During the initial calcination, the sintering, under the driving force of 88 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions a surface energy gradient, directs ions to fill inter-granular space or vacancies to transfer from smaller to larger pore space. As sintering is occurring simultaneously with calcination, C O 2 mass flow could greatly enhance sintering of CaO along the path of the offgas flow. The net results of the process are shown in Figure 4.9b. The resulting pores consist of two parts: Vx pores (<220 nm) relating to the original calcination and V2 pores (>220 nm) relating to pore growth due to fast C02-catalyzed vacancy flow. During subsequent carbonation, as portrayed in Figure 4.9c, the reaction occurs in the smaller pores, and the CaO surface of the larger pores contributes little to the carbonation because of low surface area. After the fast stage of carbonation is complete, all the Vx pores (<220 nm) are filled, whereas V2 pores are largely unfilled, serving as high-effective-diffusivity transport passages for C 0 2 during the reversible reactions. During carbonation, CaO sintering will not be considered, because the sintering-related mass flow of solid-state ions in Vx is suppressed by the fast carbonation reaction, whereas for V2 the surface energy is too small to permit appreciable sintering during carbonation. The next sintering occurs during the next calcination stage. This sintering process is similar to the previous cycle, but with Vx pores being partially occupied during the last carbonation. As C O 2 is released and transported outward, the resulting pores, which should reproduce all the pore volume of Vx if there were no sintering, could split into pores <220 nm (Vx) and those >220 nm (V2) because of the sintering. The net results in Figure 4.9d, show further reduction in Vx and growth in V2 compared to the previous cycle portrayed in Figure 4.9b. The reduced pore volume in Vx is used to accommodate carbonate during the next cycle. The pore volume in V2 is additive during cycling. As Vx continues to decrease, CaO utilization decreases accordingly. Note that the 89 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions intermediate stage of sintering also involves a decrease in V2 porosity, causing a decrease in total porosity (German, 1996). But equation (4) eliminates the need to model total porosity change, and it is also found that the decrease in total particle porosity is modest (to -0.46 for samples after 15 calcination/carbonation cycles at 850°C compared to 0.54 for the theoretical maximum porosity). It is believed that particle shrinkage only becomes dominant during the later state of the intermediate and final sintering stages. Therefore shrinkage of particles is neglected when considering calcination below. 4.4.2 Reactor model for calcination Based on the measurements discussed above, most CaO sintering occurs during calcination. To describe experimental cyclic calcination histories, a fixed bed was used for the thermogravimetric reactor. The superficial velocity in the reactor was 0.1 m/s (850°C), but the actual in-bed gas velocity in the bed should be lower due to bypassing, resulting in very low particle Reynolds numbers in the packed bed. For gas-solid flows at low Reynolds numbers (<1), the Peclet number (based on particle diameter) is small (Szekely et al. 1976), and axial dispersion has to be considered. Radial dispersion is not considered in this work. The control equation is written as, dC(t,z) d2C(t,z) . .dX(t,z) / x » w I -D, V = Pcaco3 (")0 - g t) I ' (4.5) oz dz dt The final term is related to the calcination rate for CaC0 3 . In this equation, the time-derivative of C(t,z) is omitted as it is usually negligible compared to other terms (Szekely et al. 1976). The bed voidage sb was chosen as 0.5 for the randomly packed bed of uniformly sized particles. ubed is obtained by fitting the calcination time needed for the initial calcination. Once it is determined, it is assumed to apply also to the later cycles. In equation (4.5), time t is calcination 90 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions time. As discussed above, for each calcination/carbonation cycle sintering accompanies calcination. The time needed to complete each calcination stage is termed the calcination time for that cycle. The boundary conditions are, At z=0, CO, z) = 0 (4.6) dz In the reactor, the gas flowed downward from the top. At the top inlet, the C O 2 concentration is always zero. In equation (4.5), at a certain vertical position z, dX I dt describes the solid calcination rate to produce CaO at any bed location. Because calcination of limestones proceeds in a shrinking-core manner (Borgwardt, 1985; Dennis and Hayhurst, 1987), a shrinking core model (SCM) was used to express dX I dt. The SCM model was derived with the following rate law for calcination, r = kc[l-C(t,z)/Ke] (4.7) where, the equilibrium constant Ke is based on the correlation of Baker (1962), K, = 10 ( 8 3 0 8 / r + 9 0 7 9 ) / (R 1 T) (4.8) The equilibrium constant Ke has units of mol/m3 when the concentrations are expressed in mol/m3. Kinetic data for limestone calcination suggested by Borgwardt (1985) were adopted here: £ c =3.013x l0 7 xexp[ -200 / (RiT) ] (mol/m2 s) (4.9) The final expression for the SCM is then, at vertical position z, 91 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions x C(z) _ = -1 (4.10) 3A. A„ k. 5DK„ e e where fcm=mass transfer coefficient (= ShD0 /R), with the Sherwood number Sh calculated from the correlation suggested by Bird et al. (2002) for creeping flow in packed beds. Sh = 0.6(Re pxSc)U2 (4.11) The effective diffusivity, De, for C O 2 inside the porous lime layer is estimated from D.=D0xs,2 (4.12) The newly formed lime layer has a theoretical local porosity e,, assumed to equal to the theoretical maximum porosity s . For calcined nonporous limestone, sg is around 0.54 based on a pore volume balance. As the current TGR calcinations have been found to be mainly controlled by gas-film mass transfer, the intraparticle mass transfer and surface reaction are not rate-limiting. Therefore, the assumption of this constant local porosity does not lead to appreciable errors. The correlation suggested by Levenspiel (1999) for packed beds was adopted for the axial dispersion coefficient Dz. The extent of decaying carbonation with cycling is reflected in the skeleton density of carbonate, i.e., if the CaO-CaC03 conversion achieved after the nth carbonation is Xcarb(n), during the (n+lUl) calcination, the density is taken as: PcaCOl (" + 0 = P C a C O 3 , 0 X c a r b (") ( 4 1 3 ) 92 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions For the initial calcination of fresh limestone, pCaC03Q) = PCacoio- Note that the density of CaCCh inside a particle obtained from equation (4.13) is no longer the skeletal density, but the CaC03 density for a particle with unchanged particle size. Equation (4.5) was linearized and then solved numerically using a finite difference method. Appendix V shows the linearization method. A downstream scheme was used for all differentiation terms. The bed height was divided into 20 equal intervals. Macroscopic variables, specific surface area and CaO conversions were obtained by averaging over the cells as follows, W = j)&idli (4.14) 4.4.3 Macroscopic sintering during cyclic calcination/carbonation cycles Sintering occurs simultaneously with calcination. On the macroscopic scale, the CaO surface area evolves as a result of both calcination and sintering. Hence dt dt a^c*na*'on s sintering \^' ^ An empirical correlation was adopted to describe CaO surface evolution in the absence of sintering: S = SgXcarb(n)X (4.16) This gives S = S Xcarb (n) for complete calcination and zero for no calcination. For the n t h cycle, dS _ dX at calcinati0" « carb ~a7f (4-17) ds The CaO surface area reduction (—) s i n t e r i n g involves the sintering kinetics. i^-)^g=K(S-Sa)2 (4.18) with 93 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions ks = 2.45(1 +10.3/> 0.67 )exp(-29000/7T (4.19) co 2 Equation (4.19), based on surface area reduction, has been adopted in studies on CaO sintering during calcination (Mai and Edgar, 1989; Silcox et al. 1989; Milne et al. 1990; Fuertes et al. 1993; Ghosh-dastidar et al. 1995; Mahuli et al. 1999). An obvious disadvantage of this equation is that when the reaction is under kinetic control, P c o l at the interface should closely approach the bulk C O 2 concentration, zero in the current TGA test. Although calcination was completed more quickly, resulting in a shorter sintering time for each successive cycle, the cyclic performance showed similar decay behaviour as in other TGR tests (Figure 4.10), indicating that sintering is still significant during the TGA test, probably due to accelerated calcination under the TGA test conditions. Therefore, in this work, equation (4.19) is modified to include calcination rate in the sintering rate expression, giving where the activation energy of equation (4.19) is retained. Once the sintering rate constant was determined by equation (4.20), equation (4.18) was solved by a 4 t h order Runge-Kutta method with the initial condition that at the start of calcination, the CaO surface area is zero. At the end of the n t h calcination, the Vi(n) pore volume can be estimated by relating it to the specific surface area as described by equation (4.4). Vi (n) is then used to predict CaO conversions for the next carbonation based on a mass balance, where, 56 is the molecular weight of CaO, in g/mole. A M A T L A B 6.5 based program (see Appendix V I ) was employed in the overall calculation. k, =2.45(1 + a[ ]*}exp(-29000/r) (4.20) X c a r b (") = K (") X 5 6 X PcaCOXO f ( l ~ Z ) (4.21) 94 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions AAA Results and discussion Figure 4.10 shows the reversibility under various test conditions. The general trends for different cyclic runs are seen to be very similar. As described above, these tests involve different particle sizes and sample sizes for both the TGR and T G A tests. Figure 4.10 indicates somewhat better reversibility at a lower temperature, e.g. 720°C, and for a small sample size, 4 mg for the TGA run, but the trend is not strong enough to allow clear conclusions. Results from the literature (Salvador et al. 2003; Abanades and Alvarez, 2003) indicate that different natural limestones give similar reversibility. Table 4.1 summarizes the fitted results for Ubed based on the time needed for the initial calcination. The velocities in Table 4.1 are much smaller than the superficial velocity based on the cross-section area of the reactor (0.1 m/s). As a result, the particle Reynolds number is only about 0.04, and the Peclet number is only -0.03 for a Schmidt number of ~0.7, justifying the inclusion of axial dispersion in the model. Table 4.1 Fitted results for Ubed (in equation 4.5) Test No. Test conditions Ubed (m/s) 1 TGR, 850°C/212-250um 0.012 2 TGR, 900°C/212-250 urn 0.013 3 TGR, 800°C/212-250 um 0.011 4 TGR, 850°C/38-45 um 0.0105 5 TGR, 720°C/38-45 um 0.009 6 TGA, 850°C/38-45 um Not needed Once Ubed has been estimated in this manner, it is assumed to be unchanged for the following cycles. The fitted results for calcination history are shown in Figure 4.11 for three typical cycles where 212-250 pm Strassburg limestone was subjected to calcination/carbonation 95 Chapter 4 Investigation of effect of sintering on cyclic C02 capture under FBC conditions cycles at 850°C in the TGR. Although some deviation can be seen along the calcination history, the predictions of the current model generally fit the experimental data reasonably well. Values of a and b, determined by least square fitting using all experimental data from the six runs are 1.26 and 0.59, respectively. Comparisons for all runs with these values of a and b are shown in Figures 4.12 and 4.13. The predictions are seen to agree reasonably well with the experimental data. The general decay trend of calcium utilizations is predicted by the current model. As described above, the degree of sintering is affected both by the duration of the calcination stage and by the sintering rate expressed by equation (4.20). These two effects could counterbalance each other to give similar sintering for different tests. For instance, the TGA test involved fast calcination or very brief exposure under sintering conditions, but a much faster rate of C O 2 release which should greatly accelerate the sintering rate, resulting in a similar sintering degree as seen in Figure 4.13. The fitted specific surface area for the run with 212-250 um particles in the TGR is compared with measured data in Figure 4.8. The predicted specific surface areas are generally higher but with a similar decay trend. Given that the specific surface area by mercury porosimetry is smaller than that obtained by the BET method, as discussed above, the model gives reasonable predictions. Predicted times for completion of the fast stage of calcination were in reasonably good agreement with experimental data as shown in Figure 4.14. When S was varied from 60 to 80 m2/g, the model prediction was not very sensitive to S as shown in Figure 4.14. In a practical process for cyclic C O 2 capture, the sorbent would have to endure many more cycles than the number tested here. Despite losing capacity during cycling, if the highly sintered sorbent could maintain a residual C O 2 capture ability, it could be used in an affordable for C O 2 removal process. Alvarez and Abanades (2005) showed a -10% residual C O 2 capture ability in a 96 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions 100-calcination/carbonation-cycle study. In Figure 4.15, the model is extrapolated to 1000 cycles for three cases, TGR/850°C/231 pm, TGR/900°C/231 pm, TGA/850°C/41 pm, respectively. They generally give CaO conversion, 8.5-14% after 100 cycles, 4-6% after 500 cycles and 2.5-4.1% at 1000 cycles. Given the similarity for limestone reversibility, empirically regressed equations based on the extended prediction data are also shown in Figure 4.15, all with an exponential decay factor -0.5. Considering that the current model overestimated the TGA/850°C/41 pm run and underestimated the TGR/850°C/231 pm mn, the one based on TGR/900°C/231 pm test gives a reasonable estimation for long-term limestone sorbent capture. ^ ( » ) = 1 0 5 » - 0 4 9 (4.22) The current sintering model provides guidance on optimal C O 2 removal. As longer sintering time and higher calcination rate jointly control the degree of sintering, better CaO reversibility over long-time cycling may be achieved by adopting a lower degree of carbonation for each cycle. In this way the cumulative sintering time will be greatly shortened. Figure 4.16 shows the effect of carbonation time on the cyclic performance over 20 cycles. Under the given test conditions, 8 minutes is not enough to complete the fast stage of carbonation until after the 5th cycle. On the other hand, for the 3-minute run where all carbonation step lasted 3 minutes, the slower stage of carbonation did not occur over the entire 20 cycles, except for the 18th cycle that was deliberately extended to complete the fast carbonation stage. Given our finding that extent of the carbonation is determined by the pore volume of <220 nm pores, it is clear that for the same number of calcination/carbonation cycles, a shorter carbonation time retains more useful pore volume of <220 nm pores, most likely because of the reduced cumulative calcination time (Recall that each calcination step was terminated when the weight stabilized). More experimental efforts are needed to investigate the performance over >20 cycle. 97 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions Another possibility would be to use a very long carbonation time to reactivate highly cycled samples. This is because in most cases during cyclic calcination/carbonation, Vx pore volume monotonically decreases, whereas V2 increases over cycling, since V2 pores are barely involved in the fast portion of carbonation. When the carbonation is allowed to proceed for such a long time that V2 pores are gradually filled through slow product-layer-controlled carbonation, the subsequent calcination will rearrange the pore size distribution as if it were the initial calcination. Barker (1973) demonstrated the effectiveness of one-day long carbonation on cycling, and this has recently been confirmed by Salvador et al. (2003). However, the reactivated sorbents after a long saturated carbonation would again show a normal decay trend when subjected to normal cyclic C O 2 capture where the fast stage of carbonation is completed for each cycle. 4.5 Conclusions Investigation of samples obtained after different numbers of calcination/carbonation cycles showed that carbonation with the fast stage completed for each cycle has no influence on calcine pore size distribution. Cycling gives pore size distributions which differ greatly from those for one-time calcination, even if there is a similar cumulative holding time. All cycled samples showed bimodal pore size distributions, with a clear division between pores <220 nm and 220-610 nm pores. The volume of <220 nm pores decreased monotonically with cycling. A sintering model formulated based on the experimental observations relates specific surface area to the volume of <220 nm pores. This portion of pores determines the calcium utilization achievable for the next fast stage of carbonation. A packed bed model and shrinking core model were used to predict cyclic calcination history, with a modified sintering kinetic model incorporated to describe specific surface area evolution. The predictions are in reasonably good agreement with experimental data. There is little sensitivity to specific surface area, Sg. The 98 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions model reasonably explains the very similar reversibility of sorbents for the range of conditions tested. The balance between shorter cumulative sintering time and higher calcination rate can explain the similar degree of sintering. Extrapolation of the model suggests that the calcium utilization slowly decreases over cycling, eventually approaching zero, but this needs to be tested experimentally. 4.6 Nomenclature a, b Coefficients in equation (4.20) C(t, z) C0 2 concentration mol/ m3 D0 C O 2 molecular diffusivity in N 2 m2/s De Effective diffusivity of calcined product layer Dz Axial dispersion coefficient m2/s I Cell number kc Calcination rate constant mol/(s/m2) Ke Equilibrium CO2 concentration mol/m3 km Mass transfer coefficient m/s ks Rate constant for sintering g/s/m2 L Total bed height m /,• Height of the i"1 bed cell m N Total number of cells R Particle radius m 99 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions Ri G a s constant, 8 . 3 1 4 x lfj 3 in equations ( 4 . 8 ) , ( 4 . 9 ) m3 kPa/mol/K R e p Particle Reynolds number, Re p = 2Rubed Iv S Specific surface area for CaO m2/g Sa Asymptotic specific surface area for CaO, taken as 1.5 m2/g Sc Schmidt number, vl D0 Sg Specific surface area for CaO with zero sintering m2/g Sh Sherwood number, defined in equation ( 4 . 1 1 ) t Calcination or sintering time s T Calcination temperature K ubed Absolute velocity in packed bed m/s V Total specific pore volume m3/ m3 Vx Specific pore volume of pores <220 nm m3/g Vi (n ) Specific pore volume of pores <220 nm after n* calcination m3/g V2 Specific pore volume of pores >220 nm m3/g Vg Specific pore volume at green state m3/g Va Asymptotic pore volume, 0 m3/g . Conversion of calcination when C a C 0 3 is calcined, in equation X ( 4 . 5 ) x carb (n) CaO conversion after n t h carbonation z Distance from the top of bed m Z Molar volume ratio of calcine to carbonate for limestones, 0 .46 100 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions Greek letters CD Macroscopic variables m3 pore/ m3 £ 0 Pore volume available for carbonation space sb Bedvoidage £ g Theoretical maximum porosity for complete calcination, 0.54 £1 Theoretical local porosity of product layer v Kinematic viscosity of C O 2 m2/s pCa0 Skeleton density of CaO, 3.34 x 106 g/m3 Pcacm(n) Molar density of CaC0 3 at start of calcination mol/m3 Pcaco3,o Molar density of CaC03 in original limestone mol/m3 101 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions CD E 3^ o > 0.08 0.06 O CD '</) OT~~ 5 e •_- O 2 x a —' E CD i _ O 0.04 0.02 10 - • - Calc ine after 15 mins of carbonation - • — C a l c i n e after 8 mins of carbonation Calc ine after 60 mins of carbonation 1000 100 Pore s ize (nm) Figure 4.1 Pore size distribution: effect of carbonation time. Experiments. 850°C for calcination and carbonation in the TGR. 212-250 pm Strassburg particles. Calcination with 100% N 2 ; carbonation with 100% C O 2 , fast stage completed. 102 Chapter 4 Investigation of effect of sintering on cyclic CQ2 capture under FBC conditions CD E _=j o > c: . O O ) (I) 2 E — o iS x CD E CD L _ O a 0.12 0.08 0.04 - B ~ Ground carbonate —A— Carbonate " • - After initial calcination 10 100 1000 10000 Pore s ize (nm) Figure 4.2 Pore size distribution, carbonate before and after mild grinding. Same test conditions as in Figure 4.1. Figure 4.3 X-ray carbon mapping for carbonated Strassburg limestone (a) Relative position of particles; (b) Carbon distribution (Light white points represent carbon; the background contains carbon-rich material) Same test conditions as in Figure 4.1. 103 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions 10 100 1000 Pore s ize (nm) Figure 4.4 Pore size distribution: effect of calcination time or mode. Same test conditions in Figure 4.1. 0.12 10 100 ! ! 1000 Pore s ize (nm) Figure 4.5 Pore size distribution: calcines after various number of calcination/carbonation cycle. Test conditions as in Figure 4.1. 104 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions 11 : : : : : : H i 1 5 pm | Figure 4.6 S E M pictures of cycled Strassburg calcine samples after 15 cycles of 850°C calcination/carbonation cycles. Test conditions as in Figure 4.1. TJ CD co CO CD -Q £0 .8 TJ =J S O o > TJ CD £ c3 0.6 C L Q . CO "O . i § £ co 0.4 5 2 0.2 0 Trendl ine, based on total pores y = 0.3536X + 0.5162 Trendl ine: based on pores<220 nm y = 1.0091X - 0 . 0 0 1 7 • Calculation based on total pores O Calculation based on pores<220 nm 0.2 0.4 0.6 0.8 C a O convers ions by cycl ic experiments Figure 4.7 Conversion of CaO to CaC03: Experiments vs Predictions with pore volume. Experimental conditions: 850°C for calcination and carbonation in the TGR. 212-250 pm Strassburg particles. Calcination with 100% N 2 , carbonation with 100% C 0 2 , fast stage completed. 105 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions Cycle number Figure 4.8 Specific surface area after each cycle of calcinations: experimental results vs. predictions. The predictions show the sensitivity to Sg. Strassburg limestone, TGR test, 850°C calcination in 100% N 2 , 850°C carbonation in 100% C0 2 . Fast stage of carbonations completed. 106 Chapter 4 Investigation of effect of sintering on cyclic CQ2 capture under FBC conditions a. 1ST Calcination with no sintering b. After sintering, bimodal PSD c. After carbonation, smaller pores filled d. After re-calcination and sintering, pores further developed Figure 4.9 Schematic of sintering progression during cyclic calcination and carbonation. 107 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions CJ o o o C 3 a 1 0.9 0.8 0.7 0 0.6 0.5 \-0.4 0.3 0.2 0 • • T G R , 212-250 nm, calcine at 8 5 0 ° C A T G R , 38-45 urn, calcine at 8 5 0 ° C O T G R , 212-250 um, calcine at 9 0 0 ° C X T G R , 212-250 um, calcine at 8 0 0 ° C • T G R , 38-45 um, calcine at 720&C • T G A , 38-45 um, calcine at 8 5 0 ° C 6 8 10 Cycle number 12 14 16 Figure 4.10 Reversibility under different test conditions in the TGR or TGA. All with Strassburg limestone, calcination in 100% N 2 , 850°C carbonation in 100% C0 2 . Fast stage of carbonation finished for each cycle of carbonation. c o in S3 > o o 0.6 h 0.4 f 0.2 Initial calcination, experiment After the 1st carbonation,experiment After the 15th carbonation,experiment Initial calcination, prediction After the 1 st carbonation, prediction After the 15th carbonation,prediction 20 T i m e (min) 25 Figure 4.11 CaO conversion profiles for several calcination cycles: experimental results vs. predictions. 212-250 pm Strassburg limestone, TGR test, 850°C calcination in 100% N 2 , 850°C carbonation in 100% C0 2 . Fast Stage of carbonation finished for each cycle of carbonation. 108 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions 0.8 o o o O re! U 0 0.6 0.4 h 0.2 • Experiment, 212-250 nm, calcine at 850°(fl O Experiment, 212-250 um, calcine at 900°(t X Experiment, 212-250 um, calcine at 800"C Prediction, 212-250 nm, calcine at 850°(J - - - Prediction, 212-250 nm, calcine at 900°4 212-250 u.m, calcine at 800"(l Prediction, X x X •9 6 8 10 Cycle number 12 14 16 Figure 4.12 Reversibility: experimental results vs. predictions for 212-250 pm Strassburg limestone in the TGR. Calcination in 100% N 2 , 850°C carbonation in 100% C0 2 . Fast Stag of carbonations is allowed to finish for each carbonation cycle. 109 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions 0.8 o 5x6 53 I o o 90-4 U 0.2 0 A • O Experiment, 38-45 um, calcine at 850°C Experiment, 38-45 um, calcine at 720°C Experiment, TGA, 38-45 um, calcine at 850°G Prediction, 38-45 L i m , calcine at 850°C Prediction, 38-45 L i m , calcine at 720°C Prediction, TGA, 38-45 nm, calcine at 850°C 6 8 10 Cycle number 12 14 16 Figure 4.13 Reversibility: experimental results vs. predictions for 38-45 pm Strassburg limestone in the TGR or TGA. Calcination in 100% N 2 , 850°C carbonation in 100% C0 2 . Fast Stage of carbonations is finished for each carbonation cycle. 110 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions Figure 4.14 Calcination time: experimental results vs. predication. Calcinations: in 100% N : carbonation: 850°C, in 100% C O 2 , Fast Stage of carbonation is finished for each carbonation step Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions TGA, 4mg, 850°C, 41 um, fit: y=1.59x ( 0 5 l 3 4 ) TRG, 1 g, 900°C, 231 um, fit: y=1.05x (0489) TGR, 1 g, 850°C, 231 um, fit: y=0.76x(049) 200 400 600 Cycle number 800 1000 Figure 4.15 Predicted CaO utilizations for 1000 cycles. Same calculation conditions as in Figure 4.12 and 4.13 for each case. 112 Chapter 4 Investigation of effect of sintering on cyclic CO 2 capture under FBC conditions £ 0.8 T 0.6 CD c cn to O -•—' CD i _ CM O O CD 0.4 0.2 • • • O O • O 8 minute O 3 minute • F S F Carbonat ion ex tended \ until fast s tage f in ished j 0 5 ~ i — 10 15 20 Cyc le number Figure 4.16 Effect of carbonation time on cyclic CO2 capture performance: experimental results. Starting from 850 mg of 212-250 urn fresh Strassburg limestone. Calcinations: in 100% N 2 , 850°C; carbonation: 850°C, in 100% C0 2 . Carbonation time at each carbonation stage: FSF-Fast stage finished, comparing with 3-minute and 8 minute for each cycle. 113 Chapter 5 Simultaneous CO2 and S02 capture atfluidized-bed combustion temperatures CHAPTER 5 SIMULTANEOUS C0 2 AND S02 CAPTURE AT FLUIDIZED BED COMBUSITON TEMPERATURES A version of this Chapter has been published in the proceeding of the 18th International Conference on Fluidized-bed Combustion, Toronto, 2005, Paper no. FBC2005-78125. The authors are P. Sun, J. R. Grace, C. J. Lim and E. J. Anthony. 5.1 Introduction Growing concerns with respect to greenhouse gas emissions have encouraged research on C O 2 capture. Calcium-based materials have recently attracted renewed attention (Silaban and Harrison, 1995; Silaban et al., 1996; Shimizu et al., 1999; Abanades et al., 2003; Anthony and Wang, 2003) as possible sorbents for cyclic CC^-capture processes due to their potential for regeneration. In the proposed systems, calcium oxide serves as a carbon carrier passing back and forth between the combustor and the regenerator (i.e., calciner). The concept has recently received further impetus as a result of effective C O 2 capture in fluidized bed reactor tests (Salvador et al., 2003; Abanades et al., 2004a; Johnsen et al., 2006). Given that it is constrained by thermodynamics, the CCh-CaO reaction benefits by higher C O 2 partial pressure to achieve a higher extent of C O 2 removal. For a typical flue gas stream (e.g. containing 15% C O 2 volumetric fraction) from a fossil fuel-fired fluidized bed combustor, Anthony and Wang (2003) showed that a total pressure as high as 1.5xl03 kPa is thermodynamically required to capture 80% or more C O 2 at typical FBC temperatures (e.g. 850°C) using CaO. However, the requirement for high pressure can be relaxed when the combustor is allowed to operate at 800°C or lower when burning low-heating-value feedstocks, e.g. biomass or woodchips. 114 Chapter 5 Simultaneous C02 and S02 capture atfluidized-hed combustion temperatures A pressurized fluidized bed combustor with in-situ C O 2 removal by regenerable calcium-based sorbents (Abanades et al., 2003) could be a competitive fossil-fuel-fired energy system, when C O 2 emissions are regulated in the future. To achieve this goal, several important issues need to be investigated. To reduce sorbent cost and compete with other C O 2 capture technologies, the reversibility of sorbents during cyclic sorption/calcination must be excellent (Abanades et al., 2004b). Among the candidate sorbents, dolomites have been found to be less subject to loss of effectiveness than limestones (Silaban et al., 1996), likely due to MgO enhancing porosity and reducing sintering. A key issue in this C O 2 removal concept is to be able to obtain favourable simultaneous C 0 2 and S O 2 removal with calcium oxide in the combustor. Abanades et al. (2003) believed that the presence of S O 2 in a gas stream should not be a limiting factor because large quantities of calcium required for the carbonation would result in such high Ca/S ratios that virtually complete S O 2 capture should occur. However, detailed investigation into simultaneous C O 2 and S O 2 capture is required to test this hypothesis and to determine the effect of sulphur on Ca-utilization efficiency for carbonation. Furthermore, in pressurized fluidized bed combustors, there exist regions where calcination occurs (Iisa et al., 1991; Hansen et al., 1993) (with low local C O 2 concentrations). Hence investigation of simultaneous C O 2 and S O 2 capture by calcium oxide is of special interest for PFBC systems. The objective of this work is to investigate simultaneous S O 2 and C 0 2 removal with calcined limestone and dolomite in an ambient pressure reactor, but with augmented C O 2 partial pressures in order to extend the range of applicability of the results. 115 Chapter 5 Simultaneous CO2 and S02 capture atJluidized-bed combustion temperatures 5.2 Experimental details An atmospheric pressure thermogravimetric reactor (TGR) was employed in our experiments. Detailed descriptions of the system are provided elsewhere (Laursen et al., 2000; 2001). Ultra-pure C O 2 , 10% S O 2 diluted with balance N 2 , pure N 2 , and medical grade air supplied from gas cylinders were controlled by mass flow controllers to maintain the desired inlet gas concentrations. Two commercially available sorbents were investigated, Arctic dolomite and Strassburg limestone. Their chemical analyses appear in Table 1.2. Two sieve size fractions were employed: 212-250 and 500-600 urn. For the TGR tests, the temperatures were measured by K-type thermocouples located immediately below the sample holder. To gain information on the reversibility, the two sorbents were first subjected to cyclic carbonation and calcination in pure C O 2 and pure N 2 , respectively, at 850°C. (This test is designated below as the "100% C 0 2 test"). In the tests for simultaneous C O 2 and S O 2 capture, 65% and 80%> C Q 2 mole fractions, well above those corresponding to AFBC conditions, were used to simulate typical non-calcining conditions in PFBCs. S O 2 concentrations were chosen as 1600, 2900 and 4100 ppm. The oxygen volume fraction was kept at 3%. Baselines were provided by tests with 80% C O 2 balanced by 20% N 2 for the sorption step. Fixed durations of 3, 8 and 30 minutes were employed in the sorption steps. Sulphation and carbonation temperatures were set at 750, 800 and 850°C. Calcinations were conducted in a pure N 2 environment. For all runs, calcinations were deemed to be complete when no further change in mass was observed. 850 mg ± 2 mg of fresh sorbents were used in all tests. The total gas flow rate was maintained at 1600 ml/min for both the calcination and carbonation stages. 116 Chapter 5 Simultaneous C02 and S02 capture at jluidized-hed combustion temperatures The sorbents were first calcined at 850°C in a pure nitrogen stream. Then the temperature was adjusted by the temperature controller to the desired reaction temperature, and the specified gas stream was then introduced. Sulphation and carbonation proceeded until the calcination was complete. For the cyclic tests, when the carbonation/sulphation lasted for pre-set duration, the temperature controller and gas valves were simultaneously and promptly switched to the calcination conditions. 850°C and pure N 2 . Cycles were then repeated until the run was complete. 5.3 Results and Discussion: 5.3.1 Baseline runs: carbonation test with no SO2 in gas stream The performance of both sorbents in cyclic calcination/carbonation tests are shown in Figure 5.1. Temperatures for both the calcination and carbonation were 850°C. In the tests with pure C O 2 in the sorption stage, the sorption stage was initiated as soon as the mass change was observed, as illustrated in Figure 5.2, corresponding to the turning point from the fast carbonation to the much slower product-layer-diffusion-controlled stage. Cyclic C 0 2 capture experiments in 100% C O 2 environment were usually conducted with, sorption time allowing the completion of each fast stage. This sorption time is designated as Fast Stage Finished or "FSF". It typically required 3-5 minutes to detect an obvious plateau, with the shorter time for later cycles. It is believed that the micropores are filled when the fast stage carbonation is completed, whereas the slower stage is controlled by diffusion of C O 2 through the product-layer (Bhatia and Perlmutter, 1983; Abanades and Alvarez, 2003), so that the turning point from the fast to the slow stage is usually sharp and readily detectable. This test supports the earlier conclusion (Silaban and Harrison, 1995; Silaban et al., 1996; Abanades and Alvarez, 2003) that for limestone, grain size growth caused by sintering decreases the reversibility of the sorbents when there are multiple cycles of calcination/carbonation. For 117 Chapter 5 Simultaneous C02 and SO2 capture at fluidized-bed combustion temperatures dolomites, the presence of MgO significantly reduces the rate of sintering, thus retaining the pore structure and facilitating better reversibility. The baseline data with carbonation in 80% C O 2 and 20% N 2 , operating for 3- and 30-minute durations are also shown in Figure 5.1 (open symbols). For the 30-minute reaction time, the data for three cycles follow the same trend as for 100% C O 2 tests. However, for the 3-minute reaction time, there was no obvious decay in the reversibility for either sorbent, likely due to the low calcium utilization. The reaction rate during the initial stages of carbonation experienced little change with cycling. To find whether a longer period of carbonation could alter this trend, the 18th cycle in the 3-minute reaction time carbonation test was extended to complete its fast stage (extended to 12.5 minutes for this cycle). As shown in Figure 5.1, no unexpected consequences were observed for the following two (19th and 20th) cycles. 