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Biomass/fossil fuel co-gasification with and without integrated CO2 capture Masnadi-Shirazi, Mohammad Sadegh 2014

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 BIOMASS/FOSSIL FUEL CO-GASIFICATION WITH AND WITHOUT INTEGRATED CO2 CAPTURE   by Mohammad Sadegh  Masnadi-Shirazi   B.Sc., Iran University of Science & Technology, Tehran, Iran, 2007 M.A.Sc., University of British Columbia, Vancouver, Canada, 2009  A THESIS SUBMITTED IN PARTIAL FULFILLMENT OF THE REQUIREMENTS FOR THE DEGREE OF  DOCTOR OF PHILOSOPHY  in   THE FACULTY OF GRADUATE AND POSTDOCTORAL STUDIES (Chemical & Biological Engineering)  THE UNIVERSITY OF BRITISH COLUMBIA (Vancouver)  May 2014  © Mohammad Sadegh  Masnadi-Shirazi, 2014 ii  ABSTRACT  Biomass/fossil fuel co-gasification could be a bridge between energy production based on fossil fuels and energy production based on renewable fuels. In this work, CO2 co-gasification of switchgrass with coal and fluid coke was conducted in a thermogravimetric analyzer at 800°C. Coal gasification kinetics were inhibited or enhanced, depending on the switchgrass concentration in the mixture: for low K/Al and K/Si molar ratios in the mixture, it seemed that the coal ash sequestrated the biomass potassium needed for Kalsilite (KAlSiO4) formation, and thus, no catalytic effect was observed until the biomass to coal mass ratio reached 3:1, where the switchgrass ash supplied enough potassium to more than satisfy the minerals in the coal ash. For high K/Al and K/Si molar ratios in the mixture, the non-reacted residual potassium acted as catalyst and enhanced the coal gasification. The fluid coke contained much lower Al and Si relative to the coal. Hence, the CO2 gasification kinetics of fluid coke were significantly augmented by blending it with switchgrass, due to the rich presence of potassium in the biomass.  A low-ash coal and switchgrass rich in potassium were selected to steam gasify it as a single-fuel and in 50:50 wt% coal:switchgrass mixtures in a pilot scale fluidized bed with silica sand as the bed material at ~ 800 and 860°C and 1 atm. With biomass added to the coal, the hydrogen and cold gas efficiencies, gas yield and HHV of the product gas were enhanced relative to single-fuel gasification. The product gas tar yield was also reduced considerably due to decomposition of tar catalyzed by switchgrass alkali and alkaline earth metals.  In the quest for a more sustainable process, coal/switchgrass steam co-gasification was integrated with in-site CO2 capture with limestone. Five gasification/carbonation and calcination cycles were performed in a pilot scale fluidized bed. Hydrogen production was enhanced due to partial adsorption of CO2 by the CaO sorbent particles (bed material). The sorbent particles decayed and lost their utilization efficiency in the course of cycling due to sintering.  A simple equilibrium model and an empirically kinetically-modified equilibrium model were also presented to predict syngas composition.            iii  PREFACE  By the time of submitting this thesis, two research papers have been published, both related to  Chapter 2: • M. S. Masnadi, R. Habibi, J. Kopyscinski, J. M. Hill, X. Bi, C. J. Lim, N. Ellis, and J. R. Grace, “Fuel characterization and co-pyrolysis kinetics of biomass and fossil fuels,” Fuel, vol. 117, no. Part B, pp. 1204–1214, 2013. • R. Habibi, J. Kopyscinski, M. S. Masnadi, J. Lam, J. R. Grace, C. A. Mims, and J. M. Hill, “Co-gasification of biomass and non-biomass feedstocks: Synergistic and inhibition effects of switchgrass mixed with sub-bituminous coal and fluid coke during CO2 gasification,” Energy & Fuels, vol. 27, no. 1, pp. 494–500, 2012. For the first paper, the author carried out all stages of the research described in the paper and submitted it after receiving input from the other authors, prepared responses to the reviewers and handled correspondence with the journal editors. For the second paper, Rozita Habibi was the main lead of the paper and performed and analyzed the experiments. My role was to discuss techniques and results with her and to read and comment on the paper at the draft stage. The author has also had several presentations at international, national and local conferences and scientific/technical forums: • M. S. Masnadi, R. Habibi, J. Kopyscinski, J. M. Hill, J. R. Grace, “Co-gasification of biomass and non-biomass feedstocks. Part 1: Fuel characterization, temperature effect and co-pyrolysis kinetics,” in Proceedings of the 3rd international symposium on gasification and its applications, 62nd Canadian Chemical Engineering Conference, October 2012, Vancouver, Canada. • R. Habibi, J. Kopyscinski, J. M. Hill, M. S. Masnadi, J. R. Grace, “Co-gasification of biomass and non-biomass feedstocks. Part 2: Synergistic and inhibition effects of switchgrass mixed with sub-bituminous coal and fluid-coke during CO2 gasification,” in Proceedings of the 3rd international symposium on gasification and its applications, 62nd Canadian Chemical Engineering Conference, October 2012, Vancouver, Canada.  iv  • M. S. Masnadi, J. W. Butler, J. R. Grace, “Sorbent Enhanced Steam Gasification of Biomass: Batch Studies in a Bubbling Fluidized Bed,” in IEAGHG 4th High Temperature Solids Looping Cycles Network Meeting, August 2012, Tsinghua University, Beijing, China. The XRD results presented in  Chapter 2 were performed by Dr. Rozita Habibi (Chemical and Petroleum Engineering Department, University of Calgary). The pilot scale experiments presented in Chapters 3 and 4 could not have been accomplished without technical advice and help from Dr. Yonghua Li (Highbury Biofuel Technology Inc.). The lime-enhanced gasification (LEG) research presented in  Chapter 5 was performed with the assistance of Dr. James Butler (Chemical and Biological Engineering Department, University of British Columbia).                  v  TABLE OF CONTENTS  Abstract……. .................................................................................................................................. ii Preface……… ................................................................................................................................ iii Table of Contents ............................................................................................................................ v List of Tables. .............................................................................................................................. viii List of Figures ................................................................................................................................. x Abbreviations and Nomenclature ................................................................................................ xvi Acknowledgments......................................................................................................................... xx Dedication…. ............................................................................................................................... xxi Chapter 1. Introduction .............................................................................................................. 1 1.1 Concept and Significance of Gasification ......................................................................................... 2 1.2 Concept of CO2 Capture ................................................................................................................... 5 1.3 Catalytic Gasification ........................................................................................................................ 6 1.3.1 Catalytic effects in coal gasification ..................................................................................... 6 1.3.2 Catalytic gasification mechanisms ....................................................................................... 9 1.3.3 Catalytic activity ................................................................................................................. 11 1.3.4 Black liquor gasification ..................................................................................................... 12 1.4 Co-Gasification of Biomass and Fossil Fuels ................................................................................. 13 1.5 In-Situ CO2 Capture ........................................................................................................................ 18 1.6 Scope of This Thesis ....................................................................................................................... 23 Chapter 2. Fuel Characterization and Thermogravimetric Analysis of Biomass and Fossil Fuels ....................................................................................................................... 25 2.1 Experimental ................................................................................................................................... 25 2.1.1 Feedstocks .......................................................................................................................... 25 2.1.2 Sample characterization ...................................................................................................... 26 2.1.3 Experimental apparatus ...................................................................................................... 28 2.2 Results and Discussion ................................................................................................................... 31 2.2.1 Characterization of fresh samples ....................................................................................... 31 2.2.2 Characterization of chars .................................................................................................... 34 2.2.3 Co-pyrolysis thermogravimetric analysis: Experimental ................................................... 45 2.2.4 Co-pyrolysis thermogravimetric analysis: Modeling ......................................................... 52  vi  2.2.5 Co-gasification thermogravimetric analysis ....................................................................... 58 2.2.6 Effect of operating conditions ............................................................................................ 77 2.3 Summary ......................................................................................................................................... 81 Chapter 3. Steam Gasification of Coal and Switchgrass in a Bubbling Fluidized Bed Reactor ................................................................................................................................ 82 3.1 Fuel Characterization ...................................................................................................................... 83 3.2 Experimental ................................................................................................................................... 85 3.2.1 Experimental apparatus and operation ............................................................................... 85 3.2.2 Product gas analysis ........................................................................................................... 90 3.2.3 Tar sampling ....................................................................................................................... 90 3.2.4 Condensed water analysis ................................................................................................... 91 3.2.5 Operating procedure ........................................................................................................... 92 3.2.6 Gasification indices ............................................................................................................ 93 3.3 Results and Discussion ................................................................................................................... 94 3.3.1 100% Quinsam mine coal gasification ............................................................................... 94 3.3.2 100% spring switchgrass (SP-SG) gasification .................................................................. 99 3.3.3 100% fall switchgrass (F-SG) gasification ....................................................................... 103 3.4 Summary ....................................................................................................................................... 108 Chapter 4. Steam Co-Gasification of Coal and Switchgrass Mixtures in a Bubbling Fluidized Bed Reactor .......................................................................................................... 110 4.1 Results and Discussion ................................................................................................................. 110 4.1.1 50:50 coal:spring switchgrass (SP-SG) co-gasification ................................................... 110 4.1.2 50:50 coal:fall switchgrass (F-SG) co-gasification .......................................................... 116 4.1.3 Comparison of single-fuel and co-gasification ................................................................. 120 4.2 Summary ....................................................................................................................................... 126 Chapter 5. Lime Enhanced Steam Co-Gasification of Coal and Switchgrass in a Bubbling Fluidized Bed Reactor .......................................................................................... 128 5.1 Experimental ................................................................................................................................. 129 5.1.1 Experimental apparatus and operation ............................................................................. 129 5.1.2 LEG indices ...................................................................................................................... 134 5.2 Results and Discussion ................................................................................................................. 134 5.3 Summary ....................................................................................................................................... 148 Chapter 6. Equilibrium Modeling .......................................................................................... 150  vii  6.1 Single-Fuel and Co-Gasification Equilibrium Modeling .............................................................. 150 6.1.1 Modeling results ............................................................................................................... 151 6.1.2 Comparison of model predictions with experimental results ........................................... 156 6.1.3 Modified equilibrium model ............................................................................................. 160 6.2 LEG Equilibrium Modeling .......................................................................................................... 164 6.2.1 Modeling results ............................................................................................................... 164 6.2.2 Comparison of model predictions with experimental results ........................................... 166 6.3 Summary ....................................................................................................................................... 167 Chapter 7. Conclusions and Recommendations ..................................................................... 169 7.1 Conclusions ................................................................................................................................... 169 7.2 Recommendations for Future Work .............................................................................................. 173 References…. .............................................................................................................................. 174 Appendices… .............................................................................................................................. 198 Appendix A. TGA Heat and Mass Transfer Calculations ..................................................................... 198 Appendix B. HBTI Bubbling Fluidized Bed Calibration Curves ......................................................... 206 Appendix C. TSS Standard Method [309] ............................................................................................ 212 Appendix D. Temperature, Pressure Drop, and Steam Flow Rate Profiles of Experiments in HBTI Fluidized Bed Gasifier .................................................................................................. 213 Appendix E. Water Treatment on the 50:50 Coal:SP-SG  Gasification Condensed Water .................. 217 Appendix F. Lime Enhanced Gasification (LEG) Gasification/Carbonation Results ........................... 219 Appendix G. Lime Enhanced Gasification (LEG) Calcination Results ................................................ 223 Appendix H. Lime Enhanced Gasification (LEG) Nitrogen Isotherms of Different Samples .............. 226 Appendix I. Lime Enhanced Gasification (LEG) SEM Images of Calcined Samples .......................... 227 Appendix J. Lime Enhanced Gasification (LEG) Equilibrium Model Results ..................................... 229 Appendix K. UBC Dual Fluidized Bed Reactor Flow Diagram ........................................................... 232        viii  LIST OF TABLES  Table  2-1. Ultimate, proximate and ash analysis of fresh feedstocks. .......................................... 32 Table  2-2. BET and DRK surface areas of fuel fresh and char samples produced at different temperatures. The char samples were prepared using 500 mlN min-1 nitrogen flow with a 25°C/min heating rate at atmospheric pressure. The target temperature was held for 2 h. All values in m2 g-1. ............................................................................... 33 Table  2-3. Best fitting mechanism functions with TGA data of this work for modeling three stages of pyrolysis process. ....................................................................................... 56 Table  2-4. Activation energies and reaction mechanisms for different stages of pyrolysis for fresh and mixture samples of coal, fluid coke, switchgrass and sawdust. For mechanisms A to E, see Table  2-3. ..................................................................................................... 57 Table  2-5. Elemental amounts of C, K, Si, and Al in SG and coal chars produced at 750, 800 and 900°C (in mmol g-1). ................................................................................................... 67 Table  2-6. Potassium/carbon, potassium/aluminum, and potassium/silicon molar ratios of SG and coal samples at 750, 800 and 900°C (calculated from Table  2-5). ............................. 67 Table  2-7. Coal ash leaching conditions. ...................................................................................... 69 Table  2-8. Ash elemental analysis after ash removal treatment (all in wt%). .............................. 70 Table  2-9. Elemental amounts of C, K, Si, and Al produced by pyrolysis at 750, 800 and 900°C (in mmol.g-1) for SG char and FC char. ...................................................................... 75 Table  2-10. Potassium/carbon, potassium/aluminum, and potassium/silicon molar ratios for different SG and FC samples at 750, 800 and 900°C (calculated from Table  2-9).75 Table  3-1. Ultimate, proximate, and ash analysis of fresh feedstocks for pilot-scale experiments...................................................................................................................................... 84 Table  3-2. Fluidization bed material analysis. .............................................................................. 87 Table  3-3. Material and hydrodynamic properties (based on steam at 525°C and 1 atm). ........... 90 Table  3-4. 100% coal proximate analysis of entrained particles captured by external cyclones. . 