5.3.2 Simultaneous sulphation and carbonation: Effect of SO2 on CO2 capture (a) Calcination rate Typical results for simultaneous sulphation and carbonation are shown in Figure 5.3. Each cycle includes a sorption (sulphation and carbonation) stage and a calcination stage. The sulphate and carbonate contents are obtained from the sorbent mass profiles. During each cycle, the mass gain due to carbonation and sulphation could be separated when the carbonate was subsequently subjected to complete calcination assuming that none of the sulphate was lost. For a sorption time of 30 min, the sorbents usually exhibited a strong decay in calcination rates by the end of the third cycle. On the other hand, the decay in calcination rates for the 8-minute cyclic tests was appreciable only after 7-8 cycles. There was no such decay in calcination rates in the 3-minute tests, even after 15 cycles. 118 Chapter 5 Simultaneous CO2 and SO2 capture atJluidized-bed combustion temperatures The decay in the initial calcination rate is illustrated in Figure 5.4. The initial calcination rate for the n t h cycle in min"1 is defined as, _ d(Mass of C0 2 released) d(time) m at start of n calcinatiai For comparison, we also show the initial calcination rates in the reversibility tests with pure CO2 as the gas stream during the sorption stage. For the simultaneous tests with 3-minute sorption time, the decay in calcination rate was much slower. However, for SO2/CO2 sorption tests of longer duration, a sharp decrease was observed in the initial calcination rate as the number of cycles increased probably due to the build-up of a sulphate layer which could increase the resistance for CO2 diffusion into the bulk or impede intra-particle heat transfer driving the calcination. Retardation of calcination due to sulphate shells was also observed in the dolomite tests (not shown here). (b) Effects of reaction time and effect of particle size The influence of reaction time on CO2 capture is shown in Figure 5.5 for both sorbents. Each data point corresponds to one cycle. The reaction time on the abscissa-axis is the cumulative sorption time obtained by summing the duration of each individual sorption time stage. For comparison, we have also plotted the baselines without any S0 2 in the inlet gas stream, but with an identical CO2 concentration (80% by volume). A major drop in the C 0 2 capture ability is observed compared to the baseline. Unlike the results in Figure 5.1, the dolomite performed no better than the limestone, although the limestone decayed more sharply in some cases under comparable conditions. Under FBC and simulated conditions, there are three possible reactions involving calcium: Carbonation: CaO+C0 2 ^ CaC0 3 Aff29gJt.=-178'kJ/mol (5.1) 119 Chapter 5 Simultaneous C02 and S02 capture atfluidized-bed combustion temperatures Sulphation: CaO+'/202+S02 = CaS0 4 A / / 2 9 8 K =-426.5 kJ/mol (5.2) Direct sulphation: CaC0 3+'/20 2+S0 2^ CaS0 4+C0 2 Atf 2 9 8 A : =-323.6 kJ/mol (5.3) The rates for reactions (5.2) and (5.3) have been found to be comparable (Tullin and Ljungstrom, 1989), provided that the specific surface areas are similar for the two sorbents (CaO and CaCOs). Carbonation of CaO is much faster than sulphation in typical PFBC flue gases, with the result that the sulphation of CaO proceeds to a very limited degree when both reactions are occurring (Iisa et al., 1991; Tullin and Ljungstrom, 1989). During the first few minutes of reaction, carbonation dominates, producing a compact product shell of carbonate, thereby greatly retarding both reactions (5.1) and (5.2). Exposure of the carbonate product favours reaction (5.3). Compared to the CaO-S0 2 reaction, reaction (5.3) has been reported to suffer less product-layer-related decay because of the release of gaseous C0 2 , causing a more porous product layer and a higher effective diffusivity in the nascent sulphate product (Tullin and Ljungstrom, 1989; Krishnan and Sotirchos, 1993). If enough time were allowed for the carbonation to become product-layer-diffusion controlled, direct sulphation could become dominant. The cumulative loss of calcium due to sulphur capture with each cycle, reducing the availability of solid reactant and reacting surface, thus partially causes the observed drop in C 0 2 capture capability. The reduction of calcium due to sulphur capture cannot account for the entire loss of reversibility as suggested in Figure 5.6, where the ordinate gives the moles of lime determined by subtracting the amount of calcium consumed by sulphation from the total calcium. The calcium utilization due to C 0 2 removal is also shown to decrease sharply with each successive cycle. The free-lime-based calcium utilization decays more for longer-sorption-time cycles, with the least decay occurring for the 3-minute-sorption runs. Based on a comparison with the baseline tests shown in Figure 5.6 and also in Figure 5.1 where there was no S0 2 120 Chapter 5 Simultaneous C02 and S02 capture at jluidized-bed combustion temperatures involved, we attribute this to the presence of S O 2 in gas stream in the present circumstance for the sharp decay of the sorbents in their ability to cyclically capture C 0 2 . As there might be more engineering interest in carbonation efficiency based on the free-lime-content at different sulphation conditions, the free-lime basis is adopted for the following discussion of C02-related calcium utilization. The likely explanation for the decrease of C 0 2 capture ability is that the sulphated envelope becomes thicker with longer sorption time. The thicker sulphate layer not only reduces the calcination rate, as shown in Figure 5.4, but also adds to the resistance for the subsequent carbonation. For 3-minute runs, this effect is weak, probably due to the thinness of the sulphate layer. As seen in Figure 5.5, the dependence of the decay in reversibility on the particle size is weak for the dolomite and for shorter sorption times. For the dolomite, almost all runs with larger (500-600 pm) particles experienced an increase of C 0 2 uptake during the first two cycles, probably due to a change in pore structure. For the limestone, however, there was little difference between the larger and smaller particles as shown in Figure 5.5(b), with almost the same performance for both sizes of particles in 8-minute tests. These results suggest that pore diffusion resistance is more important for calcined dolomite than for limestone, especially during the first few cycles of sorption in the 8-minute tests, (c) Effect of reaction temperatures Higher temperatures give higher sulphation rates; the direct sulphation rate-dependence on temperature shows no obvious maximum compared to sulphation of CaO (Podolski et al., 1995). The dependence of rate of carbonation on temperature is not well-understood, with some researchers (Bhatia and Perlmutter, 1983) concluding that the activation energy of reaction (5.1) is zero. The decrease in reaction temperatures results in an increase in driving force for reaction 121 Chapter 5 Simultaneous CO2 and SO2 capture atfluidized-bed combustion temperatures (5.1). In our tests, when the temperature decreased from 850 to 800°C and then to 750°C, the driving force, estimated from the correlation of Baker (1962), increased from 32 to 58 kPa and then to 71 kPa with the CO2 partial pressure maintained constant at 80 kPa. Figure 5.7 shows a strong effect of temperature for both sorbents, although the dolomite shows slightly less decay. Runs at lower temperature, e.g. 750°C, led to higher CO2 capture, especially for early cycles, due to the higher absorption rates provided by the higher driving force. However, a sharp decrease is immediately observed, indicating that SO2, or more precisely the sulphate layer, has a stronger impairing effect on the CaO-CC»2 reaction. Although higher temperatures lead to increased sulphation rates, so that one would expect the effect of S0 2 on CO2 capture to be more appreciable at higher temperature, the sharpest decreases in sorbent reversibility are observed at lower temperatures, indicating that under the current temperature conditions the effect of changes in sulphation rates is surpassed by the decrease in carbonation rates. These trends are similar for both sorbents. (d) Effect of SO2 concentration Higher S0 2 concentration allows more calcium to be sulphated. The resulting thicker sulphate shell increases the resistance to CO2 diffusion, causing a sharper decay in the rate of carbonation. This is confirmed in Figure 5.8 for both sorbents and for different durations of sorption. The influence of SO2 concentration becomes more appreciable for later cycles, consistent with thicker layers of sulphate causing greater decay in the reversibility of carbonation. (e) Effect of C O 2 concentration As shown in Figure 5.9, the effect of SO2 to diminish the carbonation reversibility was appreciably greater for high CO2 concentrations than for low CO2 concentrations. Although the 122 Chapter 5 Simultaneous CO2 and SO2 capture at fluidized-bed combustion temperatures higher driving force for the same temperatures should lead to greater CaO utilization, this was only true for the early cycles, with the advantage being greatly reduced for later cycles, presumably because of the presence of a compact sulphate layer which exerts more influence on the CaO-C02 reaction for higher C O 2 concentration and longer sorption times. 5.3.3 Simultaneous sulphation and carbonation: Effect of C O 2 on S O 2 capture Figure 5.10 shows that the presence of C O 2 in the gas stream enhances SO2 capture. Direct sulphation of the carbonates could proceed more quickly than sulphation of CaO because of the porous nature of the sulphate product (Tullin and Ljungstrom, 1989; Krishnan and Sotirchos, 1993). However, this phenomenon only became appreciable at higher conversions, e.g. calcium utilization higher than 30-40%, because at lower conversions kinetics tend to be the limiting step, and the CaO-S02 reaction was probably faster than the CaC03-S02 reaction due to CaO having a higher surface area than CaC03. In Figure 5.10, however, enhancement is observed even at the very low conversions in all cases, especially for runs of shorter duration. In other words, although CaC03 is usually lower in surface area than CaO and is supposed to be slower than CaO when reacting with SO2 at low conversions, it actually exhibits more reactivity when cyclic carbonation/calcination proceeds in parallel with sulphation, as shown in Figure 5.10. This suggests that cyclic sorption could enhance sulphation by producing more reactive C a C 0 3 compared to the CaO-S02 reaction. Cycles of shorter duration show the greatest enhancement effect. By comparing the performance of different sorbents and particle sizes, we find that the dolomite shows a significantly greater ability to capture S 0 2 than limestone. Particle size seems to have little effect on the rate of SO2 capture or on the enhancement effect of CO2. 123 Chapter 5 Simultaneous C02 and S02 capture at fluidized-bed combustion temperatures 5.4 Conclusions Simultaneous carbonation and sulphation were investigated under simulated F B C conditions in an atmospheric-pressure thermogravimetric reactor (TGR). SO2 was found to impede CO2 capture, with calcination rates diminished by the sulphate outer shells enveloping the sorbent particles. The decrease in rate was less for shorter sorption durations, but these led to less CO2 capture. Higher temperatures resulted in less decay, but also in reduced CO2 capture. Total calcium utilization decreased with multiple cycles, especially for longer sorption times. Higher SO2 concentrations led to a sharper decay in reversibility. Sharper decays in reversibility occurred at higher CO2 conditions. CO2 enhanced the capture of SO2, especially for runs of brief sorption duration. Due to the high molar ratio of CO2 to SO2, Abanades et al. (2003) believed that the high Ca/S molar ratio for cyclic carbonation/calcination would enable efficient SO2 capture in a proposed PFBC-based CO2 removal process (Abanades et al., 2003; 2004a) such that CO2 removal would not be influenced significantly by calcium loss due to reaction with SO2. However, the results of this work show that the presence of SO2 in the gas stream greatly impedes the reversibility of even potentially good sorbents such as dolomite. This may lead to difficulties for practical CO2 capture based on limestone cycles. We recognize, however, that the current results are only based on simulated non-calcining conditions and bench scale tests. Further tests under more realistic conditions and with macro-scale reactors are required. If the adverse effect of SO2 on CO2 capture is confirmed, revisions of the present concept, such as separate removal of S 0 2 and CO2 or steam reactivation would need to be tested in the search for a practical process. 124 Chapter 5 Simultaneous CO2 and SO2 capture atfluidized-bed combustion temperatures ro O to o 0.8 E ~B o "O Gi-ro o CM O O CO a> o 0.6 0.4 0.2 A • O Strassburg, 100% C 0 2 , F S F Arctic, 100% C 0 2 , F S F Arctic, 80% C 0 2 , 20% N 2 , 3 min Strassburg, 80% C 0 2 , 20% N 2 , 3 min Arctic, 80% C 0 2 , 20% N 2 , 30 min A A A A A A A A " l | I a „ A A A A -extended cycle at the 18 t h cycle • • • • 10 15 20 Number of reaction cycles 25 30 Figure 5.1 C 0 2 cyclic capture performance. 212-250 pm Arctic dolomite and Strassburg limestone. Calcination: 850°C in 100% N 2 ; Carbonation 850°C with 80 or 100% C 0 2 (no S02). Sorption time is 3, 30 min for each sorption stage or FSF. 51 53 49 Time (min) Figure 5.2 Illustration for turning point selection in a 100% C0 2 capture test with 212-250 pm Strassburg limestone. 125 Chapter 5 Simultaneous C02 and SQ2 capture at fluidized-bed combustion temperatures 6 0 1.2 1.1 a 1 (3 o ••e o OT 0.9 0.8 (a) 30 min each sorption " Mass gain due to S 0 2 A sorption in 2" cycle 1.2 M 1.1 c I 0 .9 0.8 50 100 150 Time (min) 30 60 Time (min) 200 90 250 120 0 30 60 Time (min) 90 120 Figure 5.3 Test results for 212-250 um limestone, 850°C calcination and 850°C sorption with 2900 ppm S0 2 and 80% C0 2 . 126 Chapter 5 Simultaneous C02 and SQ2 capture at fluidized-bed combustion temperatures I 3 cu -4-t a C o a *c5 U 0.1 1 0.08 0.06 A 2 0.04 4 0.02 -•— 30 minute test -O— 3 minute test -A— 8 minute test 6 . 8 10 Cycle number 12 14 16 Figure 5.4 Comparison of initial calcination rate of each cycle for the limestone, 850°C calcination/850°C carbonation, 212-250 pm particle. 127 Chapter 5 Simultaneous CQ2 and SQ2 capture at fluidized-bed combustion temperatures 20 40 60 80 Cumulative sorption time (min) B - 212-250 um, 30 min —&- 212-250 um, 8 min — 5 0 0 - 6 0 0 um, 8 min •"A-" 212-250 tun, 3 min - ± - 500-600 um, 3 min 212-250 um, 3 min, baseline Figure 5.5 Effect of cumulative reaction time. 850°C calcination/850°C sorption, gas compositions for S0 2 /C0 2 sorption: 80% C0 2 , 2900 ppm S0 2 , 3% 0 2, balance N 2 . (Baseline, 80% C 0 2 20% N2). 128 Chapter 5 Simultaneous C02 and SO2 capture at fluidized-bed combustion temperatures 20 40 60 80 Cumulative sorption time (min) —0212-250 um, 30 min - • -500-600 um, 30 min —0-212-250 um, 8 min —A—212-250 um, 3 min —•-500-600 um, 8 min — ± - 5 0 0 - 6 0 0 um, 3 min 100 20 40 60 80 Cumulative sorption time (min) _g_ 212-250 um, 30 min Q 212-250 urn, 8 min 500-600 urn, 30 min ••A—- 212-250 um, 3 min -hr~ 500-600 um, 3 min 100 Figure 5.6 Total calcium utilization change with cycles for 850°C calcination/850°C sorption, gas compositions for sorption: 80% C0 2 , 2900 ppm S0 2, 3% 0 2, balance N 2 . 129 Chapter 5 Simultaneous CO2 and SO2 capture atfluidized-bed combustion temperatures Figure 5.7 Effect of reaction temperature C0 2 concentration for each cycle, 212-250 pm, 3-minutes sorption time, 850°C calcination, gas compositions for S0 2/C0 2 sorption: 80% C0 2 , 2900 ppm S02, 3% 02, balance N2.(Base lines, 80% C0 2 , 20% N2). 130 Chapter 5 Simultaneous C02 and S02 capture atfluidized-bed combustion temperatures 0.6 CJ O C+-< C « O O O o 0.4 0.2 _ H _ 2 9 0 0 ppm SO z, 30 min _ A _ 4 1 0 0 ppm S0 2, 30 min 4100 ppm SOj, 3 min -O- 2900 ppm S0 2, 3 min NoSCh, 3 min 40 80 Cumulative sorption time (min) 120 Figure 5.8 Effect of S0 2 concentration for successive cycles for 212-250 pm particle 850°C calcination/850°C sorption, gas compositions for SO2/CO2 sorption: 80% C0 2 , 3% 0 2, balance N 2 . (Baseline, 80% C 0 2 20% N2). 131 Chapter 5 Simultaneous CO2 and SO2 capture atfluidized-bed combustion temperatures 1 0 20 40 60 80 100 Cumulative sorption time (min) Figure 5.9 Effect of C 0 2 concentration for successive cycles with 212-250 um limestone. 850°C calcination/850°C sorption. Gas compositions for S0 2 /C0 2 sorption: 80% C 0 2 , 2900 ppm S0 2 , 3% 0 2 , balance N2.(Base lines, 80% C0 220% N2). 132 Chapter 5 Simultaneous CO2 and S02 capture atfluidized-bed combustion temperatures • 30 min A 8 min O 3 min Cumulative sorption time (min) 0.4 U c/n CD £ o H o co o CO 0.3 0.2 0.1 0 0 (b) 212-250 um Strassburg limestone Baseline, no C 0 2 • 30 min A 8 min O 3 min 40 80 Cumulative sorption time (min) 120 133 Chapter 5 Simultaneous CO2 and SO2 capture at fluidized-bed combustion temperatures O a o O U in 0.4 0.3 0.2 V 0.1 40 80 Cumulative sorption time (min) 120 a U in CD "o £ -*-> o H o ccj O CD "o 0.4 0.3 0.2 0.1 (d) 500-600 um Strassburg limestone A O s m i n 3 min Baseline, no C 0 2 o 40 80 Cumulative reaction time (min) Figure 5.10 Effect of C 0 2 on S0 2 capture. 850°C calcination/850°C sorption. Gas compositions or S02/C0 2 sorption: 80% C 0 2 , 2900 ppm S0 2, 3% 0 2 (Base lines: 2900 ppm S0 2 , 3% 02). 134 Chapter 6 Removal of CO 2 by calcium-based sorbents in the presence of SO 2 C H A P T E R 6 R E M O V A L O F C 0 2 BY C A L C I U M - B A S E D SORBENTS IN T H E P R E S E N C E O F S 0 2 A version of this chapter has been published in Energy and Fuels, 2006, 21, 163-170. Authors are: P. Sun, J. R. Grace, C. J. Lim and E. J. Anthony. 6.1 Introduction The use of calcium-based sorbents as candidates to remove CO2 in-situ from reactors has been receiving increasing attention, not only for steam reformers, steam gasifiers and water-gas shift reactors to improve hydrogen yields (Han and Harrison, 1994; Lin et al. 2001; 2002; Ortiz and Harrison, 2001; Johnsen et al., 2006), but also for C O 2 removal from fluidized bed combustors (FBC) (Shimizu et al., 1999; Gupta and Fan, 2002; Hughes et al., 2004; Abanades et al., 2003, 2004a; 2005; Salvador et al., 2003). The decline in sorbent capability in multiple cycles, i.e. reversibility, is a key factor affecting process economics for calcium-based CO2 sorbents (Abanades et al., 2004b). It has been reported that dolomites are generally more reversible than limestones in such cycles (Silaban et al., 1996). Natural limestones usually follow similar decay trends when exposed to cyclic capture conditions (Abanades, 2002). Sintering of unreacted CaO is believed to be the major cause of deactivation, as evidenced by the change of sorbent surface texture after multiple cycles (Abanades, 2002; Abanades and Alvarez, 2003; Alvarez and Abanades, 2005). In Chapter 5, it was found that the SO2, a common pollutant in coal-fired fluidized bed combustors, appreciably decreases the CO2 capture capacity for multiple cycles, despite the fact that SO2 concentrations are several orders of magnitude lower than CO2 concentrations. This adverse 135 Chapter 6 Removal of CO 2 by calcium-based sorbents in the presence of SO 2 effect of SO2 was recently confirmed by Ryu el al. (2006) in a bubbling fluidized bed with batches of sorbent. However, the mechanism of deactivation by low levels of SO2 is still unclear. The objective of this work was to investigate cyclic co-capture of SO2 and CO2 in greater detail in order to clarify the mechanism of deactivation, to test the effect of sorbent type, and to test possible ways to improve the reversibility of calcium-based sorbents. 6.2 Experimental details Seven commercially available sorbents were tested: Strassburg limestone (US), Dahyang limestone (Korea), Havelock limestone (Canada), Kelly Rock limestone (Canada), Cadomin limestone (Canada), Arctic dolomite (Norway) and GS dolomite (US). Table 1.2 gives chemical analyses. For all tests in this study, the particle size range was screened to 212-250 pm. Two fixed-bed thermogravimetric reactors were used for the experiments, one operating at atmospheric pressure and the other under pressurized conditions up to 2.4 MPa. Details of the atmospheric pressure thermogravimetric reactor (ATGR) are provided elsewhere (Laursen et al., 2000; 2001). Mass flow controllers were utilized to achieve desired inlet gas concentrations. A total gas flow of 1600 ml/min was maintained for both calcination and carbonation. A typical FBC combustor temperature, 850°C, was chosen for all runs. During the sorption stage, 80% CO2 was typically used at 850°C. The inlet concentrations of SO2 and O2 were maintained at 2900 ppm and 3% by volume respectively. Fresh samples weighing approximately 850 mg were loaded into the basket for each ATGR run so that enough sorbents were generated for pore size analyses, chemical analyses or SEM studies. Under such conditions, both interparticle and intraparticle mass transfer can be significant. The sorption time for each cycle has been found to be important in previous work (Barker, 1973; Alvarez and Abanades, 2005). Each cycle includes one step of calcination followed by a 136 Chapter 6 Removal of C02 by calcium-based sorbents in the presence of S02 sorption (carbonation w/o sulphation) step. Two durations for the sorption step were employed for ATGR tests. The turning point from an initial fast stage of carbonation to a slower stage, easily detected under the current test conditions (see Chapter 5), was used to determine the stopping point for each cycle. This duration, denoted "fast stage finished" (FSF), is predominantly used in tests with no S O 2 present. Alternatively, a fixed duration of 8 minutes was selected, primarily for co-capture tests in the atmospheric TGR. During the first few cycles, a time of 8 minutes was not long enough to complete the fast stage of carbonation. However, this duration was sufficient for later cycles. The pressurized thermogravimetric analyser (PTGA), consisting of a Cahn 100 balance of 1 ug sensitivity and a reactor column made of Inconel 600 alloy, was used to simulate pressurized fluidized bed combustion (PFBC) conditions. Restricted by the operating temperatures utilized in FBC combustors, high removal efficiencies could only be achieved at high C O 2 partial pressures. A PFBC is therefore ideal for capture of C 0 2 by calcium-based sorbents. The PTGA was used to test the effects of total pressure and partial pressure of C 0 2 and S0 2. Strassburg limestone and Arctic dolomite were selected for these tests. A pressure regulator adjusted the pressure level. Mass flow controllers allowed desired mixtures of reactant gases to be delivered. A C O 2 concentration of 8% by volume was selected as the base condition. As shown below, varying the C O 2 concentration from 14%v to 20%v had negligible effect on the reversibility of sorbents when C O 2 was the only reactant gas. In PTGA runs, 50±2 mg of fresh sorbents screened to a 212 to 250 um size range were introduced to the reactor at the start of each run. For the PTGA runs, the sorption time was 4 minutes for tests with and without S0 2, enough for the fast stage of carbonation to be completed for each cycle. Each sorption stage was conducted under high-137 Chapter 6 Removal of CO2 by calcium-based sorbents in the presence of SO2 pressure conditions, whereas calcination could be conducted under either atmospheric or pressurized conditions. For the tests in both reactors, temperatures for calcination and carbonation were maintained at 850°C except where specified. Only the initial calcination was operated non-isothermally, increasing the temperature from room temperature, whereas all subsequent cycles shifted between calcination and carbonation by switching gas compositions while maintaining the temperature. After switching, the gas took around 1 minute to reach the sample surface, making no difference to the cycling time given the cyclic operational manner. All calcinations were performed in an atmosphere of pure N 2 . The collected samples were analysed by several techniques. SEM/EDX was carried out using HITACHI S-3000. For sulfur mappings, samples were first embedded in epoxy and then carefully ground and polished to expose cross-sections. For high-resolution SEM measurements, all samples were gold-coated to make them conductive. A Micromeritics 9300 Poresizer determined the pore size distributions of sorbents. A Siemens Diffraktometer D5000 provided phase detection for some samples. 6.3 Results and Discussion 6.3.1 Tests with no S O 2 among the gaseous reactants, referred to hereafter as "no S O 2 " tests The cyclic performance of the seven sorbents during the 850°C sorption/calcination cycles, based on total calcium utilization, is shown in Figure 6.1. The tests were performed at ATGR. 80% CO2, 20% N 2 were used for the sorption stage and 100% N2 for the calcination stage. Calcium utilization is calculated on a molar basis. 138 Chapter 6 Removal of C02 by calcium-based sorbents in the presence of S02 The dolomites showed both higher calcium utilizations and better reversibility than the limestones. The superior reversibility of the dolomites is probably due to a lack of sintering-related CaO crystal coarsening, presumably due to the MgO part resisting the mass flow of sintering-stress-driven ions and vacancies in this ionic compound. For limestones, the Ca utilization decreased to 20-30% of the original value after 15 cycles, with relatively minor differences among the five limestones investigated. High-resolution SEM images for calcines of Strassburg limestone, Arctic dolomite and Kelly Rock limestone after cycling are shown in Figure 6.2 (that of Danyang limestone is shown in Figure AVTJ.1 in Appendix VH). Strassburg and Kelly Rock limestones showed similar decay trends, probably related to structural changes during the sintering cycles. The microstructures for their calcines, however, differed significantly. Strassburg limestone showed a typical intermediate sintering pattern (German, 1996), characteristic of shrinking small pores and growing larger pores, leaving behind clusters when some parts of CaO coarsen. Kelly Rock limestone, showed a more grainy structure with clear neck development between grains, but with more pore volume between grains. Despite differences in their microstructures, the pore size distributions of Strassburg and Kelly Rock limestones in Figures 6.3a and 6.3b showed similar sintering patterns, reducing pore volume in the <~200 nm range in favour of pores of 200 nm or larger. This caused the previous mono-modal pore distribution to undergo a transition to a bimodal distribution, with the degree of reversibility appearing to be very similar for these two limestones. In Figure 6.2c, cycled Arctic dolomite calcines did not contain visible pores from the SEM photo, and the grain size could not be viewed clearly. The extent of sintering was much lower than for the limestones. Given that dolomites may have higher surface-area-related sintering stresses, the reduction in sintering was likely due to retarded sintering kinetics or to interspersed 139 Chapter 6 Removal of CO 2 by calcium-based sorbents in the presence of SO 2 MgO reducing diffusion of CaO ions. Figure 6.3c compares the evolution of pore size distribution from initial calcines and cycled samples. The unique pores resided in the 10-100 nm ranges with a peak at 15.8 nm, smaller than for most of the limestones, but with lower volume. After 20 cycles the pores were not appreciably retained as one might expect based on the excellent cyclic performance for this sorbent. However, there were no signs of larger pores being generated. Silaban et al.(1996) and Silaban and Harrison (1995) also found that the porosity of their dolomite decreased, even after the first cycle. Given that Arctic dolomite performed relatively well, as shown in Figure 6.1, it can be deduced that dolomite indeed sinters, but at much reduced rates. In FBC where temperatures of 850-900°C are preferred, high CO2 removal efficiency can only be achieved under pressurized conditions, given the thermodynamics of the CaO-C02 reaction. In the present work, Strassburg limestone and Arctic dolomite were selected for the PTGA tests. Two sets of results for cyclic performance of Strassburg limestone are compared with the ATGR result in Figure 6.4. It can be seen that the PTGA and the ATGR generated similar decay behaviour for this sorbent. For calcination at elevated pressures, augmented sintering or further decay was expected based on reported (Borgwardt, 1989) enhanced CaO sintering in higher concentrations of C02. However, the reversibility actually improved slightly as shown in Figure 6.4. Because under pressurized calcinations, the actual N 2 mass flow and superficial velocity were much lower, actual carbonation duration was longer than the prescribed 4 minutes, responsible for the extra CaO conversions observed. It is notable that pressurized calcination under 100% N 2 conditions did not accelerate sorbent deterioration. This was confirmed by SEM investigation, which revealed no difference in sorbent surface microstructure. This evidence indicates that a pressurized calciner is a possible option for practical applications. Further PTGA 140 Chapter 6 Removal of C02 by calcium-based sorbents in the presence of S02 tests (not plotted here) show that changes of total pressure up to 2.4 MPa or varying the C 0 2 partial pressures from 8% to 14%v and 20%v did not appreciably change sorbent reversibility. Contrary to the results for the Strassburg limestone, the PTGA test with Arctic dolomite at 1.8 MPa showed reduced reversibility as portrayed in Figure 6.5. SEM images for calcines again revealed little difference. Pressurized carbonation probably altered the sorbent, eliminating pores via enhanced ionic movement for either CaO or MgO. Further tests of Arctic dolomite with total pressure varying from 1.3 to 2.4 MPa, revealed little or no change in reversibility. 6.3.2 Simultaneous S O 2 and C O 2 capture In fluidized bed combustors, sulphur is invariably present to some degree at least so that simultaneous sorption of SO2 and CO2 need to be achieved in order to capture CO2 in situ. However, previous tests (Sun et al, 2005; Ryu et al., 2006) (also see Chapter 5) have indicated that S 0 2 impedes both cyclic C 0 2 capture and subsequent calcination under typical FBC operating conditions. The cause of this adverse influence of SO2 needs to be addressed. During co-capture, it is believed that three reactions occur simultaneously: Carbonation: C a O + C 0 2 C a C 0 3 Atf 2 9 g J i : =-178 kJ/mol (6.1) Sulphation: CaO+'/202+S02 = CaS0 4 Atf 2 9 8 J C =-426.5 kJ/mol (6.2) Direct sulphation: CaC0 3 + , /202+S02^CaS0 4 +C0 2 A H 2 9 g / . =-323.6 kJ/mol (6.3) Under current ATGR test conditions (e.g. with 80%v C 0 2 at 850°C and 101 kPa), based on thermodynamic equilibrium, the MgO-C02 reaction is thermodynamically inhibited, but the MgO-S0 2 reaction is thermodynamically possible at higher PSo2(Hartman and Svobods, 1985). After initial full calcination at 850°C, Arctic dolomite calcines were subjected to simultaneous carbonation and sulphation for about 30 minutes at both 850 and 750°C, and then sent directly for XRD analysis without calcination. In another run, the sorbent was directly taken after seven 141 Chapter 6 Removal of C02 by calcium-based sorbents in the presence of S02 cycles of co-capture at 850°C and after calcination with gas conditions for baseline test, i.e. 2900 ppm S0 2 , 80% C 0 2 , 3% 0 2 and balance N 2 at 850°C and 101 kPa pressure. Neither MgS0 4 nor MgC0 3 was detectable in any sample, indicating that MgO did not react with C0 2 , as expected based on thermodynamic grounds, while for S0 2 , the MgO was either inert, or reaction proceeded too slowly to be significant. It is also of interest that very little CaO was detected in the non-calcined sorbents, indicating that for Arctic dolomite calcium utilization could be close to unity in the initial cycle. For calcines from multi-cycle co-capture, there was only a tiny CaC03 peak, indicating that very little remained uncalcined. As it is thermodynamically possible for MgO to react with S0 2 at the pressures of the PTGA tests, XRD phase analysis for samples from a cyclic S0 2 /C0 2 capture test at 2.4 MPa pressure (triangular points in Figure 6.10), showed no appreciable MgS04 phase, probably because of the much lower reactivity of MgO to S0 2 compared to CaO. Sorbents were calcined after each cycle of co-capture, and mass differences were used to determine the S0 2 and C 0 2 capture during each cycle (see Chapter 5), based on the assumption that all of weight loss during calcination is due to C 0 2 being driven off. The mass difference method has been used by previous researchers who investigated simultaneous carbonation and sulphation of calcined limestones under PFBC conditions (Tullin and Ljungstrom, 1989; Iisa et al., 1991). Chemical analyses for total carbon and sulfur for ATGR samples confirmed that the difference by the two methods was less than 5%. Figure 6.6 shows the cyclic performance of all seven sorbents during ATGR co-capture with 8-minute sorption periods. The corresponding baselines for Arctic dolomite and Strassburg limestone for no-S02 tests with 80% C 0 2 are also shown. The calcium utilizations for all sorbents dropped sharply to about 20% over the first 10 cycles and continued to decrease beyond this. Dolomites did not appear to be any better than limestones when S02was present. In fact, 142 Chapter 6 Removal of CO 2 by calcium-based sorbents in the presence of S02 Arctic dolomite, which had shown the best cyclic reactivity in the "no S O 2 present" test, dropped even more sharply than the other sorbents. The GS dolomite performed in a similar manner. As more sulfur was captured with each successive exposure, the behaviour of all sorbents appeared to be controlled more by the sulfur-altered properties than by the original reactivities. "Once-through" tests were conducted in the ATGR to investigate the importance of reactions (6.1), (6.2) and (6.3). Strassburg limestone and Arctic dolomite were tested repeatedly, each sample starting from the very beginning and being stopped at a certain point of the sorption from 3 to 90 minutes and then calcined. The results appear in Figure 6.7 for the Strassburg limestone. For comparison, sorbents were also subjected to S O 2 and 0 2 in the absence of C 0 2 , and to C O 2 in the absence of S O 2 , with identical gaseous reactant concentrations as for the cyclic co-capture runs. During the first 14 minutes of co-capture, an almost identical amount of C O 2 was captured as in cases where there was no S O 2 for both sorbents. The CaO was very reactive initially and reacted much more quickly with C O 2 than with S O 2 , as in earlier PFBC S02-capture studies (Tullin and Ljungstrom, 1989; Iisa et al., 1991). After completing the fast stage of carbonation, sorption became much slower. As indicated by the mass loss after 20 minutes, carbonate was obviously consumed through reaction (6.3) and no further C O 2 could be captured. This indicated that direct sulphation had become dominant in the later stage of sorption, whereas CaO sulphation could only be appreciable at the beginning, and was immediately surpassed due to the formation of a carbonate shell, which was, however, readily available for direct sulphation. Direct sulphation can result in greater calcium utilization than CaO sulphation because of a less compact sulphate layer (Snow et al, 1988; Krishan and Sotirchos, 1993). The overall sulphation was therefore higher than when no C O 2 was present. Similar trends have been reported in Chapter 5 where cyclic co-capture conditions significantly enhance S O 2 capture. Note that the 143 Chapter 6 Removal of CO 2 by calcium-based sorbents in the presence of SO 2 carbonate content of Arctic dolomite was consumed more quickly than that of Strassburg limestone, indicating that direct sulphation was faster for the dolomite. It is also notable that in Chapter 5, longer sorption times led to faster decreases in calcination rate, due to the sulphate envelopes that impede subsequent calcination. Figure 6.8 shows the decrease in calcination rates as the calcium utilization for sulphur capture increase for these sorbents - Arctic dolomite, Strassburg limestone and GS dolomite. The calcination rate is defined as _ d(Massof C 0 2 released) R c . O -TT-. ~ I IP A ) d(time) t h at start of n calcination. The rate at which Arctic dolomite calcined dropped off sharply, whereas for other sorbents, the calcination rates also decreased as cycles proceeded, but less strongly. Calcination rates decreased as the sulphate content increased. The faster decay for Arctic dolomite may be attributed to higher reactivity toward S O 2 capture, promoted by higher surface area as well as cracks. The GS dolomite also showed strong ability to capture S O 2 , resulting in a high calcium conversion to CaS04, but its decline in calcination rate was much slower than for the other sorbents. SEM images suggest that GS dolomite particles have a higher tendency to fragment than the other sorbents tested, generating large numbers of fractures, resulting in faster calcination and ensuring that it is more reactive during co-capture. PTGA co-capture test results for Strassburg limestone and Arctic dolomite appear in Figures 6.9 and 6.10, respectively (with the conversions of CaO to CaS0 4 shown in Figures AVH.5 and AVH.6 in Appendix AVE). As for the ATGR tests, S0 2 clearly impeded C 0 2 capture. However, variations in total pressure did not appreciably change the sorbent reversibility, as discussed below. 