96 Table  3-5. Summary of 100% coal gasification parameters at different steam-to-fuel ratios (SF) and at atmospheric pressure. ....................................................................................... 98 Table  3-6. Summary of 100% SP-SG gasification parameters at different steam-to-biomass ratios at atmospheric pressure. ............................................................................................ 102 Table  3-7. Proximate analysis of particles captured by external cyclones 1 and 2 for 100% SP-SG. ............................................................................................................................ 103 Table  3-8. Summary of 100% F-SG gasification parameters at different bed temperatures at atmospheric pressure. .............................................................................................. 104 Table  4-1. Coal:SP-SG 50:50 wt% run proximate analyses of particles captured by external cyclones. .................................................................................................................. 112 Table  4-2. Summary of coal:SP-SG 50:50 wt% co-gasification parameters for different steam-to-coal ratios and atmospheric pressure. ....................................................................... 115  ix  Table  4-3. Summary of coal:F-SG 50:50 wt% co-gasification parameters at different bed temperatures at atmospheric pressure. .................................................................... 117 Table  4-4. Summary of steam gasification data for 100% coal, 100% SP-SG, 100% F-SG, 50:50 coal:SP-SG, and 50:50 coal:F-SG mixtures in the atmospheric pressure bubbling fluidized bed for SF≈2.4. .......................................................................................... 121 Table  5-1. Operating conditions in CERC gasifier. The hydrodynamic parameters based on steam at 525°C and 1 atm. ................................................................................................... 132 Table  5-2. Strasburg limestone material properties (wt%). ........................................................ 133                     x  LIST OF FIGURES  Figure  1-1. Diversified gasification products. ................................................................................ 3 Figure  1-2. Schematic of reaction pathways at the carbon/catalyst junction (adapted from Rao et al. [57]) ...................................................................................................................... 10 Figure  1-3. Equilibrium partial pressure of CO2 for calcium carbonate. ...................................... 19 Figure  1-4. Schematic of lime-enhanced gasification process. ..................................................... 20 Figure  2-1. Schematic of pressurized TGA set-up (modified from Butler [174]). ....................... 30 Figure  2-2. Dimensions of quartz TGA basket. ............................................................................ 30 Figure  2-3. Typical TGA result of two-stage gasification. ........................................................... 31 Figure  2-4. Char yield of fuel samples as a function of pyrolysis temperature. The char samples were prepared using 500 mlN min-1 nitrogen flow with a 25°C/min heating rate at atmospheric pressure. The target temperature was held for 2 h. ............................... 34 Figure  2-5. Ultimate analyses of carbonaceous chars provided at 750, 800 and 900°C: FC=fluid coke; SG=switchgrass; SD=sawdust. (a) carbon (b) hydrogen (c) nitrogen and (d) sulfur content. ............................................................................................................ 35 Figure  2-6. Mineral matter content in ash of: (a) potassium; and (b) calcium. ............................ 37 Figure  2-7. SEM images of (a) fresh coal (b) coal char prepared at 800°C (c) char of 50:50 wt% coal:switchgrass mixture prepared at 800°C after heating at 25°C/min heating rate and atmospheric pressure in 500 mlN min-1 nitrogen flow. ....................................... 39 Figure  2-8. SEM images of (a) fresh swithgrass, (b) and (c) swithgrass char prepared at 800°C (d) char of 50:50 wt% coal:switchgrass mixture prepared at 800°C after heating at 25°C/min heating rate and atmospheric pressure in 500 mlN min-1 nitrogen flow. ... 40 Figure  2-9. SEM images of (a) fresh fluid coke (b) fluid coke char prepared at 800°C after heating at 25°C/min heating rate and atmospheric pressure in 500 mlN min-1 nitrogen flow. .......................................................................................................................... 41 Figure  2-10. SEM images of (a) fresh sawdust (b) sawdust char prepared at 800°C after heating at 25°C/min heating rate and atmospheric pressure in 500 mlN min-1 nitrogen flow................................................................................................................................... 42 Figure  2-11. SEM image of 75:25 wt% coal:switchgrass mixture prepared at 800°C after heating at 25°C/min heating rate and atmospheric pressure in 500 mlN min-1 nitrogen flow................................................................................................................................... 43 Figure  2-12. X-ray diffraction patterns of parent (fresh) samples and corresponding chars and ash produced at 750°C for (a) SG, (b) Coal, and (c) FC. Powder diffraction files (PDF) for SiO2 and Kaolinite are no. 85-0695 and no. 72-2300, respectively. Performed by Rozita Habibi [150].................................................................................................. 44 Figure  2-13. Effect of relative proportions of fossil fuel/biomass binary mixture on char yield of different fuels prepared at 800°C after heating at 25°C min-1 in N2. Symbols are  xi  experimental data, and the solid lines are predicted if separate fuels acted independently. .......................................................................................................... 46 Figure  2-14. (a) Weight and temperature versus time during pyrolysis of coal, switchgrass, coal/SG mixtures and coal/SG ash mixture in nitrogen (b) Isolated plot of Coal:SG=25:75 wt.% showing three pyrolysis stages. .......................................... 48 Figure  2-15. Rate of devolatilization of coal, switchgrass, coal/SG mixtures of different proportions and coal/SG ash mixture in nitrogen. Heating rate at 25°C min-1 in N2................................................................................................................................ 49 Figure  2-16. (a) Weight and temperature versus time during pyrolysis of coal, sawdust and coal/SD mixtures in nitrogen (b) Rate of devolatilization versus temperature of fresh coal, sawdust and coal/SD mixtures in nitrogen at 25°C min-1 heating rate. 50 Figure  2-17. (a) Weight and temperature versus time during pyrolysis of fluid coke, switchgrass and FC/SG mixtures in nitrogen (b) Rate of devolatilization versus temperature of fluid coke, switchgrass and FC/SG mixtures in nitrogen at 25°C min-1 heating rate................................................................................................................................... 51 Figure  2-18. Experimental arithmetic mean of mixture devolatilization rate data (over 200°C) presented in Figures 2-15, 2-16(b) and 2-17(b) versus corresponding linear combination calculated ones based on equation ( 2-4). ........................................... 52 Figure  2-19. Comparison of experimental coal and SG pyrolysis data (symbols) and model (solid lines) from room temperature to 800°C at a heating rate of 25°C min-1 and atmospheric pressure in nitrogen. ............................................................................ 55 Figure  2-20. (a) CO2 gasification conversion vs. time, and (b) rate vs. conversion for coal, switchgrass and coal/SG mixtures at 800°C and 1 atm. All proportions are in wt%. ................................................................................................................................ 60 Figure  2-21. (a) CO2 gasification conversion vs. time, and (b) rate vs. conversion for coal, switchgrass and coal/SG mixture at 750°C and 1 atm. All proportions are in wt%. ................................................................................................................................ 61 Figure  2-22. (a) CO2 gasification conversion vs. time and (b) rate vs. conversion for coal, switchgrass and coal/SG mixtures at 900°C and 1 atm. All proportions are in wt%. ................................................................................................................................ 63 Figure  2-23. (a) CO2 gasification conversion vs. time and (b) rate vs. conversion for coal, switchgrass and coal/SG mixtures at 800°C and 1 atm. All proportions are in wt%. ................................................................................................................................ 64 Figure  2-24. X-ray diffraction patterns of partially gasified (with CO2 at 750°C to 50% conversion) (a) Coal, (b) SG/Coal 50:50 mixture, (c) SG ash/Coal 50:50 mixture, and (d) SG ash. Powder diffraction files (PDF) for SiO2 and KAlSi3O8 are nos. 85-0695 and 72-0077, respectively. Performed by Habibi [150]. ............................... 69 Figure  2-25. SEM images of fresh coal after leaching. ................................................................ 71  xii  Figure  2-26. (a) CO2 Gasification conversion vs. time and (b) rate vs. conversion for treated and untreated coal, and coal/SG mixtures at 800°C and 1 atm. All proportions are in wt%. ......................................................................................................................... 72 Figure  2-27. (a) CO2 gasification conversion vs. time, and (b) rate vs. conversion for FC, SG and FC/SG mixtures at 800°C and 1 atm. All proportions are in wt%. .......................... 74 Figure  2-28. X-ray diffraction patterns of FC parent, 20% gasified FC, and 50% gasified SC/FC at 800°C. Powder diffraction files (PDF) for SiO2 and K2CO3 are nos. 27-0605 and 01-1001, respectively. Performed by Habibi [150]. ................................................ 76 Figure  2-29. (a) CO2 gasification conversion vs. time and (b) rate vs. conversion for coal, SD and coal/SD mixtures at 800°C and 1 atm. All proportions are in wt%. ........................ 78 Figure  2-30. Effect on conversion vs. time profiles of (a) CO2 flow rate (b) initial sample load (c) basket size (d) CO2 concentration, for CO2 gasification of coal, SG, and coal/SG mixtures at 800°C and 1 atm. All proportions are in wt%. ...................................... 80 Figure  3-1. Particle size distributions of (a) coal (b) spring and fall switchgrasses. .................... 85 Figure  3-2. Schematic of Highbury Biofuel Technologies Inc. (HBTI) pilot-scale bubbling fluidized bed reactor (adapted from HBTI). ............................................................ 87 Figure  3-3. Highbury Biofuel Technologies Inc. unit flow diagram (adapted from HBTI). ........ 88 Figure  3-4. Detailed dimensions of bubbling fluidized bed gasifier (not to scale) (adapted from HBTI). ....................................................................................................................... 89 Figure  3-5. Atmospheric and iso-kinetic sampling train for tar [230]. ......................................... 91 Figure  3-6. Gasification results for dry nitrogen-free product gas composition and local bed temperature for coal run with different coal dry basis (db) feed rates at atmospheric pressure. ................................................................................................................... 95 Figure  3-7. Effect of steam-to-fuel (coal) ratio on gas composition at different temperatures at atmospheric pressure. ................................................................................................ 95 Figure  3-8. Coal gasification carbon efficiency, hydrogen efficiency, cold gas efficiency, and product gas higher heating value vs. time for different SF (see Figure  3-6) in atmospheric pressure gasification. ........................................................................... 97 Figure  3-9. Coal gasification reactor lower and upper pressure vs. time. .................................... 97 Figure  3-10. Gasification results for a dry nitrogen-free product gas composition and local bed temperature of 100% SP-SG run with different biomass feed rates (db) at atmospheric pressure. ............................................................................................ 100 Figure  3-11. Effect of steam-to-fuel (biomass) ratio on steady-state dry nitrogen-free gas composition at different temperatures and atmospheric pressure. ...................... 100 Figure  3-12. 100% SP-SG carbon efficiency, hydrogen efficiency, cold gas efficiency and product gas higher heating value vs. time for different SF (see Figure  3-10) at atmospheric pressure. .......................................................................................... 101 Figure  3-13. SEM image of biomass char particles caught by first external cyclone. ............... 103 Figure  3-14. Gasification results for a dry nitrogen-free product gas composition and bed temperature of 100% F-SG run with SF=2.4 at atmospheric pressure. .............. 105  xiii  Figure  3-15. 100% F-SG gasification carbon efficiency, hydrogen efficiency, cold gas efficiency, and product gas higher heating value vs. time for SF=2.4 at atmospheric pressure................................................................................................................................. 105 Figure  3-16. Effect of temperature on a dry nitrogen-free gas composition of 100% F-SG gasification for SF≈2.4 at atmospheric pressure. ................................................ 107 Figure  3-17. Effect of temperature on hydrogen efficiency, carbon efficiency, cold gas efficiency, and gas yield for 100% F-SG gasification with SF≈2.4 at atmospheric pressure. .............................................................................................................. 108 Figure  4-1. Co-gasification results for dry nitrogen-free product gas composition and bed temperature of coal:SP-SG 50:50 wt% run with different fuel feed rates (db) at atmospheric pressure. ............................................................................................ 112 Figure  4-2. Effect of steam-to-fuel (coal:SP-SG 50:50 wt%) ratio on gas composition at different temperatures and atmospheric pressure. .................................................................. 113 Figure  4-3. Coal:SP-SG 50:50 wt% co-gasification carbon efficiency, hydrogen efficiency, cold gas efficiency, and product gas higher heating value vs. time for different feed rates (see Figure  4-1) at atmospheric pressure. ................................................................ 113 Figure  4-4. SEM images of particles captured by first external cyclone for 50:50 coal:SP-SG (a) captured solids (b) coal char (c) SP-SG char. ......................................................... 114 Figure  4-5. SEM images of particles captured by second external cyclone for 50:50 coal:SP-SG (a) captured solids (b) coal char (c) SP-SG char. .................................................... 114 Figure  4-6. Large rounded agglomerates found at bottom of first condenser near TDV1 valve (see Figure  3-3). ...................................................................................................... 116 Figure  4-7. Co-gasification results for dry nitrogen-free product gas composition and bed temperature of coal:F-SG 50:50 wt% with SF=2.41 at atmospheric pressure. ...... 118 Figure  4-8. Coal:F-SG 50:50 wt% co-gasification carbon efficiency, hydrogen efficiency, cold gas efficiency, and product gas higher heating value vs. time for SF=2.41 at atmospheric pressure. Bed temperature is presented in Figure  4-7. ....................... 118 Figure  4-9. Effect of temperature on dry nitrogen-free gas composition of coal:F-SG 50:50 wt% co-gasification for SF≈2.4 at atmospheric pressure. ............................................... 119 Figure  4-10. Effect of temperature on hydrogen efficiency, carbon efficiency, cold gas efficiency and gas yield of coal:F-SG 50:50 wt% co-gasification for SF≈2.4 at atmospheric pressure. ................................................................................................................. 120 Figure  4-11. Comparison of steam gasification results for 100% coal, 100% SP-SG, 100% F-SG, 50:50 coal:SP-SG and 50:50 coal:F-SG mixtures in the atmospheric bubbling fluidized bed for SF≈2.4. ....................................................................................... 122 Figure  4-12. Product gas tar yield (N2 included and N2-free) and average bed temperature for steam gasification of 100% coal, 100% SP-SG, 100% F-SG, 50:50 coal:SP-SG, and 50:50 coal:F-SG mixtures in a bubbling fluidized bed (atmospheric pressure). ... 124  xiv  Figure  4-13. Condensed water TOC and TSS for steam gasification of 100% coal, 100% SP-SG, 100% F-SG, 50:50 coal:SP-SG, and 50:50 coal:F-SG mixtures in atmospheric bubbling fluidized bed. Temperatures are given in Figure  4-12. ........................... 126 Figure  5-1. Schematic of CERC fluidized bed experimental apparatus (modified from Sakaguchi [275]). ...................................................................................................................... 131 Figure  5-2. Limestone particle size distribution, determined by Mastersizer 2000. ................... 133 Figure  5-3. CO2 concentration, difference between equilibrium and actual CO2 partial pressure and temperatures during calcination of sorbent material in the CERC BFB. ......... 135 Figure  5-4. Dry N2-free product gas compositions and temperatures during gasification/carbonation in the CERC BFB at 1 atm: (a) Cycle 1 (b) Cycle 2. For cycles 3 to 5, see Appendix F. ......................................................................... 136 Figure  5-5. Initial and final reactor pressure drops during gasification/carbonation periods for five LEG cycles. ..................................................................................................... 138 Figure  5-6. Char and tar accumulation at entrance of downstream filter causing blockage. ...... 138 Figure  5-7. Cycle 1 co-gasification/carbonation hydrogen efficiency, carbon efficiency, cold gas efficiency, and HHV vs. time. Temperatures are shown in Figure  5-4, other operating conditions in Table  5-1. ........................................................................... 140 Figure  5-8. Agglomeration of char, lime and ash after 5 cycles of gasification/carbonation and calcination. .............................................................................................................. 140 Figure  5-9. Dry product gas compositions, difference between equilibrium and actual CO2 partial pressure and temperatures during calcination in CERC BFB: (a) Cycle 1 (b) Cycle 2. For cycles 3 to 5, see Appendix G. .......................................................................... 141 Figure  5-10. Comparison of experimental concentrations and performance indices for LEG cycles with 50:50 coal:SP-SG mixture in the CERC atmospheric fluidized bed for SF≈2.6. Averages were calculated over entire period of co-gasificaiton. ........... 143 Figure  5-11. Effect of sorbent cycling on peak product gas composition (dry, N2-free) and HHV of 50:50 coal:SP-SG steam co-gasification at atmospheric pressure. .................... 144 Figure  5-12. BET surface areas and BJH cumulative pore volumes of sorbent samples. .......... 145 Figure  5-13. BJH N2 adsorption pore volume distribution of sorbent samples. ......................... 145 Figure  5-14. SEM images of calcined bed sorbents after first and fifth cycles. ......................... 146 Figure  5-15. Product gas tar yield and average bed temperature for 1st, 2nd and 5th cycles of lime enhanced steam co-gasification of 50:50 coal:SP-SG mixture in atmospheric CERC BFB. .......................................................................................................... 