144 Chapter 6 Removal of C02 by calcium-based sorbents in the presence of S02 Figure 6.11a shows EDX sulfur distributions for the 12th cycle of A T G A co-capture before calcination for Strassburg limestone. Sulphation clearly occurred along the outer rim of particles. Sulfur permeated the particles probably by sohd-state ion diffusion, crossing both the sulphate and carbonate layers. Direct sulphation of CaCC>3 produces a more porous sulphated layer than CaO sulphation due to the outwards diffusion of C0 2(Snow et al, 1988; Krishan and Sotirchos, 1993). As a result, it is easier for S0 2 to continue to diffuse inwards. Cyclic calcination also leads to periodic generation of pores with CaO at the surface, subsequently carbonated to provide more interface at which CaC03 can react with S0 2. Pore size distributions for cyclic co-capture tend to be bimodal, as shown in Figure 6.3. However, pore volumes were much reduced, with the reduction in the <~200 nm range being due to CaO sintering, as discussed for the case when these was no S0 2 , whereas the volume reduction for larger pores was due to sulphate filling. Gullett and Bruce (1987), based on N 2 adsorption, showed that sulphation generally fills pores of 10-60 nm or larger. Other studies have shown that S0 2 tends to block pore entrances by occupying the outer rims of particles featuring mesopores (Simons and Garman, 1986; Sotirchos and Zarkanitis, 1992) due to the expanding product. Once pore mouths are filled, the effective diffusivity is greatly reduced for both CaO carbonation and subsequent calcination. Figure 6.11b presents the sulfur distribution for Kelly Rock limestone. Although previous SEM photos from the test with no S0 2 revealed that Kelly Rock limestone calcines have a grainy structure and should sulfate more uniformly, the sulfur is present in an unreacted-core pattern. The reaction between S0 2 and less porous CaC03, must have taken place at the outside of the particles, blocking inter-grain channels. As a result of the co-capture, the pore size distribution in Figure 6.3b followed a similar trend as for Strassburg limestone, due to the joint effects of CaO sintering and sulphate filling. (The EDX sulfur distribution for Danyang calcines after 2 hours of sulphation and 15 cycles of S0 2 /C0 2 capture/calcination are shown in Figure AVTI.2 in 145 Chapter 6 Removal of C02 by calcium-based sorbents in the presence of S02 Appendix A V E with the pore size distribution shown in Figure AVT1.3; the pore size distribution for Havelock calcines, either fresh, after co-capture and no S02 cycling tests are shown in Figure AVTI.4). For Arctic dolomite calcines, more sulfur was found along fractures or cracks as shown in Figure 6.11c. This explains the pore volume generation at pore sizes larger than 200 nm in Figure 6.3c for co-capture. Arctic dolomite must be weak mechanically, perhaps because of discontinuities between the magnesium and calcium parts in fully calcined dolomites, making it possible to generate cracks during cyclic capture, especially for co-capture due to the molar volume of the sulphate product being larger than for the carbonate. But the sulfur cannot permeate uniformly throughout the sorbent, so that sulfate product easily blocks pores and channels, thereby preventing further co-capture. High-resolution SEM images of co-capture calcines for Strassburg limestone, Kelly Rock limestone, and Arctic dolomite appear in Figure 6.12. Compared with the calcines in the "no-SO2" studies (Figure 6.2a), the Strassburg sorbent generated similar large pores through sintering, as seen in Figure 6.12a, covered by a layer of sulphate. The sulphate envelope, covering the pores, caused overall loss of pores as indicated in Figure 6.3a. For Kelly Rock limestone, the sorbent appeared to be much less porous after co-capture, compared with its counterpart in the "no SO2" test due to sulphate product filling inter-grain pore volume. For the Arctic dolomite, the 10-100 nm pores in Figure 6.3c cannot be clearly viewed in SEM pictures. The cracks indicate that this sorbent was easily broken. The better reversibility of Strassburg limestone during co-capture, relative to Arctic dolomite, may be due to larger effective diffusivities because of sintering-generated macrospores. SEM photos (not shown here) from the PTGA Strassburg limestone co-capture tests show patterns very similar to the ATGR ones. 146 Chapter 6 Removal ofC02 by calcium-based sorbents in the presence ofS02 In summary, a CaC0 3 product layer, once formed, greatly decreases the effective diffusivities for further carbonation, but facilitates direct sulphation, which subsequently seals macropores, reducing the effective diffusivities, both for further carbonation and sulphation. Because of the formation of an outer layer of carbonate, direct sulphation becomes dominant in the later stage of co-capture. During calcination of Strassburg limestone, released C O 2 first generates small pores in bulk carbonate, then finds exit paths through partly sulphate-filled macropore channels. The calcination rate is reduced to a variable degree, depending on the extent of filling by the sulphate product. Thus, during co-capture, two patterns of pore evolution are important: sintering occurs as in cyclic no-SCh tests while sulphate product irreversibly fills the macropores (Gullet and Bruce, 1987), reducing pore volume and blocking channels to smaller pores important during fast stage carbonation (Bhatia and Perlmutter, 1983; Abanades, 2002). 6.3 .3 Change of operating conditions in an effort to improve co-capture Several measures were next taken to see whether the co-capture performance could be improved. One factor was varied at a time to see whether the decline in calcium utilization with cycling could be arrested. In A T G A tests, pre-treatment of calcined Strassburg limestone with liquid water to give hydrated lime showed the greatest improvement of all methods tried not only for C 0 2 capture, but also for S O 2 retention. The enhanced co-capture of hydrated lime is due to its pore structure, allowing for expansion to accommodate more solid product, as for h-CaO sulphation (Gullet and 147 Chapter 6 Removal of CO2 by calcium-based sorbents in the presence ofSC>2 Bruce, 1987; Stouffer and Yoon, 1989). However, when both SO2 and C 0 2 capture are enhanced, the ability of sorbents to capture and retain C 0 2 decreases. (A). With other conditions unchanged, calcining at a lower temperature, 750°C, gave a small improvement in co-capture performance probably because the lower temperature retarded sintering, but low-temperature calcination is unlikely to be practical. (B). One test was also performed with 93% CO2 the balance N2 balance. The increase of CO2 concentration helped increase calcium utilization, but only for the first few cycles. As determined in PFBC studies (Snow et al., 1988; Tullin and Ljungstrom, 1989; Iisa et al., 1991; Krishnan and Sotirchos, 1993), direct sulphation is inhibited by increasing bulk CO2 concentrations. The lack of significant reduction in direct sulphation is probably because of limited capacity to increase CO2 concentrations above 80%v in the ATGR. (C). Both C 0 2 and S0 2 capture were enhanced somewhat by steam addition during co-capture. The water vapour might dissociate and interact with ionic surface enhancing mass transfer or kinetic behavior, e.g. via a short-lived surface hydroxyl group as reported for steam enhancement of CaO or MgO sintering (Anderson and Morgan, 1964; Borgwardt, 1989). (D). A run with 800 ml/min flow rate was conducted in an attempt to increase Ca/S molar ratio in the fixed-bed tests, both C 0 2 and S0 2 capture improved for the first few cycles compared to the 1600 ml/min run, though with lower utilizations later. When Arctic dolomite was subjected to similar ATGA tests, no clear improvement was observed as for the limestone. In PTGA tests, the total pressure and CO2 partial pressure were varied. As shown in Figures 6.9 and 6.10, changes in total pressure had little influence on CO2 removal for either sorbent tested. Similarly there was or no effect on S0 2 capture. (Figures AVn.7-9 in Appendix AVTf also show the lack of dependence of cyclic performance in no SO2 tests on total pressure, 148 Chapter 6 Removal of CO 2 by calcium-based sorbents in the presence of S02 temperature and Pco2 for Strassburg limestone, which provide comparisons to the results of co-capture.) There is a lack of agreement in the literature on the effect of total pressure on direct sulphation. Iisa and Hupa (1990) believed that an increase of total pressure increased Pso2, resulting in increased sulphation, whereas Qiu and Lindqvist (2000) reported the opposite in the range of 0.6-1.3 MPa. Iisa et al. (1991) tested simultaneous carbonation and sulphation over extended durations and found that sulphation was slower initially, but could reach higher conversions when the total pressure increased from 0.1 to 0.8 MPa. However, there was only a small dependence on pressure between 0.8 and 2 MPa. The effect of total pressure on co-capture probably depends on multiple factors. When the total pressure is increased, both Pco2 and Pso2 increase proportionally. For bulk diffusion, the diffusion coefficient varies inversely with pressure so that the effect of increased concentrations tends to be neutralized. With increasing total pressure, increasing PCo2 far above its equilibrium value would not appreciably increase the rate of carbonation (Kyaw et al., 1996), but CaO sulphation and direct sulphation might accelerate further because of rate dependence. In addition, direct sulphation is inhibited by increasing PCo2 (Tullin et al., 1993; Liu et al., 2000). If the pore dimensions are small enough for Knudsen diffusion to be rate-controlling, the diffusion coefficient is independent of pressure. However, the mass flux increases with total pressure. Hajaligol et al. (1998) suggest that direct sulphation is Knudsen-diffusion controlled. In the current study, as shown in Figures 6.9 and 6.10, these effects added up to negligible net change in CO2 and SO2 removal. Variation of C 0 2 partial pressure while holding other conditions constant affected reversibility. Figure 6.13a shows that a higher CO2 partial pressure greatly improved sorbent reversibility, corresponding to a much-reduced SO2 capture in Figure 6.13b. Increasing of Pco2, 149 Chapter 6 Removal of CO 2 by calcium-based sorbents in the presence of S02 supposedly increased both interparticle and intraparticle diffusion flux and intrinsic carbonation rates, thus speeding up carbonation. Accelerated carbonation, cannot change the final conversion of CaO to carbonate, as revealed in the no-S02 tests, but it makes direct sulphation dominant at an earlier stage during co-capture. More importantly, direct sulphation is inhibited by increasing C O 2 partial pressure (Tullin et al., 1993; Liu et al., 2000), consistent with Le Chatelier's principle applied to reaction (6.3), resulting in less S O 2 being absorbed during cycling as shown in Figure 6.13b. Therefore, increased P c o 2 resulted in less direct sulphation, less filling of macropores by the sulphate product and hence higher reactivity for C O 2 capture. SEM photos (not shown here) of calcines supported this view by indicating differences in pores for varying Pc02-The effect of total pressure and C O 2 partial pressure were confirmed with Arctic dolomite. All co-capture tests gave a very sharp decay for Arctic dolomite. Changing the total pressure was found to have little effect on the reversibility, as shown in Figure 6.10. However increasing P c o 2 greatly improved the reversibility of this sorbent, as indicated by Figure 6.14, mainly due to reduced direct sulphation. 1 6.4 Conclusions Five limestones and two dolomites were tested two thermogravimetric reactors, one at atmospheric pressure and the other at pressure up to nearly 2.4 MPa. At atmospheric pressure with no S0 2 , the various limestones performed similarly in terms of sorbent reversibility, despite appearing to sinter differently. Dolomites performed somewhat better. High-pressure tests showed similar results for Strassburg limestone, but more rapid decay for Arctic dolomite. Sorbent reversibility was generally insensitive to total pressure and to the C 0 2 partial pressure during carbonation. 150 Chapter 6 Removal of C02 by calcium-based sorbents in the presence of S02 Conversions decayed more quickly decayed conversions when S O 2 was present for all sorbents. SEM, EDX and pore size distributions revealed that direct sulphation becomes dominant after completion of an initial fast stage of carbonation, fdling larger pores by sulphation from the outside, enveloping the sorbents with an impermeable shell and thus inhibiting further carbonation and impeding subsequent calcination. For the atmospheric pressure tests, pre-hydration improved co-capture reversibility; steam helped co-capture; increasing the molar Ca/C ratio and increasing the C O 2 molar fraction from 80% to 93% had little effect on C O 2 and S O 2 co-capture. Although changing the total pressure did not change sorbent reversibility appreciably, increasing the C O 2 partial pressure was helpful. Reversibility approaching that in tests with no S O 2 present could be achieved by increasing the C O 2 partial pressure, inhibiting direct sulphation. The C O 2 / S O 2 ratio is important in determining the degree of sorbent reversibility. 151 Chapter 6 Removal of CO2 by calcium-based sorbents in the presence of SO2 co O "55 o tf) a> o E •B a» C "55 -«—* a> 1 i ex o O a> 0.8 0.6 1 g> 0.4 i 0.2 A 0 X • A O O G S dolomite Arc t ic Dolomi te St rassbun Kel ly R o d Danyang Have lock C a d o m i n • • • t u s e 5 10 Number of reaction cycles 15 Figure 6.1. Calcium utilization over 15 cycles for all seven sorbents with 212-250 urn particles and no S0 2 present. ATGR tests. Carbonation: 80% C0 2 , 20% N 2 , 850°C and 101 kPa, Fast Stage Finished. Calcination: 100% N 2 , 850°C and 101 kPa. 152 Chapter 6 Removal of CO 2 by calcium-based sorbents in the presence of SO 2 (c) Figure 6.2 High-resolution SEM pictures of calcines, with no S 0 2 present. Same test conditions as in Figure 6.1. (a) Strassburg limestone, after 15 cycles (b). Kelly Rock limestone, after 15 cycles; (c) Arctic dolomite, after 20 cycles. 153 Chapter 6 Removal of C02 by calcium-based sorbents in the presence ofSQ2 c o in 3 .2 +-» c £ 0) 0.1 0 .08 | 0 .06 After initial calcination N o S 0 2 , 15 cycles C 0 2 / S 0 2 capture, 12 cycles C 0 2 / S 0 2 capture, 1 cycle 10 100 1000 Mean pore d iamete r (nm) (a) 0.14 .1 0.12 </> •—• E ^ 0.1 -I • E -§0.08 -i £ E0.06 C 3 £ "50.04 i t > g 0.02 Q —A-No SQ,, 15 cycles •e-COj/SQ, capture, 15 cycles n After initial calcination 10 100 1000 Mean pore diameter (nm) (b) tn 3 .5 ">3 c i After initial calcination C0 2 /S0 2 capture, 7 cycles! —A— No S0 2, 20 cycles 10 100 1000 Mean pore diameter (nm) (c) Figure 6 . 3 Evolution of pore size distribution with calcination/carbonation cycling at 850°C. Test conditions: same as in Figure 6.1 for no S0 2 test and in Figure 6 . 6 for co-capture.(a) Strassburg limestone; (b) Kelly Rock limestone; (c) Arctic dolomite. 154 Chapter 6 Removal of CO 2 by calcium-based sorbents in the presence of SO 2 SS U © Si T 3 o u — "0 1 A 0 . 8 H 0 . 6 0 . 4 A 0 . 2 0 8 5 a A • A Reactor Calcination Carbonation Pt N 2 Pt C O j N 2 (MPa)(%v) (MPa) (%v) (%v) A T G R 0.1 100 0.1 80 20 P T G A 0.1 100 1.8 8 92 P T G A 1.8 100 1.8 8 92 A 6 ^ A A A A A A A A / A 0 2 0 5 1 0 15 N u m b e r of react ion cycles Figure 6.4 Cyclic performance with no S O 2 present: effect of calcination type. Calcination/carbonation cycling at 850°C with 212-250 pm Strassburg particles. </) a> o E •5 0) C 08 « o OJ O O o tf) 0 2 0.6 0.2 Reactor Calcination Carbonation ' P t N 2 Pt C 0 2 N 2 (MPa)(%v) (MPa)(%v) (%v) • A T G R 0.1 100 0.1 80 20 • P T G A 0.1 100 1.3 8 92 Q ^ n ^ P T G A 0.1 100 1.8 8 ! ! • • • • 92 n 5 10 15 Number of reaction cycles 20 Figure 6.5 Cyclic performance with no S O 2 present: effect of total pressure. Calcination/carbonation cycling at 850°C PTGA test: 212-250 pm Arctic dolomite 155 Chapter 6 Removal of C02 by calcium-based sorbents in the presence of SQ2 re *-» O (A a> o E d) B o a . re o O O 0.8 0.6 0.4 A 0.2 0 • Danyang X Kel ly R o c k A G S dolomite St rassburg , no S 0 2 • Arc t ic • Have lock O C a d o m i n O Strassburg Arc t ic dolomi te, no S 0 2 5 10 Number of reaction cycles 15 Figure 6.6 Performance of all seven sorbents for co-capture. ATGR tests, calcination/carbonation cycling at 850°C and 101 kPa with 212-250 pm particles. Sorption: 80% C0 2 , 2900 ppmv SO2, 3%v 0 2 and balance N 2 , 8 minutes for each cycle. Calcination: 100% N 2 . (Lines show corresponding results with no S0 2 present for two of the sorbents.) 156 Chapter 6 Removal of CO2 by calcium-based sorbents in the presence of SO2 0 20 40 60 80 100 Time (min) Figure 6.7 ATGR once-through tests, at 850°C and 101 kPa with 212-250 pm Strassburg limestone. Top and bottom curves and for limiting case where there was no S O 2 or C O 2 respectively. Points are for co-capture case showing total (squares) calcium utilization, utilization for C O 2 capture (triangles) and utilization for S O 2 capture (circles). 157 Chapter 6 Removal of CQ2 by calcium-based sorbents in the presence of S02 c E "5) aT s c o +•» ro c a O 0.06 0.05 0.04 0.03 0.02 0.01 0 o • o o • G S dolomite O Straussburg O Arctic O ° o o • o o • 0.1 0.2 0.3 CaSGytotal Ca • • 0.4 Figure 6.8 Relation between sulfate content and calcination rate for three sorbents. ATGR co-capture tests, same conditions as in Figure 6. (A _CD O E •6 a> C « ro O £ 3 d o o ~ o V) 0.8 0.6 "S ™ 0.4 0.2 • • 1.8 M P a V • 1.3 M P a • * • 2.4 M P a - - - Baseline, 1.8 M P a , no S 0 2 • s • *"" ^ ^ • • • A • t 0 15 5 10 Number of reaction cycles Figure 6.9 Cyclic C 0 2 retention performances in 850°C PTGA tests: effect of total pressure. Co-capture with 212-250 pm Strassburg limestone. Sorption: 8 %v C0 2 , 1125 ppmv S0 2, 3% 0 2, and balance N 2 , 4 minute for each cycle. Calcination: 101 kPa, 100% N 2 . 158 Chapter 6 Removal of CO 2 by calcium-based sorbents in the presence of SO 2 0 0.9 in a> o £ 1 ca 0.6 g 2 « o O ~ 2 0.3 to • 1.8 M P a A 2.4 M P a • 1.3 M P a - - Basel ine, 1.8 M P a , no s o 2 • • • 5 10 Number of reaction cycles 15 Figure 6.10 Cyclic C O 2 retention performances in 850°C PTGA tests: effect of total pressure. Co-capture with 212-250 pm Arctic dolomite. Sorption: 8 %v C O 2 , 1125 ppmv S O 2 , 3% 0 2 , and balance N 2 , 4 minute for each cycle. Calcination: 101 kPa, 100% N 2 . 8%v C 0 2 , 850°C. (a) (b) (c) Figure 6.11 Sulfur mapping for non-calcined ATGR samples from varied co-capture cycles, same test conditions as in Figure 6.6. (Light points mark sulfur) (a) Strassburg limestone after 12 cycles (b) Kelly Rock limestone after 15 cycles (c) Arctic dolomite, after 7 cycles 159 Chapter 6 Removal of CQ2 by calcium-based sorbents in the presence of S02 160 Chapter 6 Removal of CO 2 by calcium-based sorbents in the presence of S02 o E T 3 <D C « J a> re 6 s O 0> 0.8 0.6 X m 0.4 0.2 H 0 0 - - No S 0 2 , 8%v C O z • 8%v C 0 2 A 1 4 % v C 0 2 - 20%v C 0 2 I • A • • • _ ± ± A 10 15 Number of reaction cycles (a) —t 20 o E a? c re re O + J _ a> re >— 4-1 (/) O CO CD 0.2 42 0.1 0 0 • 8%v C 0 2 A 14%vC0 2 _ 20%vCO 2 • • * • A ± X A • • • A * J A A A (b) 5 10 Number of reaction cycles 15 Figure 6.13 Cyclic performances during PTGA co-capture tests: effect of Pco2- 212-250 pm Straussburg limestone. Sorption at 850°C and 1.8 MPa, with 1125 ppmv S O 2 , 3%v 0 2 and balance N 2 , 4 minutes for each cycle and calcination at 850°C and 101 kPa, 100% N 2 . (a) For CaO conversion to CaCOs; (b) For CaO conversion to CaS04. 161 Chapter 6 Removal of CQ2 by calcium-based sorbents in the presence of S02 w _a» o E CD C '5 £ CN O o o </> o 0.9 .« 0.6 0.3 • 8%v C 0 2 • 20%v C 0 2 - - No S0 2 , 8%v C 0 2 A A A A A 5 10 Number of reaction cycles 15 (a) 0 .2 T (A 22. o E •3 cu c « O ~ in cn cu oS 75 0.1 H • 8%v C 0 2 A 20%v C Q 2 • A A 0 5 10 Number of reaction cycles (b) Figure 6.14 Cyclic performance during PTGA co-capture tests: effect of PCo2- 212-250 pm Arctic dolomite particles. Sorption at 850°C, and 1.8 MPa, with 1125 ppmv S0 2 , 3%v 0 2 and balance N 2 , 4 minutes for each cycle and calcination at 850°C and 101 kPa, 100% N 2 . (a) For CaO conversion to CaC03; (b) For CaO conversion to CaS0 4. 162 Chapter 7 An investigation of attempts to improve cyclic CO 2 capture by sorbent hydration C H A P T E R 7 A N INVESTIGATION O F A T T E M P T S T O I M P R O V E C Y C L I C C 0 2 C A P T U R E BY SORBENT H Y D R A T I O N AND MODIFICATION A version of this chapter has been prepared for submission to Industrial & Engineering Chemistry Research. The authors are P. Sun, J. R. Grace, C. J. Lim and E. J. Anthony. 7.1 Introduction Due to the potential of calcium-based sorbents to capture C0 2 at high-temperatures, cyclic performance of sorbents has been investigated in Chapter 6 and in the literature for a variety of applications (Han and Harrison, 1994; Lin et al, 2001; Ortiz and Harrison, 2001; Abanades et al, 2004; Jukkola et al, 2005). The findings generally indicated that different types of limestones show similar trends, involving a loss in sorbent utilization as cycling progresses. In most earlier studies, C O 2 was the only gaseous reactant to react with CaO. Chapter 6, where CaO serves as a C 0 2 carrier between a fluidized bed combustor (FBC) and a calciner, demonstrated that the presence of S0 2 greatly impeded sorbent reversibility due to inhibited solid-state diffusion by a sulphate shell, formed mostly by direct sulphation. In this chapter, several other issues with respect to sorbent performance for FBC processes are investigated in detail. Steam and liquid water could conceivably be utilized in at least four ways to enhance C 0 2 capture: 1. Carbonation in the presence of steam. Although under most operating conditions, steam is not directly involved in reactions with sorbents, enhancement due to steam has been reported for CaO recarbonation (Dobner et al, 1977; Han et al, 1994). 2. Steam and liquid water can be utilized to alter CaO pore structure by producing hydroxide-derived CaO (designated here as h-CaO) through one-time hydration. It has long been known that h-CaO can achieve a higher extent of high-temperature sulphation than c-CaO (carbonate-derived CaO) (Gullett and Bruce, 163 Chapter 7 An investigation of attempts to improve cyclic C02 capture by sorbent hydration 1987; Stouffer and Yoon, 1989) because its plate-like pore structure is able to accommodate more S O 2 . Somewhat improved cyclic performance for h-CaO has been reported for cyclic C O 2 capture (Hughes et al., 2004) by Kelly Rock and Cadomin limestones. 3. Calcination of sorbents in a pure steam environment could be a convenient way to generate a relatively pure C O 2 stream (after condensation of the steam), as a key component of an overall sequestration scheme. However, the extent to which CaO sintering could be enhanced by steam needs to be investigated because steam has been reported to promote CaO sintering (Borgwardt, 1989b), thus causing decay in sorbent carbonation reactivity. 4. Similar to hydration reactivation technology applied for sulphation enhancement in FBCs (Laursen et al., 2000; 2001), hydration may also be applicable to C O 2 capture. As the sulphate layer formed during co-capture of S O 2 and C O 2 is mostly responsible for accelerated decay in cyclic C O 2 removal ability (see Chapter 6), regeneration of CaO from sulphate by applying reducing conditions could improve sorbent performance. In practice, a reducing (CO-containing) environment could be realized by feeding a low excess of oxygen in a calciner or in a char-fired furnace in a topping-cycle-based process. Preliminary data need to be collected to investigate the feasibility of this option. Moreover, a naturally occurring sorbent, e.g., a limestone, could also be artificially modified in order to achieve better reversibility. Work is needed to gain better insight on possible improvements in C O 2 capture using solid sorbents. These issues are investigated in this chapter. 7.2 Experimental details: Two commercially available sorbents, Strassburg limestone and Arctic dolomite were 164 Chapter 7 An investigation of attempts to improve cyclic CO 2 capture by sorbent hydration used in the current work. Strassburg limestone was utilized in most of the tests, whereas Arctic dolomite was only used for tests involving steam/water and in some tests involving sorbent modification. ICP chemical analyses are shown in Table 1.2 for both sorbents. An atmospheric pressure thermogravimetric reactor (TGR) was used in all tests. Details of this system are provided elsewhere (Laursen et al., 2000). All sorbents were characterized based on cyclic calcination/carbonation tests. A total flow of 1600 ml/min was maintained for all calcination and carbonation tests. A typical FBC combustor temperature, 850°C, was chosen for both calcination and carbonation in most runs. In sorption tests with no S O 2 present, 100% C O 2 was fed for carbonation, whereas in tests with co-capture of S O 2 and C O 2 , or tests with steam vapour, concentrations are given below. Except where otherwise specified, samples were exposed to 100% N 2 during each calcination stage. Two sorption durations were employed in the current work. Sorption was allowed to complete its fast stage (denoted as "fast stage finished" or FSF) in all tests, except for those involving co-capture of S O 2 and C O 2 in which a fixed time (8 or 12 min) was used for each sorption stage. Sample size was 850 mg for tests of the effect of steam or liquid water, 500 mg for CO regeneration tests and 400-500 mg for sorbent modification tests. In the CO regeneration tests, a quartz basket was used, whereas a stainless wire mesh was used in all other cases. The variation of sample size in sorbent modification tests is believed to have negligible effect on the reversibility results for the range of conditions studied. In sorbent modification tests, 212-250 um particles of calcined or uncalcined Strassburg limestone were employed as the starting material. Calcium acetate was also tested in some runs because of its good solubility in water. The particles were greatly reduced to fine powders, probably because of CaO hydration, particle breakdown, dissolution of Ca(OH)2 and subsequent 165 Chapter 7 An investigation of attempts to improve cyclic C02 capture by sorbent hydration calcination etc., except when starting with uncalcined limestone. The modifiers, all in highly pure powder or crystalline form, included Y -AI2O3, Si0 2 gel, SiC«2 sands, Ti0 2 , Zr0 2 , MgO, CaCl 2 2H 2 0, NaCl, MgS0 4-7H 20, M g C l 2 6 H 2 0 and Al(NO) 3 9H 2 0, all supplied by Fisher Scientific. High-purity precipitated calcium carbonate (PCC) powder and calcium acetate, as CaO precursors, were obtained from the same source. Kaolinite, as a modifier, was from the Quinsam Company with 45wt% Al203, 54wt% Si02 and 1% Fe203 when fully calcined and negligible CaO content. The procedure for inert doping was as follows: The CaO and modifier precursors were pre-weighed (usually < 1.5 g in total), then mixed in 70 ml of de-ioned water in a beaker. 5 ml of ethanol were next added to the beaker as a dispersant. The suspension was next stirred over a heating plate at 150°C until dry. The dried samples were then transferred into the TGR for 60 minutes of calcination at 850°C. The prepared samples were mainly characterized by cyclic C O 2 capture, with both calcination and carbonation conducted at 850°C by alternating between 100% C 0 2 and 100% N 2 feed gas. Selected samples were analyzed by EDX element distribution mapping, XRD phase detection and free calcium measurement using the ASTM C25 "fast sugar method" (ASTM C25-90, 1990). EDX element mapping was determined on a HITACHI S-3000 instrument. Before the tests, samples were first embedded in epoxy and then carefully ground and polished to expose cross-sections. A Micromeritics 9300 Poresizer was used to determine the pore size distributions of sorbents. X-Ray diffraction phase detection for the modified samples was conducted on a Siemens Diffraktometer D5000. 166 Chapter 7 An investigation of attempts to improve cyclic CO 2 capture by sorbent hydration 7.3 Results and Discussion . 7.3.1 Tests involving steam or liquid water (a) Comparison of performance of h-CaO and c-CaO in no-S02 tests. 212-250 um Strassburg limestone and Arctic dolomite were selected to study the effect of pre-hydration. Room temperature liquid water was used to hydrate the nascent calcines. Figure 7.1 shows different pore structures for c-CaO and h-CaO, which originated from 212-250 Strassburg limestone particles. h-CaO calcines clearly show plate-like pore structures, as reported by Gullett and Bruce (1987). However, as shown in Figure 7.2, h-CaO (derived from Strassburg limestone) performed only a little better than c-CaO when subjected to calcination/carbonation cycles (Figure AVTII.l in Appendix AVUI presents similar result for calcines derived from Arctic dolomite). The pore size distribution in Figure 7.3 reveals that the initial h-CaO calcines (derived from Strassburg limestone) contained somewhat more pore volume than c-CaO. After 15 cycles, both lost most pore volume for <220 nm pores while gaining pore volume for >220 nm pores. Given that sintering is responsible for the decay of cyclic performance and change of pore structure during cycling (see Chapters 4 and 6), it appears that both types of CaO sinter similarly during calcination/carbonation cycles. The slight improvement for the h-CaO is probably due to its pore structure allowing better expansion when producing a swelling product, as reported for S 0 2 capture (Stouffer and Yoon, 1989; Gullett and Bruce, 1987). ' . , c-CaO and h-CaO derived from Arctic dolomite showed negligible improvement in reversibility and no difference in pore size distribution (Figure AVm.2 in Appendix AVUI presents the pore size distribution results for calcines derived from Arctic dolomite) between the 167 Chapter 7 An investigation of attempts to improve cyclic C02 capture by sorbent hydration two types of sorbents. The lack of improvement was probably due to the interspersed content of MgO. (b) Calcination in steam in no-SOi tests The affinity between H 2 O molecules and CaO sites could affect calcination of limestone (Young, 1966) because steam could easily adsorb on the CaO surface, thereby reducing C 0 2 adsorption, and hence lowering the incipient calcination temperature. In this study, steam was used to cyclically calcine Strassburg limestone. In the TGR tests, the steam volumetric fraction was maintained above 95%, with only a small flow of N 2 to protect the top of the reactor. Calcination was conducted at both 850 and 900°C. The cyclic performance in Figure 7.4 shows no appreciable effect of sweep gas ( N 2 or steam) or steam temperature. The pore size distributions (Figure 7.5) generally showed similar pore evolution history for N 2 and steam. No enhancement of steam vapour on sintering could be detected. This does not contradict the finding of Borgwardt (1989b) that steam catalyzed CaO sintering to the same degree as C O 2 , because, regardless of whether steam or N 2 was used as the sweep gas during CaC03 calcination, the sites for C O 2 or steam adsorption on the nascent CaO sites are fixed. This suggests that steam could be used as the sweep gas in a calciner. (c) Effect of steam on carbonation with no S O 2 present. No appreciable enhancement by steam was observed for either sorbent (Figures AVIII.3 and 4 in Appendix AVTII presents the results for the effect of steam on cyclic carbonation for Strassburg limestone and Arctic dolomite, respectively). This is consistent with the multiple-cycle performance results presented by Han et al.(1994), although they found some enhancement for the first carbonation. Note that the possible adsorptive effect of steam on CaO sites discussed above may not be applicable to carbonation, because the fast carbonation reaction supplants any adsorption of water vapour on CaO sites. 168 Chapter 7 An investigation of attempts to improve cyclic CO 2 capture by sorbent hydration (d) h-CaO and steam in co-capture of C O 2 and S O 2 Given the rapid decay of sorbents during co-capture of S O 2 and C O 2 (see Chapter 6), tests were also conducted starting with h-CaO adding steam during carbonation/sulphation. Results are shown in Figure 7.6 for both sorbents. The baseline test conditions for this co-capture study starting from fresh sorbents were as described in Chapter 6, i.e., sorption at 850°C with 80% C O 2 , 2900 ppm S O 2 , 3% 02 and balance N 2 , with 8 minutes for each cycle, and calcination at 850°C with 100% N 2 . Strassburg-derived h-CaO showed improvement not only for C O 2 capture (Figure 7.6a), but also for S O 2 . The enhanced co-capture by h-CaO is believed to be due to the better expansion ability of the plate-like pores of h-CaO relative to c-CaO. Adding steam during the co-capture tests improved the reversibility of Strassburg limestone. It is speculated that this improvement is due to an interaction of H 2 O with the sorbent surface, possibly through transient surface hydroxyl group formation, commonly invoked to explain water vapour enhancement of CaO or MgO sintering (Anderson, 1964; Borgwardt, 1989b). Arctic dolomite gave similar results. Its performance is shown in Figure 7.6b for C O 2 removal. Hydroxide-derived CaO showed better performance for both C O 2 and S O 2 capture, probably for the same reason as for the limestone. However, the presence of steam only appreciably helped the co-capture performance over the first few cycles. Figure AVTO.5 presents the results of the effect of steam or hydration treatment on S O 2 part capture during the cyclic co-capture tests. 169 Chapter 7 An investigation of attempts to improve cyclic CO 2 capture by sorbent hydration (e) Effect of hydration on C O 2 capture with S O 2 present Since steam reactivation can fracture calcium sulphate layers (Laursen et al., 2000; 2001) thereby improving sulphur capture, hydration was also applied in an effort to improve C 0 2 capture. Strassburg limestone, classified as an unreacted-core type limestone (Laursen et al., 2000) under typical FBC sulphation conditions, was found to have significantly improved S0 2 capture when reactivated by hydration. Steam hydration was first used to fracture the sulphate shell formed during co-capture. In the current studies, steam with a partial pressure of ~95 kPa at 300°C with the balance N 2 , was employed in all hydration steps for 30 minutes. Three sets of tests were carried out to reveal the effect of hydration on subsequent cyclic C 0 2 capture. In Test A, 1 h of simultaneous sulphation and carbonation was followed by a series of steps: complete calcination in N 2 , 8 min of co-capture (one cycle including both sorption and calcination), a complete calcination, a steam hydration step and two cycles of co-capture. Test B differed from test A in that, after the first 8 min of co-capture, the samples were quickly withdrawn from the reactor and not returned until the reactor temperature reached 300°C, so that steam hydration was applied before any appreciable calcination; In test C, the initial one-hour co-sorption in A was replaced by a sulphation step (same S0 2 /0 2 concentration, but with no C 0 2 present, balanced by N2), followed by 8 min of co-capture, while the other steps were the same as for test A. The conversions of CaO to CaC0 3 and CaS0 4 after each sorption step are compared in Figure 7.7, where the baselines are also plotted on a cumulative sorption time basis for a typical cyclic S0 2 /C0 2 co-capture test. Figure 7.7b shows that the initial one-hour and the first 8-minute sorption in all these tests led to a similar extent of sulphation, but Figure 7.7a reveals that the sulphate layer formed in test C served as a greater barrier to carbonation during the following 8-minute co-capture than in the 170 Chapter 7 An investigation of attempts to improve cyclic C02 capture by sorbent hydration other two tests. This probably occurred because the calcination step after 1-hour co-capture in tests A and B could expose more CaO surface area as revealed by the lowest carbonation extent achieved during 8 min co-capture. However, after reactivation, test C showed greater improvement for C 0 2 capture than test A, although a similar degree of improvement was observed for S O 2 removal. This indicates that the reactivation effect of CaO by hydration is more effective in test C than in test A, probably because of the different porosity of sulphate layers formed in the two tests. For test B, it is not surprising that the hydration effect was very limited because the hydration reaction for the calcine core must be greatly inhibited by the presence of an outer carbonate layer if it is not calcined prior to hydration. However, it is also shown in Figure 7.7a that, after reactivation by hydration, the sorbent performance for C O 2 removal continued to decay with further cycling. (f) Effect of hydration on C O 2 capture with no S O 2 present As noted above, the effect of hydration is mainly to break the sulphate shell formed during co-capture. Hydration can also be utilized to break a carbonate shell or to counteract the meso-pore-eliminating sintering process by generating pores for C O 2 capture. To test the effect of hydration on a carbonate shell, non-cycling carbonation tests were conducted. During continuous carbonation (e.g., after 10 or 60 min), the carbonated sorbents were either subjected to in-situ hydration in the TGR with 95% pure steam (balance N2) at 300°C or removed from the reactor and subjected to water hydration. Water hydration was performed outside the reactor with room temperature de-ioned water poured continuously over the sorbents held in a stainless wire mesh basket for 10 minutes. The hydrated sorbents from both tests were then heated to 850°C i n a pure C O 2 stream. During heating, dehydration should proceed first, producing CaO at a lower temperature. The CaO was then carbonated during further heating for 171 Chapter 7 An investigation of attempts to improve cyclic CO 2 capture by sorbent hydration a pre-determined duration. The results are plotted based on cumulative carbonation time in Figure 7.8. Compared with continuous carbonation, CaO conversion after hydration treatment experienced no improvement. This implies that, unlike the hydration of partially sulphated samples, the carbonate layer prevented any hydration reaction on the unreacted CaO core. As noted above, hydration is difficult in the presence of a carbonate shell. This confirms that CaC03 product differs from CaS0 4. As demonstrated in Chapter 6, <220 nm pores are of dominant importance for the fast stage of carbonation. However, these are continuously eliminated by sintering during calcination/carbonation cycling. Hydration can also be used as an alternative way to re-generate pores in this range for further calcination/carbonation cycles. The results are summarized in Figure 7.9. The baseline test shows the cyclic performance of 212-250 pm Strassburg particles resulting from typical calcination/carbonation cycles. After 15 cycles, the calcined sorbents were treated for 10 or 30 minutes with steam at 300°C, steam at 150°C or water at room temperature. It can be seen that, for all methods examined here, 10 minutes of water hydration (filled circles) gave the greatest improvement in cyclic C O 2 capture. As a result of the second water hydration (open circles), the sorbent's performance increased further. This indicates that hydration, if effective, is able to crack highly sintered sorbents and re-generate the <220 nm pores, which are favourable for subsequent carbonation; it also reveals that repeated hydrations could reduce sintering to some degree and improve sorbent reversibility, probably due to reduced particle size and reduced calcination time, as revealed in Chapter 4. Other tests starting from long-time-sintered (24 hours in an 1100°C oven) 212-250 um Strassburg particles gave similar results, as indicated in Figure 7.9. The order for the effectiveness of hydration conditions can be ranked as. liquid water >150°C steam >300 °C steam. 172 Chapter 7 An investigation of attempts to improve cyclic CO 2 capture by sorbent hydration 7.3.2 Regeneration of CaO from CaS04 using C O during calcination Given that the CaS04 shell is the major reason for decayed reversibility during co-capture of S O 2 and C0 2 , CO was explored as a possible way to reduce CaS04 to CaO to improve C 0 2 capture. In this test, CO was fed during each CaO regeneration step. When CaS04 reacts with CO, two routes are possible, CaS04+4CO=CaS+4C02 (Undesired reaction) (7.1) CaS0 4 + CO = CaO + S0 2 + C O 2 (desired reaction) (7.2) Chen and Yang (1979), Mattisson and Lyngfelt (1999) and Okumura et al. (2003) found that the fraction of solid products CaS or CaO vary during CaS04 reduction in CO, with CaO production favoured by higher temperature and lower CO / C O 2 ratio. For the cyclic tests, pure CaS04 supplied by Fisher Scientific was used as the starting material. The reduction temperature was selected as 950°C. Gas concentrations are shown in Table 7.1. After complete reduction, the feed gas was switched to air. If any CaS was produced, it was then oxidized immediately into CaSO4, thus causing a mass increase. On the other hand, CaO would not undergo any mass change in the presence of air. The results are summarized in Table 7.1, showing that a reducing condition with a C O / C 0 2 molar ratio of 0.12 produced negligible CaS. This ratio was therefore selected for further testing. Table 7.1 Regeneration conditions and CaS/CaO test Temperature C 0 2 N 2 CO CO CO /CO2 CaS product ml/min ml/min ml/min (%v) molar ratio 85G°C, - - - . .. _ _ too slow 950 ° C 0 2100 180 8%v Infinite Yes 950 ° C 314 1700 90 4%v 0.28 Yes 950° C 314 1700 40 1.9%v 0.12 No 173 Chapter 7 An investigation of attempts to improve cyclic COj capture by sorbent hydration This CO regeneration study, unlike the other tests, was performed in a quartz basket with a platinum suspension wire. The change in basket could lead to a change in mass transfer, affecting the kinetics. Therefore new tests conditions were selected, with 500 mg fresh Strassburg limestone at the start of each run and 12 minutes for each sorption stage. Three runs were preformed as summarized in Table 7.2: (1) cyclic C 0 2 capture with no S O 2 present (denoted no-S02 test); (2) C 0 2 / S 0 2 co-capture with all calcinations in pure N 2 (designated normal co-capture); (3) C 0 2 / S 0 2 co-capture with CO, C 0 2 and N 2 all present during the regeneration step (Co-capture with CO reduction). Except for the 10th regeneration where complete calcination in N 2 was performed first, followed by reduction using a mixture of C O 2 , CO and N 2 , all other regeneration steps involved mixed C O 2 , CO and N 2 so that CaC03 and CaS04 decomposed simultaneously. In the co-capture test with CO reduction, reduction of CaS04 was much slower than that of the calcination of CaC03, probably because of the low CO concentration (One cycle of co-capture followed by a CaC0 3 calcination in N 2 then by CO reduction is presented in Figure AVTfI.6 in Appendix AVTJT). Figure 7.10 summarizes all three runs. Because in all CaO regeneration steps (except for the 10th regeneration step), CaC03 calcination occurred simultaneously with CaS04 decomposition, it is difficult to distinguish the calcium utilization involving C O 2 from that involving S O 2 . Therefore, Figure 7.10 plots the ratio of mass increase after each sorption cycle to the initial calcine mass. As the mass increase due to S0 2 capture is much smaller that that due to C 0 2 capture, as revealed by the 10th regeneration step, this plot roughly reflects the sorbent performance for cyclic C 0 2 removal. It shows that, due to the presence of S0 2 , the reversibility decays as a result of cycling. When CO was applied during each cycle to regenerate CaO, there was only a small improvement in reversibility. This is probably because the inhibiting effect of S O 2 on C O 2 capture was not altered, even when CO 174 Chapter 7 An investigation of attempts to improve cyclic CO 2 capture by sorbent hydration was present in the regeneration step, although the reduction of CaSCU diminished accumulation of the sulphate layer. Table 7.2 Test conditions of CaSC^ regeneration by CO Sorption (850°C) Calcination/regeneration (950°C) Condition C 0 2 S 0 2 Other components C 0 2 CO N 2 (%v) (ppmv) (%v) (%v) (%v) (%v) 1. No S 0 2 test 80 0 Balance N 2 - ' • - 100% 2. Normal co-capture 80 2900 3 0 2 , balance N 2 - - 100% 3. Co-capture with 80 2900 3 0 2 , balance N 2 15% 1.9 Balance CO reduction 7.3.3 Modification of sorbents (no S O 2 present) Sorbent reversibility is the key to successful processes utilizing sorbents for C O 2 removal. Chapter 4 explains that limestones such as Strassburg unavoidably experience similar decay behaviour in calcination/carbonation cycles due to enhanced sintering during calcination. CaO sintering is known to be due to solid-state diffusion of ions (Borgwardt, 1989a, b). For CaO sintering, volume diffusion is usually the most important mechanism, ion mobility among lattices being a decisive factor (Borgwardt, 1989a; German, 1996). Aliovalent ions addition usually enhance sintering in ionic compounds (German, 1996; Borgwardt, 1989a); C O 2 diffusing outwards during the calcination of sorbents can catalyze sintering (Chapter 4). As discussed in Chapter 6, dolomitic sorbents perform more reversibly than calcitic ones, probably due to the inhibiting effect of well-dispersed inert MgO on reactant ionic diffusion. An attempt was made to artificially modify natural sorbents by doping inter atoms/ions to reduce 175 Chapter 7 An investigation of attempts to improve cyclic CO 2 capture by sorbent hydration ionic mobility and improve thermal stability. Another possible technical route for sorbent modification is to physically or chemically bond some CaO sites by adding various reactive modifiers. The tests on modified sorbent reversibility are summarized in Table 7.3. Note that the CaO fraction has been converted to a fully calcined sorbent at 850°C, with all possible gaseous products ( C O 2 , H 2 0, N0 2, N 2 O ) driven off. (Figure AVTH.7 presents the EDX dopant element mapping, showing good mixing of dopants with calcium sorbents.) Cyclic performance using fresh Strassburg limestone (212 pm) is also shown, and that for fresh Arctic dolomite appears in Figure 7.lie. Tests using A1(N03) 39H2O encountered difficulties during stirring over the heating plate (~150°C) because of the low melting point of the nitrate (73°C). The final modified product showed little carbonation reactivity probably, because of pore-blocking by intermediate melts. However, good reversibility was reported by Li et al (2005) using a similar method. The discrepancy from the current study indicates that the synthetic method is sensitive to conditions at each preparation step. Results for sorbents modified using Y-AI2O3 are displayed in Figure 7.11a. Most of the sorbents showed reduced C O 2 capture capacity, but run 4 with 1:1 molar ratio of CaO to A I 2 O 3 showed good reversibility. XRD phase detection (see Figure 7.12) for calcined sorbent from run 4 indicated that the major solid-state reaction product is Mayenite (Cai2Ali 4033). A similarly good reversibility was achieved by synthetically mixing pure CaC0 3 and A1(N03)3'9H20, producing the same solid product Cai2Ali4033 (Li et al, 2005). Results from runs 4 and 5 based on free calcium measured by the fast sugar method (ASTM C25-90, 1990) are also presented in Figure 7.11a. This shows that when the part of CaO consumed by reaction is considered, the residual CaO is a reversible sorbent, although still subject to decay. (The calcium utilization exceeding 100% is due to analytical errors.) Lowering AI2O3 content in the 176 Chapter 7 An investigation of attempts to improve cyclic CO2 capture by sorbent hydration mixture did not yield satisfactory data. Tests starting with calcium acetate did not generate good reversibility. Thus an optimal mixing ratio may exist. Some results of tests on two forms of Si0 2 -silica gel and silica sand, are shown in Figure 7.11b. Silica sand is crystalline, and non-porous, whereas silica gel is non-crystalline and highly porous. Thus silica gel develops more reactive OH groups in a water solution. The cyclic capture results showed that those starting with silica gel (e.g. 1:1 ratio of Si0 2 to CaO) had very low carbonation ability, whereas the tests with the same composition, but starting with sand, showed a higher capture capacity. XRD analysis (Figure 7.12) indicated that the solid-state product was predominantly silicate, i.e. Ca3Si04. This implies that different proportions of CaO must be consumed by solid-state reactions between CaO and Si0 2. The different performances of the different forms of Si0 2, could indicate different degrees of basicity in water, and hence different interactions between CaO surfaces and hydroxyl groups. Tests where no carbonation occurred may be due to over-consumption of CaO by solid-state reaction, thus necessitating an optimized mixing ratio. The test starting with calcium acetate and silica sand showed somewhat better reversibility, but decay was not arrested. The run with Kaolinite (run 20) was conducted because Kaolinite is a rich source of Al203 and Si0 2. However, this modification is not ideal. XRD analysis only shows a large peak of free lime and small peaks for Anorthite CaAl 2 (Si0 4) 2 and Mullite 3A1203 2Si02. Note that fully calcined Kaolinite could also be amorphous and thus XRD-transparent. Figure 7.1 lc shows test results using Zr0 2 , which has been reported to be able to improve the thermal stability of CuO as an oxygen carrier in chemical looping combustion (Vazquez et al., 2005). However, no improvement was observed with this modifier in our study. XRD analysis in Figure 7.12 indicates that no solid-state reaction occurred between CaO and Zr0 2 . 177 Chapter 7 An investigation of attempts to improve cyclic CO 2 capture by sorbent hydration Figure 7.1 Id presents results for a test with MgO added. XRD analysis in Figure 7.12 does not indicate any solid-state reaction between CaO and MgO. Modified sorbent showed somewhat better reversibility, probably because the mechanical mixing reduced CaO ionic mobility to some degree. However, the reversibility compared to that of calcined dolomite (Figure 7. lie) was still poor. This implies that only good dispersal on an atomic level, such as that found in naturally occurring dolomites, is capable of resisting ionic volume diffusion effectively, thereby leading to improved reversibility. Figure 7.1 le presents test results for calcined Artie dolomite and T i 0 2 mixed with Strassburg limestone. The calcined dolomite/limestone mixture not surprisingly shows cyclic sorbent performance between the baseline tests for the dolomite and the limestone alone. With T i 0 2 present, however, much of the CaO was consumed (run 28) by solid-state reaction to produce C a T i 0 3 as revealed by XRD analysis (Figure 7.12). Decreasing T i 0 2 content (run 29) could necessarily release more free CaO for C 0 2 capture, but not result in reversible capture. However CaTi03 as a modifier was reported to be able to stabilize natural limestones (Aihara, et al., 2001), it was probably the Sol-Gel preparation method employed in the reference work that made the different modification effect. Precipitated calcium carbonate (PCC) has been proposed as a reversible sorbent for C 0 2 capture (Gupta and Fan, 2002). Commercial PCC was therefore tested. It is anticipated that a PCC of high CaC0 3 purity should lack impurities-derived lattice defects, and thus be low in lattice diffusivity and sintering rate. However, the decay trend in this case was similar to that for the other modifications, as shown in Figure 7.1 le (run 33). This disappointing result probably arose because high-temperature-induced lattice defects still dominate, again enhancing volume diffusion during sintering. Note that there is likely to be an effect of the PCC preparation method, accounting for the good reversibility reported by Gupta and Fan (2002). 178 Chapter 7 An investigation of attempts to improve cyclic CO 2 capture by sorbent hydration Figure 7.1 If provides results for a variety of other ionic compounds with the same valence as calcium (except for the NaCl). Our previous test (Sun et al., 2004) with dopants of an aliovalent ionic compound N a 2 C 0 3 showed that the sorbent became very non-porous, with greatly diminished carbonation ability. NaCl was chosen for the test is because of improvement reported by doping NaCl in natural limestones (Salvador et al., 2003). The results show that, except for the run using MgS04, where normal reversibility was observed, all other dopants tested greatly reduced the C 0 2 capture capacity. However, a number of runs indicated good reversibility, but at the expense of much reduced early capture capacity (e.g. runs 39 and 40). From the XRD analysis, no solid-state reaction was observed between CaO and the other modifiers in these cases. Measurement for free lime with the fast sugar method indicated that the free calcium content was very close to that of the original mixture, further confirming the lack of reactions. Therefore, the decline in C 0 2 capture capacity is likely due to pore-blocking, considering that the temperature of operation (850°C) exceeded the melting points of all the chlorides (675°C for CaCl 2 , 800°C for NaCl, 714°C for MgCl 2 ) (Barin and Knacke 1973; Barin etal., 1977). 7.4 Conclusions Strassburg limestone and Arctic dolomite were investigated at a typical fluidized-bed combustor temperature of 850°C. Addition of steam during carbonation with no S0 2 present did not arrest the decay caused by cycling, but reduced the rate of decay somewhat for C0 2 /S0 2 co-capture. One-time treatment of calcined limestone could produce h-CaO, which showed enhanced C 0 2 capture ability with or without S0 2 present, but appeared to be dominated by CaO sintering. 179 Chapter 7 An investigation of attempts to improve cyclic CO2 capture by sorbent hydration Calcination with steam as the fluidizing gas appears to be appealing as it did not augment sintering, and steam calcination can easily produce a relatively pure C O 2 stream in practice after condensing the steam. Steam and water hydration for partially carbonated sorbents did not show enhanced carbonation, probably because of a compact carbonate layer preventing the diffusion of steam. Addition of steam and water to calcined and highly sintered samples showed improved C O 2 capture ability, attributable to regeneration of <220 nm pores. Among the hydration methods, water achieved better results than steam. Intermittent water hydration during cycling could lead to improved reversibility, probably related to particle size reduction. Steam/water hydration can also be utilized to fracture sulphate layers formed during co-capture of CC^and S O 2 . The sulphate layer can also be fractured by using CO to regenerate CaO from CaS04. However, CO regeneration is very slow, in part because the CO concentration has to be kept low to obtain CaO rather than CaS as the by-product. CO regeneration can improve sorbent reversibility somewhat, but not significantly. Various salts, which are inert to C 0 2 capture, were added to natural limestone in an effort to find a dopant which could improve sorbent reversibility. Most of the additives did not achieve good results. A -1:1 molar ratio of CaO to Al203 showed the most promising results on a free-lime basis. 180 Chapter 7 An investigation of attempts to improve cyclic CO 2 capture by sorbent hydration Table 7.3 Summary of experiments to test possible additives to improve sorbent reversibility. Run number Precursors Figure showing the result Original mixture CaO mass fraction (converted to basis of fully calcined sorbent (wt%) Measured CaO mass fraction after being fully calcined by fast sugar method (wr%) 1 Limestone +A1(N0 3 ) 3 -9H 2 0 Low reactivity, not shown 36.3 2 Calcined limestone +A1(N03) 5 -9H 2 0 Low reactivity, not shown 36.8 3 Limestone +A1 20 3 Figure 7.1 l a 36.3 4 Calcined lime+Al 20 3 Figure 7.1 l a 37.0 14.0 5 Calcined lime+Al 20 3 Figure 7.1 l a 53.7 31.0 6 Calcined lime+Al 20 3 Figure 7.1 l a 83.1 37.0 7 Acetate + A1 2 0 3 Low reactivity, not shown 37.3 8 Acetate + A1 2 0 3 Figure 7.1 l a 81.7 9 Limestone+SiOj gel Low reactivity, not shown 49.6 10 Calcined lime+Si0 2 gel Low reactivity, not shown 49.6 11 Calcined lime+Si0 2 gel Figure 7.1 lb 96.1 12 Aeetate+Si02 sand Low reactivity, not shown 50.5 13 Aeetate+Si02 sand Low reactivity, not shown 81.8 14 Calcined lime+Si0 2 sand Figure 7.1 lb 49.1 15 Calcined lime+Si02 sand Figure 7.1 lb 81.8 16 Acetate+Si02 sand Figure 7.1 lb 48.2 40.0 • 17 Acetate+Si02 sand Figure 7.1 lb 65.1 18 Acetate+Si02 sand Figure 7.1 lb 31.8 19 Acetate+Si02 sand Figure 7.1 lb 81.7 20 CaO+Kaolinite Figure 7.1 lb 90.0 21 CaO+Zr0 2 Figure 7.1 l c 30.0 22 CaO +Zr0 2 Figure 7.11c 77.6 23 Acetare+Zr02 Figure 7.1 l c 31.2 24 Acetare+Zr02 Figure 7.11c 82.0 25 CaO+MgO Figure 7.1 Id 54.6 26 CaO+MgO Figure 7.1 Id 28.2 27 Acetare+MgO Figure 7.1 Id 58.3 28 CaO+Ti0 2 Low reactivity, not shown 42.0 29 CaO+Ti0 2 Figure7.11e 77.3 30 Acetate+Ti02 Low reactivity, not shown 45.0 31 Limestone+calcined dolomite Figure 7.1 le 54.5 32 Calcined limestone+calcined dolomite Figure 7.1 le 69.2 33 Pure P C C Figure 7.1 le 100 34 Limestone+MgS0 4.7H 20 Figure 7.1 I f 73.6 35 Calcined limestone+NaCl Figure 7.1 I f 91.8 36 Calcined limestone+NaCl Figure 7.1 If 95.7 37 Limestone+MgCl2 Figure 7.1 I f 61.0 38 Limestone+MgCl2 Figure 7.1 I f 76.0 39 Limestone+CaCl2 Figure 7.1 I f 72.3 68.5 40 Limestone+CaCI2 Figure 7.1 I f 85.1 81.0 41 Limestone+C aC l 2 Figure 7.1 I f 94.3 181 Chapter 7 An investigation of attempts to improve cyclic CQ2 capture by sorbent hydration o c u 0.8 0.6 { g 0.4 0.2 • c-CaO • h-CaO • U u • • • • • p 5 10 Number of reaction cycles 15 Figure 7.2 Cyclic performance: comparison of c-CaO and h-CaO (no S 0 2 present, sorbent derived from 212-250 um Strassburg limestone). Test conditions: 850°C calcination and carbonation, carbonation in 100% C 0 2 , calcination in 100% N 2 . Fast stage of carbonation finished. 182 Chapter 7 An investigation of attempts to improve cyclic CQ2 capture by sorbent hydration o > G O 0.16 H 0.12 fi^ 0.08 a es o c 0.04 10 - B ~ c-CaO, initial calcination -A -c-CaO, 15th cycle h-CaO, initial calcination -^-h-CaO, 15th cycle 100 tvban pore diarreter (nm) 1000 Figure 7.3 Pore size distribution: comparison of c-CaO and h-CaO (no S O 2 present, sorbent derived from Strassburg limestone). Samples are the same as in Figure 7.2. o o o 0.6 4 0.4 4 ~ 0.2 0 + • Calcine in N 2 , 8 5 0 ° C • Calcine in steam, 8 5 0 ° C • Calcine in steam, 9 0 0 ° C D 1 • * o B * * • • A 10 15 Number of reaction cycles Figure 7.4 Effect of steam calcination on cyclic capture (No S O 2 present, 212-250 pm Strassburg limestone Strassburg limestone). Test conditions: 850°C calcination and carbonation, Carbonation in 100% C O 2 . Calcination in 95% steam, balance N 2 . Fast sta of carbonation finished. 183 Chapter 7 An investigation of attempts to improve cyclic CO2 capture by sorbent hydration Figure 7.5 Pore size distribution: effect of steam calcination (No S0 2 present, Strassburg limestone). Same samples as in Figure 7.4. 184 Chapter 7 An investigation of attempts to improve cyclic C02 capture by sorbent hydration 73 1 .a ca \—i O" O o o s 0.8 T 0.6 4 0.4 0.2 4 Baseline, no S 0 2 A Co-capture with 14% steam O Co-capture, typical conditions • h-CaO 5 10 Number o f reaction cycles (a) Strassburg limestone 15 CD O 03 O 0.9 •£ 0.6 o E TJ CD CD o o o c/) _cu o 0.3 — Baseline, no S 0 2 A Co-capture with 14% steam O Co-capture, typical conditions • h-CaO 5 10 Number of reaction cycles 15 (b) Arctic dolomite Figure 7.6 Calcium utilization for C 0 2 capture: effect of varying operating conditions, 212-250 pm particles Test conditions: 850 °C calcination and sorption. Sorption in 80% C0 2 . 3% 0 2, 2900 ppm S0 2 and balance N 2 or with steam. Calcination in 100% N 2 . 8 minutes for each sorption. 185 Chapter 7 An investigation of attempts to improve cyclic C02 capture by sorbent hydration U <4-l O o -a G "3 o u CD 1 0.8 0.6 -i 0.4 0.2 0 -f 0 A • O • o X X Test A,before hydration Test A, after hydration Test B, before hydrarion Test B,after hydration Test C, before hydration Test C, after hydration Baseline, cyclic co-capture X X X X X • • x x o x x X 20 40 60 80 100 Cumulative sorption time (min) (a) "o 6 73 o d" CJ 73 0.3 0.2 0.1 4 A Test A , before hydration • Test A , after hydration O Test B , before hydration • Test B , after hydration • O • Test C , before hydration Test C , after hydration • X x x I X Baseline, cyclic co-capture x x l X X x x • X 20 40 60 80 Cumulative sorption time (min) (b) 100 Figure 7.7 Effect of hydration on co-capture (CO2+SO2). 212-250 pm Strassburg limestone. Procedure: 1-h co-capture or S 0 2 sorption followed by 30 min hydration and 8 min cycles of re-capture: (a).Conversion of CaO to CaC0 3 ; (b) Conversion of CaO to CaS0 4 . 186 Chapter 7 An investigation of attempts to improve cyclic CO 2 capture by sorbent hydration Figure 7.8 Effect of intermediate hydration of carbonates on further carbonation (no S0 2 present) A: 10 min carbonation plus 30 min steam hydration at 300°C plus 30 min carbonation; B: 60 min carbonation plus 30 min steam hydration at 300°C plus 30 min carbonation; C. 60 min carbonation plus 10 min liquid water steam hydration at room temperature plus 30 min carbonation. 850°C for both calcination (100% N2) and carbonation (100% C0 2). 187 Chapter 7 An investigation of attempts to improve cyclic C02 capture by sorbent hydration 1.2 o S 0.8 0.6 0.4 0.2 Q X x 8 X Baseline, before hydration X After 15 cycles+30 min 300°C steam hydration • After 15 cycles+30 min 150°C steam hydration 0 After 15 cycles+10 min water hydration 0 After 15 cycles+10 min water hydration+2 cycles+10 min water hydration ^ After long-sintering+10 min water hydration K After long-sintering+300°C min water hydration After long-sintering+150°G min water hydration O o o o o o o o o 0 '• — 1 — — i — ' i . . . i _ J _ i I 0 2 4 6 8 10 12 14 16 Cycle number Figure 7.9 Effect of intermediate steam (95%v) or water (100%) hydration at various temperatures of sintered calcines on cycling (no S0 2 present). 850°C for both calcination (100% N2) and carbonation (100% C0 2). 188 Chapter 7 An investigation of attempts to improve cyclic CO 2 capture by sorbent hydration 0.8 O > o o 3 0.6 O S 0.4 S -*-» S 0.2 IJie.mas5.jalio.fQr.lD.Q%i.CaD..c.Qnyersioa.to.,CaC.0.3... • Co-capture, with C O reduction 0 Co-capture, with no C O reduction - X - No S0 2 present • - ® ° 9 • *~-*---x I 0 0 0 S 1 L 1 i 5 10 Cycle number 15 Figure 7.10 Variation of mass increase due to sorption divided by total initial mass of with cycling. Test conditions are provided in Table 7.3. 189 Chapter 7 An investigation of attempts to improve cyclic CO2 capture by sorbent hydration o u 0 2 0.6 S 0.4 0.2 A A A A A A A A o v o 0 o ^ o 0 ° ° O o - B a s e l i n e , f r e s h l i m e s t o n e O R u n 2 9 A R u n 3 2 O — — B a s e l i n e , fresh d o l o m i t e • R u n 3 1 O R u n 3 3 Cycle number 0.6 0.4 0.2 (f) — — B a s e l i n e , fresh l i r n e s t o r i e O . R u n 3 5 0 R u n 3 7 • R u n 3 9 • R u n 4 1 • : 1 ; • * * & S i X - 0 • A X k R u n 3 4 R u n 3 6 R u n 3 8 R u n 4 0 0 0 o s i 10 Cycle number Figure 7.11 Effect of various dopants on CaO reversibility in cyclic calcination/carbonation. Cycling conditions: 850°C for calcination (in 100% N2) and carbonation (in 100% C0 2) (a) y-Al203;(b) Si0 2 and Kaolinite; (c) Zr0 2; (d) MgO; (e) dolomite, Ti0 2 , precipitated calcium carbonate; (f) other dopants as identified in Table 7.3. 190 R O i s o >> S a. <3 O o g a. 5 c o •43 .5 fc 6 c C For all the runs 0 1 0 0 1 0 Run 20 CaO+ Kaolinite (SiCV A I 2 O 3 peak cannot be clearly viewed) Run 35 CaO+NaCl (NaCl peak cannot be clearly viewed) Run22CaO+ZrQ2 Run 15 CaO+ SiO? sand (Ca 2 Si04 peak cannot be clearly Run28CaO+Ti02 Run 11 CaO+SiOj Gel (Ca2Si04 peak cannot be clearly viewed) CaTiQj — ^ Run25CaO+MgO Run4CaOfAl203 CaO JL X i t I I 1 1 I I I I I I I Figure 7.12 XRD analyses results for typical runs Chapter 8 Sequential capture of CO 2 and SO 2 under FBC conditions C H A P T E R 8 S E Q U E N T I A L C A P T U R E O F C 0 2 AND S 0 2 UNDER F B C CONDITIONS A shortened version of this chapter has been accepted by Environmental Science and Technology, in press. Authors: P. Sun, J. R. Grace, C. J. Lim and E. J. Anthony. 8 .1 Introduction Initiated by climate change concerns and the need to reduce CO2 (greenhouse gas) emissions by the power industry, several novel CO2 capture concepts have been proposed based on calcium-based sorbents as CO2 carriers between a combustor (also acting as a carbonator) and a calciner (acting as a sorbent regenerator) (Shimizu et al., 1999; Abanades et al., 2003). Given the thermodynamics of the CaO-CCh carbonation reaction, higher C02-removal could be achieved in pressurized fluidized bed combustors (PFBC) than at atmospheric pressure. The cyclic CO2 capture ability during calcination/carbonation cycles, referred to as sorbent reversibility in this study, is a key factor affecting process economics for calcium-based CO2 sorbents (Lin et al., 2001). Most studies have focused on the reversibility of natural sorbents in the presence of a single gaseous reactant, CO2 (Abanades, 2002; Abanades and Alvarez, 2003; Abanades et al., 2004; Alvarez and Abanades, 2005). However, the work in Chapters 5 and 6 using an atmospheric TGR, and previous work in a bubbling fluidized bed (Ryu et al, 2006) under FBC conditions have demonstrated that SO2 greatly decreases the reversibility of calcium-based sorbents compared to a C02-only environment. An important factor for minimizing the loss of sorbent reversibility is to use as large a molar ratio of CO2/SO2 as possible in combustors. However, the total SO2 is pre-determined for a combustor of a fixed fuel type and demand. Therefore, methods of separately removing SO2 and CO2 in sequential reactors need to be 192 .