147 Figure  5-16. Particle size distribution of original limestone and calcined bed material after 1st, 3rd and 5th cycles. ................................................................................................ 148 Figure  6-1. Aspen Plus equilibrium simulation diagram for steam gasification process. .......... 151 Figure  6-2. Effect of temperature and pressure on H2, CO, CH4 and CO2 equilibrium compositions for 50:50 coal:F-SG steam gasification with SF=2.41. Colour bars are related to Y-axis. ............................................................................................ 152  xv  Figure  6-3. Effect of temperature and pressure on HHV of equilibrium product gas for 50:50 coal:F-SG steam gasification with SF=2.41. Colour bars are related to Y-axis. ... 153 Figure  6-4. Effect of temperature and pressure on SO2, H2S, NH3, CS2, NO, HCN, and COS equilibrium compositions for 50:50 coal:F-SG steam gasification with SF=2.41. Colour bars are related to Y-axis. ........................................................................... 154 Figure  6-5. Equilibrium product dry gas composition and HHV of 100% SP-SG (solid lines), 100% coal (long-dashed lines), and 50:50 coal:SP-SG mixture (short-dashed lines) steam gasification vs. temperature with SF≈2.4 and 1 atm. ................................... 155 Figure  6-6. Comparison of experimental data and equilibrium simulation predictions for SF≈2.4 and 1 atm (experimental data same as in Table  4-4 of  Chapter 4). ........................ 157 Figure  6-7. Comparison of experimental data and equilibrium modeling predictions: effect of steam-to-fuel mass ratio on dry gas composition and HHV for 100% coal run (same experimental data as in Figure  3-7 of  Chapter 3). .................................................. 158 Figure  6-8. Comparison of experimental data and equilibrium modeling predictions: effect of temperature on dry gas composition and HHV for (a) 100% coal (experimental data same as in Figure  3-6 of  Chapter 3) (b) 50:50 SP-SG (experimental data same as in Figure  4-1 of  Chapter 4). ....................................................................................... 159 Figure  6-9. Modified Aspen Plus equilibrium simulation diagram for steam gasification process. ................................................................................................................................. 161 Figure  6-10. Parity plot comparing experimental and modified equilibrium model product gas compositions for 100% coal, SP-SG, F-SG, 50:50 coal:SP-SG and 50:50 coal:F-SG steam gasification experiments (experimental data same as in Figure  6-6). .. 162 Figure  6-11. Comparison of experimental data and modified equilibrium modeling predictions: effect of temperature on dry gas composition and HHV for 100% coal (experimental data same as in Figure  3-6 of  Chapter 3). ..................................... 163 Figure  6-12. Comparison of experimental data and modified equilibrium modeling predictions: effect of temperature on dry gas composition and HHV for 50:50 coal:F-SG mixtures. ............................................................................................................... 163 Figure  6-13. Predicted equilibrium product dry gas composition, HHV and calcium utilization efficiency for 50:50 coal:SP-SG steam co-gasification vs. temperature with SF≈2.6 at 1 atm. HHV calculated based on equation ( 5-1). ............................................. 165 Figure  6-14. Calcium utilization efficiency vs. temperature for different pressures predicted by the equilibrium model for LEG of 50:50 coal:SP-SG mixture. See Table  5-1 for the operating conditions. ............................................................................................ 166 Figure  6-15. Comparison of experimental data for five cycles and equilibrium modeling predictions: effect of temperature on dry gas composition of 50:50 coal:SP-SG LEG with SF≈2.6 and 1 atm. .............................................................................. 167    xvi  ABBREVIATIONS AND NOMENCLATURE  Abbreviations AAEM Alkali and alkaline earth metals ASTM  American Society for Testing and Materials BET  Brunauer-Emmett-Teller BFB  Bubbling fluidized bed BJH  Barrett-Joyner-Halenda model BL  Black liquor BLC  Black liquor char CCS  Carbon capture and storage CERC  Clean Energy Research Centre daf  Dry ash free  db  Dry basis DME  Dimethyl ether DRK  Dubinin-Radushkevich-Kaganer F  Flotsam FC  Fluid coke F-SG  Fall switchgrass g  Gas GC  Gas chromatograph GHG  Greenhouse gas GTI  Gas Technology Institute  HBTI  Highbury Biofuel Technologies Inc. HHV  Higher heating value IC  Internal combustion IGCC  Integrated gasification combined cycle  xvii  J  Jetsam l  Liquid LEG  Lime-enhanced gasification MEA  Monoethanolamine Mt  Million tons of oil equivalent NPOC  Non-purgeable organic carbon PAH  Polycyclic aromatic hydrocarbon PDU  Pilot display unit PLOT  Porous layer open tubular  s  Solid s,l  Solid-liquid molten phase SD  Sawdust SEM  Scanning electron microscope SF  Steam-to-fuel mass ratio SG  Switchgrass SP-SG  Spring switchgrass  TGA  Thermogravimetric analyzer  TG-MS Thermogravimetric-mass spectrometer TOC  Total organic carbon TSS  Total suspended solids UBC  University of British Columbia wt  Weight basis XRD  X-ray diffraction  Nomenclatures A  pre-exponential factor (s-1) C   BET adsorption constant  xviii  ds  solid particle size (m) E  activation energy (J mol-1) f(α)  reaction mechanism function (-) fbiomass  mass fraction of biomass (-) fCaO   mass fraction of CaO in sorbent (-)   g(α)  integral form function defined by equation ( 2-8),  Chapter 2 (-) k  reaction rate constant (s-1) k1   constant characterizing the Gaussian distribution (-) mbiomass  mass of biomass (g) mcalc   expected mass of fuel based on calculation (g) mf  final mass of fuel (g)  mfuel,dry  dry fuel feeding rate (g h-1)   mfuel,wet  wet fuel feeding rate (g h-1) mgas  dry gas mass flow rate (g h-1) mi  mass of fuel at time t (g) mnon-biomass mass of non-biomass (fossil fuel) (g) msteam    steam mass flow rate (g h-1) m0  mass of fuel at t=0 (g) m0, dry  dry mass of fuel at T=200°C (g) m'i  ash free mass of fuel at time t (g) M  molecular weight (g mol-1), mixing index (-) nCaO   moles of CaO (mol)  nCaCO3  moles of CaCO3 (mol)   P  pressure of gas adsorption (atm)  P0   saturation pressure of analysis gas at analysis temperature (atm)  xix   Peq,CO2  CO2 equilibrium partial pressure (atm)  PCO2  CO2 partial pressure (atm)  q  heating rate constant (K s-1)  Qgas   product gas heating values (MJ g-1)  Qfuel   fuel gas heating values (MJ g-1)  R  universal gas constant (J mol-1K-1)   t  time (s or min)  T  temperature (K)  Tpeak  temperature at the highest rate of devolatilization of fuel (K)  Umf  minimum fluidization velocity (m s-1)  Ut  terminal velocity (m s-1)  vgas   dry gas flow rate (Nm3 h-1)   Va   volume of gas adsorbed (cm3 g-1)  Vm   adsorbed monolayer volume (cm3 g-1)  xash   ash content in feed (-)  X  calcium utilization efficiency (-)   α  chemical conversion (-)  ηC  carbon conversion efficiency (-)  ηH  hydrogen conversion efficiency (-)  ηcg  cold gas efficiency (-)  λ  X-ray wavelength (Å)  ρb  solid bulk density (kg m-3)    xx  ACKNOWLEDGMENTS  I would like to acknowledge everyone who had a share in helping me in the course of my studies at UBC. First and foremost, I would like to express my genuine gratitude to my outstanding, supportive and kind supervisor, Dr. John R. Grace, and my wonderful co-supervisors, Dr. Xiaotao Bi, Dr. Jim Lim and Dr. Naoko Ellis for their continuous support and patience during my stay at UBC. I consider myself very privileged for having the opportunity to work with them. I gratefully acknowledge Dr. Paul Watkinson and Dr. Yong Hua Li from Highbury Biofuel Technologies Inc. for the use of their fluidized bed reactor and for their valuable technical advice during the course of the project. My regards also to my friend and colleague, Dr. James Butler. He was always generous with his time for discussion, no matter how busy he was; without him, I could not have accomplished part of this work. I thank Dr. Josephine Hill and her catalysis group from the University of Calgary, Dr. Charles Mims from the University of Toronto, and Cliff Mui from Nexterra Systems Corp. for productive discussions and collaboration during the project. I am grateful to the members of the UBC Fluidization Research Centre, with whom I was able to exchange ideas and request assistance. I would also like to thank the faculty and staff of UBC Chemical & Biological Engineering Department for providing such a great environment for me to learn and grow, and for being my mentors.   These years in Vancouver would have never been this precious without great friends and a graceful community. Best regards and prayers to Dr. Sayyed Mohammad Solieman-Panah and his family.   And last but not least… Dear father, mother and brothers (Mostafa and Amir Hossein),  I love you…       xxi  DEDICATION            1  Chapter 1. INTRODUCTION  Every day we use energy, and every day we ask for more. With global population set to rise from ~ 7 to ~ 9 billion by 2050 [1], world energy consumption is expected to increase by ~ 56% over the next 30 years alone [2]. The global surface temperature increase by the end of the 21st century is likely to exceed 1.5°C relative to the 1850 to 1900 period for most scenarios, and is likely to exceed 2.0°C according to many scenarios [3].  The world is far from achieving a sustainable energy future. There are persistent concerns about the long-term balance of energy supply and demand, while CO2 emissions have increased dramatically over the last decade. If the future aligns with present trends, GHG emissions of CO2, CH4 and N2O, and oil demand will continue to grow rapidly over the next 25 years [4, 5]. Canada’s total greenhouse gas (GHG) emissions in 2011 were 702 megatonnes (Mt) of carbon dioxide equivalent (CO2 eq), 19% (111 Mt) above the 1990 emissions of 591 Mt. Steady increases in annual emissions characterized the first 15 years of the intervening period, followed by fluctuating emission levels between 2005 and 2008, and a steep decline in 2009 (mostly due to the global recession) with emissions somewhat stabilizing in 2010 and 2011 [6]. The increasing GHG concentrations and concern over the effect on climate is a commanding incentive for development of new advanced energy cycles. As part of a solution, renewable energy technologies currently supply ~ 18.5% of the world’s primary energy supply, with bioenergy the largest contributor to that. In 2011, total biofuels supply was ~ 1311 Mtoe1, accounting for 10% of the world’s total primary energy supply and ~ 54% of its renewable energy supply [7]. Among bioenergy technologies, "gasification" has potential to be utilized as a promising and diversified clean energy option. Canada enjoys the third largest share (6%) of the world’s forest resource [8] and is a leader in biomass utilization [9]. Although energy utilization from renewable sources will climb, fossil fuels will remain our main source of energy and emissions for decades to come. Stabilizing CO2 emissions will require a number of different reduction strategies [10]. “Carbon capture and storage (CCS)”, mostly                                                  1 Million tons of oil equivalent   2  originating from power plants, is among the strategies. To meet CO2 reductions by CCS alone would require 600 Gt of carbon storage [11]. CCS offers a near term option for CO2 emission reductions using available technology. It is needed to move from successful small-scale CCS projects in operation today to building ~ 3400 commercial scale projects worldwide by 2050, if CCS is to provide 20% of the needed CO2 reductions [12].   Bio-CCS (combining biomass as the energy source with carbon capture and storage) is one of the largest scale technologies that can remove CO2 from the atmosphere. Bio-CCS combines sustainable biomass conversion with CO2 capture and storage – e.g. in biofuels and bioenergy production – and is already being deployed at industrial scale in the US (ADM bioethanol-CCS project) [12].  1.1  CONCEPT AND SIGNIFICANCE OF GASIFICATION The manufacture of combustible gases from solid fuels is an ancient art, but by no means a forgotten one. Gasification involves the conversion of a solid feed (e.g. woody biomass, coal, municipal wastes, petroleum coke) into a synthesis or fuel gas (containing CO, H2 and CH4) in the presence of an oxidizing agent gas (e.g. air, oxygen, steam, CO2). Gasification is sometimes called "staged combustion", since the gas is usually produced with the intent to burn it later [13]. This raises the question of why gasifying first and then burning the gas is advantageous over direct combustion of the fuel. There are several essential advantages of gasification over direct combustion including higher potential overall energy efficiencies, cleaner processing, ability to transport in pipelines, ease of control and continuous operation [14]. Alternatively, the gaseous product can be converted to hydrogen or liquid fuels, e.g. dimethyl ether (DME), methanol, ethanol, transport fuel, as shown in Figure  1-1. The distribution of gasification capacity by application is shown elsewhere [15].  From a chemical point of view, gasification is a complex process. The conversion of gasification feedstocks can be divided into several stages: (1) decomposition of the original feedstock into volatile matter and char; (2) conversion of volatiles by secondary reactions (such as combustion and reforming); (3) conversion of the char by "char gasification" reactions with H2O and CO2 to produce fuel gases (CO, H2, CH4), in addition to char combustion when oxygen is present. Devolatilization produces a spectrum of products ranging from light gases to tars.   3  Figure  1-1. Diversified gasification products.  A simplified mechanism of gasification can be represented by the following array of reactions1 [16]:  (1) Pyrolysis:  2 2 2 4, , , , , , , ,...n mFuel Heat Char H CO CO H O CH C H Tars+ →  ( 1-1) (2) Tar cracking:  2 2 ...Tar Heat H CO CO+ → + + +  ( 1-2) (3) Combustion (oxidation): 21 / 2 110.6 /C O CO kJ mol+ → +  ( 1-3) 2 2 393.8 /C O CO kJ mol+ → +  ( 1-4) 2 21 / 2 283.2 /CO O CO kJ mol+ → +  ( 1-5)  (4) Boudouard reaction: 2 2 172.6 /C CO CO kJ mol+ ↔ −  ( 1-6)                                                   1 Heats of reaction were obtained from a library of Aspen HYSYS V7.2 software at 25°C. Methanol Synthesis ElectricityNatural GasHydrogenMethanolDimethyl Ether (DME) Dehydration Turbine, Engine or Fuel Cell Methanation Membrane Separation Fischer-Tropsch Gasification Coal Biomass Coke Bitumen Municipal Wastes  Clean Synthesis Gas (CO, H2) Liquid             Transportation Fuel  4  (5) Steam-carbon reactions: 2 2 131.2 /C H O H CO kJ mol+ ↔ + −  ( 1-7) In commercial reactors, the above reaction takes place above 825°C [27].    2 2 22 2 89.8 /C H O H CO kJ mol+ ↔ + −  ( 1-8) (6) Methanation: 2 42 74.9 /C H CH kJ mol+ ↔ +  ( 1-9) (7) CO shift (water-gas shift): 2 2 2 41.4 /CO H O CO H kJ mol+ ↔ + +  ( 1-10) (8) Steam-methane reforming: 4 2 23 206.1 /CH H O CO H kJ mol+ ↔ + −  ( 1-11) Water-gas shift and methanation reactions take place at or above about 425°C [17]. Lowering of gasifier temperatures can be achieved by the use of catalysts. Operation of gasifiers at lower temperatures and heat integration between gasifier and shift-methanator are desirable economically, and they can improve the thermal efficiency and increase the heating value of the product gas.  The pyrolysis/gasification products are strongly dependent on the nature of the feedstock and process conditions, mainly temperature and pressure. Partial reforming of the products in contact with components of the char bed may result in improved gas quality. For example, if fuel gas is the desired product, such conversion could preserve methane while reforming undesirable tars. The progress of such reforming reactions depends on the nature of the char, the inorganic (ash) components and the type of reactor. Conversion of the feedstock to fuel gases by gasification is generally endothermic, with air or oxygen typically needed to balance the process thermally. Various gasifier technologies have been developed over many decades, tailored to suit specific needs. These processes operate at pressures from atmospheric to >20 bars and at temperatures of ~700-1000ºC in fixed bed, fluidized bed and transported bed reactors. Overviews of the effect of operating parameters and feed properties on gasification reactions have been provided [18, 19]. Reviews of industrial gasification activities can be found in Knoef [13] and Li [20].   5  1.2  CONCEPT OF CO2 CAPTURE Various CO2 capture and separation processes have been considered for managing CO2 emissions, particularly from fossil-fuel combustion units and large industrial processes. CO2 capture systems are characteristically identified as post-combustion, pre-combustion or in-situ systems. In-situ capture involves both process reaction (e.g. gasification) and CO2 capture in a single integrated environment. Integrating these reaction steps involves selecting suitable reaction conditions under which each of the processes can be carried out. Of the available techniques for CO2 capture, membrane and cryogenic separation require high pressures or low temperatures for efficient CO2 capture [21]. Adsorption of CO2 with carbon-based adsorbents is limited to low temperatures (150–250ºC) [22]. Hydrotalcite compounds (Mg6Al2(CO3)(OH)16.4H2O) have a significantly diminished capacity for CO2 capture beyond 300ºC [23]. Separation processes that cannot be operated effectively under conditions suitable for the gasification reactions are not suitable for the proposed process. Past studies have demonstrated the high reactivity of metal oxides such as CaO under conditions suitable for gasification (i.e., 550–750°C, 1 atm) [21, 24–26]. This allows the solid sorbent to be used in-situ in the reaction chamber of a packed or fluidized bed conversion system, eliminating the need for a separate absorbing reactor or for sub-cooling the flue gas. These sorbents operate on a cyclic basis, with CO2 absorbed from a low concentration stream in one reactor and stripped in a second to produce a nearly pure stream of CO2. Another benefit of lime-based capture systems is the ability of lime to co-capture harmful sulfur dioxide emissions. From an economic point of view, metal oxides, which are abundant in naturally occurring rocks, present the cheapest sorbent option. For example, CaO capture of CO2 is reportedly 50% more cost-effective than absorption based on MEA (Monoethanolamine), a liquid sorbent, on a $/mol-CO2 basis [26]. This economic advantage is particularly significant for process operations if the addition of fresh sorbent is needed to maintain capture effectiveness. Abanades et al. [27] argued that the performance of lithium-based sorbents must be proven for up to ~10,000 reaction cycles to be economically competitive with sorbents derived from naturally occurring limestone. There are a number of lime-based solid sorbents, both natural and synthesized particles, being considered for enhanced gasification process. Due to their lower cost, as well as increased  6  availability, natural sorbents are currently more viable for CO2 capture than synthetic sorbents. There are two naturally-occurring lime-based sorbents: limestone (typically at least 95% CaCO3 by wt) and dolomite (~50% CaCO3 by wt, with the balance being mostly MgCO3). Dolomite has shown higher capture capacity (moles of CO2 captured per moles of CaO present) for early carbonation/calcination cycles, and almost the same performance in long term cycling [28], but its lower mass fraction of CaO gives limestone the advantage in terms of CO2 captured per unit mass of sorbent [29]. The carbonation reaction of lime-based sorbents is as follows: 2 3( ) ( ) ( ) 178 /CaO s CO g CaCO s kJ mol+ ↔ +  ( 1-12) The reverse reaction is known as calcination. The exothermic forward reaction (carbonation) has an added benefit; providing heat to the endothermic gasification reactions. However, this creates a heat demand in the calciner, creating challenges for implementation of lime-based CO2 capture on an industrial scale. Fluidized bed gasification is likely to be best-suited to in-situ CO2 capture due to increased mass and heat transfer and the ability to provide the high recycle rates required to circulate large quantities of lime-based sorbent between carbonation and calcination conditions [19]. 1.3  CATALYTIC GASIFICATION 1.3.1 Catalytic effects in coal gasification This section describes the reported effects of inorganic materials on coal gasification. These effects can be considered as kinetic effects on reaction rates and physical effects. CO2 and steam gasification are discussed. The carbon reaction with oxygen differs from reactions with oxidative CO2 and H2O agents in that it is exothermic. The rate of the un-catalyzed reaction of carbon with oxygen is five orders of magnitude faster than with CO2 and H2O [30]. Catalytic gasification with CO2 and H2O as gasifying agents between 600 to 1000°C has become of great interest. Oxygen and air coal catalytic gasification will not be discussed here, but are covered elsewhere [31]. Catalysts which are active for the Boudouard reaction (CO2 gasification) have similar reactivity with steam [17]. Effective catalysts fall into three main groups: alkali, alkaline earth and transition metals. The review below focuses on alkali and alkaline earths, as biomass is typically rich in these metals.  7  Alkali metals    McKee [32] compared carbonates of lithium, sodium and potassium and showed that K2CO3 and Na2CO3 were similar in activity, but Li2CO3 was considerably more active for the gasification by CO2 and H2O of pure graphite (a pure and well-ordered carbon). The sequence of activity of alkali metal carbonates was shown to be Li2CO3>Cs2CO3>Rb2CO3>K2CO3>Na2CO3 for pure graphite CO2 gasification on a weight % basis [33, 34]. The same trend was found for steam gasification.  Alkali carbonate activities in coal char gasification, however, differed from those for graphite gasification: In CO2 gasification of Pittsburgh char, K2CO3 was the most active salt and Li2O3 the least active [35]. The order of catalytic activity of alkali metal carbonates was the same for H2O as for CO2: Cs2CO3>K2CO3>Na2CO3>Li2O3 [36]. Activation energy was independent of CO concentration for catalyzed graphite, but dependent on CO pressure for coal char CO2 gasification [35]. Results indicated that the catalytic activity does not depend significantly on whether the (potassium) catalyst is added before or after charring [17].   The activity of potassium salts depends on the anion. For a series of potassium salts in graphite-CO2 reaction, McKee [32] showed that the carbonate, sulphate and nitrate are more effective catalysts than silicates and halides. K2SO4 was significantly less active than K2CO3 [36]. Alkali metal hydroxides exhibited a comparable effect to carbonates in promoting CO2 and H2O gasification [37]. It seems likely that an oxygen-containing anion, or an anion which is converted to an oxygen containing species, is necessary for effective catalysis [38]. The chlorides (i.e. LiCl, NaCl, KCl, RbCl and CsCl) appeared to behave very similarly and to act as inhibitors during the early stages of steam gasification, although the rate accelerated later [37].   Walker et al. [39] compared the catalytic activity of a lignite char with steam for a number of inorganic cation species, demineralized chars and raw coal chars. The order of activity was K>Ca>Na>Fe>raw coal>Mg>demineralized coal. For steam gasification of highly caking coal, Tomita et al. [40] found the catalytic order was K>Ba>Ni>Fe>coal ash, for an ash containing quantities of iron, calcium and magnesium. Veraa and Bell [37] also found potassium to be the most effective cation in a study of a series of alkali salts for the steam gasification of a sub-bituminous coal char.     8  Rates of coal char catalyzed CO2 gasification have been found to increase with increasing temperature and pressure [36]. Steam gasification cation also showed dependence on temperature, but no pressure effects were observed [36]. Physical changes during gasification have also been observed. Gasification proceeds by loss of carbon atoms from faces of basal planes, steps and crystal edges. If catalysts promote gasification by the former method, the formation of pits, which become progressively deeper and more circular, is observed. Catalysts acting on crystal edges produce a channelling effect [41–43]. Marsha et al. [43] found that while isotropic components are more susceptible to gasification, anisotropic components (e.g. carbon prismatic edges) are more susceptible to attack by alkali and therefore to catalyzed gasification. In the hot stage, microscopy studies with alkali metal catalysts, the catalysts appeared to melt and be absorbed into the coal chars pores. In the case of graphite, spherical channels of roughly equal breadth and depth were formed on the surface [44]. Similar microscopic evidence of channelling and pitting by alkali salts in steam was observed as in carbon dioxide [44].        Alkaline earth metals Alkaline earth salts are generally less active than alkali salts in oxidative gasification. As far as metals present in fuels are concerned, the presence of calcium in coal is very significant and accounts for increased reactivity of low rank coals. Spiro et al. [45] reported that activation energies of alkaline earth salts are higher than alkali catalyzed and even un-catalyzed reactions for coal char gasification. However, apparent activation energies for the graphite CO2 reaction were considerably lower. The catalytic activity of the cations was in the order Ba>Sr>Ca for both coal char and graphite CO2 and steam gasification [32, 45, 46, 47]. Carbonates of these elements are more active. Magnesium showed very little activity for graphite gasification [48]. Catalyst effects were more pronounced in a CO2/CO gas mixture than in pure CO2. The presence of calcium was found to decrease the inhibiting effect of CO [49]. When compared with the alkali metal catalysts, CaO was approximately equal in activity to Na2CO3 [36].             Even though calcium is not such a good catalyst for steam gasification as potassium, it does give higher rates at low loadings because it reacts less with coal minerals than alkali metals [50].    9  Hydrogasification (C-H2) Coal and coke have less reactivity to hydrogen than to steam or CO2. There is interest in catalysts which are effective at lower operating temperatures and pressures. Salts and oxides of the alkali metals are less effective in hydrogasificaton than transition metals [17]. Alkaline earth metals are only poor catalysts for hydrogasification [35, 51]. Therefore, alkali and alkaline earth metals (AAEM) are mostly effective on catalyzing the Boudouard, equation ( 1-6), and steam-carbon, equation ( 1-7), reactions.    1.3.2 Catalytic gasification mechanisms Catalyzed C-CO2 reaction The Boudouard reaction is a simpler reaction to study than steam gasification, which is complicated by further reactions involving products. In laboratory scale experiments (e.g. thermogravimetric analysis), therefore, CO2 is often used for kinetic studies.  Mechanistic studies (e.g. [32]) suggest that alkali metal catalysis operates by an oxidation-reduction sequence. However, the nature of the intermediate oxidized and reduced species is uncertain. Observations have shown that K2CO3 melts and spreads into the micropore structure of the char, facilitating contact between carbon and catalyst [38]. The molten film then serves as an oxygen transfer medium between the gaseous reactant (CO2 or H2O) and the char. In other words, the catalyst extracts oxygen ions from the gasifying agent. The O2- ions are the mobile in the molten phase. McKee [32] proposed a sequence of redox reaction to explain the mechanism:   2 3 ( , ) 2 ( ) 2 ( ) 3 ( )M CO s l C s M g CO g+ ↔ +  ( 1-13) 2 22 ( ) ( ) ( , ) ( )M g CO g M O s l CO g+ ↔ +  ( 1-14) 2 2 2 3( , ) ( ) ( , )M O s l CO g M CO s l+ ↔  ( 1-15) where M denotes an alkali metal. The summation of the above equations yields the Boudouard reaction, equation ( 1-6). Reactions ( 1-14) and ( 1-15) will be rapid and ( 1-13) is likely to be rate-controlling, consistent with the observation [32] that gasification proceeds slowly below, but rapidly above, the carbonate melting point. Figure  1-2 presents a clear picture of how the   proposedalkali meFigure  1 A mechaearth elehigher teoxides isCatalyzeWood et twice themechanisChatterji CO2 gasisteam invThe sum( 1-16) is [35]. Oth mechanismtal salt catal-2. Schematnism analogments of thmperatures  likely [31, 4d C-H2O reaal. [38] repo rate of COms were fo[33] and Mfication to eolving the i2 of the aboinhibited byer similar a operates. Systs [52–56ic of reactioous to that e Boudouarthan group 8]. Other prction rted that un2, and addund to be cKee  [32] pxplain the cntermediate2 3 ( , )M CO s l( ) 2M g H+2 ( , )MOH s lve reactions increasing pproaches himilar mech]. n pathways for the alkald reaction. IA monoxioposed mecder identicaition of K2operating froposed a satalytic effe formation o2 ( )C s+ ↔2 ( ) 2O g M↔( )CO g+ ↔ is the steaamounts ofave been reanisms werat the carboal. [57])  i metals hasHowever, ades, reductihanisms canl conditionsCO3 enhancor both COequence of ects of potasf hydroxide2 ( ) 3M g C+( , )OH s l +2 3 ( , )M CO s lm-carbon r CO, reactioported in the proposed bn/catalyst ju been propos group IIAon of the c be found e, steam gasied the rate2 and stealementary rsium, sodium:    ( )O g  2 ( )H g  2H+  eaction (equn ( 1-16) is e literature y different nction (adapsed for cata oxides ararbonates olsewhere (efied char an proportionm gasificateactions par and caesiation ( 1-7)likely to be for alkali mauthors for  ted from Ralysis by alke more stabnly as far a.g. [58]).  d carbon at aately. The ion. McKeeallel to thosum carbonat( 1-16)( 1-17)( 1-18)). Since rearate-determetals [32, 3710 other o et aline le at s the bout same  and e for es in ction ining , 59,  11  60]. Also mechanisms comparable to those of the alkali metals have been proposed for a cyclic sequence of redox reactions for alkaline earth elements [46, 47].     1.3.3 Catalytic activity Catalyst dispersion Catalyst activity partly depends on dispersion properties. Potassium compounds are mobile and more easily dispersed than calcium compounds [17]. Unlike calcium, potassium appears to be able to diffuse through the char to form active gasification sites [17]. The solubilities of alkali sulphates and calcium salts in steam are much lower than for K2CO3 [61]. K2CO3 is less mobile during CO2 gasification compared to steam gasification of coal [36]. It is suggested that dissolving the catalyst in the gasifying medium will aid dispersion on the coal surface [17]. The activity of calcium was found to decrease during gasification [50], whereas potassium maintained its activity, reflecting its relative mobility and ability to regenerate active sites. Wood et al. [38] found that the performance of alkali catalysts did not depend on the method of addition: simple dry mixing was as effective as impregnation with aqueous solution of catalyst, followed by drying.   Catalyst loading The effectiveness of potassium is sensitive to catalyst loading. A minimum loading of about 2 wt% is needed with high ash coals because of deactivation by minerals [62]. High catalytic activity of potassium can be only achieved with high catalyst concentration up to 30 wt% [63]. An excessive amount of catalyst may cause plugging of coal pores. No general simple correlation between rate and catalyst loading could be found. Catalysts which are effective only in generating reaction sites, not lowering activation energy, have to be added in large amounts to increase the specific rate per m2 especially when the surface area is large. Furthermore, catalyst efficiency depends on its ability to spread out thinly to maintain contact with the carbon reactant. It was shown that a coal char surface is more efficiently catalyzed by alkaline earths than is graphite [17].    12  1.3.4 Black liquor gasification In the pulp and paper industry, black liquor char (BLC) is the pyrolysis product of kraft black liquor solids (BLS). Kraft black liquors are dark viscous liquids resulting from digestion of wood by an aqueous solution of sodium hydroxide and sodium sulfide. In general, during kraft pulping, about 50% of the wood enters the solution. This dissolved wood originates primarily from hemicellulose and lignin and is present in the liquor as, respectively, hydroxy acids and alkali lignin. A black liquor char bed is formed at the bottom of a chemical recovery furnace after drying, pyrolysis, and partial gasification of a spray of concentrated black liquor droplets. The porous char has a very high sodium content, and also contains inorganic sulfur compounds like Na2SO3, Na2S2O3, and Na2SO4. Gasification of the char provides the reducing atmosphere required to convert the inorganic sulfur compounds to Na2S, one of the active pulping chemicals. Because black liquor char consists mainly of a mixture of carbon, sodium carbonate, and sodium sulfide, alkali-catalyzed carbon gasification studies provide useful information for gasification of black liquor char.  Despite the similarities in chemical composition, there is a significant difference between black liquor char and carbonaceous chars doped with alkali metal salts. In the latter case, the chars are prepared either by impregnation or by pyrolysis of the char precursor with catalyst powder. Black liquor char, on the other hand, is produced from a liquid phase in which sodium is mixed and chemically bound with the char precursor on a molecular scale. This leads to a three-dimensional, and presumably, fine dispersion of the alkali metal catalyst.  Another difference is the several times larger initial sodium/carbon atomic ratio for black liquor char. The only study on CO2 gasification kinetics of "slow pyrolysis" BLC was reported by Li and van Heiningen [64]. They found that the gasification rate of BLC was more than an order of magnitude higher than for coal char with an optimum alkali metal catalyst loading. Also, the BLC gasification rate remained high when the sodium/carbon ratio far exceeded the optimum ratio for doped coal chars [65]. Li and van Heiningen [66] performed BLC CO2 gasification at 600-800°C in a TGA. The fine dispersion of sodium in the carbon structure was confirmed by SEM-EDS. They reported that the high loading and fine distribution of sodium are responsible for a gasification rate at least one  13  order of magnitude higher than for activated carbon or coal chars impregnated with an optimal loading of Na2CO3. Fine dispersion of sodium in the bulk of the carbon structure of BLC, rather than solely on the internal surface, is probably responsible for the sustained high gasification rates at different Na/C ratios. However, the rate of coal char gasification with impregnated Na2CO3 was greatly reduced due to pore plugging and sintering of the catalyst. In another study, Jaffri and Zhang [67] studied the CO2 gasification of Fuijian anthracite using black liquor (BL) as a catalyst. Experiments were performed in a TGA at temperatures ranging from 750 to 950°C at ambient pressure to evaluate the catalytic effects of 3%, 5%, 8% and 10%  Na-BL-loading on conversion of carbon. Better catalytic activity was found with 8% Na-BL-loading than with the other concentrations. The activation energy ranged from 76 kJ/mol to 104 kJ/mol at 8% Na-BL-loading, significantly less than the 151 kJ/mol to 185 kJ/mol found for a non-catalyst case. This clearly demonstrated that BL could be the source of an inexpensive and effective catalyst for coal gasification. 1.4  CO-GASIFICATION OF BIOMASS AND FOSSIL FUELS The major challenge to achieve commercialization of biomass gasification is the lack of fuel supply security. There is heavy dependence on the location of the gasifier relative to the source of the bio-feedstock, while coal can be transported more easily and economically due to its higher density and much greater concentration at source [68]. Catalyst can be added as an additional solid component to coal gasifiers. Some gasification catalysts, such as alkali salts, can be added to the feed solids. The presence of catalytic action in the char bed can aid operation, decreasing tar yield. Other catalysts can also be used, such as dolomite, olivine, nickel and magnesium oxides, zinc oxides, cobalt, and molybdenum oxides [69]. Cost and recovery of the catalysts are major issues in catalytic gasification. One possibility to solve the above problems related to feedstock of biomass gasification and coal catalytic gasification is to co-gasify a biomass and coal. Biomass/coal co-gasification could be a bridge between energy production based on fossil fuels and energy production based on renewable fuels [70]. Several synergistic benefits might be realized by biomass/coal co-feeding: (1) The addition of biomass to an energy production system lowers the CO2 footprint for that process. (2) Some components, such as alkali and alkaline earth metals in the biomass, may act  14  as catalysts, promoting gasification of coal. (3) The presence of the biomass may reduce and/or modify the properties of the waste products of the gasification process, in particular the tar and char. (4) Co-feeding coal with biomass may help to overcome some of the feeding difficulties commonly associated with biomass feeding. (5) The overall feedstock will be drier than for biomass alone, thereby avoiding or reducing the need for predrying of the wetter biomass materials. (6) In cases where feeding a single waste material is uneconomic, for example due to high transportation costs, combining different hydrocarbon sources could result in processes that are economically viable. Whether such synergetic effects can be realized depends on gasification operating and fuel conditions such as feedstock type, direct particle contact, temperature and pressure, reactor type, etc. [71]. Co-gasification of coal and biomass is currently being conducted in IGCC1 electricity generating power plants like the Willem-Alexander power station in Buggenum (Netherlands), where residual wastes from the agricultural sector, such as sawdust, grape and sunflower seeds, and peanut shells are co-gasified with coal. Similarly, at ELCOGAS in Puertollano (Spain), the world’s largest IGCC facility which uses coal and petcoke as feedstocks, there is an ongoing project aimed at evaluating the effects of adding a small percentages of biomass (up to 10 wt%) on the plant performance [72]. The notable features of mixed-feed gasification can be divided into kinetic/catalytic effects of one material on the conversion of either the char or volatiles from the other, and reactor behaviour, particularly in fluid beds, when mixed feedstocks are employed.   Co-gasification literature review Considerable excellent work in co-gasification from bench scale to large pilot scale has been undertaken during the last two decades. Some of this work has been summarized in APAS Clean Coal Technology [73] and Carvalho et al. [74]. Davidson [75] provided a very good summary of pertinent efforts prior to 1997. Below is a brief review of the literature related to co-gasification of different fuels. Thermogravimetric analysis (TGA) Brown et al. [76] performed a novel study on catalytic gasification of coal char using switchgrass-derived potassium salts, with CO2 as the medium in a TGA. The results obtained                                                  1 Integrated Gasification Combined Cycle  15  with switchgrass ash were especially impressive, with an almost eight-fold increase in coal char gasification rate at 896ºC in a 10:90 by weight mixture of coal char and switchgrass ash. These results indicate that biomass ash could provide inexpensive coal gasification catalysis. On the other hand, Pan et al. [77] studied the pyrolytic behaviour of mixtures of low-rank coals with pine chips in an atmospheric pressure thermogravimetric analyzer, heating samples at 100ºC min-1 to temperatures between 108 and 900ºC. No interactions were observed between the two fuels. Instead, the pyrolytic behaviour of blends in any proportion consisted of the weighted average behaviour of the two individual samples. In view of the different temperature ranges required to devolatilize coals and lignocellulosic biomass materials [78, 79], the utility of slow heating experiments in characterizing fuel blend properties appears to be limited. Under very low heating rates, the temperature ranges for pyrolysis of biomass and coal differs considerably, revealing that the two processes are separated and therefore synergies cannot be found.  