• Chapter 8 Sequential capture of C02 and S02 under FBC conditions considered. In the current study, the technical feasibility of alternative configurations is investigated. In addition to favourable economics, several criteria have to be met to achieve excellent capture of both S O 2 and C O 2 in fluidized bed combustion. Firstly, when capturing C 0 2 , sorbents should retain a high level of reversibility to minimize sorbent consumption. Secondly, the kinetics of carbonation and calcination should be fast to prevent the equipment from being too large Furthermore, there should be no obvious impact of the sulphate layer or carbonate layer on downstream capture. Other requirements, such as flexible operation, may also be important. Calcium sites needed for S O 2 removal account for only a small portion of those required for C 0 2 removal given the relative number of moles of the two species. For instance, a boiler firing coal containing 60%wt C and 2%wt S has a molar C/S ratio of 79. However, sulphur capture is irreversible, whereas carbonation is, at least in principle, reversible. To realize sequential capture, there are at least four candidate processes, each based on existing technology, as shown in Figure 8.1: Option A. Fresh sorbent is sent to an atmospheric fluidized bed combustor (FBC) first for S O 2 capture before being cycled to a carbonator/calciner system. The carbonator could be operated at atmospheric pressure and a relatively low temperature (e.g. 600°C) to achieve higher C 0 2 removal efficiency. Option B: Fresh sorbent is cycled in the carbonator/calciner system for C 0 2 capture until its reactivity is lost. The residues are then sent for S0 2 capture to an atmospheric FBC, where the sorbents are maintained as long as needed to maximize overall or calcium utilization for S-capture, before being discharged. Option C: Sorbents are first sent to a pressurized fluidized-bed combustor (PFBC) for S O 2 capture before being cycled within a carbonator/calciner system. A recent economic study 193 Chapter 8 Sequential capture of CO 2 and SO 2 under FBC conditions (MacKenzie et al., 2006) shows that a PFBC for C O 2 capture compares favourably with amine-based capture. Choosing practical operating temperature and pressure levels for the carbonator and calciner requires system integration, as well as removal efficiency, thermodynamics and transportation of CCvrich off-gas. The separation of the C O 2 removal system from the combustor allows a lower carbonator operating temperature. Option D: Sorbents are cycled first in a carbonator/calciner system for C O 2 capture, and then sent to a PFBC for S O 2 capture. In the PFBC, the sulphation could be realized either through CaO sulphation or CaC03 direct sulphation. The often-cycled sorbents could be sent to the PFBC, with or without a final calcination. Sending calcined CaO to a PFBC may not be attractive, but in this chapter this condition provides comparisons with direct sulphation using CaC0 3 . The operating temperatures and pressures in Figure 8.1 are for illustrative purpose. Actual test conditions for these options are specified where they are discussed: Several reactions are important in this chapter, Carbonation and calcination: CaO+C0 2 ^ CaC0 3 A i / 2 9 8 J C =-178kJ/mol 8.1) Lime sulphation: CaO+S0 2 +l/20 2 ^CaS0 4 Afl" 2 9 S K =-502 kJ/mol (8.2) Direct sulphation: CaC0 3 +S0 2 +l/20 2 ^ CaS0 4+C0 2 AH2gSK =-324 kJ/mol (8.3) Note that the carbonator and calciner in each of the above options have considerable flexibility in their operation, in particular in temperature and pressure, because of their separation from the combustor. In the experimental work presented below, however, in order to facilitate comparison of the options, carbonation and calcination conditions were all conducted at the same temperature for the cyclic C O 2 capture. 194 Chapter 8 Sequential capture of CO 2 and SO 2 under FBC conditions Practical heat supply issues for the calcination are not stressed in this study. For example, the calciners in the various options of Figure 8.1 are not necessarily oxy-fired, but could utilize multiple heat sources, such as steam or a solid heat carrier, depending on the available technology and cost. The pressure level of the carbonator/calciner could be varied to take advantage of thermodynamic limits and system requirements. 8.2 Experimental Details The pressurized T G A (PTGA) system, consisted of a Cahn 100 balance of 1 pg sensitivity, a reactor and a control system. The reactor is made of Inconel 600 alloy, allowing high temperatures and high pressures. Dwyer mass flow controllers connected to a computer adjusted the gas flow rates to obtain desired gas concentrations of mixed gases directed into the bottom of the reactor. The PTGA pressure level was set by a pressure regulator, with a bypass to allow atmospheric pressure operation. Operating temperatures were 850°C for both calcination and sorption. Table 8.1 Gas compositions for the four options examined. All reactions were conducted at 850°C. S 0 2 sorption in FBC Cyclic C 0 2 sorption stage Option Pressure S 0 2 Other components Carbonation Calcination (MPa) (ppmv) Composition Pressure (MPa) Composition Pressure (MPa) A-B 0.1 0.1 5000 5000 3%0 2 , balance N 2 3%0 2 , balance N 2 . 14%C0 2 , balance N 2 14%C0 2 , balance N 2 1.82 1.82 100% N 2 100% N 2 0.1 0.1 C 1.82 1500 or 5000 1 5 % C 0 2 , 3 % 0 2 , balance N 2 14%C0 2 ; balance N 2 1.82 100% N 2 0.1 D 1.82 5000 15% C 0 2 , 3 % 0 2 , balance N 2 14%C0 2 , balance N 2 1.82 100% N 2 0.1 195 Chapter 8 Sequential capture of CO 2 and SO 2 under FBC conditions Gas concentrations for the various tests are summarized in Table 8.1. As described above, all options involved: a SO2 sorption stage and a CO2 sorption stage. The SO2 sorption stage was realized through sulphation alone, direct sulphation, or simultaneous sulphation and carbonation. The CO2 sorption was simulated by cycling sorbents between a carbonator and a calciner. During these tests, except where indicated, the carbonation operated at 1.82 MPa and 850°C with 14% CO2 balanced by N2, whereas the calcination was at 0.1 MPa and 850°C in a pure N2 atmosphere. Note this cycling condition is different from that proposed in Figure 8.1. The selection of the current cycling condition is only for convenience to compare tests with different histories before or after cycling. A 4-minute sorption time was adopted for each cycle, enough for fresh sorbents to reach the slow stage of carbonation, whereas the calcination portion of each cycle was prolonged until further mass loss was negligible. The chemical analyses of the 212-250 pm Strassburg limestone and Arctic dolomite tested are shown in Table 1.2. All runs started with 50±2 mg of fresh sorbent. 8.3 Results and discussion To simulate option A, calcined Strassburg limestone was sulphated for 10 minutes and then sent for cyclic carbonation and calcination. Cyclic carbonation/calcination of this fresh limestone without sulphation was also performed at 850°C and 0.1 MPa, with pure N 2 for calcination, 14% C 0 2 and the balance N 2 for carbonation, to provide a baseline, shown as the upper line in Figure 8.2. The pre-sulphated PTGA tests show the following features during carbonation/calcination cycles. Carbonation still consisted of fast and slow stages. The conversions corresponding to the onset of the slow stage decreased sharply with cycling. The first calcination following the first carbonation lasted much longer before leveling off than the initial calcination and following 196 Chapter 8 Sequential capture of CO 2 and S02 under FBC conditions cycles. After 10 minutes of sulphation, -12% calcium utilization was achieved. The partially sulphated sorbents still showed high reactivity to C0 2 . The above findings suggest interaction between the pore structure and the relevant gas-solid reactions. Strassburg limestone has been found to sulphate in a shrinking core manner (Laursen et al., 2000) and can only achieve -35% calcium utilization during long-term (usually >24 h) sulphation due to pore-mouth plugging. After long-term sulphation, the sulphate layer was compact, as evidenced by low effective diffusivity in sulphation (Krishan and Sotirchos, 1994) and surface texture (Duo et al, 2000). A separate run indicated that the Strassburg limestone, calcined after prolonged sulphation, did not carbonate significantly further. It is believed that the residual carbonation in Figure 8.2 is due to an intermediate degree of sulphation. The remaining pore passages could still provide some surface area for carbonation, while the blocked or the partially blocked pore mouths were responsible for the much lower CaO conversion to carbonate. It is important to note that CaO carbonation and sulphation have different dependence on pore passages. Previous studies, mostly with liquid N 2 adsorption (Gullett and Bruce, 1987; Stouffer and Yoon, 1989; Sotirchos and Zarkanitis, 1992; Mahuli et al, 1997; Wu et al, 2002), indicate that CaO sulphation tends to fill pores of diameter 10-60 nm or larger, easily blocking pore mouths of this size range. Higher effective diffusivity in the sulphate layer is usually the key to improve sulphation. On the other hand, the work in Chapter 4 and in the literature (Bhatia and Perlmutter, 1983; Alvarez and Abanades, 2005) show that the fast stage of CaO carbonation depends on surface area residing in smaller pores and is less likely to experience pore blockage. The slow calcination after the first carbonation must be due to constricted passages, with sulphate causing blockage, resulting in lower effective diffusivity. With further cycling, during which sintering eliminates smaller pores while creating larger ones (see Chapters 4 and 6), the larger pores increase the effective diffusivity, helping to speed up calcination in later cycles. 197 Chapter 8 Sequential capture of C02 and S02 under FBC conditions Figure 8.3a shows a Scanning Electron Microscope (SEM) image for the sulphated, then cycled calcine. A porous texture is seen, indicating the start of macropore growth due to sintering that also provides passages for C0 2 . Smooth coverage by a sulphate layer can be clearly identified. This study found that a potential problem of this option is poor control of the extent of sulphation, which may severely reduce carbonation ability and retard calcination. In addition, reversibility was far from ideal. Hence option A is not considered further. In option B, the sorbents are first sent to a carbonator/calciner to concentrate C0 2 . After the initial, 2nd, 7 th and 15th calcination, sorbents were sulphated for 2400 s under atmospheric conditions. The subsequent sulphation results for the cycled sorbents appear in Figure 8.4. Compared with the initial calcines, the cycled sorbents showed a decreased extent of sulphation. However, after 15 carbonation/calcination cycles, calcines demonstrated better sulphation ability than after 2 and 7 cycles. This must be due to the dependence of sulphation on surface area and pore volume. Sulphation suffers from pore-mouth blocking because of the much higher molar volume of C a S O a product relative to the original limestone. This is especially true for Strassburg limestone, which sulfates as an unreacted core type (Laursen et al., 2000). Improved performance after the 15th calcination is believed to be due to enlarged macropores generated during cycling. Chapter 4 and previous studies (Alvarez and Abanades, 2005) have generally shown that, as a result of the intermediate stage of CaO sintering, carbonation/calcination cycling eliminates pores in the <220 nm mesopore range and generates macropores of diameter >220 nm. Growth in macropore volume could increase effective diffusivity for subsequent sulphation, resulting in higher calcium utilization than for fewer cycles. Therefore, when using cycled limestones as S0 2 sorbents, after the first few cycles of carbonation/calcination, reduction in surface area or pore volume may result in less sulphation, 198 Chapter 8 Sequential capture of CO2 and SO2 under FBC conditions but with further cycling, the change of pore structure or improved effective diffusivity due to macropore growth benefited sulphation, again demonstrating different pore dependence of the carbonation and sulphation processes. Whereas the fast stage of carbonation relies primarily on pores in the <220 nm size range (see Chapter 4) and slows down greatly as these pores are filled, sulphation utilizes macropores larger than -220 nm to achieve higher conversion, presumably with increased internal diffusivity. EDX (Energy-Dispersive X-ray) analyses in Figure 8.5 reveal that both the initial calcines and those after 15 cycles sulphated either in an unreacted-core manner or with the pore-plugging mechanism. AJthough macropores were generated during cyclic sintering, the unreacted-core behaviour did not change, either because the larger pores may still reside at the outer layer of particles, or SO2 molecules could still not enter the smaller pores. Arctic dolomite was also tested based on option B. The results appear in Figure 8.6. Unlike the Strassburg limestone, the calcium utilization after 40 minutes of sulphation decreased monotonically with cycling. Calcines from the 15th cycle did not perform significantly differently from those after the 7th. Different sintering degrees for the Arctic dolomite and Strassburg limestone appear to have affected re-sulphation behaviour. As revealed in Chapter 6, sintering of calcined Arctic dolomite eliminates <100 nm pores, while not generating larger pores because of the presence of MgO. Therefore, unlike Strassburg limestone, cycling did not improve sulphation of Arctic dolomite. Overall option B is more attractive than option A in that it makes better use of the sorbents for both carbonation and sulphation. The highly cycled sorbents, considered as spent for further CO2 removal, still capture SO2 due to a more favourable pore structure. A PFBC-based system can also be applied in a similar manner to the AFBC-based systems discussed above. This was tested in Options C and D. A key difference is that in PFBC-based 199 Chapter 8 Sequential capture of C02 and S02 under FBC conditions systems, sulphation may either be through direct sulphation of carbonate, or by sulphation of CaO, possibly accompanied by parallel carbonation. During the tests, a batch of fresh sorbent was first sulphated directly for 600 or 2400 s at high C O 2 partial pressure to prevent calcination. The sorbents were then sent to carbonation/calcination cycles. The total pressure was preset to 1.82 MPa Except where specified, gas concentrations for the PFBC sorption step were maintained as 15%v C O 2 , 5000 ppmv S O 2 and 3%v O 2 , with N 2 as the balance. The sharp peaks in Figure 8.7 are related to gas switching, picked up by the sensitive load cells and should be disregarded. Figure 8.7a shows the test history with 600 s direct sulphation In another set of tests, fresh sorbents were calcined first, and then exposed to the same test conditions, so that simultaneous carbonation and sulphation could occur. Figure 8.7b shows the results. The conditions and durations tested are listed in Table 8.2. The extents of sulphation are based on recorded mass changes. The mass change in direct sulphation only involves reaction (8.3). The mass change in simultaneous sulphation and carbonation is due not only to carbonation reaction (8.1), but also to reactions (8.2) and (8.3), CaO conversions were calculated with the aid of a full calcination following sorption, as shown in Chapter 5. Note all sulphations were performed here before carbonation/calcination cycling. Several findings require discussion. First, Figure 8.7a indicates that the first calcination after direct sulphation was slower than the initial calcination of fresh limestone. This is due to the sulphate layer along the outer rim of the particles because direct sulphation occurs (Hajaligol et al., 1988; Tullin et al., 1993; Qiu and Lindqvist, 2000) in a typical unreacted core manner. This implies that overly sulphated sorbents may not show favorable cyclic performance compared to fresh sorbents. 200 Chapter 8 Sequential capture of CQ2 and SQ2 under FBC conditions Table 8.2 Sulphation extents after direct sulphation Sulphation type Conditions (850°C, 1.82 MPa) (all with 15% C 0 2 , 3 % 0 2 , N 2 balance) Duration for sorption step (s) Sulphation extent (based on total Ca) Direct' Strassburg limestone,1500 ppm S 0 2 600 2.1% Direct Strassburg limestone, 1500 ppm S 0 2 2400 6.1% Direct Strassburg limestone, 5000 ppm S 0 2 600 4.8%(Figure 8.3c)* Direct Strassburg limestone, 5000 ppm S 0 2 2400 13.3%(Figure8.3d)* With carbonation2 Strassburg limestone, 1500 ppm S 0 2 600 4.7% With carbonation Strassburg limestone, 5000 ppm S 0 2 600 7.6%(Figure 8.3b)* Direct Arctic dolomite, 5000 ppm S 0 2 600 9.5 % With carbonation 1^-Arctic dolomite, 5000 ppm S 0 2 600 15.3% Direct: means direct sulphation is occurring With carbonation: means sulphation occurring simultaneously with carbonation. *Corresponding to the same samples as shown in Figure 8.3b-d. Secondly, Table 8.2 shows that 600 s of simultaneous carbonation and sulphation of CaO gave higher CaO conversion to CaS04 than 600 s of direct sulphation. The porous nature of the calcines applied in simultaneous capture was responsible for the higher sulphation extent and the subsequent faster calcination seen in Figure 8.7. However, the effect of the mode of sulphation on the calcination rate became less significant with further cycles. In Figure 8.8, after the sulphation stage Strassburg limestone is seen to have exhibited similar calcination/carbonation reversibility as in the baseline cycling tests with fresh limestone. They all appeared to have no memory of the direct sulphation histories. Note that the results plotted here are on a free calcium basis, where free calcium refers to the part of calcium existing in any form other than sulphate. Similar trends are observed in Figure 8.9 for simultaneous sulphation and carbonation, except that the mn with 5000 ppm S0 2 experienced somewhat lower reactivity for the first cycle, presumably because of incomplete recovery of effective diffusivity. The results imply that in a PFBC system, S0 2 removal could be achieved by either direct 201 Chapter 8 Sequential capture of CO 2 and SO 2 under FBC conditions sulphation or simultaneous sulphation and carbonation. If the extent of sulphation does not exceed those in Table 8.2, subsequent carbonation/calcination cycles give similar cyclic performance as for fresh limestone. High-resolution SEM images of calcines from the above runs appear in Figure 8.3. Figure 8.3b shows a grainy structure compared with that after CaO sulphation in Figure 8.3a or after direct sulphation in Figures 8.3c and 8.3d. The grain structure is probably inherited from the initial calcination, with sintering commencing after a few cycles. As the sulphation extent is not high (7.6%), the sulphate layer probably did not completely cover the outer surface, resulting in faster calcination as discussed above. For direct sulphation, however, complete coverage occurred as seen in Figures 8.3c and 8.3d for different durations, indicating that direct sulphation proceeds more like the unreacted core case. The original grain boundaries are still visible after 600 s of direct sulphation (Figure 8.3c), whereas blocks of sulphate layers appear in the 2400 s runs (Figure 8.3d). Both calcines show pores acting as passages through which C O 2 can permeate. These pores were generated originally by calcination after direct sulphation, and then by sintering-related large-pore growth, probably related to more reactive CaO pores coated with a sulphate layer. Dolomites are commonly used as sorbents in PFBCs. Here Arctic dolomite was tested in the same manner as the Strassburg limestone, based on the most promising option, option C. As seen in Chapter 6, Arctic dolomite was more reactive than Strassburg limestone for sulphation, probably because of its high surface area, with 600 s of direct sulphation resulting in 9.5% calcium utilization, whereas 600 s of simultaneous reaction with S0 2 and C 0 2 gave 15.3% calcium utilization, as seen in Table 8.3. The enhanced sulphation via simultaneous exposure to S0 2 and C 0 2 greatly affected subsequent cyclic C 0 2 capture. As seen in Figure 8.10, reversibility deteriorated greatly, whereas the 9.5% calcium conversion to sulphate through 202 Chapter 8 Sequential capture of CO 2 and SO 2 under FBC conditions direct sulphation had relatively little effect on subsequent carbonation/calcination cycles as shown in Figure 8.11. As the volume of pores <100 nm is especially important to carbonation, it can be concluded that simultaneous carbonation and sulphation exposure greatly reduced the pore volume in this size range. (All calcium utilizations are again based on free calcium oxide). SEM photos for calcines from both of these runs are compared in Figure 8.12. The one which experienced direct sulphation appeared to have more porous grains than the one subjected to co-capture. As indicated by the much slower calcination in the co-capture case, the sulphate layer greatly reduced the effective diffusivity, probably due to the uniform distribution of sulphate in the co-capture product, whereas in the case of direct sulphation, the entire sulphate layer was an outer envelope so that further sulphation proceeded to a much lesser extent. After calcination, the pore channels could be emptied more easily by the CO2, generating sorbents with higher effective diffusivity and higher CaO availability for subsequent CO2 capture. As for option B, sorbents after being cycled in the carbonator and calciner could also be used as SO2 sorbents in a PFBC. To test this option D, Strassburg calcines after initial calcination and after 7 th and 15lh cycles of calcination/carbonation cycles were subjected to simultaneous carbonation and sulphation for 2400 s in an atmosphere containing 5000 ppm SO2, 15% CO2 at 850°C and 1.82 MPa. After that, the conversion was determined by calcining until there was no further mass change. A typical run with 7 calcination/carbonation cycles is shown in Figure 8.13. After cycling, 2400 s of co-capture generated a compact sulphate layer that extended the final calcination period (more than 2000 s required to level off). As plotted in Figure 8.14, the conversion of CaO to CaC0 3 decayed with cycling, presumably because of CaO sintering. However, the conversion of CaO to CaS0 4, after 15 calcination/carbonation cycles was higher than after 7 cycles. Recall that for AFBC-based option B, the sintered pore structure or growth of macropores enabled a higher degree of sulphation due to increased effective 203 Chapter 8 Sequential capture of CO 2 and SO 2 under FBC conditions diffusivity. Similar factors are believed to be responsible for the somewhat enhanced extent of sulphation in this co-capture case. In Figure 8.14, the sulphation extents for 2400 s sulphation based on option B, labeled as AFBC, are plotted for comparison. The conversion of CaO to CaS04 is similar for PFBC and AFBC conditions, whereas that for PFBC co-capture exceeds that for AFBC sulphation, consistent with findings in Chapters 5 and 6 that under comparable conditions, co-capture gives a greater extent of sulphation. As shown in Figure 8.15, tests with Arctic dolomite did not show any enhancement of sulphation by using highly cycled sorbents. This confirms the trend found for option B, and is believed to reflect the monotonic elimination of pores for Arctic dolomite during cycling as discussed above. The cycled sorbent can also be directly sent to a combustor for sulphation before being calcined. In one run, Strassburg limestone after 15 calcination/carbonation cycles, after the carbonation stage, was immediately exposed to PFBC conditions for 2400 s of sorption followed by complete calcination. It was found that 2400 s of further sorption in CO2/SO2 gave a 10.3% calcium utilization for sulphation, less than the 19% achieved by calcined sorbents after 15 cycles. Given that actual sulphation requirements are low when CaO is in large excess to meet the carbonation demand, the highly cycled sorbents without calcination are likely to be able to achieve the desired SO2 removal efficiencies. In all the options discussed above, the flue gases after being scrubbed of SO2 in an AFBC or PFBC, are then sent for downstream CO2 removal. There is still concern that the lack of efficient SO2 capture upstream would cause further decay, reducing CO2 capture as observed in Chapters 5 and 6. To test the tolerance of sorbents to residual sulfur, results of a test with 14v% C 0 2 and only 100 ppmv S0 2 are shown in Figure 8.16. This again shows some decay due to the presence 204 Chapter 8 Sequential capture of CO 2 and SO2 under FBC conditions of S O 2 for the Strassburg limestone, indicating that high-efficiency S O 2 removal is essential to protect the sorbents downstream. 8,4 Conclusions In fluidized bed combustion operations, calcium-based sorbents (limestones and dolomites) are commonly used for sulphur capture. There is also interest in using these sorbents for cyclic calcination/carbonation in order to provide a concentrated CO2 steam suitable for regeneration. This chapter addresses the question of whether the sorbents should be subjected to sulphation before, after or simultaneously with calcination/carbonation cycling. Options B and D, where the sorbents are exposed to calcination/carbonation cycle, before being sent to fluidized-bed combustion to capture S O 2 , are clearly better than other two options where sulphation occurs before calcination/carbonation, or where sulphation and carbonation occur simultaneously. (The basic features of a process based on Option B are presented in Appendix ATX through Aspen simulation.) The dolomite tested showed similar trends to the limestone in most cases. Even a small amount of S O 2 present can cause rapid deterioration in the ability of calcium-based sorbents to undergo cyclic calcination/carbonation cycles, due to pore blockage and formation of a relatively impermeable outer shell. 9 205 Chapter 8 Sequential capture of CQ2 and SQ2 under FBC conditions Flue gas H tree Fresh limestone/dolomite Fuel Air F B C 850<t 0.1 MPa 650°C 0.1 MPd J Carbonator Calciner Stack gas, SO2 free • 90(PC 0.1 MPli Residue sorbent OxygeiJ Ash drain Air Separation Unit -N2 Spent sorbent Fuel (a) Option A Fuel A i r Flue gas SOz free C F B 850<€ 0.1 M P a 65CPC 0.1 M P i J Carbonator balciner Stack gas. SO2 free 900°C 0.1 MPb Residue sorbent . Fresh limestone ^/dolomite Ash drain Spent sorbent Oxygen A i r Separation Unit Fuel (b) Option B 206 Chapter 8 Sequential capture of C02 and S02 under FBC conditions Fresh limestone/dolomite Fuel A i r Flue gas S u i tree Stack gas, P F B C 8 5 0 C ~ 2 M P a 6 S 0 ° C ~ 2 M P a J S O i free arbonatqr Calciner 9 0 0 t O . l M P a Residue sorbent Ash drain A i r Separation Unit (c) Option C O xygen Fuel Spent sorbent Fuel Air Flue gas S d free Stack gas, PFBC 850°C ~2 MPa parbona|tor 6S(PC ~ 2 M P i 1 900°C 0.1 M P i S O J tree Fresh limestone Calcine^ ^dolomite Fuel Residue sorbent Ash drain Spent sorbent Oxygen Air Separation Unit -N2 • (d) Option D Figure 8.1 Candidate processes for CO2 and SO2 removal with calcium-based sorbents; Options A and C. sulphation before calcination/carbonation; Options B and D: calcination/carbonation before sulphation; Options C and D involve pressurized fluidized-bed combustion (PFBC). 207 Chapter 8 Sequential capture of CO 2 and SO 2 under FBC conditions (0 © o E "B o c <o « O © o *1 £ 0.4 o -o </> o 1 0.8 H 0.6 0.2 0 O —e- Unsulphated • Sulphated sorbents G "G-O- •^ ~e~0"©-©_.Q.^ _e-o~-©-^ .^ ...i G 5 10 15 N u m b e r of react ion c y c l e s 20 Figure 8.2 Sorbent performance: conversion history based on option A. 212-250 pm Strassburg limestone. Sulphated first for 10 minutes at 850°C and 0.1 MPa. 850°C calcination/carbonation cycles. Baseline conditions for fresh limestone: see Table 8.1 for C O 2 sorption stage. 208 Chapter 8 Sequential capture of CQ2 and SQ2 under FBC conditions Figure 8.3. S E M photos for calcines from different runs; for (b), (c) and (d), see Table 8.2, for test conditions, (a) CaO sulphation, followed by 4 calcination/carbonation (c/c) cycles, 12% calcium utilization for sulphation (on molar base). Test conditions, as for solid points in Figure 8.2. Option A. (b) Co-capture for 600 s followed by 4 c/c cycles, Option C. (c) Direct sulphation, 5000 ppm S0 2 , for 600 s followed by 4 c/c cycles, Option C. (d) Direct sulphation, 5000 ppm S0 2 , for 2400 s followed by 4 c/c cycles, Option C. 209 Chapter 8 Sequential capture of CQ2 and S02 under FBC conditions 0.18 1 Time (s) Figure 8.4 Sulphation of cycled sorbents after different number of calcination/carbonation cycles. 212-250 pm Strassburg limestone. (Option B) Figure 8.5 E D X sulfur-mapping of calcines for 212-250 pm Strassburg limestone. Test conditions as in Figure 8.4. (a) Sulphation of Strassburg limestone for 2400 s (Fresh sorbent) (b) Sulphation of Strassburg sorbent after 15 calcination/carbonation cycles. 210 Chapter 8 Sequential capture of C02 and S02 under FBC conditions 0.3 0.2 O O 0.1 After initial calcination After 2 cycles 500 1000 1500 Time (s) 2000 2500 Figure 8.6 Sulphation of cycled sorbents after different number of calcination/carbonation cycles. 212-250 pm Arctic dolomite (Option B). For conditions, see Table 8.1. 211 Chapter 8 Sequential capture of CO 2 and SO 2 under FBC conditions 50 45 40 Switch to N2 for calcination Switch to carbonation introduction of SO? 0 1000 2000 3000 4000 5000 6000 Time (s) 7000 (a) 0 1000 2000 3000 4000 Tirre(s) (b) Figure 8.7 Mass change profiles during typical runs for Option C with 212-250 urn Strassburg limestone, (a). Direct sulphation for 600 s. (b). Simultaneous sulphation and carbonation for 10 minutes. Test conditions: Both direct sulphation and simultaneous sulphation and carbonation: 5000 ppm S0 2 , 15% C O 2 , 3% O2, N 2 balance before calcination/carbonation cycling, 1.82 MPa, 850°C. 212 Chapter 8 Sequential capture of CQ2 and S02 under FBC conditions CO u D CO "o £ c '3 +-< <u o u 0.8 H 0.6 * 0.4 0.2 0 • x A O Fresh unsulphated limestone Sulphated sorbents, sulphation for 600 s in 1500 ppm SO2 Sulphated sorbents, sulphation for 2400 s in 1500 ppm SCh Sulphated sorbents, sulphation for 600 s in 5000 ppm SQ2 Sulphated sorbents, sulphation for 2400 s in 5000 ppm SO2 0 15 5 10 Number of reaction cycles Figure 8.8 Cyclic calcination/carbonation performance of212-250 pm Strassburg limestone after direct sulphation (Option C). See Table 8.1 for calcination/carbonation cycling conditions. o I O a 0.8 1 0.6 A 0.4 0.2 Fresh unsulphated limestone O Sulphated sorbents, sulphation for 600 s in 1500 ppm S 0 2 • Sulphated sorbents, sulphation for 600 s in 5000 ppm S 0 2 o o 5 10 15 Number of reaction cycles 20 Figure 8.9 Cyclic performance of 212-250 pm Strassburg limestone after simultaneous sulphation and carbonation, and a complete calcination. Option C with 212-250 um Strassburg limestone. See Table 8.1 for calcination/carbonation cycling conditions. 213 Chapter 8 Sequential capture of CO 2 and SO 2 under FBC conditions U "o a -a g "3 CD CS o u CO CU 1 0.8 0.6 0.4 H 0.2 Fresh unsulpriated sorbent • Sulphated sorbent, sulphation for 600 sin 5000 ppm S Q 0 —1— 10 5 10 15 20 Number of reaction cycles Figure 8.10 Cyclic performance of 212-250 pm Arctic dolomite after simultaneous sulphation and carbonation, and calcination. Option C. Baseline conditions for fresh sorbent: See Table 8.1 for calcination/carbonation cycling conditions. O cS "o . 6 C •a o u CO 1 0.8 0.6 « 0.4 0.2 0 Fresh unsulphated sorbent Sulphated sorbent, sulphation for 600 s in 5000 ppm S Q 0 15 20 5 10 Number of reaction cycles Figure 8.11 Cyclic performance of 212-250 pm Arctic dolomite after direct sulphation and calcination. Option C. Baseline test conditions for fresh sorbent: See Table 8.1 for calcination/carbonation cycling conditions. 214 Chapter 8 Sequential capture of CO 2 and SO 2 under FBC conditions Figure 8.12 Surface texture of cycled 212-250 um Arctic particles different sulphation history (a) Same test conditions as in Figure 8.10 (b) Same test conditions as in Figure 8.11. 215 Chapter 8 Sequential capture of CO 2 and S02 under FBC conditions 35 A 0 2000 4000 6000 8000 10000 Time (s) Figure 8.13 Mass change profiles with simultaneous carbonation and sulphation after 7 cycles of carbonation and calcination. Option D with 212-250 pm Strassburg particles 0.8 o 1/1 •a g B o C/3 O o - M )-i o o « <s o 0.6 0.4 0.2 • PFBC, conversion of CaO to C a S 0 4 O PFBC, conversion of CaO to C a C 0 3 A AFBC, conversion of CaO to C a S 0 4 o CP CD IS 6 9 Number of reaction cycles 12 15 Figure 8.14 Sulphation and carbonation extents during co-capture with cycled 212-250 pm Strassburg limestone particles. Option D vs. Option B. See Table 8.1 for conditions. 216 Chapter 8 Sequential capture of CQ2 and S02 under FBC conditions o o 1) <1> g O C/3 O o o O 0.8 0.6 h 0.4 0.2 m A P F B C , conversion of CaO to C a S 0 4 O P F B C , conversion of CaO to C a C 0 3 • A F B C , conversion of CaO to C a S 0 4 • 0 21 6 9 Number of reaction cycles 12 15 Figure 8.15 Sulphation and carbonation extents during co-capture with cycled 212-250 pm Arctic dolomite particles. Option D vs. Option B. See Table 8.1 for conditions 217 Chapter 8 Sequential capture of CO 2 and SO2 under FBC conditions Figure 8.16 Effect of small residual S0 2 on calcination/carbonation cycling. 212-250 um Strassburg limestone particles. Points: Sorption with 14% C O 2 , 100 ppmv S02,3% O 2 and balance N 2 at 850°C and 1.82 MPa. Calcination in 100% N 2at 850°C and atmospheric pressure. Cycling conditions for fresh sorbent with no S O 2 present: see Table 8.1 for carbonation conditions. 218 Chapter 9 Co-capture ofH2Sand CO2 in a pressurized-gasifier-basedprocess C H A P T E R 9 C O - C A P T U R E O F H 2 S AND C 0 2 IN A PRESSURIZED-GASIFIER-BASED PROCESS A version of this chapter has been accepted by Energy and Fuels, in press. Authors are: P. Sun, J. R Grace, C. J. Lim and E. J. Anthony. 9.1 Introduction Calcium-based sorbents have been receiving widespread attention as candidates to remove the greenhouse gas C O 2 in-situ from reactors in combustors (Shimizu et al., 1999; Abanades et al., 2003), gas shift reactors (Han and Harrison, 1994) and steam reformers (Balasubramanian et al., 1999; Ortiz and Harrison, 2001; Ziock et al., 2004a; 2004b; Johnsen et al., 2006), while improving higher hydrogen yields in the latter two cases. They may also be attractive to integrate C O 2 capture and enhanced H 2 production in fossil fuel gasifiers in an attempt to produce H 2 in a more economical and environmentally advantageous manner. The overall Pressurized Gasifier/calciner process utilizes calcined limestone as a C O 2 carrier, circulating between a pressurized gasifier and a calciner. As an example of this process, Lin and coworkers (Lin et al., 2002a; b) proposed and tested a HyPr-RING process where C 0 2 is cyclically absorbed in high-pressure gasifiers by hydrated lime. Another advantage of using calcium-based sorbents in gasifiers is thatH2S can be removed simultaneously due to the parallel reactions: Carbonation C a O + C 0 2 ^ C a C 0 3 Atf2 9 8 / :=-178 kJ/mol (9.1) Sulfidation CaO+H 2 S^ CaS+H20 &H29*K =" 5 9- 4 kJ/mo\ (9.2) Direct Sulfidation CaC0 3+H 2S ^ CaS+H 20+C0 2 A / / 2 9 8 / : =1 18.7 kJ/mol (9.3) 219 Chapter 9 Co-capture ofHjS and CO2 in a pressurized-gasifier-based process Many tests have focused on sulfidation mechanisms and sorbent performance (Efthimiadis and Stotirchos, 1992; Fenouil et al., 1994; Krishnan and Sotirchos, 1994; Fenouil and Lynn, 1995a; b; c; Yrjas et al., 1996; Zevenhoven et al., 1996; 1998; Chauk et al., 2000; Garcia-Labiano et al., 2004). The CaO-F£2S reaction was found to be fast and to achieve high, or even complete, conversions (Borgwardt et al., 1984; Efthimiadis and Stotirchos, 1992; Fenouil and Lynn, 1995a; Chauk et al., 2000), whereas the CaC03-H2S reaction is much slower and achieves much lower conversions (Krishnan and Sotirchos, 1994; Fenouil and Lynn, 1995b; Zevenhoven etai, 1998). Although particle-size-dependent, both sulfidation and direct sulfidation have usually been reported to be controlled by product layer diffusion (Fenouil and Lynn, 1995a; b; c; Zevenhoven etai, 1998; Chauk etai, 2000; Garcia-Labiano etai, 2004). To the best of our knowledge, simultaneous removal of H 2S and C 0 2 has never been explicitly studied before. Recent studies (Sun et al, 2005, 2006; Ryu et al, 2006) on simultaneous S0 2 and C 0 2 removal based on a PFBC/ calciner process (Shimizu et al, 1999; Abanades et al , 2003) have shown that S0 2 inhibits the reversibility for C 0 2 capture. The objective of the current study was to investigate whether H 2S has a similar deleterious effect on sorbent reversibility in a pressurized gasifier (PG)/calciner process. 