Fixed bed reactor Fermoso et al. [80] performed high-pressure (15 atm) steam/oxygen co-gasification in a bench-scale fixed bed reactor with a mixture of bituminous coal, olive pulp and pine sawdust. The addition of a small amount of biomass (up to 10 wt.%) led to ~9% more carbon conversion than gasification of coal alone and surprisingly greater H2 and CO production [81]. The authors believed that complementary effects may take place due to volatile-volatile and volatile-char interactions, and possibly also resulting from mineral matter catalytic effects.  A mixture of woody biomass (Japanese cedar) and coal (Indonesian) was gasified by Kumabe et al. [82] in an atmospheric downdraft fixed bed gasifier to investigate the effect of mixtures with varying contents of biomass and coal on the H2, CO and CO2 products using air and steam. The conversion to gas and H2 increased with the proportion of biomass, whereas the yields of char, tar and CO2 decreased. No apparent synergy was observed in the carbon distribution of products. Based on the product gas compositions (H2/CO ratio), the authors also reported that a low biomass/coal ratio led to the production of a gas favourable for the syntheses of methanol and hydrocarbon fuel, whereas a high biomass/coal ratio facilitated production of a gas favourable for DME synthesis.   16  Fluidized bed reactor Pinto et al. [69, 83, 84] comprehensively investigated the effect of different bio/non-bio wastes on coal gasification in a 3.7 m diameter atmospheric bubbling fluidized bed operating at 845ºC. Higher proportions of pine in the feedstock led to a marked increase in carbon oxides and a decrease in the tar in the fuel gas. The authors passed the gaseous product through two-stage catalytic cracking reactors to produce clean gas with very low contents of tar, sulphur, nitrogen and halogen compounds, suitable for use in motors, turbines or fuel cells.  McLendon et al. [71] co-gasified furniture manufacture sanding waste with coal in a pilot-scale ash agglomerating fluidized bed at an operating pressure of 3.03 MPa using a rotary shaft feeder. Synergies with sub-bituminous coal/biomass mixture were not readily apparent. However the transport (rheological) properties of all the coal/biomass mixtures greatly improved relative to coal alone. As a result, plugging was greatly reduced, and solids handling was much easier.    Sjostrom et al. [81] reported synergies in fluidized bed oxygen co-gasification of wood and coal mixtures at small particle sizes with maximum feed rates of 5.2 kg/h and a maximum pressure of 15 bar. They reported higher reactivity with the mixed fuel, with both tar and ammonia yields being reduced. Their group reported similar results for co-gasification of mixtures of coal and wood [85, 86]. Madsen and Christensen [87] carried out a series of tests in an air-blown fluidized bed and entrained bed with co-gasification of coal and straw. Pressure in the larger unit (based on GTI U-GAS design) was up to 14.2 bar, while feed rates of the feedstock were as high as 720 kg/h. Feeding presented problems, but some synergies were noted. Reinoso et al. [88] reported on a comprehensive experimental program including a series of tests in larger pilot scale air-blown circulating fluid bed gasifiers at near atmospheric pressure. Feedstocks were waste coal, lignite, and pine chips at solids feed rates up to 800 kg/h. The results indicated that gasifying waste coals in circulating fluid beds of the type tested would not be viable because of economic, operational and design constraints. Modeling comparisons indicated synergies between the coal and biomass.  17  De Jong et al. [89] found that there were synergies with co-gasification of coal, Miscanthus (a type of grass) and straw in an air-blown fluidized bed gasifier operating at up to 5 bar. Pan et al. [77] co-gasified low-quality coals with pine chips at atmospheric pressure in an air-blown bench-scale fluidized bed gasifier. They confirmed the feasibility of the co-gasification process. Kurkela et al. [90] reviewed co-gasification of coal and biomass and suggested that biomass wastes could be used as an additional feedstock in a large pressurized coal gasification IGCC process or synthesis gas plant. Furthermore, they studied co-gasification of coal with different biomass species in a PDU-scale pressurized (5 bar) fluidized bed and confirmed the technical feasibility of a simplified IGCC system using blends of wood wastes and coal. Up to 50% of the coal on a weight basis could be replaced by wood wastes without problems of ash sintering, tar loading or operational difficulties. Co-gasification of wood residues and coal has also been successfully carried out in the 18 MWt Pilot Plant of Enviropower in Finland [90]. Comparison of fluidized and fixed bed reactors Only one paper [91] was found which compared the co-gasification inside fluidized and fixed-bed reactors. The experiments were carried out at 850ºC and 1000ºC and pressures up to 25 bar. The authors concluded that neither gasification in the fixed-bed reactor nor in the fluidized-bed reactor during pyrolysis appeared to affect the total release of volatiles from the mixed fuels relative to independent feedstocks. Although co-gasification in CO2 showed similar trends to those with a single fuel, relative combustion reactivities of Daw Mill coal silver birch wood mixture chars prepared at 1000ºC were almost twice those for coal char alone, despite the total gasification of silver birch at this temperature. This suggests that the silver birch mineral matter may have had a catalytic role during combustion of the coal char. Apart from the possible catalytic effect of biomass ash at high temperatures, it was speculated that the observed synergistic effects were not of a magnitude that might influence process design. Entrained flow reactor Working with an entrained flow reactor, Rudiger et al. [92] reported an absence of synergistic effects among different fuels. They explained their findings in terms of the different temperature ranges within which devolatilisation of coal and biomass take place (because of differences in particle size and chemical structure of the two fuels). The spatial separation of sample particles  18  in this apparatus does not appear to have been considered. It has also been reported that short residence times (<1 s) in entrained flow reactors could limit total weight loss, because of incomplete pyrolysis [93]. In any case, entrained flow reactors appear to be unsuitable for investigating synergistic effects with blended feedstocks. This finding could well also apply to wire-mesh reactors, where the design requires that the sample particles be separated. In summary, based on the fuel, type of reactor and operating conditions, some previous work has reported significant synergies in co-gasification, whereas other authors have observed none. Ricketts et al. [94] discussed technical, economic, governmental and social aspects of co-gasification technology for electricity generation. They concluded that there are environmental, technical and commercial reasons why coal might be co-gasified with biomass and wastes. Addition of biomass to fossil fuels reduces greenhouse gas CO2 emissions, leading to real and tangible environmental benefits. Based on their analysis, the economics of small scale gasification power systems are unattractive, whatever fuel is used. However, large scale IGCC technologies lend themselves to co-gasification, and this is likely to become attractive as coal-based IGCC is adopted by the market. 1.5  IN-SITU CO2 CAPTURE The CO2 capture is a vast field to review. There have been a number of studies, both experimental and computational, of CO2 capture using lime-based sorbents. Therefore, this section presents an abbreviated review of some key elements of in-situ CO2 capture. A useful extensive summary of recent research on biomass gasification with in-situ CO2 capture was provided by Florin and Harris [19].  Several have examined the performance of lime-based sorbents over many absorption/desorption cycles [29, 95–106]. However, little attention has been paid to re-creating the actual operating conditions of a gasification/calcination system. A thorough understanding of the cyclic behaviour will be crucial to implement this technology. To date, studies of biomass gasification integrated with CO2 capture have involved almost entirely computer-based modeling of the reaction process. The majority of models are thermodynamic, based on Gibbs free energy minimization with a few models based on reaction kinetics [107–109]. Experimental studies are needed to prove the feasibility of the process, validate models and investigate potential problems.  19  Carbon dioxide capture using a lime-based sorbent is dependent on the difference between the CO2 partial pressure in the reactor (PCO2) and the equilibrium pressure of CO2 (Peq,CO2) in the calcium carbonate (CaCO3) matrix. Hu and Scaroni [110] fitted the relationship between the equilibrium partial pressure of CO2 and temperature by  27.,19,680( ) 1.826*10 exp( )( )eq COP atmT K−=  ( 1-19) The region above the equilibrium line in Figure  1-3 represents conditions where carbonation is favoured, whereas calcination is favoured below the equilibrium line. Cyclical absorption and desorption of CO2 is accomplished by cycling the operating conditions between the carbonation and calcination regions. Temperature swing cycling, represented by the horizontal line in Figure  1-3, is difficult in industrial applications due to the large thermal load required to heat the solids, including inerts, to the calcination temperature, as well as to provide the heat required for the endothermic calcination reaction. A less-energy intensive and more industrial viable option would be to swing the CO2 partial pressure while keeping the temperature nearly constant, as represented by the vertical line in Figure  1-3.  Figure  1-3. Equilibrium partial pressure of CO2 for calcium carbonate. Sorbent-enhanced gasification A typical lime-enhanced gasification process with gasification and CaO carbonation taking place in one reactor and CaCO3 calcination in the other is shown in Figure  1-4. The first description of Temperature, ºC 500 600 700 800 900 1000 1100CO2 equilibrium partial pressure, Peq,CO2, atm1e-51e-41e-31e-21e-11e+01e+11e+2CarbonationCalcinationPressure swingTemperature swing  calcium-bduring cytheoretic[111]. Hstabilizatenhancedet al. [1sorbent tGibbs freequilibriuGasificatbased onBretado e(based owater-gayields wselected carbonatiased sorbencling occural capacity oowever, doion by Mg gasification12] performo investigate energy mim hydrogeion efficienc the LHV oft al. [113] pn Gibbs fres shift reactere predicteas the prefeon/regeneraFigts for hydrs due to sinf 0.79 g(COlomite shoO [111]. W is gaining ed equilibrie the effect nimization. n yield byy was pred the producterformed che energy mion. Three sd to be 98%rred sorbenttion cycles. ure  1-4. Schocarbon contering duri2)/g(CaO),ws higher ith its poteinterest, andum modelinof a CO2 soThe CaO so ~ 19% anicted to incr gas. emical equinimization)orbents we, 81% and  due to its  ematic of liversion wasng high-temhigher than multi-cyclential benef a number g of ethanrbent on hyrbent was fod to reducease from 6ilibrium ana to determire examined95%, respehigh H2 yieme-enhance in 1886 [1perature refor dolomit conversioits over staof models hol gasificatidrogen yielund to havee CO2 in t2.9% witholysis using ne the best : CaO·MgOctively, on ld and thermd gasificatio11]. Loss ogeneration. e, 0.46 g(COn levels dndard gasifave been pron in the pd based on  the potentihe productut CaO to 7HSC ChemiCO2 sorben, Li4SiO4 aa molar basal stabilityn process. f sorbent acLimestone h2)/g(CaO·Mue to strucication, soresented. Maresence of Aspen Plusal to enhanc gas by 502.1% with stry 5.1 softts for imprnd Na2ZrOis. Na2ZrO3 during rep20 tivity as a gO) tural bent-hishi CaO  with e the .2%. CaO, ware oved 3. H2  was eated   21  Weimer et al. [114] performed equilibrium calculations and condensed material balances demonstrating the influence of process conditions and sulfur concentration on CaO carbon capture. Solids must be purged from lime-enhanced gasification (LEG) processes to remove sulfated CaO and ash. If LEG is used in conjunction with cement production, it was predicted that electrical efficiencies up to 42% could be achieved using an integrated gasification combined cycle (IGCC). Gasifier operating conditions of 750°C and 2.0 MPa were found to be optimal, with effective capture of 80% of the carbon. A purge fraction of 0.061 (solid purged / regenerator solid output) was predicted to be optimal.  There have been few small-scale enhanced biomass gasification experiments. One, conducted by Florin and Harris [115] in a thermogravimetric-mass spectrometer (TG-MS) system, showed that the use of CaO in methyl cellulose gasification (100% Ar atmosphere) led to the formation of H2, whereas, no H2 was formed in the absence of CaO. CaO hydration to Ca(OH)2 was found to be a dominant reaction, with the exothermic hydration, and, to a lesser extent, carbonation reactions driving secondary decomposition reactions and particle swelling, facilitating the escape of evolved species from the reaction vessel. Addition of water vapour caused a substantial reduction in char formation and a 14.4% increase in hydrogen yield in the presence of CaO. Formation of low-temperature melts The formation of melts involving mixtures of CaO, CaCO3 and Ca(OH)2 at temperatures significantly below the melting temperatures for the pure components (2927°C for CaO; 1339°C for CaCO3; and Ca(OH)2 decomposes at 385°C, [116]) have been reported [24, 117–119]. The formation of low-temperature melts may hamper operations by causing blockages, whilst the agglomeration of sorbent and fuel particles is expected to limit gas-solid interactions, reducing CO2 capture and fuel conversion efficiencies. Thus, the conditions under which melts form impose additional constraints on reaction systems. Sorbent particle sintering Selecting operating conditions which minimize the sintering of CaO is important for maintaining reactivity during multiple capture and release cycles. Sintering is the process by which solid particles fuse together when heated at temperatures below their melting points [120, 121]. Sintering leads to a reduction in both surface area and pore volume, in turn adversely affecting  22  the rate and extent of gas–solid reactions. Sintering of CaO tends to be accelerated by increased temperature and the presence of CO2 and H2O in the gas phase [120, 121]. Borgwardt [120] studied the sintering rate of CaO in a N2 atmosphere at temperatures between 700 and 1100°C. CaO derived from pure CaCO3 was compared with CaO derived from naturally occurring limestone and Ca(OH)2. The principal impurities amounted to ~ 4 wt% of the parent material and included Mg, Si, Al and trace amounts of Fe. Borgwardt’s analysis indicated that the transport occurred via diffusion. Consistent with this mechanism, the reported sintering rates of pure CaO were much slower than for CaO derived from natural limestone. The rate of diffusion in a solid is known to increase with an increase in the number of defects in the crystal structure (which may be associated with the presence of foreign atoms). Thus, a greater rate of sintering is expected, when a solid-state diffusion mechanism applies, for the sintering of impure CaO, compared to pure CaO. Due to the likely ranges of impurities and their levels in different limestones, sintering behaviour must be assessed on a case-by-case basis. Limestones with very high levels of impurities may be unsuitable for the integrated capture process due to enhanced susceptibility to severe sintering. Effect of tar on sorbent deactivation  Interaction between CaO and biomass conversion products, in particular tar and coke, is expected to hamper CO2 capture due to sorbent deactivation. However, the deactivation mechanism is not well understood. The elimination of tar from the product gas is a major technical challenge for biomass conversion processes. Elevating the reaction temperature (>800°C) and the steam-to-biomass ratio is expected to increase the rate and extent of tar elimination by promoting endothermic cracking and reforming tar species [122–125]. These elimination reactions lead to coke formation. Build-up of coke on sorbent particles is expected to affect sorbent deactivation [123–126]. Elevated temperatures are also expected to promote the decomposition of coke via steam or CO2 gasification [124]. Thus, there is a trade-off between the optimal temperatures for eliminating tar, decomposing coke and maximizing CO2 capture by CaO due to the kinetically-limited tar decomposition and thermodynamic constraints on the carbonation reaction. If it is assumed that CaO deactivation depends on the concentration of the tar and subsequent deactivation due to coking. (For example, Corella et al. [126] suppose a tar concentration of 2 g/Nm3 to be a critical upper limit for minimizing sorbent deactivation.) The  23  tar concentration is then likely to be critical for determining CaO loading and cycling rate in the context of a continuous regeneration process. In addition, CaO is well known to have a catalytic effect on the decomposition of tar species during biomass gasification [122–124, 127–129]. The use of CaO or calcined dolomite in situ or downstream from biomass gasifiers is discussed in detail in three reviews [125, 130, 131].  Delgado et al. [123] reported that the order of catalytic activity for tar conversion is calcined dolomite> CaO>MgO, based on bench-scale experiments using a downstream fixed bed reactor for the gasification of pine sawdust. The lowest temperature tested was 800°C, and catalytic activity was assessed in terms of the minimum temperature required for tar elimination and the rate of deactivation of the tar decompositions catalysts. Simell and Leppälahti [129] reported effective tar destruction using calcined dolomite in the temperature range from 800 to 900°C. Furthermore, they reported that an increase in temperature, residence time and the Ca:Mg ratio improved the extent of tar decomposition. In a later study, Simell et al. [122] investigated the influence of CO2 partial pressure on the catalytic activity of CaO and calcined dolomite for the decomposition of toluene, as a model tar species. A loss of activity was reported when the CO2 partial pressure exceeded the decomposition pressure of CaCO3 (reversing the driving force for CO2) for both limestone and dolomite. Clearly there are operating limitations of CaO for both CO2 capture and tar elimination. Pfeifer et al. [132] reported effective tar destruction (<2 g/Nm3) for the gasification of wood pellets in a circulating fluidized bed gasifier (100 kWth pilot plant) using CaO and olivine as bed material at temperatures from 600 to 700°C. Further research is required to better elucidate the mechanisms of tar elimination using CaO or calcined dolomite, particularly at moderate temperatures suitable for in situ CO2 capture. 1.6  SCOPE OF THIS THESIS As an old concept and technology, gasification has attracted extensive research and development. In this work, it is intended to study co-gasification of biomass and fossil fuels with the following primary objectives: (1) To help understand the interactions between biomass and fossil fuels and their kinetic behaviour in a two-stage CO2 co-gasification lab scale process. This is addressed in  Chapter 2 with the aid of thermogravimetric test data.  24  (2) To gain data at pilot plant scale which demonstrates the key characteristics of co-feed gasification. This is covered in Chapters 3 and 4. (3) To study the integration of co-feed gasification with in-situ CO2 capture via lime-based sorbents in a pilot scale facility using temperature swing. This is presented in  Chapter 5.  (4) To examine the reliability of equilibrium modeling to predict the product gas compositions for the co-gasification and lime-enhanced gasification processes. This is addressed in  Chapter 6 with the assistance of Aspen Plus simulator.       Chapter 7 summarizes the main conclusions of the thesis and makes recommendations for future work.  This project is strategic because it addresses key issues related to a major clean energy option, potentially contributing to significant avoidance or replacement of fossil fuels (hence to reductions in greenhouse gas emissions), and improving understanding of complex high-temperature reactors. The key intended strategic benefit of this project is increased flexibility and efficiency in the conversion of solid resources to useful fuels and chemicals, while playing a role in accelerating the transition from fossil fuels to renewable fuels.             25  Chapter 2. FUEL CHARACTERIZATION AND THERMOGRAVIMETRIC ANALYSIS OF BIOMASS AND FOSSIL FUELS   As presented in the previous chapter, synergistic interactions during co-pyrolysis/gasification of coal and biomass have been widely investigated [133–138]. Although some previous studies showed catalytic effects of alkali and alkaline earth metals on the reactivity and gas product distribution during pyrolysis and gasification [139–142], there is still debate in the literature over whether there are kinetic and product distribution effects when biomass is co-feed with coal [143]. Full understanding of co-feed gasification is not possible without studying the physical and chemical properties of raw and charred fuels  [144–148]. In this chapter, physical and chemical properties of two fossil fuels (sub-bituminous coal and fluid coke) and two types of biomass (switchgrass and sawdust) are first investigated at different pyrolysis temperatures. The surface area of each of the fuels was analyzed using Brunauer-Emmett-Teller (BET), Dubinin-Radushkevich-Kaganer (DRK), scanning electron microscope (SEM), and X-ray diffraction (XRD) methods. Next, the pyrolysis kinetics of fresh and blended samples were studied in a thermogravimetric analyzer (TGA). A kinetic model is developed to predict the fuel conversion, and to determine the pyrolysis activation energy of single-fuel and mixture reactions [149]. Finally, the gasification kinetics of blended biomass and fossil fuels and the interaction between biomass and fossil fuel minerals  are studied based on TGA measurements [150]. 2.1  EXPERIMENTAL 2.1.1 Feedstocks Two types of Canadian biomasses with widely differing ash compositions, Manitoba switchgrass (SG) and beetle-killed BC pine sawdust (SD) from British Columbia, were chosen. Panicum virgatum, commonly known as switchgrass, is a perennial warm-season bunchgrass native to North America, where it grows naturally from Canada southwards into the United  26  States and Mexico. Switchgrass has also been identified as having potential as an energy crop for Eastern Canada [151] and for the US [152–154], where it was selected as a model herbaceous crop for the Oak Ridge National Laboratory’s Biofuel Feedstock Development Program. The fossil fuels investigated were an Alberta Genesee coal, classified as sub-bituminous and suitable for IGCC systems for electricity generation [155, 156], and fluid coke (FC) from Syncrude Canada Ltd. in Fort McMurray, Alberta, a low-ash fuel. The fluid coke is a by-product in the conversion of oil sands bitumen to synthetic crude oil. Large quantities of this coke are being stockpiled in northern Alberta [157]. Containing over 85 wt% carbon, the oil sands fluid coke represents a significant energy resource. 2.1.2 Sample characterization All samples were crushed and sieved to a particle size of 300-355 µm (US mesh #45-50) before characterization analysis and pyrolysis/gasification experiments. This particle size range is typical of particles fed to bubbling fluidized bed gas-solid reactors, and also of what is used in this study for larger scale experimental tests (see the next chapter). A C.S. Bell Co. hammer-mill with 1/16 inch (1.59 mm) screen and a FRITSCH disc-mill were used for biomass and fossil fuel crushing, respectively.  The ASTM D346 and D346M-11 standard method were followed for feedstock sampling. All fresh samples were then characterized according to ASTM D3176 and D5373 standards for carbon, hydrogen, nitrogen and sulfur (CHNS) contents (based on Elementar vario MACRO); ASTM D4239 for total sulphur content (using a Leco S632); ASTM D3174 to determine ash content (using a Carbolite AF1100 ashing furnace); ASTM D3302/D3173 for moisture content (via Carbolite MFS/1); ASTM D3175 for volatile matter (using Leco S632) and ASTM D5865 for heating value (using a Leco AC600). The “blend samples” were well-mixed mechanically by hand prior to being added to the reactor basket. Elemental carbon, hydrogen, nitrogen and sulphur analyses were performed using a Perkin-Elmer 2400 Series II CHNS/O analyzer operated in the CHNS mode [158]. Measurements (with both nitrogen and carbon dioxide adsorption) were performed for fresh samples, as well as for chars derived from these samples, to determine surface areas using a Micromeritics ASAP 2020 surface area and porosity analyzer. The main source of adsorption  27  measurement errors came from fuel gravimetric measurements, as well as from electrostatic charges acquired during loading of the samples into the measurement tubes. Prior to the adsorption measurements, the samples were degassed at 105°C and 0.15 mbara for 24 h. N2 surface areas were calculated using the Brunauer-Emmett-Teller equation [159]: 0 01 1( )( )a m mP C PV P P V C V C P−= +−  ( 2-1) whereas the Dubinin-Radushkevich-Kaganer (DRK) equation, [160, 161], 201ln ln (ln )a mPV V DP= −  ( 2-2) where 2 21 1606D k R T=  ( 2-3) was used to determine surface areas based on CO2 adsorption, as it is more appropriate than the BET equation for the CO2 adsorption experimental pressure range [144, 162]. Here P is the pressure of gas adsorption (atm), P0 the saturation pressure of the analysis gas at the analysis temperature (atm), Va the volume of gas adsorbed (cm3 g-1), Vm the adsorbed monolayer volume (cm3 g-1), C the BET adsorption constant, k1 a constant characterizing the Gaussian distribution, R the universal gas constant (J mol-1K-1) and T the temperature (K). The DRK equation is based on the Polanyi thermodynamic theory of adsorption assuming that the adsorption in micropores involves a volume-filling process. On the other hand, in the BET equation, it is assumed that adsorption is a layer-by-layer process on the surface of a solid, with all active sites having the same energy. DRK is a reliable method for calculating micropore volumes and equivalent surface areas [163, 164]. However, it provides a complement to N2 adsorption to characterize fresh and charred samples with a complex structure of micropores and macropores. CO2 (at ~ 0°C) diffuses faster than N2 (at ~ –196°C) because, at the higher temperatures CO2 gas molecules penetrate more deeply with higher kinetic energy and acidic properties through narrow pores. As a result, higher CO2 adsorption, coupled with low N2  28  adsorption, indicates a microporous structure adsorbent, while similar CO2 and N2 adsorption indicates a larger pore (meso and macropore1) structure [165–168].  In order to study the surface of samples, they were first coated by gold and palladium to make them more conductive, then examined by a Hitachi S-3000N Scanning Electron Microscope.  The XRD tests were performed and the results analyzed by Rozita Habibi, University of Calgary, and published in Energy & Fuels [150]. XRD powder patterns of the fresh, char, and ash samples of switchgrass, coal, and fluid coke were obtained on a Multiflex X-ray diffractometer (Rigaku, Woodlands, TX, USA) using Cu/Kα radiation (λ= 1.54056 Å), with a 40 kV tube voltage, a 20 mA tube current, and a scan rate of 2° min−1. Partially gasified samples of the single and mixed feeds were also analyzed [150]. 2.1.3 Experimental apparatus As shown in Figure  2-1, a Thermax500 high-pressure TGA, capable of temperatures up to 1100ºC and pressures to 100 bar, was used for char preparation and for the pyrolysis and gasification tests. Fuel samples were loaded into a small hemispherical quartz basket (17 mm ID and 20 mm in height) (Figure  2-2), connected to the load cell via a thin metal wire. A non-metal basket was chosen to minimize catalysis by the basket during the experiments.  Nitrogen (for char preparation and pyrolysis) and CO2 (for gasification) were introduced from the bottom of the metal chamber into the reactor. The reaction between CO2 and carbon (the Boudouard reaction) is important in biomass and fossil fuel gasification as CO2 is a product of steam gasification. The effectiveness of particular catalysts for CO2 gasification is generally the same as for steam gasification [17].    The output gas passed through a solvent tank (containing 100 ml of acetone) held in the ice bath and a Genie membrane separator (Model 101) to remove tar and moisture. The temperature inside the reactor was monitored by an immersed thermocouple, fixed immediately underneath at the bottom of the basket. The balance has a maximum allowable weight differential of 10 g                                                  1 A micro, meso and macro porous materials are materials containing pore with diameters <2 nm, 2-50 nm and >50 nm, respectively.   29  between the two balance arms and a sensitivity of 10 μg. Weight, temperature and pressure data were recorded on a computer using TherMax DAQ software.  Fresh samples were charred using a few mg of fuel and 500 mlN min-1 nitrogen flow (99.99% purity, Praxair), starting from room temperature to the target temperature (750, 800 or 900°C, typical in gasification industrial plants) with a 25°C/min heating rate. The target temperature was then held for 2 h. An intermediate heating rate was chosen to minimize systematic errors in temperature measurement due to thermal lag during pyrolysis [169] and to maximize char yield [170]. The pyrolysis heating rate also affects reactivity: Walker [171] showed that reactivity of chars produced by rapid heating of lignite is greater than that of chars produced by slow heating. This is due to surface annealing which reduces the number of nascent carbon sites, and thus the active surface of the char. The operating conditions (i.e. nitrogen flow rate, particle size, sample thickness), of this work were very similar to those of Thurner et al. [172, 173], appropriate for avoiding significant external heat and mass transfer limitations. In previous co-gasification studies [150], it has been very common to complete the pyrolysis stage, preparing fuel chars separately, then mixing the chars in different proportions to investigate the co-gasification stage afterward. However, since in real applications, raw materials are introduced to the reactor (and not fuel chars) and both pyrolysis and gasification stages take place in the same environment, in this work these two stages were carried out consecutively in the same reactor. Therefore, the raw materials were mixed and loaded into the basket. As mentioned above, the pyrolysis stage was performed for 2 h; then the reaction gas was switched from N2 to CO2, and gasification proceeded until the total weight of the basket and its contents reached a constant value. For either single or biomass/fossil fuel blend samples, the initial mass of loaded fuel was calculated based on having 15 mg char at the end of 2 h of pyrolysis. Figure  2-3 presents a typical TGA result based on the above procedure.  All tests were performed at atmospheric pressure. The pyrolysis experiments for coal, fluid coke, sawdust and switchgrass were replicated three times at each temperature to verify the reproducibility of the data. At a 95% confidence level, the measurements for each case were found to be statistically identical. The TGA load cell was calibrated and tared before any experiment using a 1000 mg standard reference weight.         NN2Figure  2-1. CO22SolenoidValveSolenoidValveSchematic oFigure  2FMFC4FMFC3FMFC2FMFC1f pressurized-2. Dimens TGA set-uions of quarPurge Gap (modifiedtz TGA basBalances from Butle ket.  P-13ComputerGMembrSample BasketIce Bathr [174]). Product gasenie ane FilterBack PresRegulatoVSolvent30  sure rent 31   Figure  2-3. Typical TGA result of two-stage gasification.  2.2  RESULTS AND DISCUSSION 2.2.1 Characterization of fresh samples The ultimate, proximate and ash analyses of each fresh fuel sample are provided in Table  2-1. The sub-bituminous coal contains a considerable proportion of ash (30.5 wt%). As expected, the biomass samples contained much higher proportions of volatiles (87.6 wt% for SD and 76.9 wt% for SG) than the fossil fuels (31.3 wt% for coal and 6.9 wt% for FC). The biomass samples mainly consist of carbon, hydrogen, oxygen and nitrogen, showing their organic nature. Switchgrass contains a higher proportion of alkali metals than coal, whereas sawdust has less (wt. basis), enabling investigation of the catalytic effect of alkalis on coal and fluid coke gasification. Although the sawdust contains the highest calcium content in its ash (14 wt%) which can catalyze gasification, its catalytic effect may not be significant because of the very low ash weight proportion (only 0.4%). The switchgrass has a higher ash content (6.3 wt%) and time, s0 5000 10000 15000Weight, mg0102030405060Temperature, °C0200400600800CO2 gasification Pyrolysis with N2 15mg  32  is rich in potassium (16.8 wt% in its ash), and is expected to have the greatest catalytic effect on gasification. As discussed in  Chapter 1, the inorganic species in fossil fuels are partly responsible for their relative reactivity. The silicon and aluminum contents of fuels are also important, as they can react with alkali metals in blended biomass/coal samples to form crystals, e.g. potassium aluminosilicates, which inhibit gasification [62, 150]. The fluid coke has the highest carbon content (83.7 wt%) and a low ash content (2.0 wt%), and therefore a higher heating value (32.4 MJ kg-1) than the other samples. The analyses are generally consistent with previously reported data (e.g. [175–177]). Table  2-1. Ultimate, proximate and ash analysis of fresh feedstocks. Sample Sub-bituminous coal Fluid coke (FC) Sawdust (SD) Switchgrass (SG)  Ultimate analysis (wt%), daf* Carbon, C  73.1   83.7  50.1   47.9 Hydrogen, H 4.3 1.9 6.2 6.2 Nitrogen, N 1.0 2.2   0.02 0.8 Sulfur, S 0.4 7.5   0.04 0.1 Oxygen, O (diff**)   21.2 4.8   43.7   45.0  Proximate analysis (wt%) Moisture 17.5 1.3 8.0 6.0 Ash (db*) 30.5 2.0 0.4 6.3 Volatile (db) 31.3 6.9   87.6   76.9 Fixed Carbon (db) 38.3   91.2   12.0   16.8 Higher heating value (MJ/kg) 20.1   32.4   20.4   19.6  Ash analysis (wt%) Si   26.9   16.2   21.1   24.6 Al   12.5   13.2 7.0 1.1 Ti   0.32 2.4 0.2   0.01 Fe 2.0 7.0 1.6 0.2 Ca 4.0 3.8   14.0 4.5 Mg 0.8 1.4  2.8 3.9 Na 1.9 1.6  1.0 1.2 K 0.6 1.2  5.8   16.8 P   0.04 0.2  0.6 2.2 S 0.9 0.8  0.8 1.0 Undetermined***   50.0   52.1    45.0   44.4 * daf = dry and ash free, db = dry basis; ** calculated by difference, *** Mainly oxygen, since the inorganic elements are present as oxides (e.g., Al2O3, SiO2, K2O).    33  The amounts of undetermined species in Table  2-1 were less than 3 wt% for all samples, except for sub-bituminous coal which had ~ 15.2 wt% (50.0% of 30.5% db ash) of undetermined species, including vanadium, nickel, manganese, chromium, cadmium, zinc, bromine and copper [175]. The upper part of Table  2-2 shows the N2 and CO2 adsorption surface areas of fresh samples. It is clear that fresh biomass fuels are less porous (e.g. 0.6 m2 g-1 versus N2 and 35 m2 g-1 for CO2 adsorption of switchgrass) than coal (e.g. 12 m2 g-1 for N2, 151 m2 g-1 for CO2 adsorption). Furimsky [178] reported the same BET surface area for fluid coke as listed in Table  2-2 based on our measurements. The higher surface areas measured using CO2 than for N2 adsorption, suggest that the samples had microporous structures. More detailed analysis of isotherm plots and pore size distribution of these feedstocks can be found elsewhere [179].   Table  2-2. BET and DRK surface areas of fuel fresh and char samples produced at different temperatures. The char samples were prepared using 500 mlN min-1 nitrogen flow with a 25°C/min heating rate at atmospheric pressure. The target temperature was held for 2 h. All values in m2 g-1.  Fresh Samples  BET Surface area DRK Surface area Fresh switchgrass 0.60 35.01 Fresh sawdust 0.03 0.40 Fresh sub-bituminous coal 12.31 151.41 Fresh fluid coke 2.50 52.08 BET Surface Area of Char Samples  750°C  800°C 900°C Switchgrass char 6 5 4 Sawdust char 442 335 32 Sub-bituminous coal char 138 134 102 Fluid coke char 20 12 7 DRK Surface Area of Char Samples  750°C  800°C 900°C Switchgrass char 353 353 310 Sawdust char 575 574 565 Sub-bituminous coal char 246 243 238 Fluid coke char 243 62 20   34  2.2.2 Characterization of chars Figure  2-4 presents the effect of pyrolysis temperature on the char yield, determined by weighing the samples before and after pyrolysis. Fluid coke had the highest char yield (~ 92 wt%) among the samples, consistent with its low volatile content of 6.9 wt%, as shown in Table  2-1 and as reported elsewhere [175, 179]. The lowest char yields were for sawdust (<20 wt%). The char yields did not vary significantly with pyrolysis temperature, as observed previously [180].   Figure  2-4. Char yield of fuel samples as a function of pyrolysis temperature. The char samples were prepared using 500 mlN min-1 nitrogen flow with a 25°C/min heating rate at atmospheric pressure. The target temperature was held for 2 h.  Ultimate and ash analyses of char can be used to identify changes in chemical composition after pyrolysis. Figure  2-5 provides CHNS analyses for the chars produced at 750, 800, and 900°C. The char samples were prepared using 500 mlN min-1 nitrogen flow with a 25°C/min heating rate at atmospheric pressure. The target temperature was held for 2 h. Comparison with Table  2-1, indicates that the carbon content for most samples increased from the fresh sample to the char prepared at 750°C. As the pyrolysis temperature increased, the carbon content increased in the Char yield (wt.%)0 20 40 60 80 100Fluid cokeCoalSawdustSwitchgrass750°C800°C900°C 35  (a)  (b)  (c)  (d)   Figure  2-5. Ultimate analyses of carbonaceous chars provided at 750, 800 and 900°C: FC=fluid coke; SG=switchgrass; SD=sawdust. (a) carbon (b) hydrogen (c) nitrogen and (d) sulfur content.  FC Coal SG SDCarbon content, wt% daf020406080100750oC 800oC 900oCFC Coal SG SDHydrogen content, wt% daf0.00.51.01.52.0750oC800oC900oCFC Coal SG SDNitrogen content, wt% daf0.00.51.01.52.02.5750oC800oC900oCFC Coal SG SDSulfur content, wt% daf02468750oC800oC900oC 36  fluid coke char, but decreased in the char from sub-bituminous coal, switchgrass and sawdust. The hydrogen proportion decreased in the samples as they were pyrolyzed and as the pyrolysis temperature increased. Compounds containing more hydrogen tend to be more volatile. Nitrogen and sulfur mass fractions did not change much from the fresh samples to the product chars generated at all three temperatures. Figure  2-6 summarizes the contents of potassium and calcium, the elements most likely to have positive catalytic effects on pyrolysis and gasification. The potassium and calcium contents of fresh fluid coke and fluid coke char were almost unchanged after pyrolysis. For the coal, the calcium content decreased upon pyrolysis, but was independent of the pyrolysis temperature for the limited range covered. The potassium contents of switchgrass and sawdust decreased when the pyrolysis temperature increased from 750 to 900°C, possibly due to evaporation. Wigmans et al. [181] and Sams et al. [182] reported that potassium carbonate evaporated at >727°C. The calcium concentration in the sawdust ash increased from 14 wt% for the fresh sample to 21 wt% for the char produced at 900°C, possibly due to alkali metals (e.g. potassium) evaporating more quickly than calcium.  