9.2. Experimental Details The pressurized T G A (PTGA) system consists of a Cahn 100 balance of 1 pg precision, a reactor column and a control system. The reactor is made of Inconel 600 alloy, allowing high temperatures and high pressures to be used. The desired gas concentrations were achieved by means of computer controlled of the mass flow controllers. Mixed gases were directed into the reactor from the bottom. The total gas flow rate was normally -500 ml/min for both calcination and sorption. The PTGA pressure was set by a pressure regulator. A bypass gas line allows 220 Chapter 9 Co-capture ofH2S and C02 in a pressurized-gasifier-basedprocess atmospheric operation. Typical operating temperatures were 850°C for both calcination and sorption. • Strassburg limestone, Arctic dolomite and Kelly Rock limestone were studied. Particles were screened to 212-250 pm except where specified. Table 1.2 shows the chemical analyses for all these sorbents. All runs started with 50±2 mg of fresh sorbent in a platinum sample basket. Strassburg limestone was intensively tested in parametric tests. The cyclic co-capture test procedure is illustrated in Figure 9.1 for a typical run. After calcination at atmospheric pressure, the reactor was pressurized to the desired level with N 2 . The N 2 stream was quickly replaced by a pre-determined syngas stream containing C0 2 , H 2S, CO, H 2 and N 2 . About 20 s later, sorption started when the syngas front reached the sorbent. A time lag of 20 s was therefore excluded from the timing of all runs. Switching to N 2 immediately finished the sorption stage and started the calcination and depressurization. When no further mass change was observed, the next cycle was started. For each cycle, as shown in Figure 9.1, the mass increase due to H 2S capture was obtained by comparing the mass at the completion of calcination for that cycle with that for the previous cycle. The mass increase due to both H 2S and C 0 2 capture was derived from the total mass increase during the pressurized sorption stage. Thus calcium utilizations due to H 2S and C 0 2 could be estimated separately. This method of separating the mass increase due to carbonation and to sulfidation was also used in the "once-through" tests below. Gasifiers can be operated under various conditions, for example air-, oxygen- or steam-fired; pressurized or atmospheric; and at widely varying temperatures. Thermodynamic calculations by the HSC4 package (Ronie, 1997) show that higher temperatures and lower pressures favour higher equilibrium hydrogen yields. Higher temperatures and higher pressures 221 Chapter 9 Co-capture of H2S and CO2 in a pressurized-gasifier-based process are advantageous for tar cracking and solid conversions. For carbonation, lower temperatures and higher pressures are preferable from an equilibrium viewpoint. In this work, baseline co-capture operating conditions were selected as 850°C and 0.76 MPa. The inlet gas flow rate was controlled by mass flow controllers to obtain a mixture of l%v F£2S, 20%v C 0 2 , 12.6%v H 2 , 1.5%v CO, with the balance N 2 . H 2S and C 0 2 concentrations were also varied for parametric studies. Over 10%v Ff2 was added to prevent decomposition of H 2S. CO addition was used to prevent the reaction of CaS solid product with C0 2 . Oxidation of CaS by C 0 2 has been observed at both atmospheric and pressurized conditions, with the final product being CaC03 (Qiu et al., 2001; Anthony et al., 2003). Fenouil and Lynn (Fenouil and Lynn, 1995a) suggested that the CO:C0 2 molar ratio should exceed 1% to prevent the CaS-C0 2 reaction in the temperature range of 600-900°C. In addition, thermodynamics predict that under the current test conditions, the reverse water-gas shift reaction generates H 2 0 and CO. Chemical equilibrium calculations were also performed for various temperatures at a total pressure of 0.76 MPa for the above initial concentrations, i.e. l%v H 2S, 12.6%v H 2 , 20%v C0 2 , 1.5%v CO and the balance N 2 . The calculated equilibrium C 0 2 partial pressure is about 14%v for 700°C, and 12%v for 850°C. Equilibrium predictions show that a maximum of 0.05%v COS is expected, with the majority of sulfur existing as H 2S. Neither decomposition of H 2S nor oxidation of CaS was appreciable for the current tests. During cyclic co-capture tests, the sorption time was 3 minutes, except for one test where it was 8 minutes. 222 Chapter 9 Co-capture of H2S and C02 in a pressurized-gasifier-based process 9.3 Results and Discussion 9.3.1 "Once-through" tests. Before the cyclic co-capture tests, a set of tests, denoted "Once-through tests", was conducted with 50±2 mg Strassburg limestone to evaluate the relative importance of different reactions. These tests included sulfidation with no C 0 2 present, carbonation with no H 2S present, and a set of duplicate co-capture tests stopped at varying times. All tests were performed at 850°C and 0.76 MPa. Gas concentrations were maintained as similar as possible to those in the baseline co-capture mn. For example in the "no C 0 2 present" test, a mixture of l%v H 2S, 12.6%v H 2 , l%v CO, with the balance N 2 , was introduced so that all reactive components remained the same as in the co-capture tests, except for the absence of C 0 2 . Similarly for the "no H 2S present" test, the H 2 S portion was replaced by pure N 2 . The calculated conversions based on mass changes are plotted in Figure 9.2. Co-capture slowed down C 0 2 capture to some degree, especially before the carbonation reached a near-plateau (slower stage). The carbonation portion of the co-capture, stopped at 3, 9, 30 and 60 minutes, gave lower conversions of CaO to CaC03 than for continuous C 0 2 capture with no H 2S present. For the sulfidation portion, prolonged co-capture resulted in a similar extent of sulfidation as for continuous sulfidation with no C0 2 . In our high-pressure measurements, the sulfidation reaction does not appear to have been as fast as for atmospheric studies described in previous work (Borgwardt et al., 1984; Efthimiadis and Stotirchos, 1992; Agnihotri et al, 1999; Chauk et al., 2000). Garcia-Labiano et al.(2004), Agnihotri et al. (1999) and Hu et al. (2006) found that increasing total pressure, with the volume fraction of H 2S held constant, decreased the sulfidation rate, likely because of the decrease in effective diffusivity in the product layers (Garcia-Labiano et al., 2004). 223 Chapter 9 Co-capture of H2S and CO2 in a pressurized-gasifier-basedprocess A significant observation is that, unlike results for S O 2 / C O 2 co-capture (Sun et al., 2006), direct sulfidation, i.e. reaction (9.3), played little role, even with the sorption time extended well beyond the slow stage of carbonation. The evidence for this is that no appreciable consumption of carbonate was found after 60 minutes duration. This probably occurred because direct sulfidation of limestone (CaC03) is much slower than sulfidation of calcined limestone (Fenouil and Lynn, 1995a; b; Yrjas et al., 1996) (i.e. CaO), so that an increase in total pressure would further decrease the direct sulfidation rate by decreasing the effective diffusivity in the product layer (Garcia-Labiano et al., 2004; Hu et al., 2006). In contrast, direct sulphation, in previous co-capture studies of S O 2 and C O 2 under excess oxygen conditions (Sun et al., 2006), became dominant during the slow stage of carbonation, consuming carbonate and causing rapid decay in carbonation during cycling. Figure 9.2 shows no sign of a decrease in the extent of sulfidation after the slow stage of carbonation. Sulfidation is known to be solid-state diffusion-controlled reaction (Borgwardt et al., 1984; Fenouil and Lynn, 1995a; Zevenhoven et al., 1996; Chauk et al., 2000). Borgwardt et al. (1984) and Chauk et al.(2000) proposed that the diffusing ions are 0 2"andS 2 +. In this case, the ions have to permeate through two product layers (CaC03 and CaS) to reach either reactive CaO sites or the bulk sorbent. Comparing the co-capture with continuous sulfidation where no C O 2 was present suggests that the effective diffusivity for the sulfidation portion of the co-capture was similar to that for continuous sulfidation with no C O 2 present. Given that CaC03 only serves as a porous network to diffusing ions and that CaO sulfidation is product-layer-controlled (Borgwardt et al., 1984; Fenouil and Lynn, 1995a; Zevenhoven et al., 1996; Chauk et al., 2000), the diffusion resistance in the CaS layer is larger than that in carbonate layer so that the mobility of ions diffusing through CaC03 plus CaS is similar to that where they diffuse through CaS alone. This is probably related to the sintering of CaS at high temperatures. 224 Chapter 9 Co-capture ofH2Sand CO2 in a pressurized-gasifier-basedprocess Zevenhoven (1998) and Fenouil and coworkers (1994, 1995a) showed that CaS sintering could be the major factor accounting for the slower sulfidation at high temperatures, e.g. at 850°C. 9.3.2 Parametric tests A series of parametric studies was carried out on Strassburg limestone. The first parameter varied was the H 2 S / C O 2 molar ratio. Figure 9.3 shows the cyclic performance for the baseline run and the effect of changing both P c o 2 and PH 2S , with a cycle time of 3 minutes for the sorption stage. General features of the co-capture were first studied by comparing the baseline co-capture run (20% C 0 2 , 1% H 2S, 12.6% H 2 , 1.5% CO, balance N2) and its counterpart without any H 2S present (i.e. 20% C0 2 , 12.6% F£2, 1.5% CO, balance N2). In the test with no H 2S present, a time of 3 minutes was sufficient to reach the slow stage of carbonation. Changing the partial pressure of C O 2 from 8% to 20% had little influence on the final carbonation conversion for each cycle, and hence on the reversibility of sorbents during cycling without H 2S, provided that the fast stage of carbonation was complete and that P c o 2 was well in excess of the equilibrium P c o 2 , which is around 50 kPa at 850°C. Thus the test with no H 2 S present also provided a baseline for comparison of runs with different inlet H 2 S and/or C O 2 concentrations. Figure 9.3a shows that the reversibility based on total Ca further decreased during co-capture compared to the test with no H 2S present. However, the additional decay was limited compared with sulphation for co-capture of C O 2 / S O 2 (Chapter 5 or Sun et al., 2005). Shown in Figure 9.3a for the run with 20% C 0 2 and 1% H 2S, a deliberately prolonged 16th cycle demonstrated that the sorbent still had good ability to react further with C O 2 . Under current test conditions, high conversions of CaO to CaS could be achieved, as revealed by Figure 9.3c. CaO conversions were also re-plotted based on free CaO, i.e. excluding the portion utilized for sulfidation. Figure 9.3b shows that on this basis, CaO retained 225 Chapter 9 Co-capture o/HiSand CO2 in a pressurized-gasifler-based process reversibility at a level only slightly less than that in the No H 2 S tests during cycling, when also capturing H 2S. This reversibility is regarded as favourable because the capture-ability of natural sorbents is almost retained. In the following discussions, all cyclic performance is assessed based on free lime content. The close-to-H2S-free performance also suggests that the cumulative CaS layer did not appreciably reduce cyclic C 0 2 capture. Below it is shown that a CaS layer slowed carbonation, but the reduction in the achievable extent of carbonation for each cycle was not appreciable after the first few cycles. An important finding is that the C O 2 removal performance observed in this study was much better than in our previous sulphation study where, under oxidizing conditions, S0 2 was found to seriously impede C O 2 co-capture (see Sun et al., 2005; 2006 or Chapters 5 and 6). The most probable reason is that CaS (molar volume 29 cm3/mol) has a much lower molar volume than CaS04 (46 cmVmol), whereas carbonate product (CaCOs) has a molar volume of 36.9 cm3/mol. All are larger than the molar volume of CaO (17 cm3/mol), suggesting that some swelling does occur for all three reactions involving CaO, i.e.: sulphation, carbonation and sulfidation. However the sulphate product is significantly less porous, resulting in much lower effective diffusivity. As a result, sulphation produces premature blocking for sorbents, especially for pore mouth diameters in the range of 10-100 nm (Gullett and Bruce, 1987; Stouffer and Yoon, 1989). Given the smaller molar product volume, sulfidation is less likely to suffer from plugging of pores than sulphation. S E M images in Figure 9.4 show major differences in surface texture for a H 2 S / C 0 2 run (Figure 9.4a) compared with a S0 2 /C0 2 run (Figure 9.4b) for the same temperature and similar gas concentrations, but a different total pressure. The calcine from the H 2 S / C O 2 test shows a much more porous surface than its S O 2 / C O 2 counterpart, due to the difference in molar volumes of the solid products. Therefore, during co-capture, the difference 226 Chapter 9 Co-capture ofH2Sand C02 in a pressurized-gasifier-based process in product layer compactness for sulphate and sulfide, produced very different degrees of inhibition for parallel carbonation and subsequent calcination. Conversion of CaO to CaS over 15 cycles is plotted on a cumulative time basis in Figure 9.3c and compared with a test with similar conditions, but with no C O 2 present (the baseline). The sulfidation performance for cyclic co-capture was much higher than for continuous sulfidation with no C O 2 present. EDX sulphur-mapping in Figure 9.5 proves that continuous sulfidation results in unreacted cores, whereas for cyclic co-capture, the sulphur penetrated more deeply into the interior of the particles. 9.3.3 Effect of C O 2 and H 2 S Partial Pressures To further elucidate the effect of varying P c o 2 and PH2S, it is helpful to study the kinetic dependence on these partial pressures. Although the sorption stage involved multiple reactions and mass transfer limitations, apparent rates can be compared, because all runs started with the same sample size. The first and the fifth cycles for relevant co-capture runs are compared in Figure 9.6 with the corresponding test where there was no H 2 S present. In Figure 9.6a, it is seen that addition of H2S slowed down the sorption process markedly. Whereas 3 minutes were sufficient to finish the fast stage of carbonation without H2S present, the mass climbed much more slowly and did not level off after 3 minutes when sulfidation was occurring in parallel. Sulfidation apparently either reduced the availability of CaO for the subsequent carbonation or added more diffusion resistance, retarding the carbonation. Increasing the C O 2 mole fraction from 20 to 40% speeds up carbonation. Kyaw et al. (1996) showed that increasing P C 02 well beyond its equilibrium level (~ 50 kPa at 850°C) did not increase the intrinsic carbonation rate, so that this accelerated carbonation is probably mainly due to enhanced intraparticle and interparticle mass transfer. It is notable that the apparent sorption rates shown here mostly reflect carbonation kinetics because carbonation is much faster than 227 Chapter 9 Co-capture of H2S and CO2 in a pressurized-gasifier-based process sulfidation as indicated in Figure 9.2. Figure 9.6b shows that the relative difference between runs with varying operating conditions in the first cycle persist at the fifth cycle, but the mass increases due to carbonation over the 3-minnute sorption period are similar, i.e. the presence of H 2 S did not affect the achievable extent of carbonation as much as in the first cycle. This is because the capture capacity is limited by the changes of sorbent properties during cycling, e.g. sintering reduced the ability to capture C O 2 , and because some CaO has been consumed by sulfidation during previous cycles. Figure 9.3 portrays the effect of increasing partial pressure of C O 2 on cyclic co-capture performance. Figures 9.3a and 9.3b show that increasing P c o 2 from 20 to 40%v improved carbonation for the first few cycles, but the improvement disappeared upon further cycling. Sulfidation, however, was seen to decrease markedly in Figure 9.3c. As discussed above, based on the evidence in Figure 9.6, it is believed that increased C O 2 can accelerate carbonate formation for all cycles. As a result, the CaO-FI^S reaction was slowed either because of increased mass transfer resistance due to the CaS layer or because of the reduced availability of CaO. From the observation that, relative to the baseline continuous sulfidation with no C 0 2 present, 20%v P c o 2 enhanced sulfidation, but the enhancement decreased when P c o 2 was increased to 40%v, it can be deduced that there should be an optimum P C o2 for maximum sulfidation. The optimum occurs probably because calcination/carbonation enhances sulfidation by periodically exposing highly-reactive CaO sites to H 2 S , but the enhancement is counteracted by a faster-growing carbonate layer with increasing Pco2-Increasing PH2s from 0 to 0.5% greatly slowed carbonation rates, as seen in Figure 9.6. This retarding effect only lasted for the first few cycles and had little influence on the extent of cyclic C O 2 capture after further cycling, as seen in Figure 9.3c. A further increase in P H 2 S from 0.5 to 1% had a much smaller effect. These results indicate that sulfidation reduced rates of carbonation 228 Chapter 9 Co-capture ofH2S and CO2 in a pressurized-gasifier-basedprocess mainly by competitively occupying CaO sites. The CaS layer did not appear to be a major barrier retarding carbonation since the test for free lime (excluding the sulphide part) gave results very similar to that without H 2 S in Figure 9.3c. Also shown in Figure 9.3c, the cyclic extent of sulfidation increased slightly with an increase of PH2S from 0.5 to 1%, with 20% C 0 2 present, due to the increased driving force for sulfidation. The total calcium utilization in Figure 9.3d, obtained by summation of the free-CaO conversion to CaC03 and that to CaS, showed little decrease with cycling for the tests with various C O 2 and H 2S partial pressures. This is unlike the sulphation for S 0 2 / C 0 2 (Chapter 5), where there was a sharp decline in the total calcium utilization. 9.3.4 Effect of temperature Gasifiers can be operated at lower temperatures than combustors. Lower temperatures are advantageous for in-situ carbonation, provided that the desired gas compositions can be obtained. Since the water-gas-shift reaction is exothermic, C 0 2 concentrations also tend to be high at lower temperature. With a high molar ratio of C 0 2 to H 2 S in a gasifier, it may be desirable to achieve a higher extent of carbonation while deliberately limiting the extent of sulfidation. Control of temperature provides a means of optimizing the fractional sorbent utilization for C 0 2 and H 2S. Figure 9.7 shows cyclic performance results for 700°C sorption conducted to compare with the base run at 850°C, with all calcinations performed at 850°C. The mass breakthrough histories during sorption for the selected cycles (first and fifth) at 700 and 850°C are also compared in Figure 9.6. Much faster sorption (mainly due to enhanced carbonation) can be observed for the lower temperature, with the rates and the achievable mass increases between the two temperatures narrowing markedly by the fifth cycle. As discussed above, the narrowing is due to the limitation of capture capacity. For further cycling, more narrowing must occur, so that the 229 Chapter 9 Co-capture ofH2S and C02 in a pressurized-gasifier-based process cyclic performance shown in Figure 9.7a gave similar results for the two temperatures except for the first two cycles. Figure 9.7b, however, confirms that the lower temperature led to much lower conversions of CaO to CaS, presumably either because of a change in sulfidation kinetics, or because of a faster-growing carbonate layer favoured by lower temperatures. The wider difference for sulfidation extents and close approach of the carbonation extents for the two temperatures, suggest that the variation of thickness of the CaS product layer did not appreciably affect cyclic carbonation performance, as long as the slower portion of carbonation process (see Figure 9.2) could be attained during each sorption cycle. 9.3.5 Effect of particle size The effect of particle size is shown in Figure 9.8. For most runs, 50+2 mg of sample was used for the 212-250 pm particles, but for the 38-45 pm particles, this was found to exceed the capacity of the basket due to the different bulk density for a closed packing, so only 23 mg were used in this test. Compared to the baseline run for the 212-250 pm particles, a slightly faster decay in C O 2 capture was found in Figure 9.8a for the finer material. On the other hand, there was a marked enhancement of CaO conversion to CaS, as shown in Figure 9.8b. This indicates that intraparticle diffusion played an important role during sulfidation, consistent with the conclusion from a previous high-pressure sulfidation study (Garcia-Labiano et al., 2004). Our findings are also consistent with a previous report that sulfidation is very sensitive to particle size (Efthimiadis and Stotirchos, 1992). For carbonation, on the other hand, the influence of particle size is minor as the cyclic calcium utilization is likely limited by sintering-related pore structural changes at the surface. Despite the much-enhanced H 2 S removal, the increasing CaS layer did not hinder carbonation appreciably, as discussed above. Compared to the findings for S O 2 / C O 2 co-capture (Chapters 5 and 6), the CaS layer is much more permeable than a CaS0 4 230 Chapter 9 Co-capture o/H2Sand CO2 in a pressurized-gasifier-based process layer for carbonation. Due to the enhanced H 2 S capture, the total calcium utilization achieved a relatively high level after 15 cycles as shown in Figure 9.8c. Note, however, that due to the irreversible nature of the sulfidation, over-sulphided sorbent would have limited ability to capture C0 2 . 9.3.6 Effect of residence time To simulate the effect of residence time on the reversibility, a run with an 8-minute sorption time was carried out, with the sorption time being the only quantity varied compared to the baseline co-capture run. Extension of the sorption time would necessarily result in higher conversions, as can be readily seen in the once-through test. During co-capture, as found in Figure 9.9a, the higher conversion for CaO to CaC03 only occurred for the first few cycles. The extended sorption time offered negligible enhancement of C O 2 removal after further cycling, probably because carbonations extended beyond 3 minutes are predominantly in the slower stages. A 9 t h cycle deliberately extended to 20 minutes still showed unrealized carbonation and sulfidation capacity, as shown in Figures 9.9a and 9.9b. In Figure 9.9b, it is found that cumulative sulfidation with 8 minutes of exposure for each cycle and 20 minutes for the 9 th cycle gave a conversion similar to that for a large number of 3-minute runs having the same cumulative time. It appears that when after calcination, when fresh CaO is cyclically exposed to gaseous reactants, sulfidation and carbonation occur simultaneously, but carbonation reaches a much slower stage after ~ 3 minutes, whereas sulfidation maintains a nearly unchanged rate until the next calcination reactivation step, in spite of the increase in sorption time from 3 to 8 minutes. 9.3.7 Effect of total pressure Changing the total pressure of the reactor from 0.76 to 1.7 MPa accelerated the mass uptake, mostly due to enhanced carbonation during each cycle, as shown in Figure 9.6 for both the first and the fifth cycles. The change presumably resulted from both enhanced external C O 2 231 Chapter 9 Co-capture ofH~2Sand C02 in a pressurized-gasifier-based process mass transfer and more favourable intraparticle mass transfer. However, the higher conversions to carbonate only applied to the first few cycles. The achievable conversions of calcium to carbonate remain similar to those for a total pressure of 0.76 MPa as seen in Figure 9.10a. As discussed above, with further cycling, the carbonation conversion would ultimately be limited by the achievable carbonation when it reaches the slower stage of carbonation, regardless of how much the rate of carbonation was enhanced. Figure 9.10b shows, however, that less H 2S was captured at the higher pressure, especially after 5 cycles. As noted above, increasing the total pressure decreases the sulfidation rate (Agnihotri et al., 1999; Garcia-Labiano et al., 2004; Hu et al., 2006). Enhanced carbonation could be another factor slowing down the sulfidation. The net effect of the two reactions is that the total calcium utilization was only a little lower at 1.7 MPa than at 0.76 MPa. 9.3.8 Effect of sorbent type The cyclic performances for Arctic dolomite and Kelly Rock limestone are compared with that of Strassburg limestone in Figure 9.11. All sorbents showed much-reduced C 0 2 capture ability, especially for the first few cycles, compared to corresponding tests where there was no H 2S. The H 2S in the mixed gas decreased the rate of carbonation, with the greatest decay being experienced by the most reactive sorbents, namely Arctic dolomite, which has been found to be more reactive in the initial cycles for both sulphation and carbonation, due to its well-preserved surface pore structure and surface area (see Chapters 5 and 6). Cyclic performance for all sorbents could approach the baseline for tests with no H 2S. The Arctic dolomite gave more conversion than the other two sorbents, both in terms of C O 2 capture and sulphur retention. Figure 9.12 compares the textures of calcines for the Arctic dolomite and Kelly Rock limestone. Figure 9.4a for Strassburg calcines indicates that all sorbents showed porous surfaces, differing greatly from the surfaces observed previously for S O 2 / C O 2 capture tests (see Chapter 6). 232 Chapter 9 Co-capture o/H2S.and C02 in a pressurized-gasifier-based process The macropores for the Strassburg limestone suggest more sintering than for the Arctic dolomite. The Kelly Rock limestone appears to be grainier, containing more pore volume between grains; it probably sinters in a similar matter to the Strassburg limestone. Sulphur-mapping in Figure 9.13 for Arctic dolomite and Kelly Rock limestone, in combination with Figure 9.5b for Strassburg limestone at the same test conditions, shows dispersed sulphur distributions for all three sorbents, despite the evidence that more sulphur is distributed on the outer area of grains or particles. As discussed above, cyclic operation could enhance H 2 S capture by cyclically exposing fresh CaO to H 2 S . 9.3.9 Cycled sorbents in sulfidation A gasifier could be integrated with a steam reformer operating with sorption-enhancement by a calcium-based sorbent. The highly cycled sorbent discharged from a downstream separate reformer could then be used as a sorbent for H 2 S capture in the gasifier. Removal of C O 2 and H 2 S would then be performed in two separate reactors. In the final test, Strassburg limestone particles after 15 cycles of carbonation and calcination at 850°C and 0.76 MPa were sent to be sulphidized under the same condition as in the sulfidation test with no C O 2 . The results (not plotted here) showed almost no difference between fresh Strassburg limestone and its cycled calcines. As cycling between calcination and carbonation is known to decrease the surface area and total pore volumes of the sorbents and to shift pores to larger sizes (see Chapter 4), this test indicates that sulfidation is not very sensitive to change of either of these properties. This may arise because the EhS-CaO reaction is product-layer controlled, but with a high enough effective diffusivity that it is less sensitive to change in pore structure. In a practical process, sorbent may experience hundreds or thousands of cycles. Therefore the results shown in this work, mostly based on 15 cycles, are not enough to represent a practical situation. However, as shown above in this chapter, in Chapter 6 and in some previous work and 233 Chapter 9 Co-capture ofHiS and CO2 in a pressurized-gasifier-based process (Shimizu et al., 1999; Alvarez and Abanades, 2005), the most severe decay in reversibility occurs during early cycles. Extended continuous fluid reactor studies are needed before practical long-term processes can be designed and executed with confidence. 9.4 Conclusions Simultaneous capture of C 0 2 and H 2S was studied under simulated gasification conditions for three calcium-based sorbents. The reversibility of sorbents was nearly as good as for cases where no H 2S was present, especially for limestone sorbents. Cyclic sulfidation led to higher extents of sulphur retention than continuous sulfidation with no C 0 2 present. In a practical system, the sulfidation will not increase the sorbent required for the C 0 2 capture very much, given the high stoichimetric ratio of carbon to sulphur in most fuels. Varying the reactant gas concentrations not only affected the physical limitations, but also caused complex interactions between solid product layers and solid-state diffusion. Parametric studies indicated that the effect of sulfidation on cyclic carbonation was usually weak. However, H 2S competes with C 0 2 for fresh CaO, so that a higher extent of sulfidation is not helpful in achieving good C 0 2 capture. Faster carbonation usually reduced the extent of sulfidation by tying up free Ca. Because of the weak effect of H 2S, co-capture of C 0 2 and H 2S in a gasifier appears to be much less problematic than co-capture of C 0 2 and S0 2 , examined in our previous work. 234 Chapter 9 Co-capture OJHTS and CO2 in a pressurized-gasifier-based process False peaks associated 3800 4200 4600 Time (sec) Figure 9.1 Illustration of operating procedure and masses used to calculate of calcium utilization for C 0 2 and H 2 S capture. Co-capture at 850°C and 0.76 MPa. Calcination: in 100% N 2 . Sorption: l%v H 2S, 20%v C0 2 , 12.6%v H 2 and 1.5 %v CO with the balance N 2 . 235 Chapter 9 Co-capture ofH2Sand CO2 in a pressurized-gasifier-based process 1 -r O 0.8 -0 5 10 15 20 25 30 35 40 45 50 55 60 Time (min) Figure 9.2 Once-through tests with 212-250 urn Strassburg limestone. Initial calcination at 850°C and 101 kPa with 100% N 2 . Co-capture at 850°C and 0.76 MPa with l%v H 2S, 20%v C 0 2 , 12.6%v H 2 and 1.5 %v CO, balance N 2 . Carbonation with no H 2S: same as for co-capture tests except with no H 2S. Sulfidation with no C0 2 : same as co-capture tests except with no C0 2 . 236 Chapter 9 Co-capture of H2S and C02 in a pressurized-gasifier-based process <JQ' c O 00 ° ^ ^  3 O ^° CD O p 3 o rr. a 0 ^ , 0 M O O 3 - T3 3 r+ „ E c CD T P re 3 2 a °- ^  OT O • • p w Q. b\ p ^ I p 3 o CD C/3 < g P C 3 O O CD O M V J o p T3 to S o g . » £t> 3 ° £ to CD K) p 3^  L/ l cT o 0 3 ' ^ T= 3 O C 2 cz; r+ H o g , . 3^ OT „ TJ OT o pj-. a* o o 3 -^ n P IO 00 O ft) O OT 3 TJ O OT O »• OT o CD TJ P 3 O-OT O TJ ' , O a- TJ O 3 to -O c p o o o % CD o. , 3 . CD P <-K 0 0 o o 3 E ED < CD C/> O o 3 _ a o-n = (D Moles of H2S retained/moles of total Ca 0 0 0 0 0 0 0 0 O i >——J- 1 1 1 ' I 1 O • o to to 0 0 0 ^? o x O O 0 O O O p O O \j\ X 0-. X X GO on (SI ro o ro Ol w o w cn o cn cn o cn cn O) o CD cn -j o Total calcium utilization © p o o p p © p O I—> t o U) L/i 'ON -~J i x H 0 •z Lft -g£ quin 0 rt P lei 0 3 3 P 0 c op. 0 " s iza eye 5' n P o o o 3 < CD O 3 O n w p OT CD O -O 3 CD CD O P (D fi) O 6 3 o >< o (D cn M o l e s of C 0 2 re ta ined /moles of total C a 0 0 0 0 0 0 0 0 0 - » N ) U * U l b ) N | ( J c 3 IT CD o p o o o 3 < CD 2. 2, o 3 o O P o o P OT CD Q . O 3 (-»-O P . n P ro o <D0 0 ^ / < » CK> i «> } 0 • 0 K J 4^ . tO O O ' O o o o o to f 0 tji o • O ~ Ln ^ g X X S tfl B OT Moles of H2S retained/moles of free Ca 0 0 0 0 0 0 0 0 o ro (A) cn b) s bo 1 1 1 1 1. i 1.. 3 CD CD O o >< o CD CD ro o <30 x -0 . o. • d — > 0 o o • o to 4^  10 z 0 O 0 0 oN ^? 0 s O O 0 ' O O 0 to O %I 0 X 10 H2S O O 237 Chapter 9 Co-capture ofH2S and CQ2 in a pressurized-gasifier-based process (a) After 15 cycles of H 2S and C 0 2 (b) After 15 cycles of S 0 2 and C 0 2 capture capture Figure 9.4 High-resolution SEM photos of co-capture 212-250 pm Strassburg limestone calcines: (a) H 2S and C02capture in l%v H 2S, 20%v C0 2 ,12.6%v H 2 , l%v CO, balance N 2 at 0.76 MPa, 3 min sorption for each cycle, (b) S 0 2 and C 0 2 capture in 1125 ppmv S0 2 , 8%v C 0 2 , 3%v 0 2 and balance N 2 at 1.8 MPa, 4 min of sorption for each cycle. Sorption and calcination temperatures for both cases were 850°C. Al l calcinations in 100% N 2 . Figure 9.5 E D X sulfur mapping for 212-250 pm reacted Strassburg limestone, (a) Calcine from 1 h continuous sulfidation at 850°C, 0.76 MPa with l%v H 2 S, 12.6%v H 2 , 1.5 %v CO, balance N 2 . (no C0 2 ) . (b) Calcine from cyclic H 2 S / C 0 2 sorption for 15 cycles at 850°C, 0.76 MPa with l%v H 2S, 20%v C 0 2 , 12.6%v H 2 , 1.5 %v CO, balance N 2 ; Calcination at 850°C and 101 kPa with 100% N 2 ; 3 min for each sorption stage. 238 Chapter 9 Co-capture ofH2Sand CQ2 in a pressurized-gasifier-based process 20 r 4 0 % C O 2 7oo°C 1.7 MPa N o H , S 60 120 Sorption time (sec) (b) The fifth cvcle 180 Figure 9.6 Mass changes during first and fifth cycles of H 2 S/C0 2 capture tests. 212-250 pm Strassburg limestone. Tests conditions except where specified: sorption at 850°C and 0.76 MPa with 20% C0 2 , 1%H2S, 12.6%v H 2 , l%v CO, balance N 2 , 850°C; Calcination at 850°C and 101 kPa with 100% N 2 . 3 minutes for each sorption stage. 239 Chapter 9 Co-capture ofH2S and C02 C02 in a pressurized-gasifier-based process 3> do" 3 >-t CD g J 8 " >° £, p ^ ^ ~ W s I . S5 -a -I o o 3 £ 3-d 81. CD ; ° 3 ' • 0 0 P ^ ° HH O w < p cr S L 5T 3 o CP o p T3 r-K C i-l CD O ° Q p o 0* ^ to i t-J <Vi O 3 P 5' 3 P 0 0 Tj<-ft 3 O 3 O C/J ° 3 p w 3, O OQ 2? CD 3 o 3 3" P 2 5" to 3 OJ P Total calcium utilization p H o o 5 ~ I* 3' Q p O 3 CD O a o p p o • IO • O •o • o IO • o • o o o o o o o o • o ~J 00 O L / i o o o o o o o o P o o o 3 < CD i - l Vi 5' 3 o p t / 3 n fa O O o 3 < CD O 3 O O P o o Moles of C O , retained/moles of free Ca o o o ON o bo a " o 3 o >-% / 10/ •e/ 1 €/ • O 1 o / 1 -J o 00 ' /I o o o o O a OK o <u O o I o 1 o o Moles of H 2S retained/moles of total Ca ON 240. Chapter 9 Co-capture qfH2Sand C02 C02 in a pressurized-gasifier-based process oo" c 0 s - VO to ON 00 W ^ P X & OT N" n> o 3 < o o P a c CD o " 5 O to P 3 CL o O 3 § " O 3 P <-+ 00 c / l _ O to O £ 5 OT O. OT o g P^ 3 PV o 3 " 3 CD 55 cr. to O S5 CD P o 3-OT o -3 j+ o' 3 OT OQ 3 00 on o o o p 3 C L . O -~J 0 \ p I. 3. 3 -to o O o Total calcium utilization o o o ON O 00 H o E o p N P O 3 p o 5 ' 3 •3 • o to o • o t j j 0 0 3 t o O o • 0 • 0 • 0 • 0 0 • 0 • 0 • 0 • 0 • 0 • 0 > 0 0 0 0 o p o o o • 3 < re o 3 r+ O O P o o o p o o o 3 < o 3 O P Moles of CO2 retaineoVmoles of free Ca o t o o o ON © bo ro p o 5 ' 3 Q o t o o o »€>* •a 0/ • O l£)i • a • O ! •o .• 01 o • 0 0 I t o t o I t o U i O "J= 3 O o to t o t-o O "J= 3 Moles of H2S retained/moles of total Ca d o o' •3 §f o o to o 241 Chapter 9 Co-capture ofH2S and C02 C02 in a pressurized-gasifier-based process Moles of C0 2 retained/moles of free Ca ON re o ' ° S" o 5 - CD CT- O P" o CD X to P O O SL O O to 3 P to CM O 00 - p CM 3 o 1:1 O ^ 3 cn O* C P CD 2. o 3 1 3 rc o C/3 o o >° 3 u) p 3 ~ 3> CD P o 3* 00 o o O P 3 P O P o o o CD O 3 o P n o o P o o o 3 < CD O 3 <-+• O O P C/3 o l-t) l-t a 5 3 Moles of H 2S retained/moles of total Ca O •o v 3 3 c p" $ ° 3 ON 00 o o O ON 242 Chapter 9 Co-capture of H2S and C02 C02 in a pressurized-gasifier-based process oo" c N J O r-K £ 8" ON r+ ?^ o < 2 N J fa o> CD < CO _ vi ° s O 3 p p £T o o SL CD — • p c 3 - i O CD 5' o ta i-+> ET. ffi O fT 3 C/5. P P r+ 3 oo a . ^ rN C J N J § N> 2 & p g ^ c« s-- ET 3* P cn 1—' cn O CT 0 C NP . 3 |^  ^ cn 1 ^  -1 o CD *1 P X5 O tt . =r o cn 3 2 P •3 ~ rt- 00 S3. <""} CD S . P o p O o o 3 < CD O 3 o p n cr O P O o o CD O 3 o p CO Moles of C O 2 retained/moles of free Ca p p p p p o o o Moles of H 2 S retained/moles of total Ca o O. LO O O L/l I© s r . " (i • o • o • o • o • o o o o o o o • o — o >-> U l to o o o 243 2 o Si I CD s 13 O o 55 • a £5. a ON CJ o -a c •a 1 T 0.8 0.6 i S, 0.4 u <+-> o "o s 0.2 i O • A No H 2 S, Kel ly Rock limestone Strassburg limestone Kelly Rock limestone Arctic dolomite No H 2 S , Strassburg limestone No H 2 S , Arctic dolomite 10 15 Number of reaction cycles (a) CaO conversion to CaC03 20 o w o 0.6 0.8 o co o E "O CD c ca o * 0.4 co 2 CN ZE 0.2 .0 O Strassburg limestone • Kelly Rock limestone A Arctic dolomite A A o o A A A A , ^ • • 0 0 0 0 0 0 CN 0 5 10 15 Number of reaction cycles 20 (b) CaO conversion to CaS Figure 9.11 Effect of sorbent type on cyclic capture of H 2S and C 0 2 . 212-250 pm. Sorption at 850°C and 0.76 MPa, with 20% C 0 2 ; 1% H 2S, 12.6%v H 2 , l%v CO, balance N 2 . Calcination at 850°C and 101 kPa with 100% N 2 . 3 min for each sorption stage. Chapter 9 Co-capture ofH2Sand CO2 in a pressurized-gasifier-based process (a) (b) Figure 9.12 SEM photos for different sorbents after 15 co-capture cycles. Same test conditions as for Figure 9.10. (a) Arctic dolomite (b) Kelly Rock limestone hi 'fa*''-.1 :- Hi. -•' •<*••• (a) (b) Figure 9.13 Sulfur mapping for calcines of different sorbents after co-capture. Same test conditions as for Figure 9.10. (a) Arctic dolomite after 15 cycles, (b) Kelly Rock after 15 cycles. 245 Chapter 10 Conclusions and recommendations for fiiture work C H A P T E R 10 CONCLUSIONS AND R E C O M M E N D A T I O N S F O R F U T U R E W O R K 10.1 Conclusions Calcium-based sorbents for the capture of CO2 capture were investigated in this thesis, covering topics ranging from reaction kinetics to application in combustors or gasifiers, the mechanism of simultaneous CO2 and SO2 removal, attempts to improve sorbent reversibility and alternative techniques for cyclic removal of C0 2 . The major conclusions drawn in this thesis work are: ( 1 ) Kinetic study of the CaO-C02 reaction found that the reaction rate is first-order with respect to (Pco2,-Pco2,eq), for (Pco2,-Pco2,eq) ^ 1 0 kPa. The order abruptly changes to zero when (Pco2,-Pco2,eq) > 1 0 kPa. Atmospheric pressure TGA results were supported by pressurized PTGA results. A possible explanation is that the intermediate complex CaO»C02 becomes immediately saturated on CaO sites when the CO2 driving force exceeds 1 0 kPa, and the reaction rate then becomes determined by the rate of formation of CaC03 through surface reaction from the complex. Based on the data obtained from the zero-order region, the activation energies were found to be 2 9 ± 4 kJ/mol and 2 4 ± 6 kJ/mol for Strassburg limestone and Arctic dolomite, respectively. ( 2 ) Experimental data obtained in both atmospheric and pressurized thermogravimetric reactors indicate that carbonation is insensitive to the C 0 2 partial pressure in terms of final conversion and apparent carbonation rate. There is a need to predict the carbonation history using a structural model. However, normal structural gas-solid models, such as the grain model and random pore model cannot predict the entire carbonation history, Previous experimental observations for this reaction (Abanades and Alvarez, 2 0 0 3 ; 246 Chapter 10 Conclusions and recommendations for future work Alvarez and Abanades, 2005a) showed that the reaction is highly dependent on pore size distribution. A new reaction model was formulated based on discrete-pore-size-distribution. This model utilizes measured rate constant and measured pore size distribution data as input. The effective diffusivity in the product layer is the only fitting parameter, dependent on the evolution of the pores. The model is able to predict atmospheric and pressurized thermogravimetric reactor carbonation data with fitted activation energies of 215 and 187 kJ/mol for the limestone and dolomite tested. (3) Sorbent reversibility, or sorbent cyclic C O 2 removal ability during calcination/carbonation cycling, is of critical importance for practical application of calcium-based sorbents. The decay in the reversibility of limestones has been found (Salvador et al., 2003; Alvarez and Abanades, 2005a) to be almost independent of reactor type and operating conditions. Similar results were found in our experiments in both an atmospheric pressure TGA unit and an atmospheric pressure thermogravimetric reactor (Chapters 4 and 6). In Chapter 4, a detailed study was conducted showing the evolution of limestone pore size distribution during calcination/carbonation cycling. It was found that when the fast stage of carbonation is completed during each cycle, sintering occurs mostly during the calcination stage. Periodic calcination causes a much higher degree of sintering than holding samples under calcination condition for the same cumulative time because the cyclic release of C O 2 during calcination enhances sintering. The evolution of pore size distribution during calcination/carbonation cycling shows that <~220 nm pores shrink, limiting the achievable conversion for the next fast stage of carbonation. A simultaneous calcination and sintering model was formulated based on a simple description of pore evolution, consistent with experimental evidence. The model can satisfactorily explain the experimental data. It correctly predicts that cyclic C O 2 capture 247 / J Chapter 10 Conclusions and recommendations for future work is not very sensitive to the particle size, sample size and operating conditions. Extrapolation of long-term cycling results predict residual C 0 2 capture ability of -2-3% after 1000 cycles. (4) Because the molar ratio of carbon to sulfur in conventional fuels, e.g. coal, is usually high, e.g. >50, it has commonly been assumed that the calcium oxide needed for S O 2 capture accounts for an insignificant fraction of the sorbent for C O 2 removal. Strassburg limestone and Arctic dolomite were investigated in an atmospheric pressure thermogravimetric reactor for co-capture of S O 2 and C 0 2 at fluidized bed combustion temperatures, e.g. 850°C. A high partial pressure of C O 2 (80%v) had to be maintained to make carbonation favourable. It was found that the sorbent cyclic performance for C O 2 removal was appreciably impeded by the presence of S0 2 , even though the C O 2 concentration was much higher than the S O 2 concentration. Parametric studies indicated that the impeding effect is sensitive to such factors as temperature, C O 2 and S O 2 concentrations, particle s i z e and reaction time. Whereas S O 2 inhibits the capture of C O 2 , carbonation was found to enhance the capture of S O 2 The calcination rate was also observed to decrease as a result of the presence of the sulphate layer which forms during cyclic co-capture. (5) To further investigate the co-capture of S0 2 and C0 2 , five limestones and two dolomites were tested at both atmospheric and elevated pressure, with a maximum total pressure up to nearly 2.4 MPa. When no S O 2 was present, the limestones performed similarly in terms of sorbent reversibility, and dolomites performed better. Conversions decayed more rapidly when S0 2 was present for all sorbents. Direct sulphation became dominant after completion of an initial fast stage of carbonation, filling larger pores by sulphation from the outside, enveloping the sorbents with an impermeable shell, inhibiting further 248 Chapter 10 Conclusions and recommendations for future work carbonation and retarding subsequent calcination. Of the various attempts to improve sorbent reversibility, increasing the CO2 partial pressure was most helpful. (6) In Chapter 7, tests with steam, CO and modified sorbents led to the following conclusions: (i) Steam addition during carbonation did not improve sorbent reversibility when no SO2 was present, but helped somewhat for CO2/SO2 co-capture tests. One-time hydration of calcined limestone improved CO2 removal to some degree with or without SO2 present, (ii) Calcination in a steam environment did not appreciably enhance sorbent sintering. Steam and water hydration were unable to reactivate partially carbonated sorbents to give higher extents of carbonation. (iii). Use of low-temperature steam or liquid water to reactivate highly cycled sorbent can improve CO2 removal. Periodic hydration of highly cycled sorbents can lead to some improvement in reversibility. Steam/water hydration was also found to be effective during cyclic CO2 and SO2 co-capture by breaking sulphate layer, (iv). Use of CO to cyclically regenerate CaO from CaS04 did not give good results because the regeneration of CaS04 is very slow since the CO concentration has to be maintained at a low level to obtain CaO, rather than CaS, as the by-product, (v). Using inerts to modify natural limestone was unsuccessful, except that a 1:1 molar CaO:Al203 mixture showed good reversibility; but low CO2 capture capacity. (7) Given that co-capture of SO2 and CO2 was found to be difficult because SO2 severely impeded CO2 removal, four FBC-based processes were investigated as possible means of sequentially capturing SO2 and C 0 2 . Test results at low and high pressure indicate that the best option involved highly cycled limestones as S02-sorbents in atmospheric FBC combustors after many cycle of CO2 capture. Highly sintered limestone remains a good 249 Chapter 10 Conclusions and recommendations for future work sorbent for SO2 because of the generation of macropores during calcination/carbonation cycling. (8) An alternative route for C 0 2 capture based on calcium sorbent was investigated in which the sorbent is applied to fossil fuel gasification, with C 0 2 removed in situ in the gasifier. Appropriate operating conditions can be chosen to produce a hydrogen-rich syngas stream in a single step. The sorbent performance was investigated with H 2S present in a pressurized TGA system. The results contrasted sharply with those under oxidizing condition with S0 2 and C 0 2 present. No obvious inhibition was found on sorbent performance due to the presence of H 2S. Parametric studies, on the effect of particle size, sorption time, total pressure, temperature and partial pressure of C 0 2 and H 2S and sorbent type showed complex interactions between the solid product layers and diffusing ions. The results generally indicate that co-capture of H 2S and C 0 2 is feasible. Combining the results for sorbent performance with thermodynamic considerations (covered in Chapter 1) shows that one-step hydrogen production from gasification should be achievable. 10.2 Recommendations for future work (1) In Chapter 7, preliminary sorbent modification work demonstrated that improving sorbent reversibility using additives is possible, but at the penalty that sorbent removal capacity decreased. Sorbent modification is a complex topic, more research effort is still needed to develop more stable sorbents. Due the sensitivity of sorbent performance to the preparation procedure, a standardized preparation approach needs to be developed and cross-checked; Investigation is needed on whether there is an optimum mixing ratio that can give best sorbent performance. More work is also needed on sorbent modification. 250 Chapter 10 Conclusions and recommendations for future work Other preparation methods employed in catalyst preparation should be tested, for example, Sol-Gel or co-precipitation; More additives shall be tested; Pelletization technology should be tested together with modifiers. At least one previous study (Gupta, et al. 2004) has shown that pelletization can enhance sorbent reversibility. (2) New sorbents other than calcium-based sorbents should be tested for the combustor and gasifier processes, such as lithium-based sorbents. Since preliminary data (Kato, 2001; Kato et al., 2002a,b) suggest that these are highly reversible sorbents. Brief test results on the carbonation and sulphation ability of L i 4 S i 0 4 are included in Appendix X . (3) Advanced analytical techniques should be used in the study of sorbent performance. For example, in-situ infrared spectroscopy technology could be used to evaluate the stabilizing mechanism by the formation of Cai2Ali4033 in CaO composite. (4) Chapter 7 showed that periodic hydration can improve sorbent reversibility. Chapter 4 also suggested that shorter residence times for each cycle of sorption could reduce sorbent sintering. These methods focusing on affordable natural sorbents deserve more testing. (5) Calcination is a key step, but little work has been done to investigate practical and affordable calcination methods. (6) Most work to date has been at the bench scale using batch reactors. More tests are needed on continuous reactors. If continuous reactors could be operated on interconnected columns, questions on long-term sorbent performance, the effects on recycling ratio of sorbents, and optimal operating methods for the calciner, could be more easily answered. (7) This work showed that calcium-based sorbents have difficulties when co-capture of S 0 2 and C O 2 is required. Applying the sorbents to capture C 0 2 under steam reforming conditions may be more practical. Fixed bed tests (Ortiz and Harrison, 2001) and 2 5 1 Chapter 10 Conclusions and recommendations for future work fluidized bed tests (Johnsen et al., 2006) have all been carried out in batch-wise mode. A challenge is to demonstrate sorption-enhanced hydrogen production in a more practical system, e.g. with interconnected fluidized beds. (8) The proposed process Option B in Chapter 8 deserves tests on a larger scale reactor and detailed economic study. (9) The feasibility of co-capture of H 2S and C 0 2 needs more research. 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(2004b) Technical progress in the development of zero emission coal technologies, www. zeca.org. 264 Appendix I About kinetic data fitting APPENDIX I A B O U T KINETIC DATA FITTING The objective of this appendix is to compare the fitting result by the use if nonlinear fitting and linear fitting for the kinetic data in Chapters 2 and 3 . Al. 1 Dataset and equation used 1. Four datasets Dataset 1: rate constant, ks measurement for Strassburg limestone in Chapter 2 . Dataset 2 : rate constant, ks measurement for Arctic dolomite in Chapter 2 . Dataset 3 : diffusivity, Dp , data for Strassburg limestone in Chapter 3. Dataset 4: diffusivity, Dp , data for Arctic dolomite in Chapter 3. 2. Two equations 56 x Sn 56 x Sn E, , . (j. Equation 1: r0 = ks k0 exp( ) (see equation 2 .10, in zero reaction order 3 3 RT range), S0 -29 m2/g for Strassburg limestone; 31 for m2/g Arctic dolomite, R=0.008314 kJ/mol/K. r, ks and Ex are the kinetic rate, rate constant, and activation energy in Chapter 2. k0 is a pre-exponential factor. Equation 2 : Dp - Dn exp( —) . Dp and E2 are the effective diffusivity and activation energy in Chapter 3 with Dp0 being the pre-exponential factor. 3. Four fittings (all fittings were preformed using M A T L A B 6.5 fitting tools) Fit 1: nonlinear equation: y =_Aexp(-B Ix); data format: ks vs T; then obtain k0 from A, and Ei from B. 265 Appendix I About kinetic data fitting Fit 2: linear equation: y = A + Bx; data format: ln(ks) vs 1/T; then obtain k0 from A, and Ei from B. Fit 3: nonlinear equation: y — Aexp(-B/x) ; data format: Dp vs T; then obtain DpQ from A, and E 2 from B. Fit 4: linear equation: y = A + Bx; data format: ln(Dp ) vs 1/T; then obtain Dp0 from A, and E 2 from B. 4. Goodness of fit parameters: . . . , SSE(n-\) a. Adjusted R-square =1 i -ssr(v-i) SSE: sum of squares of the regression: SSE = ^(j),- - ) 2 SST: total sum of squares about the mean: SST = - j , ) 2 m: number of fitted coefficients n: number of data points v = n-m =number of degree of freedom b. RMSE: Root Mean Square Error, also known as the fit standard error and the standard error of . SSE the regression, 266 Appendix I About kinetic data fitting AI.2 Fitting results The fitting results are shown in Tables A l l and AI.2. For dataset 1 and dataset 2, nonlinear fitting provided a somewhat better result in term of the goodness of fit. The nonlinear results have therefore been adopted in Chapter 2. For dataset 3 and dataset 4, nonlinear fittings failed to give physically correct values, either because of the data quality or the fitting tool used (MATLAB 6.5). The fitting results using linear fitting methods were therefore adopted in Chapter 3. Table A l l Fitting results with nonlinear equations. Values in [] are values with 95% confidence level. Values in bold are utilized in Chapter 2. Results Adjusted R 2 ID Pre-exponential factor ( m ° l or m2/s) Activation energy (kJ/mol) RMSE Dataset 1 £ 0 =0.0016 [0.0008,0.0023] E,=28.85 [25,32] 0.933 0.0025 . Dataset 2 &0 =0.0011 [0.00033,0.0017] Et=23.5 [17.7,29.2] 0.9544 0.0022 Dataset 3 D, 0=-7.07e-9 [-9.14e-8,7.72e-8] E,=0.00507 [-87.55,87.55] -4.1E5 8.65E-9 Dataset 4 £>„ 0 =-6.09 [-1.22E-7, 1.08E-7] E2=0.005 [-132, 132] -3.1E6 8.9E-9 267 Appendix I About kinetic data fitting Table AI.2 Fitting results with linear fit to logarithmic equations. Values in [] are values with 9 5 % confidence level; Values in bold are adopted in Chapter 3 . Results Adjusted R 2 ID Pre-exponential factor ( m ° l orm2/s) Activation energy (kJ/mol) R M S E Dataset 1 k0 =0.00367 [0.0019,0.00789] E,=35.3 [40.9,29.7] 0.90 0.1765 Dataset 2 &0 =0.0011 [0.0005,0.0025] E,=24.05 [17.75,29.25] 0.937 0.08 Dataset 3 D f0=0.27 [1.8E-2,4.13] E2=215 [249,183] 0.94 0.952 Dataset 4 Dp0 =0.0008 [1.67E-5,0.038] E2=187 [248,126] 0.877 1 Al. 3 Comparisons with experimental data and equations with fitted parameters The fitted kinetic data are then used to compare the experimental data. The comparisons appear in Figures AVI. 1-6. Note that for dataset 3 and 4, the prediction using nonlinear fitting results were not able to present physically correct values and therefore are not plotted in Figures AVI. 5 - 6 . 268 Appendix I About kinetic data fitting 0.00012 0.0001 V - 0.00008 ^ 0.00006 O 0.00004 0.00002 600 • Experimental points - - Nonlinear fit Linear fit 700 800 900 1000 1100 1200 T(K) 1300 Figure A l l Comparison of experimental and fitting results (Dataset 1) l o.i 0.01 _ 0.001 o ^ 0.0001 \-0.00001 0.000001 • Experimental points - - - Nonlinear fit —:—Linear fit 0.6 0.8 1 1.2 1/T* 1000 (1/K*1000) 1.4 1.6 Figure AI.2 Comparison of experimental and fitting results (Dataset 1). Same as above but i Arrhenius plot 269 Appendix I About kinetic data fitting 0.00012 r 0.0001 ^ 0.00008 ~ 0.00006 o G, 0.00004 0.00002 0 • Experimental points - - - Nonlinear fit Linear fit 600 700 800 900 1000 1100 1200 1300 T(K) Figure AI.3 Comparison of experimental and fitting results (Dataset 2) 1 J2 0.1 0.01 0.001 0.0001 0.00001 • Experimental points - - - Nonlinear fit Linear fit 0.6 0.8 1 1.2 1.4 1/T* 1000 (1/K* 1000) 1.6 Figure AI.4 Comparison of experimental and fitting equation (Dataset 2) Same as above but Arrhenius plot 270 Appendix I About kinetic data fitting l.OOE-05 1.00E-07 ^ 1.00E-09 l.OOE-11 l.OOE-13 l.OOE-15 1.00E-17 • Experimental points - - Linear fit 0.6 0.8 1 1.2 1.4 1/T* 1000 (1/K* 1000) 1.6 1.8 Figure AI.5 Comparison of experimental and fitting equation (Dataset 3) in Arrhenius plot (Data predicted by the nonlinear fitted equation is off this chart) Q 1.00E-05 1.00E-07 1.00E-09 1.00E-11 1.00E-13 1.00E-15 1.00E-17 > Experimental points - Linear fit 0.6 0.8 1 1.2 1.4 1/T* 1000 (1/K* 1000) 1.6 1.8 Figure AI.6 Comparison of experimental and fitting equation (Dataset 4) in Arrhenius plot (Data predicted by the nonlinear fitted equation is off this chart) 271 Appendix I About kinetic data fitting AL 4 Conclusions Based on the 95% confidence level data range, nonlinear fitting gives better results for the experimental rate constants for datasets 1 and 2 where the activation energy is relatively small. In these cases, the fitted results and goodness of fitting are similar between the linear fitting and the nonlinear fitting. For the diffusivities (where the activation energy is much larger), nonlinear fitting failed to yield physically meaningful results, either because of the failure of the fitting tools or because of the quality of data. Hence a linear fit was applied after taking logarithms. Based on the results in this Appendix, the best fitting methods may be related to the range of activation energy itself. As the temperature step size in studies is usually uniformly distributed by experimenters, when activation energy is large, the reaction rate increases quickly with temperature, tending to make the measured data points too dispersed to give an accurate nonlinear fit when there are not enough data points in the high temperature. 272 Appendix II Molar volume ratio for a dolomite employed in the model of Chapter 3 APPENDIX II M O L A R V O L U M E RATIO F O R A D O L O M I T E E M P L O Y E D IN T H E M O D E L O F C H A P T E R 3 (The symbols employed in this appendix are summarized in the nomenclature section of Chapter An XRD study by Engler et al. (1988) has shown that no solid solution is present for fully calcined dolomite. Thus a fully calcined dolomite can be considered as a physical mixture of MgO and CaO, with other impurities as minor constituents. CaO and MgO crystals are both cubic, with the cube edge length 4.8xl0"9 m for CaO and 4.2xl0"9 m for MgO (Boynton, 1979). Considering a dolomite with molar ratio of CaO to MgO of i\, the molar volume ratio of CaO to MgO is 4 83 8 = — rj (AIL l) For the Arctic dolomite investigated in the current work, r\ =1.1 so that 8 =1.64. Note that the surface area is also composed of both CaO and MgO. With the assumption that these are randomly interspersed, the surface area ratio of CaO to MgO is also 8. At time t, the reactant radius surface has reached Ri r and the product radius is Rt as illustrated by Figure 2.2. The solid matrix inside the annular region Vir-Vj contains not only product CaC03, but also inert MgO. In this region, the volumes of the different compartments, i.e. CaO consumed, CaC03 produced and MgO remaining are rao=<r,,r-r,.0)T^-= (An.2) 1 + 8 yCaco^(y,,r-Vifi)-^-Z (AH3) 1 + o > i 4 0 = ( ^ - ^ . o ) Y 7 J ( A H - 4 ) 273 Appendix II Molar volume ratio for a dolomite employed in the model of Chapter 3 If the annular layer Vir-Vi p is assumed to be non-porous, a mass balance on the carbonation reaction leads to, Vi.r-v^ = vCaCm+vMg0=(vir - K,o)r^z +(vi, - ^.0)777 ( A n 5) Rearrangement leads to Vip=rVi0-(Z'-l)Vir (AH.6) where Z'= Z + is the molar product-to-reactant volume ratio for the dolomite. \+S l+S Equation (A U.6) for a dolomite is identical with equation (2.32) for a limestone, only with Z', a lower molar ratio of product to reactant, replacing Z. 274 Appendix III Initial pore size distribution employed in the model of Chapter 3 APPENDIX III INITIAL P O R E SIZE DISTRIBUTION E M P L O Y E D IN T H E M O D E L O F C H A P T E R 3 (The symbols employed in this appendix are summarized in the nomenclature section of Chapter 3) Mercury intrusion is known to be able to measure the sizes of larger pores than gas adsorption, so the former was used to obtain the initial pore size distribution in the current study. It is not clear from the literature whether or not the initially measured pore volume involves overlapping, Here this issue is addressed and the initial condition from mercury intrusion data is obtained. Bhatia and Perlmutter (1980) used measured total pore volume Vm as the non-overlapping pore volume VQE. With the aid of equation (2.2), this leads to, V0=l-exp(-V0E) = \-exp(-VJ*Vm (ATJI.1) where V0 is the initial pore volume considering overlap. Gavalas (1980) derived the probability density X(Rim) by initial mercury intrusion or gas adsorption data, Vim (Rim ) , mm) = V'm{R'i R (AUI.2) where Rim (i=l to N), is obtained from the mercury intrusion data, calculated from Washburn equation (Lowell and Shields, 1991). The value of A f ^ , ) is close to /(i?, m)/2. Gavalas (1980) also showed that the initial total pore volume can be expressed as, V0 = e = 1 - exp(-X2^,^,m2) (Am.3) 1=1 275 Appendix III Initial pore size distribution employed in the model of Chapter 3 Equation (AJJI.3) indicates that the measured pore volume does not account for overlap. This is a reasonable approximation when the initial volume is small, as ^ goes to unity, V0 by (AJJI.3) is the total measured volume V0 * (j^lnX^J) = Vm , Vm = £ (/?,,m) and Vjm(Rjm) is the measured pore volume of each pore radius Rim. However usually the initial volume is not small enough for this approximation to be valid, so at time zero, V0 determined from equation (AIU.3), is usually smaller than the measured pore volume Vm. This contradicts the assumption that the measured pore volume Vm should equal the true volume V0 at time zero. To deal with this difficulty, a simplified treatment is introduced below. It is believed that the primary difficulty arises from the definition of pore radius. In mercury intrusion data, if pores are assumed to be cylindrical, pore radius is related to intrusion pressure by the Washburn equation. The measured nominal pore size is Rim, i=l to N. Pore volume is Vim(Rim), so that total pore volume is Vm = Z Vum (R,,m ) = Z " O i • (AJH.4) i=i In the current model, a different set of cylindrical pores is needed with initial pore radius RiQ (i=T to N). When considering overlap, the overall pore volume V0 is set equal to the measured Vm. 276 Appendix III Initial pore size distribution employed in the model of Chapter 3 Initial radii of pores with overlap as imput of the model ==> Measured equivalent radii represent intruded Hg volumes-=> Figure AIU. 1 Illustration of the relation between initial pore radii based on measured pore volume data The relationship between the two sets of pores is illustrated in Figure AJU.l. Each incremental pore volume corresponds to a certain pore radius, so that there should be no overlap. Rim can be regarded as an equivalent radius of voids filled at a certain intrusion pressure, whereas the pore system with radius Ri0 (i=l, N) introduced here involves overlap. The two sets of radii share the same total pore length, i.e. '(*,>) = '(*,.<>)=/, i=l to N (AJJI5) Total pore length per unit volume (m/m3), is obtained given from the measured data, KKm) = Vim(fi;) =h i=l toN , (AEI.6) In the pore system with overlap, V0 = 1 - exp(-F0 E) = 1 - exp(-Z (AEI.7) 277 Appendix III Initial pore size distribution employed in the model of Chapter 3 where V0E represents the pore volume without overlap. The two sets of pore system (Rjfi vs Rim, or V0 vs Vm ) are related by the initial total pore f volume, i.e. V0 equal Vm, so that V0=l-exp(-V0E) = Vm (Affl.8) Equation (AJJI.8) further leads to V0E=-ln(l-Vm) ( A J H 9 ) N N N N M VO,E = HVUO,E =Y J 7 d i R i f i 2 a n d vm =YJ7d,Ri,m2 , the two sets of pores can be i=l related by, » l , 0 , £ - - — K i , m ^ . m ^ , « i , m (AJJI. 1 0) m m m Since K 0 B = TZ/,./?,. 0 2 , equation (ALT. 10) finally gives "'•'-I^T1*- < A f f l l l ) 2 Pore radius Rt0 and Vi0E = TZ/ ,^, 0 are used as initial model input. Finally the directly measured pore volume Am (m3 pore/kg particle) needs to be converted to the fractional pore volume Vm (m3/m3) in equation (Affl.l 1). For a differential pore size distribution A n = Si Am (AJJI. 12) 1 = 1 where Aim (in m 3 pore/kg particle) is the differential pore volume enclosed by pores of radius ft.. 278 Appendix III Initial pore size distribution employed in the model of Chapter 3 For measured pore sizei? i m, the fractional volume Vim (m3 pore/m3 space) can be related to the measured volume fraction by: V,m = , Afs (Am. 13) N which gives the value of Vm in (AJJI.11) with Vm = ^ JVI,m . 1=1 279 Appendix IV A Fortran program employed in Chapter 3 APPENDIX IV A F O R T R A N P R O G R A M E M P L O Y E D IN C H A P T E R 3 c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c CC-A COMPUTER PROGRAMM FOR CARBONATION CALCULATION-CC-A NEW STRUCTRUAL MODEL CC-NOTE: A L L T H E OTHER RELATED CALCULATIONS CAN B E M A D E B Y CC-ADJUSTING SOME STATEMENTS C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C PROGRAM PORES PARAMETER(MR=29, MT=518000) DIMENSION RO(MR),VOBAR(MR),VO(MR),RR(MR),RP(MR),C(MR),VRE(MR), + WE(MR),\T.(MR),T(MT),X(MT),TfflCK(MT) ;CCC(MR,MT) , REAL MCAO,K0,KS,KSB REAL L0(MR), LA(MR) c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c CCCCC--VARJBLES-DEFENTTIONS— c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c CCCCC--MR: PORE GROUP NUMBERS; MT: TIME STEPS--CCCCC--R0, V0BAR,V0 RELATED MEASURED PORE PROPERTIES~ C C C C C - C IS C02 CONCENTRATION-CCCCC--VRE,VPE ARE VARTBLES WITH NO CONSIDERATION OF O V E R L A P -CCCCC--T INCLUDE TIME; X IS CONVERSION; THICK IS PORE W A L L CCCCC-THINKNESS— CCCCC--MCAO, KO, KS, KSB ARE A L L CONSTANT RELATING R A T E LAW--CCCCC--L0 PORE L E N G T H -CCCCC—V0BAR(T) the increamental volume in cc/g— C C C C C - - A L L UNITS SHOULD BE CONVERTED INTO SI UNITS; --CCCCC--DCA IS DENSITY IN KG/M3; INPUT PORE RADII IN ANSTRANM, THEN CCCCC—CHANGED to M CCCCC-INPUT PORE V O L U M E VOB AR AND VT CONVERTED F R O M C C / G TO CCCCC--M3/KG C C C C C - M O L A R WEIGHT MCAO IN KG/MOL--CCCCC--INPUT CONCENTRATION ARE PNITIALL IN BAR THEN CONVERTED TO C-C C C - M O L / M 3 -CCCCC--THERMOAL CONSTANT R EST CM3/K/MOL--C C C C C - K S AND KSB ARE RATE CONSTANT IN DIFFERENT R X N ORDER ZONE-C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C WRJTE(*, * )'RUNNING ' C C C C C C - INPUT MEASURED PORE PARAMETERS-C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C OPEN(l, FILE=,inputlimestone.txt,,STATUS^OLD1) DO 1=1, M R READ (1, *) R0(I), VOB AR(T) END DO 280 Appendix IVA Fortran program employed in Chapter 3 CLOSE (1) DR=10000*1. E-10/MR DCA=3.34*1.E3 VT=0.0 DOI=l,MR LA(r)=V0BAR(I)/(R0(I)**2.) R0(I)=(l.E-10)*R0(lj V0BAR(I)=V0B ARQ)* 1 E-3 VT=VT+VOBAR(T) END DO VOLUME=0.0 DO 1=1,MR V0(D=V0B AR(I)*DC A/(l+VT*DCA) RR(f)=R0(I) RP(1)=R0(I) VRE(D=V0(I) VPE(I)=V0(I) L0(r)=V0(T)/3.1415/(R0(I)**2.) VOLUME=VOLUME+V0(I) END DO VPRTME=-Log( 1. -VOLUME) VW=VPRIME/VOLUME DO 1=1,MR ' R0(l)=(V0(r)*VW/3.1415/L0(I))**0.5 RR(f)=R0(I) RP(I)=R0(I) VRE(I)=V0(I)*VW VPE(I)=V0(I)*VW END DO cccccccccccccccccccccccccccccccccccccccccccccccccccccccc CCCCC-INPUT DENSITY, STEP LENTH, Z V A L U E , TEMPERATURE, R A T E CCCCC-CONSTANT CCCCC-INITIAL VALUES, FITTED DffFISrVTTY, EQUTLIRrrjM CONCENTRATION cccccccccccccccccccccccccccccccccccccccccccccccccccccccc MCAO=56.*l.e-3 DTIME=0.006 Z=2.17 PC=20. TT=600 PE=10**(-8308/(TT+273)+7.079) TSTOP=MT C0=(PC-PE)/(82.05e-6*(TT+273.15)) CSTAR=0. l/(82.05e-6*(TT+273.15)) 281 Appendix IVA Fortran program employed in Chapter 3 DP=1.22e-13 KS=3.1e-5 KSB=KS/CSTAR N=2 NT=1 T(1)=0.0 X(1)=0.0 DO NN=2,MT C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C - C A L C U L A T E D NEW REACTION FRONT-C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C VRET=0.0 DO MM=1,MR 1F((1 .El 0*RP(MM)).LE. 1.) THEN C(MM)=0.0 RK=DTIME*MCAO/DCA*KS*C(MM) GOTO 100 END IF IF(NN.EQ.2) THEN C(MM)=C0 RK=DTIME*MCAO/DCA*KS ELSE C(MM)=C0-KS/DP*RR(MM)*LOG(^ IF (C(MM).LE.CSTAR)C(MM)=C0/(1 +KSB/DP*RR(MM)*LOG(RR(MM)/RP(]viM))) CCC(MM,NN)=C(MM) IF (C(MM).LE.CSTAR) THEN N=l RKl=DTIME*MCAO/DCA*KSB*C(MM) RK2=DTEVffi*MCAO/DCA*KSB*C0/(l .+KSB/DP*(RR(MM)+1 ./2.*RK1) + *L0G((RR(MM)+1./2.*RK1)/RP(>1M))) RK3=DTIME*MCAO/DCA*KSB*C0/(l +KSB/DP*(RR(MM)+1V2.*RK2) + *LOG((RR(MM)+l./2.*RK2)/RP(lvlM))) RK4=DTIME*MCAO/DCA*KSB*C0/(l +KSB/DP*(RR(MM)+RK3) + *LOG((RR(MM)+RK3)/RP(MM))) RK=1./6.*(RK1+2.*RK2+2.*RK3+RK4) ELSE C(NfM)=C0-(KS/DP*RR(MM)*LOGT^(M^ N=0 RK=DTIME*MCAO/DCA*KS END IF ENDJJF 100 CONTINUE RR(MM)=RR(MM)+RK V1^(MM)=3.1415 9*L0(MM)*RR(MM)* *2. VRET=VRET+VRE(MM) 282 Appendix IV A Fortran program employed in Chapter 3 END DO VR=1.-EXP(-1.*VRET) VP=(1 -Z)*(W-VOLUME)+VOLUME c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c C C C C C C C C C C C C C C C - C A L C U L A T E D NEW PRODUCT FRONT-c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c VPET=0.0 DOK=l ,MR EF((1 E l 0*RP(K)).LE. 1.) THEN RK=0.0 GOTO 200 END IF IT(RP(K).LE.O.) THEN RP(K)=0.0 RK=0.0 GOTO 200 EN DIF IF (N.NE.O) THEN C(K)=C0/(1 .+KSB/DP*RR(K)*LOC<RR(K)/RP(K))) • RKl=DTIME*MCAO/DCA*KSB*C(K)*(l -Z)*(1-VR)/(1 -W)*RR(K)/RP(K) IF((RP(K)+l./2*RKl).LE.O.) THEN RP(K)=0.0 RK=0.0 GOTO 200 ENDIF C(K)=C0/(1 +KSB/DP*RR(K)*LOGrRR(K)/(RP(K)+l./2.*RKl))) RK2=DTIME*MCAO/DCA*KSB*C(K)*(l -Z)*(l -VR)/(1 -VP) + *RR(K)/(RP(K)+1./2.*RK1) IF((RP(K)+l./2.*RK2).LE.O.) THEN RP(K)=0.0 RK=0.0 GOTO 200 ENDIF C(K)=C0/(1 +KSB/DP*RR(K)*LOG(RR(K)/(RP(K)+l ./2*RK2))) RK3=DTIME*MCAO/DCA*KSB*C(K)*(l -Z)*(l.-VR)/(1 -VP) + *RR(K)/(RP(K)+1./2.*RK2) IF((RP(K)+RK3).LE.0.) THEN RP(K)=0.0 RK=0.0 GOTO 200 ENDIF C(K)=C0/(l.+KSB/DP*RR(K)*LOC-(^(K)/(RP(K)+RK3))) RK4=DTlME*MCAO/DCA*KSB*C(K)*(l-Z)*(l-VR)/(1-VP) + *RR(K)/(RP(K)+RK3) RK=1./6.*(RK1+2.*RK2+2.*RK3+RK4) ELSE C(K)-C0-KS/DP*RR(K)*LOG<RR(K)/(RP(K))) 283 Appendix IVA Fortran program employed in Chapter 3 RK=DTIME*MCA0/DCA*KS*(1 -Z)*(l-VR)/(1 -W)*RR(K)/T^(K) END IF 200 CONTINUE RP(K)=RP(K)+RK VPE(K)=3.14159*L0(K)*RP(K)**2. VPET=VPET+VPE(K) TfflCK(k)=RR(k)-RP(k) END DO VP=1.-EXP(-1.