The second and third parts of Table  2-1 show BET and DRK surface areas of char produced at three temperatures. As expected, most char samples had higher surface areas than the corresponding fresh feed, indicating more porous surface structures after devolatilization. When the pyrolysis temperature was raised from 750 to 900°C, the surface areas for both N2 and CO2 adsorption decreased, possibly due to sintering at higher temperatures. The higher DRK surface area for fresh fluid coke than for fluid coke char prepared at 900°C (Table  2-1) reflects the collapse of micro-structures after pyrolysis at higher temperatures. This can be explained by softening of the fluid coke char produced at above 750°C resulting in closing of pores. Micrometer scale SEM images of coal, fluid coke and switchgrass samples are consistent with the BET results. As shown in Figure  2-7(a), the fresh coal had a rock-like appearance, with large cracks, whereas the coal char was more porous, with higher surface area, as shown in Figure  2-7(b). Note that the coal char particles did not have uniform morphology, and that the images in Figure  2-7 are representative of multiple areas surveyed.  37     (a)  (b)   Figure  2-6. Mineral matter content in ash of: (a) potassium; and (b) calcium.    FC Coal SG SDPotassium content, wt%024681012141618Fresh750oC 800oC 900oC FC Coal SG SDCalcium content, wt%0510152025Fresh750oC800oC900oC 38  As shown in Figure  2-8, the switchgrass had a fibre-shaped surface and was more amorphous than coal. However, after pyrolysis, the switchgrass char surface contained more holes and therefore had greater surface area than the fresh material. Also, during preparation of the char, biomass fibers combined with each other, forming a lump (Figure  2-8(c)). After pyrolysis at 800°C, fibrous structures were still evident, as reported by Cetin et al. [144]. Figure  2-9 shows images of fresh fluid coke and fluid coke char. The fluid coke had a rigid, smooth, non-porous surface. However, with the escape of the volatile components from the fluid coke structure, large cracks developed on the surface, as shown in Figure  2-9(b). This observation is consistent with the increase in BET and DRK surface areas at high temperatures (Table  2-2), and in good agreement with results of Furimsky [178]. Figure  2-10 presents SEM images of fresh sawdust and its char. These are very similar to the swithgrass images. The fresh sample had a fibrous structure. After devolatilization, more pores appeared on the surface, again consistent with the corresponding BET data. In the case of co-pyrolysis, char particles of biomass and coal can be easily identified via SEM. Figures 2-7(c) and 2-8(d) present SEM images of coal and biomass char surfaces of a 50:50 wt% sub-bituminous coal:switchgrass mixture, subjected to pyrolysis at 800°C. From comparisons of Figures 2-7(b) and 2-7(c), and of Figures 2-8(b) and 2-8(d), it appears that mixing the fuels did not change the surface structures of biomass and coal. Both types of the char particles still have fibrous and porous surface structures, respectively. An SEM image of 75:25 wt% sub-bituminous coal:switchgrass mixture subjected to pyrolysis at 800°C is presented in Figure  2-11. Coal and biomass char particles are readily distinguishable from each other. Solid particle-particle contacts did not change their surfaces significantly after pyrolysis.                (a) Figure  2-7. SEM images of (a)800°C after hea  fresh coal (b) cting at 25°C/mi200µm (b) oal char preparn heating rate aed at 800°C (c)nd atmospheric   50 ( char of 50:50 w pressure in 500µmc) t% coal:switch mlN min-1 nitrgrass mixture pogen flow.    50µ39   repared at m      (a) Figure  2-8. SEMmixture pr  images of (a) epared at 800°C fresh swithgras after heating a  300µm (b) (d) s, (b) and (c) swt 25°C/min heaithgrass char pting rate and at      50(c) repared at 800°mospheric pres  50µm µm C (d) char of 50sure in 500 mlN :50 wt% coal:s min-1 nitrogen   500µ40  witchgrass flow. m             (a) Figure  2-9. SEM images of (a) fresh fluid catmosoke (b) fluid copheric pressure  200µm (b) ke char prepare in 500 mlN min d at 800°C afte-1 nitrogen flowr heating at 25°. C/min heating r  200µm 41 ate and          (a) Figure  2-10. SEM images of (a) fresh sawatmosdust (b) sawduspheric pressure          200µm (b) t char prepared in 500 mlN min at 800°C after -1 nitrogen flowheating at 25°C. /min heating ra  100µm 42  te and   Figure  2at Figure  2-Rozita Hfrom the of switch16.4° andsuggestedmainly spotassium(PDF nostructure=22.5, 28SiO2, anmagnesiuclay mine0695), bu29.7, 36-11. SEM im 25°C/min h12 shows Xabibi of the fresh feedstgrass (Figur 22.4° beca by Kim emall peaks, and calci. 25-0626). s. The XRD.6, 29.2, 30d Ca4Si2O7m (see to Tral kaolinitt some unid.4, 39.9, 43age of 75:2eating rate RD results fUniversity ock, to the ce  2-12(a)) ime smaller at al. [183] , likely origum. The peIt is difficul pattern of .3, 33.1, an(CO3)2, as able  2-1). Te (i.e., Al2Sientified spe.6, 47.6, an5 wt% coal:and atmosphor fresh, chof Calgary [orrespondinllustrates thnd broaderand Keiluwinating froaks at 2θ = t, however,switchgrassd 43.6°. Sothe ash che parent c2O5(OH)4; Pcies as weld 48.8°, wswitchgrasseric pressurar and ash sa150]. It shog char and aat, by charri, indicating aeit et al. [1m the inhe29.1° and 3 to unambigash (SG asme of theseontained moal has peaDF no. 72-2l (Figure  2-hich could  mixture pree in 500 mlNmples of SGws how the sh producedng the samp decrease in84]. Switchrent minera0.6° could uously assih) also has  peaks coulainly silicoks consisten300) and qu12(b)). The correspond           500µm pared at 800 min-1 nitro, coal and crystalline s at 750°C. Tle, the carb crystallinitgrass char (l matter cobe attributegn peaks tosmall and sd corresponn, potassiut with the cartz (i.e., Scoal char h to calcium°C after hegen flow. FC performetructure chahe XRD paon peaks aty of celluloSG char) sntaining sild to K2Ca(C specific miharp peaks d to K2MgSm, calciumommonly fiO2; PDF noad peaks at  carbonate 43 ating d by nges ttern  2θ = se, as hows icon, O3)2 neral at 2θ i3O8,  and ound . 85-2θ = (i.e.,   CaCO3; Pcoal ash  Figure  2-produceKaolinDF no. 47-sample had p12. X-ray dd at 750°C fite are no. 81743) and aeaks that coiffraction paor (a) SG, (b5-0695 and t 2θ = 28.6°uld corresptterns of par) Coal, andno. 72-2300, which couond to quartent (fresh) s (c) FC. Pow, respectiveld be from z, CaAl2Si2amples andder diffractly. Performecalcium alumO8, and CaA  correspondiion files (PDd by Rozitainosilicatel2Si2O8·4H2ng chars anF) for SiO2 Habibi [1544 . The O.  d ash  and 0].  45  Silicon, aluminum, and calcium were identified in the ash analysis (refer to Table  2-1). Again, it is difficult to identify exact species due to the chemical complexity of the ash and overlapping XRD peaks. Parent fluid coke and fluid coke char had similar amorphous structures with broad carbon-related peaks at 2θ = 25° and 2θ = 44° (Figure  2-12(c)) [185]. The lack of additional peaks is consistent with the very low ash content (2 wt %) of this sample. The consistency in the XRD spectra was expected because the parent fluid coke already underwent a high temperature process. In contrast, the fluid coke ash sample had a variety of distinct peaks in the XRD profile representing the ash components of mainly quartz (i.e., SiO2; PDF no. 70-3317 at 2θ = 22.1, 28, and 39°), iron oxide (i.e., Fe3O4; PDF no. 26- 1136 at 2θ = 31.3, 36.9, 44.8, and 55.6°), and aluminosilicate (i.e., Al2SiO5; PDF no. 38-0471 at 2θ = 26, 40.9, and 60.9°). The presence of these compounds is consistent with the high concentration of aluminum, silicon and iron in fluid coke ash, as shown in Table  2-1 [150]. 2.2.3 Co-pyrolysis thermogravimetric analysis: Experimental  The pyrolysis kinetics of fossil fuel and mixed (25 wt%, 50 wt%, 75 wt% biomass char) samples were investigated after preparation by heating from room temperature to 800°C at a 25°C min-1 heating rate under a nitrogen environment (500 mlN min-1). To reach the target char mixture ratios and target weight (15 mg), the required amounts of fresh biomass and fossil fuel added to the reactor were estimated from the fresh fuel char yields (Figure  2-4). This procedure is only valid if the interaction of biomass and coal does not alter the char yield of the individual fuels, i.e. if each component acts independently. Figure  2-13 shows the linear behaviour between the weight percentage of fresh biomass in the basket and the char yield. This shows that in co-pyrolysis of the samples, the char yields of biomass and non-biomass fuels were independent of the blend mixture ratio, consistent with the results of Xu et al. [186].  Figures 2-14, 2-16(a) and 2-17(a) show the pyrolysis (in nitrogen) weight loss behaviour of pure coal, fluid coke, switchgrass, sawdust and their mixtures vs. time. One experiment was also performed with a 50:50 wt% mixture of coal and switchgrass ash. Three stages can be distinguished, as shown in Figure  2-14(b): (Ι) a drying stage under 200°C, where moisture evaporated from the fuel samples; (ΙΙ) a stage where volatiles were removed, the thermal decomposition of cellulose and hemicelluloses occurred, and the rate of devolatilization   46   Figure  2-13. Effect of relative proportions of fossil fuel/biomass binary mixture on char yield of different fuels prepared at 800°C after heating at 25°C min-1 in N2. Symbols are experimental data, and the solid lines are predicted if separate fuels acted independently.  reached its peak; and (III) a later stage where loss of heavier hydrocarbons and lignin occurred and the rate of reaction decreased with increasing temperature. Stage (II) can be distinguished from stage (III) by noting the time at which the absolute slope of the weight loss curves started to decrease. These three stages were identified in several previous studies [187–189]. During the pyrolysis, as expected, the pure switchgrass and sawdust lost a higher percentage of their original weight and lost weight more quickly than the fossil fuel samples, due to their greater volatile contents. The figures also show that for biomass samples, most devolatilization occurred in the second stage (~ 200 to 400°C), whereas fossil fuels required a higher temperature (>475°C) to lose volatiles in the third stage. Figures 2-15, 2-16(b) and 2-17(b) show the rates of devolatilization of pure and mixture samples, based on an ash-free mass (m'i), as a function of temperature. Three pyrolysis stages can again be observed in these plots. The peak rate of devolatilization of the pure biomass samples occurred Fresh biomass wt% in blend0 20 40 60 80 100Ultimate char yield %020406080100FC/SG BlendCOAL/SD BlendCOAL/SG Blend 47  below 400°C, whereas for coal the peak was around 470°C. The difference is due to the different biomass and coal physical and chemical properties, as described above. No significant peak can be seen for the fluid coke as it contained little volatiles (see Table  2-1). For the mixture samples, two peaks, one for devolatilization of the biomass portion of the mixture and a smaller one for the coal portion, can be seen. It appears that the biomass and fossil fuels pyrolyzed independently, with no significant interaction between them. Even for the 50:50 wt% coal:SG ash mixture (Figure  2-15), where the potassium concentration of the blend sample is the highest, no significant difference in the behaviour can be observed relative to pure coal.   Experimental results were compared with the calculated results if separate fuels acted independently, i.e. compared with the weighted values, (1 )calc biomass non biomass biomass biomassm f m f m−= − +  ( 2-4) where mnon-biomass and mbiomass are the masses of the original components, with fbiomass being the mass fraction of the biomass char in the mixture. Hence, mcalc is the “expected” char mass, if the components acted independently.  The parity plot in Figure  2-18 shows an experimental arithmetic means of the mixture devolatilization rate data (over 200°C) from Figures 2-15, 2-16(b) and 2-17(b) versus the corresponding values from equation ( 2-4). All mixture sample data lie close to the 45 degree parity line, revealing that, on average, the experimental mixture rate of devolatilization was close to that which would be observed if there were no interaction, i.e. no enhancement or synergy in the mass-loss pattern of co-feed pyrolysis of coal and fluid coke with switchgrass and sawdust. It should be pointed out that the flow of N2 in a TGA seemed to sweep the volatile products away from the devolatilizing fuels, and therefore, the devolatilizing gas could not easily interact with the char to enhance co-pyrolysis. The physical proximity of alkali/alkaline earth metals containing functional groups in biomass relative to fossil fuel samples might also be of great importance, and, therefore, it should be recognized that enhancement in co-pyrolysis kinetics might have been observed using smaller fuel particle sizes [143].       48      Figure  2-14. (a) Weight and temperature versus time during pyrolysis of coal, switchgrass, coal/SG mixtures and coal/SG ash mixture in nitrogen (b) Isolated plot of Coal:SG=25:75 wt.% showing three pyrolysis stages.     Time, t, min0 10 20 30 40 50 60% of Original Weight, %mi/m0020406080100Temperature, °C0200400600800Coal:SG ash = 50:50Coal:SG = 100:075:2550:5025:750:100(a)Time, t, min0 10 20 30 40 50% of Original Weight, %mi/m0020406080100Temperature, °C0200400600800Coal:SG = 25:75Stage    (I) Stage     (II)Stage (III)(b) 49     Figure  2-15. Rate of devolatilization of coal, switchgrass, coal/SG mixtures of different proportions and coal/SG ash mixture in nitrogen. Heating rate at 25°C min-1 in N2.         Temperature, °C200 400 600 800-d(m' i/m0,dry)/dt, min-10.000.050.100.150.20Coal:SG=100:0Coal:SG=0:100Coal:SG=50:50Coal:SG=75:25Coal:SG=25:75Coal:SG ash=50:50 50  Figure  2-16. (a) Weight and temperature versus time during pyrolysis of coal, sawdust and coal/SD mixtures in nitrogen (b) Rate of devolatilization versus temperature of fresh coal, sawdust and coal/SD mixtures in nitrogen at 25°C min-1 heating rate.         Time, t, min0 10 20 30 40 50 60% of Original Weight, %mi/m0020406080100Temperature, °C0200400600800Coal:SD = 100:075:2550:5025:750:100(a)Temperature, °C200 400 600 800-d(m'i/m0,dry)/dt, min-10.000.050.100.150.200.25Coal:SD=100:0Coal:SD=0:100Coal:SD=50:50Coal:SD=75:25Coal:SD=25:75(b) 51     Figure  2-17. (a) Weight and temperature versus time during pyrolysis of fluid coke, switchgrass and FC/SG mixtures in nitrogen (b) Rate of devolatilization versus temperature of fluid coke, switchgrass and FC/SG mixtures in nitrogen at 25°C min-1 heating rate.    Time, t, min0 10 20 30 40 50 60% of Original Weight, %mi/m0020406080100Temperature, °C0200400600800Fluid Coke:SG = 100:075:2550:500:100(a)Temperature, °C200 400 600 800-d(m'i/m0,dry)/dt, min-10.000.050.100.150.20FC:SG=100:0FC:SG=0:100FC:SG=50:50FC:SG=75:25(b) 52   Figure  2-18. Experimental arithmetic mean of mixture devolatilization rate data (over 200°C) presented in Figures 2-15, 2-16(b) and 2-17(b) versus corresponding linear combination calculated ones based on equation ( 2-4).  2.2.4 Co-pyrolysis thermogravimetric analysis: Modeling  The integral method of Coats and Redfern [190] has been widely used for kinetic analysis of solids decomposition. Based on this model, the kinetic equation can be written as  ( ) ( )dk T fdtα α= ( 2-5) where the f(α) function depends on the reaction mechanism. The rate constant can be defined as a function of temperature using the Arrhenius equation. Here the conversion α is defined as Calculated (-d(m'i/m0,dry)/dt)mean, min-10.010 0.015 0.020 0.025 0.030 0.035 0.040Experimental (-d(m' i/m0,dry)/dt) mean, min-10.0100.0150.0200.0250.0300.0350.040Coal:SG=25:75Coal:SG=50:50Coal:SG=75:25Coal:SD=25:75Coal:SD=50:50Coal:SD=75:25FC:SG=50:50FC:SG=75:25 53  0 ii fm mm mα −= −  ( 2-6) where m0, mi, and mf are the initial, actual and final mass of the solid sample, respectively. Assuming no temperature gradient along the sample and instantaneous rise of the sample temperature when heating in the TGA, for a constant heating rate of 25°C min-1  dTdtq= ( 2-7) where q≅ 0.42 K s-1. After substitution of the Arrhenius equation and equation ( 2-7) into equation ( 2-5):  0 0( ) exp( )( )Td A Eg dTf q RTα αα α= = −∫ ∫  ( 2-8) where A and E are the pre-exponential factor and activation energy, respectively.  There is no exact analytical solution for the right hand of equation ( 2-8). Cauchy’s rule was used to estimate the integral: 202exp( ) (1 )exp( )TA E ART RT EdTq RT qE E RT− ≅ − −∫  ( 2-9) After some rearrangement, ignoring the temperature dependence of E, and taking logarithms, we can rewrite equations ( 2-8) and ( 2-9) as  2( ) 2ln ln (1 )g AR RT ET qE E RTα ⎛ ⎞⎛ ⎞ = − −⎜ ⎟⎜ ⎟⎝ ⎠ ⎝ ⎠  ( 2-10) After choosing an appropriate reaction mechanism ((f(α) or g(α)), a plot of ln(g(α)/T2) versus 1/T gives a straight line with a slope of –E/R. Several reaction mechanisms summarized by Vlaev et al. [187] were examined, and the best-fit mechanism to TGA data (based on R2) for each pyrolysis stage was selected for non-linear parameter estimation to obtain the activation energy. Trust-Region non-linear minimization, Levenberg-Marquardt, and Gauss-Newton algorithms with 95% confidence interval were used for non-linear regression. The reaction mechanisms examined are in four classes of chemical reaction (first, second, and third order), random nucleation and nuclei growth (bi-dimensional and three-dimensional), phase boundary reaction  54  (one, two, and three dimensional), and diffusion (one- two- and three-way transport, Ginstling-Brounshtein model, and Zhuravlev model) [191].  The best correlations for different pyrolysis stages were obtained with five reaction mechanisms, as shown in Table  2-3. Table  2-4 presents the activation energy and reaction mechanism of different pyrolysis stages for fresh and mixture samples of coal, fluid coke, switchgrass and sawdust. For the first pyrolysis stage below 200°C, the Zhuravlev diffusion model [192] gave the best fit, implying that it corresponds to the reaction mechanism to evaluate the activation energies of all fresh and mixture samples. For the second stage, where the rate of devolatilization reached its peak, all of the reactions of the samples with biomass are best fitted by the first order chemical reactions, whereas fresh coal and coal:SG ash samples reactions are well fitted by the diffusion model. Walker [171] and Mahajan and Walker [193] reported diffusion limitations during coal thermo-chemical conversion. The best model to evaluate the pyrolysis of fresh fluid coke in stage II was three-dimensional random nucleation and nuclei growth. This can be explained based on Table  2-1 and Figures 2-14 to 2-18, where the fossil fuels had higher fixed carbon and were less reactive in the second stage of pyrolysis than biomasses with high volatile content. Previous studies [173, 189, 194] also applied a first-order chemical reaction to model this stage of biomass pyrolysis. Based on Table  2-4, for all investigated samples, the rate of reaction of the third stage was controlled by diffusion, presumably of species leaving the particles. This is the stage where the reaction rate decreased with increasing temperature and heavier hydrocarbons decomposed. Decomposition of solids is typically controlled by diffusion when large heat changes are involved [189, 191]. The relatively high activation energies for stage III may reflect to some extent dependence of pore structure on temperature. This implies that for the third stage, higher temperature is required to promote the forward pyrolysis reaction. The conclusions of Table  2-4 are in good agreement with those of Lu et al. [188] and Min et al. [189] who characterized biomass pyrolysis (with a heating rate close to that in this work) by three successive stages controlled by internal diffusion, chemical reaction, and internal diffusion.  