*VPET) c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c CCCCCC-UPDATE O V E R A L L CONVERSION AND SAVE D A T A -c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c c X(NN)=1 .-(1 .-VR)/(1 .-VOLUME) NT=NT+1 T(NN)=DTIME*(NN-1) IF (T(NN).GT.TSTOP) GOTO 400 END DO 400 CONTINUE OPEN(2,FILE= ,DATAX.DAT•,STATUS='OLD ,) DO J=l, 10000,10 WRTTE(2,*)T(J),X(J),CCC(14,J) END DO CLOSE(2) C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C C END 284 Appendix V Linearization of Equation (4.5) APPENDIX V LINEARIZATION O F E Q U A T I O N (4.5) (The symbols employed in this appendix are summarized in the nomenclature section of Chapte 4) The objective of this appendix is to linearize the following equation, ^^.Di^.Paaam.ti)mi> (AV,) dz dz dt The boundary conditions are, Atz=0, C(t,z) = 0 (AV.2) A t z = L , ^ ^ = 0 (AV.3) dz ! C(z) dX K„ e e (AV.4) is rewritten as dX K. dt f(X) with e e (AV.4) (AV.5) / W = P c ^ 3 ( ^ { - i — + ( 1 f " " 3 + T ^ r [ ( l - * r 1 / 3 - l ] } (AV.6) so that equation (AV. 1) can be written as l C(z) dC(t,z) ^ d2C(t,z) t w i x Ke M ^ - j ^ - D , & V -PcacoAn)(\-eb)-j^ (AV.7) At any time step, from t to t+dt, use one-order upwind schemes to linearize equation (AV.7) 285 Appendix V Linearization of Equation (4.5) ubed ~J±Lr~~L ~ Dz — ~ — 7 ^— = Paaam-^)-F^r (AV.8) Az Az f\X) After rearrangement, 1 Az Az 2 / ( ^ J ' + l l A z 2 J — , L Az A z J f(X)Ke J (AV.9) At boundary points: C,=0 (AV.10) CM=CM_, , (AV.l l ) M is the maximum number of the cells The discrete equation (AV.9) complies with the four principles proposed by Partankar (1980) order to achieve convergence during iterative solving. 286 Appendix VI A Matlab program employed in Chapter 4 APPENDIX V I A M A T L A B P R O G R A M E M P L O Y E D IN C H A P T E R 4 %%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%0/o%% %Simultaneous calcination and sintering modeling %%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%% global r rous km Ke ks fbed ubed ez stepl lfinal %%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%% %% input fitted parameters, temperature, cycle numbers, experimental data, %% initialization, reactor dimensions, gas property etc. %% xcarb: carbonation conversion at each cycle; scycle, ss: surface area; %% ecycle, ee,ee2: porosity %% units: mass in gram; rous is true density of CaC03, 2.71g/cc, converted to %% mol/m3; rounn2 is density for N2; visa is viscosity; variables are usually in SI units, if%% not specified %%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%%% maxcycle=1000; T=850; AA=1.2591; BB=0.5922; Z = 2.71/100/(3.34/56); xcarb=zeros( 1 ,maxcycle+1); scycle=zeros(l ,maxcycle); ecycle=zeros(l ,maxcycle); tcycle=zeros(l ,maxcycle); crcycle=zeros(l ,maxcycle); vbed=zeros(l ,maxcycle); porel=zeros(l,maxcycle); pore2=zeros(l,maxcycle); eporosity=zeros(l ,maxcycle); ee2=zeros(l,maxcycle); xcarb(l)=l. r= 0.000212/2.; mass=l; diabed=2.*0.015; diapan=2.*0.006; fbed=0.5; coldflow=2000/60*l.e-6; hotflow=coldflow*(T+273.5)/(20.+273.5); area=pi*(diabed/2.)A2; areapan=pi*(0.01/2)A2.; areapass=area-areapan; 287 Appendix VI A Madab program employed in Chapter 4 diapass=2. *(areapass/pi)A0.5; uO=hotflow/area; hbed=mass/areapan/(l -fbed)/(2.71 * 1 .e6); roun2=1.229*(20+273.5)/(T'+273.5);%inkg/m3 visa=3.93*l.e-5; %in Pas or kg/m/s at 850 deg C; 4.17 AT 800, 3.93 at700 deg, %% 4.4 at 900 deg C visb=visa/roun2; %in m2/s % % % % % % % % % % % % % % % % % 0 / o % 0 / o % % % % % % % % % % % % % % 5 %% Transport property in the reactor, step length and cell length % % % % % % % % % % % % % % % % 0 / o % % % % 0 / o % % % % % % % % % % % % 5 ubed-0.012; upass=(uO*area-ubed*areapan)/areapass; ergun=(150*(l-fbed)A2/fbedA3*visa*ubed/rA2+1.75*(l-... fbed)/fbedA3 *roun2*ubedA2/r)*hbed; red=upass*diapass/visb; coeff=0.96; beida=diapass/diabed; qm=coefTt:areapass*(2.*roun2*ergun)A0.5/(l-beidaA4.)A0.5; qv=upass*areapass; %—about steps 1=0.0; stept=l; lastt=1500; tfinal=floor(lastt/stept); lfinal=20; stepl=hbed/(lfinal-l); rep=ubed*2.*r/visb; d0=0.163*l.e-4*((T+273.5)/(20.+273.))A1.75; sc=visb/d0; shO=0.6*(rep*sc)A(l/2.); %for creeping flow km=sh0*d0/r; ez=2.4*10A(-5) % Obtained from charts of Chemical Reaction Engineering, 2rd ed. %LevespielP310 ez=2*r*ubed/Peclet*0.2; Peclet=2*r*ubed/ez; sa=1.5; sg=70; eg=l-Z; radius 1 =25 *l.e-9; iadius2=300*l.e-9; rouscaco3=2.71/100.*l.e6; 288 Appendix VI A Matlab program employed in Chapter 4 rouscao=3.34*l.e6; % g/m3 vg=eg/ro us cao/(1 -eg); coeffa=(sg-sa)/(vg); %—kinetics data—— Ke=l0A(-8308/(T+273)+7.079)/(82.05* 1 .e-6*(T+273.15));%Ke in mol/m3 ks=3.013*l.e7*exp(-200/0.008314/(T+273.5)); %in mol/m2/s % % % % % % % % % % % % % % % % % % 0 / o % % 0 / o % % % % % % % % % % % 0 / o % % %% Calculation of the outer loop %%%%%%%%%%%%%%%%%%% % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % 0 / o % for ncycle=l :maxcycle x=zeros (tfi nal, lfinal); c=zeros(tfinal,lfinal); creal=zeros(tfinal,lfinal); cr=zeros(tfinal,lfinal); ss=zeros(tfinal,lfinal); ee=zeros(tfinal,lfinal); fff=zeros(l, lfinal); crbed=zeros(l ,tfinal); xbed=zeros(l ,tfinal); drive=zeros(l ,tfinal); sbed=zeros(l ,tfinal); ebed=zeros(l ,tfmal); xint=zeros(l,lfinal); de int=zeros( 1, lfinal); rous=2.71/l 00. * 1. e6*xcarb(ncycle); if (ncycle > 1) for mm=l:lfinal ss(l,mm)=si; end else for mm=l:lfinal ss(l,mm)=si; end end for i=l:lfinal %de(i)=0.163 * 1 .e-4*((T+273,5)/(20.+273 .))A1.75 *e(i)A2; de(i)=0.163*l.e-4*((T+273.5)/(20.+273.))A1.75*0.54A2; end 289 Appendix VI A Madab program employed in Chapter 4 %% Call subroutine to solve concentration c— for i=l:tfinal xint=x(i,:); c(i,:)=solvec(xint,de); for j=l:lfinal rkl =stept*(l .-c(i j)/Ke)/fx(min(x(i,j),0.999999),de(j)); rk2=stept*(l .-c(i,j)/Ke)/fx(min(x(i ,j)+rkl/2.,0.999999),de(j)); rk3=stept*(l .-c(ij)/Keyfx(min(x(ij)+rk2/2.,0.999999 rk4=stept*(l.-c(ij)/Keyfx(min(x(ij)+rk3,0.999999),d rk=l/6.*(rkl+2.*rk2+2.*rk3+rk4); if(rk+x(ij)>l) rk=0.99999999-x(ij); end x(i+lj)=x(ij)+rk; creal(i+l j)=creal(i,j)+rous*rk; pre=rous*r/ks; cr(i+lj)=pre*(l.-min(e(ij)/KeJ0.9999))/f3c(min(x(y),0.999999),de(j)); mps=2.45 *(l+AA*(cr(i+l j)ABB))*exp(-29600/(T+273.5))*(ss(i j)* 1000-sa* 1000)A2 ps=(sg-sa)*rk*xcarb(ncycle); ss(i+l,j)=ss(ij)+ps-stept*mps; end %~update properties for this moment— xbed(i)=0.0; drive(i)=0.0; ebed(i)=0.0; sbed(i)=0.0; crbed(i)=0.0; forj=l:lfinal; • crbed(i)=crbed(i)+cr(i,j)*stepl; xbed(i)=xbed(i)+x(ij)*stepl; drive(i)=drive(i)+c(ij)*stepl; ebed(i)=ebed(i)+ee(ij)*stepl; sbed(i)=sbed(i)+ss(ij)*stepl; end crbed(i)=crbed(i)/hbed; xbed(i)=xbed(i)/hbed; drive(i)=l -drive(i)/hbed/Ke; sbed(i)=sbed(i)/hbed; ebed(i)=ebed(i)/hbed; ifx(i,lfinal)>0.99995 290 Appendix VI A Matlab program employed in Chapter 4 crcycle(ncycle)=crbed(i-l 0); scycle(ncycle)=sbed(i); ecycle(ncycle)=ebed(i); tcycle(ncycle)=i*stept; break end end OA 0 / 0 / 0 X 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / / O /O /O / O / O / O / O / O / O / O /O /O / O / O / O / O /O / O / O / O / O / O /O / 0 7 0 7 0 7 0 7 0 7 0 7 0 7 0 7 0 7 0 7 0 7 0 7 0 7 0 7 0 / 0 7 0 %% Update properties for this cycle, calculate VI and V2, surface area and %% porosity will be used as initial consitions for next cycle 07 , 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / . 0 / 0 / 0 / 0 / 0 / 0 / . 0 / O / O / O / O / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / /O /O /O / O /O /O /O / 0 70 70 70 70 70 / O / O 70 70 / O / O / O 7 0 / O /O 7 0 / O / O 7 0 7 0 7 0 7 0 7 0 7 0 7 0 7 0 7 0 7 0 7 0 7 0 7 0 7 0 vbed(ncycle)=ecycle(ncycle)/rouscao/(l-ecycle(ncycle)); porel(ncycle)=vg+(scycle(ncycle)-sg)/coeffa; pore2(ncycle)=vbed(ncycle)-porel(ncycle); xcarb(ncycle+l )=pore 1 (ncycle)*rouscaco3 *56/(l -Z); ee2(ncycle+l)=(ecycle(ncycle)*Z+(l-Z)*xcarb(ncycle+l)*(ecycle(ncycle)-l))/(l ecycle(ncycle)); eporosity(ncycle+l)=ee2(ncycle+l)/(ee2(ncycle+l)+xcarb(ncycle+l)*(l-Z)+Z); end % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % 0 / o % % % % % % % Subroutine for gas-solid rate model OA.OA0/0/0/0/0/0/,0/0/ 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 / 0 7 0 / 0 / 0 / 0 / 0 / /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o /o function f=fx(conve,diffu) global r rous km Ke ks fbed ubed ez stepl lfinal aa=l/3./km/Ke; bb=r/37diffu/Ke*((l-conve)A(-l/3.)-l); cc=l/3.*ks*(l-conve)A(-2./3); f=r*rous*(aa+bb+cc); % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % 0 / o 0 / o % % % % % % Subroutine for discrete fixed bed mass balance equation % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % % 0 / o % function conl=solvec (xxx,dede) global r rous km Ke ks fbed ubed ez stepl lfinal f=zeros(l ,lfinal); aip=zeros(l,lfinal); aie=zeros(l,lfinal); ai w=zeros( 1, lfinal); con=zeros(l,lfinal); sp=zeros(l,lfinal); pt=zeros(l, lfinal); qt=zeros(l, lfinal); 291 Appendix VI A Matlab program employed in Chapter %aie(l)=ez/steplA2; %aiw(l)=0.0; %con(l)=0.0; %sp(l)=-l.*ubed/stepl; %aip(l)=aie(l)+aiw(l)-sp(l); aie(l)=0.0; aiw(l)=0.0; con(l)=l.e30*0.0; sp(l)=-l.e30; aip( 1 )=aie( 1 )+ai w( 1 )-sp(l); aie(lfmal)=0.0; aiw(lfinal)=l.e30; con(lfinal)=0.0; sp(lfmal)=0.0; aip(lfmal)=aie(lfinal)+aiw(lfinal)-sp(lfmal); for H2:lfinal-1 aie(i)=ez/step 1A2.; ai w(i)=ub ed/step 1+ez/step 1A2.; con(i)=rous*(l -fbed)/fx(xxx(i),dede(i)); sp(i^l.*rous*(l-fbed)/Ke/fx(xxx(i),dede(i)); aip(i)=aie(i)+aiw(i)-sp(i); end pt(l)=aie(l)/aip(l); qt(l)=con(l)/aip(l); for i=2:lfinal denom=aip(i)-pt(i-1 )*aiw(i); if(abs(denom)<=l .e-18) denom=l.e-18; end denom=l ./denom; pt(i)=aie(i)*denom; qt(i)=(con(i)+aiw(i)*qt(i-l))*denom; end if(qt(lfinal) < 0.0) qt(lfinal)=0.0; end f(lfinal)=qt(lfinal); for ii—2:lfinal iii=lfinal+l-ii; f(iii)=f(iii+l)*pt(iii)+qt(iii); end 292 Appendix VI A Madab program employed in Chapter conl=f; 293 Appendix VII Supplementary results within the scope of Chapter 6 APPENDIX VII S U P P L E M E N T A R Y RESULTS WITHIN T H E SCOPE O F C H A P T E R 6 Figure AVTI. 1 (Comparing with other runs in Figure 6.2) High-resolution S E M pictures of calcines, with no S O 2 present. Same test conditions as in Figure 6.1. Danyang limestone. (a) S 0 2 only (b) co-capture of S 0 2 and C 0 2 Figure A W . 2 (Comparing with other runs in Figure 6.11) Sulfur mapping for non-calcined ATGR samples. 850C for both tests. S 0 2 only test: 2900 ppm S0 2 , 3% 0 2 and balance N 2 for 2 hours; Co-capture test conditions as in Figure 6.6, 15 cycles. (Light points mark sulfur). Danyang limestone. 294 Appendix VII Supplementary results within the scope of Chapter 6 Figure AVJJ..3 (Comparing with other runs in Figure 6.3) Evolution of pore size distribution with calcination/carbonation cycling at 850°C Test conditions: same as in Figure 6.1 for S O 2 test and in Figure 6.6 for co-captures. Danyang limestone. Q 10 100 1000 Mean pore diameter (nm) Figure AVU4 (Comparing with other runs in Figure 6.3) Evolution of pore size distribution with calcination/carbonation cycling at 850°C. Test conditions: same as in Figure 6.1 for no S O 2 test and in Figure 6.6 for co-captures. Havelock limestone 295 Appendix VII Supplementary results within the scope of Chapter 6 -« 0 2 13 •*-» o o ui £ f a o . i o o o • 1.8 MPa • 1.3 MPa A 2.4 MPa t • • A 5 10 Number of reaction cycles 15 Figure AVII. 5 (This is for conversion of CaO to CaS0 4, comparing that for CaO to CaC0 3 shown in Figure 6.9) Cyclic C 0 2 retention performances in 850°C PTGA tests: effect of total pressure. Co-capture with 212-250 pm Strassburg limestone. Sorption: 8 %v CO2, 1125 ppmv S0 2 , 3% 0 2, and balance N 2 , 4 minute for each cycle. Calcination: 101 kPa, 100% N 2 . o ui *o S -a a, o o co <+* o U5 0.2 CJ 13 0.1 -I -*-> o • 1.8 MPa A 2.4 MPa • 1.3 MPa 2 4 6 8 10 Number of reaction cycles —1 12 Figure AVII. 6 (This is for conversion of CaO to CaS0 4, comparing that for CaO to CaC0 3 shown in Figure 6.10) Cyclic C 0 2 retention performances in 850°C PTGA tests: effect total pressure; Co-capture with 212-250 pm Arctic dolomite. Sorption: 8 %v C 0 2 , 1125 ppmv S0 2 , 3%> 0 2, and balance N 2 , 4 minute for each cycle: Calcination: 101 kPa, 100% 8%v C0 2 , 850°C. 296 Appendix VII Supplementary results within the scope of Chapter 6 o t o *o s CD 3 cd CM O o ** u o c o CD 1 0.8 0.6 0.4 H 0.2 0 • 1.8 MPa O 2.4 MPa ' 9 9 0 • • • • • • 4 5 10 15 Number of reaction cycles 20 Figure AVE.7 Cyclic performances in no SO2 only tests: effect of total pressure. Atmospheric calcination. 850°C, 8% CO2 balanced by N2, 212-250 pm Straussburg limestone. "3 1 -0 les of' 0.8 -captured/mol Ca 0.6 -captured/mol Ca 0.4 -O O 0 0.2'-<D Mol 0 -• 850°C • 750°C 5 10 15 Number of reaction cycles 20 Figure AVTI.8 Cyclic performances in no S0 2 only tests: effect of temperature. 8%v C 0 2 balanced by N2, 212-250 pm, and Straussburg limestone 297 Appendix VII Supplementary results within the scope of Chapter 6 -*-< o -»-» O "o 1 CD >-< _. t U a o o o o c/1 0.8 0.7 0.6 0.5 -I 0.4 0.3 A 0.2 • 81. • 8%vC0 2 O 14%vC0 2 O 20%vCO 2 5 10 15 Number of reaction cycles 20 Figure AVTI.9 Cyclic performances in no S0 2 only tests: effect of PCo2. Pt=18.2 bars C 0 2 only, Straussburg, 212-250 pm 298 Appendix VIII Supplementary results within the scope of Chapter 7 APPENDIX VIII S U P P L E M E N T A R Y RESULTS WITHIN T H E S C O P E O F C H A P T E R 7 cd -t-» O -*-> C w O c/i CD "o a CD 3 '3 CD i -o o o cn CD *o u 5 10 Number of reaction cycles 15 Figure AVUI.l Performance of c-CaO and h-CaO (no S0 2 present, sorbent derived from 212-250 pm Arctic dolomite). Same test conditions as specified in Figure 7. 2. CD s > o 3 3 CD P 0.16 0.14 -0.12 -0.1 0.08 0.06 0.04 0.02 -| 0 ~EF~ c-CaO, initial calcination -A— h-CaO, initial calcination -A—h-CaO, 15th cycle c-CaO, 15th cycle 10 100 1000 Mean pore diameter (nm) Figure AVTJI.2 Pore size distribution: comparison of c-CaO and h-CaO. (no S0 2 present, sorbent derived from Arctic dolomite) 299 Appendix VIII Supplementary results within the scope of Chapter 7 -»-> o o o a CD C 'B CD UH CN o u o CD rt U 1 0.8 H 0.6 0.4 H 0.2 0 A No steam • With 14%v steam A A A A A A 4 A A 4 5 10 Number of reaction cycles 15 Figure AVTII.3 Effect of steam on cyclic capture (No SO2 present, 212-250 pm Strassburg limestone). Test conditions: 850°C calcination and carbonation; Carbonation in 86% C02.and 14% steam; Calcination in 100%N2. .»-» o <+-! o cn CD CD S rt 'rt u CD O o o cn CD 1 i 0.8 0.6 -( 0.4 0.2 0 U • • • u • n • • • • • • • rj] • No steam • With 14%v steam 5 10 Number of reaction cycles —1 15 Figure AVUI.4 Effect of steam on cyclic capture (No S0 2 present, 212-250 pm Arctic dolomite). Test conditions: 850°C calcination and carbonation; Carbonation in 86% C02.and 14% steam; Calcination in 100% N 2 . 300 Appendix VIII Supplementary results within the scope of Chapter 7 a cj 13 o C w O •a o O 0.4 r 0.2 O A • o * O o 4 0 ° o o o o • h-CaO A Co-capture with 14%v steam O Co-capture with, typical conditions 5 10 Number of reaction cycles (a) Strassburg limestone — i — 15 20 O o 0.9 C3 CJ •*-» o -*-» o X 0.6 1) •E3 0.3 t • h-CaO • Co-capture with, typical conditions O Co-capture with 14%v steam • • 8 • O O o 4 6 8 Number of reaction cycles 10 — i 12 (b) Arctic dolomite Figure AVTJI.5 Calcium utilization for S O 2 capture: effect of varying operating conditions Comparing with Figure 7.6 Calcium utilization for S O 2 capture: effect of varying operating conditions, 212-250 pm particles Test conditions: 850 °C calcination and sorption, Sorption in 80% C0 2 . 3% 0 2, 2900 ppm S0 2 and balance N 2 or with steam. Calcination in 100% N 2 . 8 minutes for each sorption. 301 Appendix VJII Supplementary results within the scope of Chapter 7 Figure AVJJI.6 Close-up view of one cycle of S O 2 / C O 2 sorption followed by a CaC0 3 calcination in N 2 and then a slow reduction of CaS0 4 at cycle 9. test conditions are described in Table 7.3 302 Appendix VIII Supplementary results within the scope of Chapter 7 (a) CaO with AI2O3, light colonAl (b) Calcium acetate with A1 20 3 , light colonAl (c) CaO with S i 0 2 , light color: Si (d)CaO +Zr0 2, light colonZr (e) CaO with MgO, light color: Mg (f) CaO with T i0 2 , light color: Ti Figure AVTJI.7 E D X element mapping for modified sorbents with the light color points marking the element of the dopants Appendix IX A case study based on Option B proposed in Chapter 8 APPENDIX IX A C A S E STUDY BASED O N OPTION B PROPOSED IN C H A P T E R 8 AIX. l Case study A case study was conducted to provide practical information on the application of calcium-based sorbents. The process simulated is shown schematically in Figure ATX. 1, based on Option B in Chapter 8. . AIX.2 Major features of the process simulation: • Steady-state operation. • Pressure drop and power needed for pure oxygen production are not considered. • The FBC is simulated as a Gibbs reactor. The reactor labeled DESUL is simulated as a conversion reactor for S 0 2 removal. CaO utilization for SO2 capture is assumed to be 50%, with the Ca/S molar ratio maintained at 2:1. This determines the split at the CAOSLT block. • The 02/Fuel stoichimetric ratio in the FBC is taken as 1.3, whereas it is assumed to be 1.1 in the Calciner. • The FBC and DESUL reactors operate at 1 atm and 850°C, whereas the CARBNER (carbonator) is at 650°C and 1 atm, and the CALCINER at 950°C and 1 atm. (Note in this Appendix, the unit of pressure takes "atm" as requested by the software.) • All cyclones are assumed to provide perfect gas/solid separation. • Solid drainage are specified as follows: After the first cyclone (CYC1), 5% of solids from the LOOPSPT block are drained, while the other 95% is recycled to the FBC; after CYC3, 10% of the solids from the CAC03SPT block are drained, the other 90% being recycled to the CARBNER. 304 Appendix IXA case study based on Option B proposed in Chapter 8 • The CaCC>3 feed stream is determined as the minimum feed needed to achieve a zero carbon emission in the S T A C K F L G stream. • F U E L 2 for the calciner is adjusted so that the net heat duty of the C A L C I N E R and D E C O M 2 reactors is zero. In other words, the fuel fed to C A L C I N E R is set to provide the heat needed to carry out the calcination reaction at 950°C. • Intended to exploration of issues relating to fuel and sorbent consumption, this simulation ignores detailed calculation of heat losses and heat transfer in all reactors. The heat duty analysis is instead based on empirism, e.g. overall heat transfer efficiency etc. 305 Appendix IXA case study based on Option B proposed in Chapter 8 A1X3 Fuel and sorbent properties • Because of the assumption of perfect separation in all cyclones, the particle size distributions for both the fuel and sorbent did not need to be specified. • The sorbents were assumed to be pure, i.e. 100% CaCC<3, treated as a conventional solid in the Aspen simulation. • In the Aspen simulation, fuel was treated as non-conventional solid. The fuel analysis (except for varying sulfur content in the sensitivity analysis) is shown in Table VUI.1. Note the fuel selected is an idealized fuel. It is only chosen for illustration purpose. • This simulation assumes fuel types. The fuel applied in the CALCINER is assumed to be similar to that in the FBC, except that it has a lower ash content and contains no sulfur. The latter is to prevent the deactivation of sorbents because of the presence of S O 2 in the calciner. In practice, the fuel could be natural gas, biomass or some and other fuels. Table ATX. 1 Proximate and ultimate analysis for the fuels simulated. Fuel in FBC Fuel in calciner low ash, low S Heat value (LHV) (kJ/kg) 36,370 36,370 Proximate Moisture 25 25 analysis FC 45.1 47 (%wt) V M 45.1 47 ASH 9.2 6 Ultimate ASH 9.2 6 analysis C 67.1 68 (%wt) H 4.8 5 N 1.1 1 S 1.3 0 0 16.4 20 AVTX.4 Typical run summary 307 Appendix IXA case study based on Option B proposed in Chapter 8 The results for the mass flow rates of major streams with CaO conversion to CaC03 are shown in Table ATX.2, with 40% the CaO conversion to CaC0 3 assumed in the carbonator. The predicted gas emissions are shown in Table AJX.3. The favourable mass balance implies that good convergence has been achieved. Overall, 1 mole of CaC0 3 can be responsible for cumulatively removing 7.2 moles of C 0 2 from the FBC over successive cycles. This is largely due to the recycle of sorbents in this system. A relatively pure C 0 2 stream can be obtained as long as just enough excess oxygen can be generated to meet the combustion requirement in the calciner. The overall stream result is shown in Table ATX.4 Table ATX.2 Summary of simulation results for overall mass balance Stream ID FUEL1 AIR F U E L 2 02 C A C 0 3 F D R A I N 1 DRATN2 C 0 2 S T A C K F L G Position in Fig. 10.1 Fuel in F B C Air in F B C Fuel in Calciner 0 2 in Calciner C a C 0 3 Feed Solid drain in F B C Solid drain in carbonator Gas product from calciner Gas product from carbonator Flow rate (kg/hr) 40000 286065 29277 46266 44405 5166 31932 160404 248515 Balance SOLIDIN=446,013 kg/hr SOLEDOUT=446,017 kg/hr C 0 2 sorbent consumption Moles of C 0 2 removed/ Moles of C a C 0 3 fed= 7.2 (mole C0 2/mole C a C 0 3 ) Table ATX. 3 Summary of gas emissions Flue gas to stack C 0 2 stream C 0 2 (%v, wet basis) 0 75.5 C 0 2 (%v, dry basis) 0 97.4 0 2 (%v, wet basis) 1.9% 1.5% NOx (ppm, wet basis) ~0 5 SOx (ppm, wet basis) 0 0 AJX.5 Capacity estimation of the simulated unit This sequential-capture unit is different from conventional units in that fuels are fed in both the FBC and the calciner. The exothermic carbonation reaction generates a large amount of heat in the carbonator. Hence heat exchangers need to be installed in all three reactors to generate steam. The capacity of this plant is therefore estimated based on the total fuel consumption with 308 Appendix IXA case study based on Option B proposed in Chapter 8 the aid of the rule of thumb for fuel consumption, e.g., -330 g of standard fuel (-29260 kJ/kg LHV) per kWh for large (e.g. 100-300 MW) pulverized-coal-fired power plants. For a fuel feed of-40 t/h in the FBC, as well as 29.4 t/h (with CaO conversion=0.4) in the calciner, the simulated unit is equivalent to a 262 MW fossil-fuel-fired conventional unit. ATX.6 Sensitivity analysis results: As. sorbent performance is a key factor, sensitivity tests were conducted with variable CaO conversion in the carbonator. The results are shown in Figures ATX.2 and A D O . The results show that: • Increasing CaO conversion leads to a decrease in the CaC03 needed, but very gradually after the CaO conversion reaches a certain level, as shown in Figure ATX.2, • In Figure ATX.3, the total heat available was calculated by adding the heat generated in the FBC and CARBNER and the sensible heat of the C 0 2 stream (e.g. 140670 kg/h at 950°C) and of the C02-free STACKFLG stream (650°C). If it is assumed that 40% of the total heat can be used to generate electricity (a typical overall thermal efficiency for large PC-fired units). At least -250 MW of power can be generated, almost the same capacity as for a unit without C 0 2 capture based on the same fuel consumption or (or heat). In other words, if 40% overall thermal efficiency can be achieved in a C 0 2 capture plant based on this process, the overall power generation per unit mass of fuel consumed is not greatly harmed. • The total CaC0 3 needed (for both S0 2 and C 0 2 removal) in the FBC depends on the sulfur content in the fuel fed to the FBC is shown in Figure ATX.4. Figure ATX.4 shows that the CaC03 requirement changes very little because of the much larger amount of CaC0 3 required for C 0 2 removal in the carbonator relative to the S0 2 removal. Given 309 Appendix IXA case study based on Option B proposed in Chapter 8 that the molar ratio of carbon to sulfur is very large (e.g. >50) and more sorbents have to be added to the system because only a fraction of the sorbents (e.g. <40%) retains its reactivity to react with C O 2 , the sorbent needed for S O 2 removal only accounts for a minor fraction of the total sorbent. 400000 T 350000 i 0 i " . . . . . . , , . . , j 0 0.2 0.4 0.6 0.8 1 CaO conversion Figure ATX.2 Sensitivity of CaC03 feed to average CaO utilization in carbonator for a 262 MW fossil-fuel-fired conventional FBC unit.. 310 Appendix IX A case study based on Option B proposed in Chapter 8 1200 T 200 A 0 -f————p——— — , — — i — — — i i 0 0.2 0.4 0.6 0.8 1 CaO conversion Figure ATX.3 Sensitivity of energy production to average CaO conversion in the carbonator for of a 262 MW fossil-fuel-fired conventional FBC unit. 311 Appendix IXA case study based on Option B proposed in Chapter 8 60000 ^ 55000 A 2 40000 ] U <3 35000 -30000 I , , , , , i 0 1 2 3 4 5 6 Sulfur content in FUEL1 (in FBC) Figure ATX.4 Sensitivity of total CaC0 3 needed in FBC on sulfur content in raw fuel fed to the FBC for a 262 MW fossil-fuel-fired conventional FBC unit AIX.7 Conclusions from case study. The case study show that a power plant using calciun>based sorbent to capture CO2 must deal with the following issues: • The average calcium utilization in the carbonator needs to be maintained at an intermediate level, e.g. 30 to 50%, to appreciably reduce the feed rates of fuel and CaC03. A further increase in sorbent performance would not appreciably decrease the feed requirement. From Chapter 7, intermittent hydration using liquid water or even exposure to ambient humid air might help achieve a high level of calcium utilization. 312 Appendix IXA case study based on Option B proposed in Chapter 8 • A process of this kind would be applicable to high-sulfur fuel in the FBC, as the sorbent needed for desulfurization accounts for only a small fraction of the total sorbent feed. However, for the calciner, the fuel must have a low sulfur content to avoid sorbent deactivation. • For calcination, it is important to select an appropriate fuel with low sulfur content, to maintain a low level of excess oxygen or to choose alternative ways to provide the heat of calcination. • Heat exchangers are needed in the carbonator and at the downstream end of the calciner to utilize the sensible heat, and improve the overall thermal efficiency. • A detailed economic study is needed to provide further information on the viability of alternative CO2 strategies. 313 Table AIX.4 Calculation results for all streams in Figure VIII. 1 (with CaO conversion to CaCOs is assumed as 0.4 in the CARBONER block) Stream ID AIR B A C K C A C 0 3 C A C 0 3 F E E D C A O C A 0 C 0 2 C A O S 0 2 C 0 2 DRAIN1 DRAIN1 FLUE1 F L U E 2 F L U E 3 F L U E 4 FUEL1 F U E L 2 INBURN E R INCALC 0 2 ' O U T C A L R E C A C O 3 R E S O L I D S S T A C H F LG Temperature C 300 849.96 650 15 950 950 950 950 849.9605 650 850 850 850 650 15 15 15 15 300 950 650 850 650 Pressure bar 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 M a s s Vfrac 1 0 0 • 0 0 0 0 1 0 0 0.991529 0.7572588 1 0.437648 0 0 0.173984 0.209815 1 0.393783 0 0 1 M a s s Sfrac 0 1 1 1 1 1 1 0 1 1 0.008471 0.2427412 0 0.562352 1 1 0.57268 0.5475 0 0.606217 1 1 0 A L L P H A S E S * " M a s s F low kg/hr Volume • Flow cum/hr 286065 472508 98161.4 29.9704 319327 108.46 44405 16.3596 246939 74.712 245514 74.2808 1425.34 0.43124 160405 430237 5166.388 1.577388 31932.7 10.8464 326065.4 1026619 425652.09 1024978.9 322329 1024947 567841.7 713809.8 40000 28.8431 29277.93 21.72804 40000 21468.58 29277.93 17012.46 46266.63 58901.67 407344 430311.2 287394.4 97.61718 103327.8 31.54776 248514.6 713701.3 Enthalpy GJ/hr 80.676 -439.6 •3430 -535.2 • -2470 -2456 -14.256 -1303 -23.136 -343 -656.91 -1116.69 -654 -3555.5 -209.3 -163 -157.56 -116.82 12.0969 -3772.9 -3087 -462.71 -125.45 Density kg/cum 0 6 0 5 4 2 3275.28 2944.1 2714.31 3305.2 3305.21 3305.21 0.37283 3275.281 2944.1 0.317611 0.4152789 0.31448 0.795508 1386.81 1347.472 1.863188 1.72097 0.671488 0.946627 2944.097 3275.281 0.348205 M a s s R o w kg/hr N 0 2 0 0 0 0 0 0 0 0.O0378 0 0 0.110253 0.1102527 0.11025 0.110253 0 0 0 0 0 0.003775 0 0 0.110253 N O 0 0 0 0 0 0 0 0.61407 0 0 10.39994 10.399938 10.3999 10.39994 0 0 . 0 0 0 0.614069 0 0 10.39994 S 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 390.2927 0 0 0 0 0 0 S 0 2 0 0 0 ' 0 0 0 0 0 0 0 732.6011 0 0 0 0 0 0 0 0 0 0 0 0 S 0 3 0 0 0 0 0 0 ' 0 0 0 0 58.94565 0 0 0 0 0 0 0 0 0 0 • 0 0 H2 0 0 0 0 0 0 0 0.00031 0 0 4.49E-05 4.49E-05 4.49E-05 4.49E-05 0 0 1441.081 1097.922 0 0.000308 0 0 4.49E-05 C L 2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 H C L 0 0 0 0 0 0 0 0 0 0 0 0 0' 0 0 0 0 0 0 0 0 0 0 C 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 20145.11 14931.74 0 0 0 0 0 -co 0 0 0 0 0 0 0 0.02095 0 0 0.000898 0.O0O8982 0.00O9 0.000898 0 0 0 0 0 ' 0.020946 0 0 0.000898 C 0 2 0 0 0 0 0 0 0 140670 0 0 73814.19 73814.187 73814.2 0 0 0 0 0 0 140670 0 - 0 0 N 2 219436 0 0 0 0 0 0 438.881 0 0 219761.2 219761.24 219761 219761.2 0 0 330.2477 439.1689 0 438.8811 0 0 219761.2 H 2 0 0 0 0 0 0 0 0 17131.3 0 0 22885.99 22885.987 22886 22885.99 0 0 • 10007.51 7319.482 0 17131.26 0 0 22885.99 0 2 66629.5 0 0 0 0 0 0 2164.12 0 0 6039.826 5856.8677 5856.87 5856.868 0 0 4923.693 4391.689 46266.63 2164.124 0 0 5856.868 C A C 0 3 0 0 167868 44405 0 0 0 0 0 16786.8 0 0 0 167868.4 0 0 0 0 0 0 151081.6 0 0 C A O 0 12969.7 141089 0 236508 235143 1365.14 0 682.6172 14108.9 0 13652.316 0 141088.9 0 0 0 0 0 236508.3 126980 13652.34 0 C A S 0 4 0 31484 0 0 0 0 0 0 1657.053 0 0 33141.069 0 0 0 0 0 0 0 0 0 33141.07 0 A S H 0 53707.6 10370 0 10431 10370.5 60.2068 0 2826.718 1036.98 2762.072 56529.914 0 10369.8 0 0 2762.072 1097.922 0 10430.75 9332.824 56534.35 0 Fuel 0 0 0 0 ' 0 0 0 0 0 0 0 0 0 0 40000 29277.93 0 0 00 o Ik a. s o e o a R Appendix X Preliminary test results for Li4Si04 carbonation and sulphation APPENDIX X PRELIMINARY T E S T RESULTS F O R Li4Si0 4 C A R B O N A T I O N AND SULPHATION AX. 1 Preparation method of Li4Si04 The preparation method was the same as described in the literature (Kato, 2001; Kato etai., 2002a; 2002b). High-purity L i 2 C 0 3 (5 grams) and Si0 2 (quartz) (2.03 grams), supplied by supplied by Fisher Scientific were blended in a molar ratio of 2:1. L i 2 C 0 3 particles are less than 30 pm and Si0 2 particles are less than 50 pm. The mixture was ground in a mortar in order to achieve a good mixing. The mixture was then maintained at 1000°C for 8 h in a sealed oven exposed to ambient air. The agglomerated product was then ground mildly before being tested. The following reaction proceeds to produce Li 4Si04, 2Li 2C0 3+Si0 2=Li 4Si0 4+2C0 2 A# 2 9 g j e =225.7 kJ/mol (AX.l) AX.2 Carbonation test The Gibbs free energy of the reaction of Li4Si04 carbonation or reaction (AX.2) is shown in Figure AX. 1 generated by use of the database of HSC 4.0. Li 4Si04+C0 2=Li 2Si0 3+Li 2C0 3 AH29SK =-141.9 kJ/mol (AX.2) It is seen that at atmospheric conditions, temperatures higher than 720°C favour calcination in the presence of 100% C0 2at 1 atm pressure. An experimental test was performed in the atmospheric thermogravimetric reactor. The carbonation temperature was chosen as 650°C with 100% C0 2 . The calcination temperature was 800°C in 100% N 2 . Each sorption stage lasted 6 minutes. 315 Appendix X Preliminary test results for Li4Si04 carbonation and sulphation 20 o Carbonation -60 300 400 500 600 700 800 900 Temperature (°C) Figure AX. 1 Ellingham plot for the reaction (ATX.2) with isobaric lines for partial pressure of Figure AX.2 presents the cyclic C O 2 capture performance over 15 cycles by plotting the mass increase due to reaction (AX.2) divided by the initial total mass of the sorbent. A very stable, ~16 wt% gain during carbonation was observed in this study. Kato et al. (2002a) reported a gain of -25 wt% upon carbonating a sorbent at 500°C in 20% C O 2 and balance N 2 where their sorbent was prepared in the same manner as in this work. If the starting solids were all Li4Si04, i.e. the solid-state reaction (AX.l) went to completion, then the conversion of reaction (AX.2) was calculated to have reached - 45% over each of the 15 cycles on a molar basis. This intermediate conversion and the lower wt% obtained in this test implies that there might be a C 0 2 , 316 Appendix X Preliminary test results for Li4Si04 carbonation and sulphation fraction of Li 2CC>3 and S i 0 2 which remained unconverted because of either non-ideal mixing or solid-state reaction kinetics. 0.4 r £ 0.3 V Sulphation of calcines for 140 mins after 15 cycles | 0.2 o SP ^ 0.1 15 cycles of calcination/carbonation 5 10 Cycle number 15 20 Figure AX.2 Cyclic C 0 2 capture performance and subsequent S 0 2 capture. Test conditions: carbonation-650°C, 100% C 0 2 ; Calcination- 800°C, 100% N 2 ; Sulphation-850°C, 2900 ppm S0 2 , 3% 0 2 , balance N 2 . Our test is compared with the literature test in Figure AX.3 where the data in Figure AX.2 are repotted with the mass increase in the 1st cycle taken as 100. The small fluctuation of the curve in this work is probably due to experimental errors. Note, however, that in the test of Kao 317 Appendix X Preliminary test results for Li4Si04 carbonation and sulphation et al. (2002) the preparation method was somewhat different in that pelletization step was performed in the presence of an inorganic binder. Also the test conditions in that case were: carbonation at 600°C in 20% C02/balance N 2 , 1 h for each sorption; calcination at 800°C in 100% N 2 for 1 hr. Figure AX.3 shows a large difference between the two tests, indicating that more work is needed to understand the effects of preparation method and test conditions. 110 ioo V CD I 90 o JH & C3D C rt JG o H 70 80 L 60 -This work Replot of Kato et al. (2002a) data 10 15 Cycle number 20 25 30 Figure AX.3 Comparison of cyclic C 0 2 capture ability with literature test. As our lithium silicate sorbent showed very good reversibility over 15 cycles and the L i -based sorbents need much lower temperature for regeneration, it appears that this Li-based sorbent is likely to be a better C 0 2 sorbent. As indicated in Figure A X . l , to achieve efficient carbonation at combustion temperatures (e.g. 850°C), pressurized conditions would be required. 318 Appendix X Preliminary test results for Li4Si04 carbonation and sulphation AX .3 Sulphation test It is also important to know whether this sorbent is reactive to SO2 and whether there is a negative effect of S0 2 on C 0 2 capture as found for the calcium-based sorbents tested in this thesis project. Since the sulphation reactivity of Li4Si04 is very scarce in the literature, a sulphation test was carried out with the Li-based sorbent after the 15 cycles of carbonation and calcinations. The sulphation test, performed with 2900 ppm S0 2 , 3% 0 2 and the balance N 2 at 850°C, gave 21.3 wt% increase due to S0 2 capture over 140 minutes test (see Figure AX.2). The overall test history plotted in Figure AX.4, indicates that sulphation was far from being finished after 140 minutes, suggesting that this sorbent may also be a good S0 2 sorbent. This implies that its good reactivity to S0 2 capture may influence its cyclic C 0 2 capture. Further tests are clearly needed to clarify the application of this sorbent to FBC systems. 4.1 4 3.9 3 3.8 Cfl I 3.7 3.6 3.5 3.4 Cyclic C 0 2 capture and calcination 140 mins S 0 2 capture 50 100 150 '200 250 300 Time (mins) 350 400 450 Figure AX.4 Test history for cyclic C 0 2 capture and S0 2 capture. Test conditions: carbonation-650°C, 100% C0 2 ; Calcination- 800°C, 100% N 2 ; Sulphation-850°C, 2900 ppm S0 2 , 3%> 0 2, balance N 2 . 319 

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