In general, reactions with higher activation energies are more temperature-sensitive [195]. In all three stages, the activation energies of coal:SG ash samples are higher than for fresh coal samples, consistent with the gasification data for the same samples [179]. As shown in Table  2-4,  55  for all three stages, the activation energies of mixed biomass and fossil fuel samples lie between the corresponding activation energies of the separate feedstocks.  The non-linear integral isoconversional method [196, 197] can also be used to estimate the activation energies. This method is able to predict the evolution of activation energy vs. conversion, without determining the reaction mechanism. The general trend of increase in activation energies in the three stages is in good agreement with this method.  In order to test the proposed model, the experimental and modeling conversions are compared for pure coal, switchgrass and a 50:50 wt% mixture of these feeds in Figure  2-19. For each sample, three different pyrolysis stages are indicated by the vertical lines. The general trend predicted by the model is in acceptable agreement with experimental data.   Figure  2-19. Comparison of experimental coal and SG pyrolysis data (symbols) and model (solid lines) from room temperature to 800°C at a heating rate of 25°C min-1 and atmospheric pressure in nitrogen. Temperature, °C100 200 300 400 500 600 700 800Conversion, α0.00.20.40.60.81.0Pure SGPure Coal50:50 Mixture 56      Table  2-3. Best fitting mechanism functions with TGA data of this work for modeling three stages of pyrolysis process. Mechanism Symbol f(α) g(α) Ref. Chemical reaction (first order) A 1- α -ln(1- α)   [191]  Random nucleation and nuclei growth  (three dimensional)  B  3(1- α)[-ln(1- α)]2/3  [-ln(1- α)]1/3  [198]  Diffusion (one-way transport, plane)  C  1/(2α)  α2  [199]  Diffusion (three-way transport, spherical)  D  (2/3)(1- α)2/3/1-(1- α)1/3  [1-(1- α)1/3]2  [200]  Diffusion (Zhuravlev equation)    E  (2/3)(1- α)5/3/1-(1- α)1/3  [(1- α)-1/3-1]2  [192]          57    Table  2-4. Activation energies and reaction mechanisms for different stages of pyrolysis for fresh and mixture samples of coal, fluid coke, switchgrass and sawdust. For mechanisms A to E, see Table  2-3.  Stage I (T<200°C) Stage II (200°C <T<Tpeak) Stage III (Tpeak<T<800°C) Sample E1/R (K) Mechanism E2/R (K) Mechanism E3/R (K) Mechanism Coal:SG=100:0 1,564 E 2,289 E 5,164 D Coal:SG=75:25 1,708 E 2,707 A 5,040 E Coal:SG=50:50 1,946 E 3,927 A 4,803 E Coal:SG=25:75 1,956 E 5,042 A 4,568 E Coal:SG=0:100 2,366 E 6,353 A 4,811 E Coal:SG ash=0:100 1,916 E 3,178 E 5,793 D Coal:SD=75:25 1,665 E 3,964 A 6,218 E Coal:SD=50:50 1,701 E 5,047 A 6,286 E Coal:SD=25:75 1,763 E 6,184 A 6,351 E Coal:SD=0:100 1,829 E 6,301 A 6,449 E FC:SG=100:0 1,443 E 1,376 B 3,148 C FC:SG=75:25 1,696 E 3,729 A 4,045 E FC:SG=50:50 1,881 E 5,463 A 4,572 E   58  2.2.5 Co-gasification thermogravimetric analysis The kinetics of CO2 gasification of coal, SG, SD, FC and their char mixtures were investigated. As explained in section  2.1.3, in order to reach the target char mixture ratios to begin the gasification, the required amounts of raw biomass and coal fed to the reactor were calculated based on the raw fuels char yield. The initial mass for all gasification experiments was 15 mg. 2.2.5.1 Coal-switchgrass co-gasification The kinetics of gasification of coal, SG and coal/SG char and ash mixtures (25 wt%, 50 wt%, 75 wt% biomass char and 50 wt% biomass ash) were investigated at different temperatures using 500 mlN min-1 flow of CO2 (99.99% Praxair). The temperature differential in the sample bed due to endothermic gasification reactions was estimated and found to be minor. Also, the mass transfer calculations showed that the internal and external resistances to mass transfer were negligible. Appendix A provides the heat and mass transfer formulations and calculations.  Figure  2-20 shows the CO2 gasification conversion vs. time and rate conversion for pure coal, SG and coal/SG mixtures at 800°C. As expected, the char gasification process is generally much slower than the initial pyrolysis. Thus it is dominant in the overall gasification process [201]. Figure  2-20 also demonstrates that pure biomass gasification is faster than corresponding coal gasification. The biomass needed less than 300 min to be fully converted, whereas 100% coal required more than 600 min. For the biomass char, the fine voids can allow more uniform gas mass transfer throughout the gasification process, helping maintenance of high gasification reactivity, increasing with elapsed time, consistent with the CO2 adsorption results (Table  2-2). For the coal char, large cracks made it possible for gas to be transferred rapidly in the early stages of the gasification. However, the compacted clusters caused high resistance for the gas transfer from the compacted clusters to the large cracks [186]. Unexpectedly, adding 25 wt% and 50 wt% switchgrass to coal had a negative effect on the gasification rate. As shown in Figure  2-20, in the initial stage, both the 25% and 50% biomass samples had greater conversion than the pure coal sample. However, from the middle stage, after about 50% conversion when the biomass was almost fully gasified, the gasification rates declined. On the other hand, when the biomass char was dominant in the mixture (75 wt % SG), the rate of conversion was faster than for coal alone.  59  The calculated conversions are weighted-average conversions of the char mixture and can be estimated based on the experimental results for the separately-reacted feedstocks, similar to equation ( 2-4), as follows: (1 )calc biomass non biomass biomass biomassf fα α α−= − +  ( 2-11) where αcalc, αnon-biomass and αbiomass are the conversions of the mixture and separate chars, whereas fbiomass is the mass fraction of the biomass char in the mixture (f =0.75, 0.5 and 0.25). Here, the conversion of the char mixture is assumed to be a linear function of each single char conversion. If the measured conversion of the char mixture is higher than the weighted-average conversion, enhancement can be assumed. A smaller measured conversion, on the other hand, implies inhibition. In all sample mixtures presented in Figure  2-20, the experimental conversion was not equal to, or higher than, the linearly predicted calculated values. Therefore not only was there no significant positive catalytic effect observed, but inhibition of the gasification appeared in two of the mixtures. In another experiment, switchgrass ash was prepared in the oven at 800°C by passing air through the sample for 2 h. Then a 50:50 wt% coal char: SG ash mixture was prepared for gasification. Figure  2-20 shows that the presence of a considerable proportion of SG ash enhanced the coal gasification rate significantly seemingly due to increase in AAEM-carbon interactions. The reaction rate is even higher than for SG alone up to ~ 80% conversion. This is presumably due to the high concentration of potassium in the fuel mixture. Therefore the switchgrass ash in this case displayed impressive catalytic activity when mixed with the coal char. Effect of temperature In order to investigate the effect of operating temperature on the reaction kinetics, experiments were carried out at both 750 and 900°C after heating up at 25°C/min. At 700°C, due to low activation energy no significant mass change was observed for either biomass or coal. Based on Figures 2-20, 2-21 and 2-22, higher temperature led as expected to faster reaction and full conversion. The general trend at 750°C was very similar to that at 800°C with the difference being that for the blend of coal and SG at 750°C, the reaction kinetics were slower from the beginning of the gasification, as shown in Figure  2-21. From Figure  2-22, at 900°C, the 50:50   60  (a)  (b)  Figure  2-20. (a) CO2 gasification conversion vs. time, and (b) rate vs. conversion for coal, switchgrass and coal/SG mixtures at 800°C and 1 atm. All proportions are in wt%. Time, t, min0 200 400 600 800 1000 1200Conversion, α 0.00.20.40.60.81.0Coal:SG = 100:0Coal:SG = 0:100Coal:SG = 50:50Coal:SG = 75:25Coal:SG = 25:75Coal:SG ash = 50:50Conversion, α 0.0 0.2 0.4 0.6 0.8 1.0Rate, dα/dt, min-10.0000.0020.0040.0060.0080.0100.0120.0140.016 61  (a)  (b)  Figure  2-21. (a) CO2 gasification conversion vs. time, and (b) rate vs. conversion for coal, switchgrass and coal/SG mixture at 750°C and 1 atm. All proportions are in wt%. Time, t, min0 200 400 600 800 1000 1200 1400Conversion, α 0.00.20.40.60.81.0Coal:SG = 100:0Coal:SG = 0:100Coal:SG = 50:50Conversion, α 0.0 0.2 0.4 0.6 0.8 1.0Rate, dα/dt, min-10.0000.0010.0020.0030.0040.005 62  coal:SG result falls between that for the coal alone and biomass conversion curves. Also, although similar to the behaviour at 800°C, the 75:25 coal:SG curve lies below that of pure coal, but the conversion is closer to the 100% coal results and approaches 100% faster than at 800°C.  As shown in Figure  2-22, adding SG ash to coal can increase the gasification reactivity dramatically at 900°C. However, it is apparent that the catalytic action is not as effective as at 800°C. This may be due to potassium evaporation at the higher temperature and/or to sintering. In Figure  2-6, the decrease in potassium concentration in passing from 750 to 800 and 900°C suggests some evaporation of catalyst at the higher temperatures.  The reactivity of char depends not only on the type of fuel, but also on the conditions under which it was prepared. Charring may take place prior to, or concurrent with, conversion at temperatures up to those of gasification, or the char may be prepared at higher temperatures than are subsequently used in gasification. To clarify this issue, another experiment was designed to gasify the 50:50 coal:SG char blend at 800°C which had previously been pyrolyzed at 900°C. Figure  2-23 shows that the gasification of this sample was much slower that for the same blend pyrolyzed and then gasified at 800°C. It even had less reactivity than the 75:25 coal:SG sample. This is thought to be because in the pyrolysis stage at 900°C, some of the alkali metals evaporated and were lost (Figure  2-6), and fuel pore structures were collapsed via sintering (Table  2-2). Kosky et al. [202], who studied the effect of char preparation heat treatment temperature on gasification reactivity of the resultant char, obtained results in agreement with ours. Their results were interpreted as being caused by annealing of active sites at high temperatures. Also Radovic et al. [203] showed how higher pyrolysis temperatures of lignite caused the reactivity of chars to decline. Discussion As explained in  Chapter 1, alkali, alkaline earth, and transition metals are effective catalysts for coal gasification. Ash analysis of the four fuels tested in this project (Table  2-1) revealed that switchgrass is rich in potassium and calcium. However, it was found, unexpectedly, that the presence of switchgrass inhibited coal gasification. On the other hand, switchgrass ash had a significant positive effect on coal gasification. The main reason for the former could be related to the coal ash content.  63   (a)  (b)  Figure  2-22. (a) CO2 gasification conversion vs. time and (b) rate vs. conversion for coal, switchgrass and coal/SG mixtures at 900°C and 1 atm. All proportions are in wt%. Time, t, min0 100 200 300 400Conversion, α 0.00.20.40.60.81.0Coal:SG = 100:0Coal:SG = 0:100Coal:SG = 50:50Coal:SG = 75:25Coal:SG ash = 50:50Conversion, α 0.0 0.2 0.4 0.6 0.8 1.0Rate, dα/dt, min-10.000.020.040.060.08 64  (a)  (b)  Figure  2-23. (a) CO2 gasification conversion vs. time and (b) rate vs. conversion for coal, switchgrass and coal/SG mixtures at 800°C and 1 atm. All proportions are in wt%. Time, t, min0 200 400 600 800 1000 1200 1400 1600Conversion, α 0.00.20.40.60.81.0Coal:SG = 100:0Coal:SG = 0:100Coal:SG = 50:50Coal:SG = 75:25Coal:SG = 50:50 (Pyrolysis at 900°C)Conversion, α 0.0 0.2 0.4 0.6 0.8 1.0Rate, dα/dt, min-10.0000.0020.0040.0060.0080.0100.0120.014 65  Previous work has shown that the catalytic activity can be lost by at least four different mechanisms: (a) vaporization of potassium [181, 182]; (b) irreversible reactions between potassium and ash components in the char; (c) formation of inactive alkali metal carbonate species; and (d) diffusion from the reaction surface into the carbon matrix [204] and intercalation [181]. Wood et al. [205] showed that only water-soluble potassium is catalytically active in gasification.  The catalytic effects of potassium compounds such as K2CO3 and KOH in coal gasification have been recognized for many years [31, 206]. However, these compounds can undergo secondary reactions with mineral matter in coal [62]. The potassium can react with coal minerals, like illite and kaolinite, to form a new mineral phase, e.g. kaliophilite (KAlSiO4). Formella et al. [207] and Chin et al. [208] reported that mineral matter in bituminous coal interacted with potassium during gasification. They examined the gasification of treated (acid-washed to remove coal minerals) and untreated chars with different loads of catalyst. It was found that the reactivity of the treated char was nearly twice that of the untreated. Moreover, the reactivity of the latter decreased with higher burn-off, probably as a result of continuous deactivation of the catalyst. The mineral kaolinite reacted with potassium carbonate to kaliophilite according to: 2 3 2 2 2 3 4 2 2Al O 2SiO 2H O+K CO 2KAlSiO +2H O+CO⋅ ⋅ →  ( 2-12) The authors concluded that for potassium-catalyzed gasification, not only is the coal ash total proportion important, but also its composition. The higher the fraction of clay minerals (mainly Al and Si), the larger the amount of catalyst that must be added to increase the reactivity.   Bruno et al. [209] also studied the reactivity of potassium carbonate with some mineral constituents of coal. Then eight different coals doped with K2CO3 were steam-gasified at 973 K and 3.6 MPa to find a correlation between potassium losses and mineral matter composition. Examination of the behaviour of eight carbonaceous matrices showed that each Al atom in the clay blocks a potassium atom, inhibiting the potassium catalytic effect. Kühn and Plogmann [210] and McCoy et al. [36] analyzed different crystalline phases formed during catalytic gasification of coal with potassium by means of microscopy, XRD, XRF and  66  microprobe analyses. They reported the existence of Osumilite (K2O4.4MgO4.6Al2O3.20.4SiO2), Sanidine (K2O.Al2O3.6SiO2), Leucite (K2O.Al2O3.4SiO2), Kaliophilite (K2O.Al2O3.2SiO2), Nepheline (Na2O.Al2O3.2SiO2) and Gehlenite (2CaO.Al2O3.SiO2) as gasification reaction products. Therefore, although potassium is a good oxidative gasification catalyst, it can interact disadvantageously with inherent mineral matter in high ash coals.   Radovic et al. [211, 212] reported that deactivation of catalytic potassium is brought about by interaction with inherent aluminosilicates. Kim et al. [213] investigated the effect of K2CO3 loading and potassium loss on the composition of product gas in K2CO3-catalyzed steam gasification of a coal char. In an attempt to discover the major mechanism responsible for the catalyst loss, they found that the proportion of potassium lost by mechanisms other than reactions of potassium with coal ash is less than 5 wt% of the total potassium added. This indicates that the potassium loss is mainly attributable to irreversible reactions of potassium with ash, forming water-insoluble, catalytically inactive potassium compounds. McKee et al. [35] studied the effect of alkali metal carbonates on coal chars using steam and CO2. They speculated that loss in the reactivity of gasification during thermal cycling is due to reaction of the alkali metals with mineral matter in the char to form stable inert alkali silicates or aluminosilicates [35]. They also reported that although the highest rank coal, anthracite, had lower reactivity compared to other coals, its reactivity with alkali metals (treated) was almost the same as for other treated samples. Therefore it can be speculated that the ash content of low rank coals might react with alkali salts, decreasing the gasification reactivity.  Table  2-5 shows the SG and coal char elemental amounts of C, K, Si, and Al produced at 750, 800 and 900°C (in mmol.g-1). Based on Table  2-5, the initial potassium/carbon (K/C), potassium/aluminum (K/Al), and potassium/silicon (K/Si) molar ratios for different SG and coal samples at 750, 800, and 900°C are presented in Table  2-6. It is seen that adding more biomass to coal caused the K/C, K/Al, and K/Si ratios to increase. For low ratios (up to 50:50 coal:SG), it seems probable that the coal ash (Al and Si)  sequestered the biomass potassium needed for Kalsilite (KAlSiO4) formation, and, thus, no catalytic effect was observed until the biomass-to-coal ratio reached 3:1, i.e. until the SG ash supplied enough potassium to satisfy the minerals in the coal ash. Only unreacted potassium can act as a catalyst for coal gasification. As explained in   67   Table  2-5. Elemental amounts of C, K, Si, and Al in SG and coal chars produced at 750, 800 and 900°C (in mmol g-1).  SG Coal Element  750°C 800°C 900°C 750°C 800°C 900°C C  58.84 57.11 54.55 43.06 38.10 36.53 K  0.98 0.90 0.78 0.07 0.07 0.07 Si  2.40 2.40 2.40 4.57 4.57 4.57 Al  0.11 0.11 0.11 2.21 2.21 2.21   Table  2-6. Potassium/carbon, potassium/aluminum, and potassium/silicon molar ratios of SG and coal samples at 750, 800 and 900°C (calculated from Table  2-5).   Sample K/C K/Al K/Si 750°C 800°C 900°C 750°C 800°C 900°C 750°C 800°C 900°C Coal:SG = 0:100 0.017 0.016 0.014 8.782 8.092 6.963 0.409 0.377 0.324 Coal:SG = 25:75 0.014 0.013 0.012 1.184 1.093 0.945 0.256 0.237 0.204 Coal:SG = 50:50 0.010 0.010 0.009 0.454 0.420 0.366 0.151 0.140 0.122 Coal:SG = 75:25 0.006 0.007 0.006 0.178 0.166 0.148 0.074 0.070 0.062 Coal:SG = 100:0 0.002 0.002 0.002 0.033 0.033 0.033 0.016 0.016 0.016 Coal:SG ash = 50:50 0.085 0.089 0.080 1.397 1.289 1.113 0.274 0.253 0.219    68  section  2.1.1, coal minerals can increase the reactivity of coal, and when they react with biomass potassium, the potassium is deactivated so that, inhibition was observed for some cases.   To analyze the formation of species, ~ 2 g samples of different compositions were partially gasified with CO2 at 750°C to 50% conversion in a fixed-bed reactor. Their XRD profiles (obtained by Rozita Habibi, University of Calgary) are shown in Figure  2-24. The XRD pattern of 50:50 SG/coal (pattern b) has a peak corresponding to KAlSi3O8 at 2θ = 29.9°. This species was not observed in the partially gasified 100% coal (pattern a), because the amount of potassium in the coal was too low (Table  2-1). The peaks at 2θ = 21.3, 27.1, 37.0, 40, 42.9, and 50.6° correspond to SiO2 and were observed in both the partially gasified 100% coal (pattern a) and the partially gasified 50:50 SG/coal samples (pattern b). The peak at 2θ = 20.4° was visible in pattern a (100% coal), and at a lower intensity in pattern b (50:50 SG:coal), whereas the peak at 2θ = 28.2° disappeared in the profile of SG/Coal (Figure  2-24, pattern b). These two peaks could not be identified, but they might be related to alumina components, because the amount of alumina is less in the char mixture than in the coal (i.e., 1.16 vs. 2.21 mmol g−1). The XRD pattern of the partially gasified 50:50 switchgrass ash mixed with coal char (SG ash/Coal) had a larger peak at 2θ = 29.9° (pattern c), which may correspond to KAlSi3O8, consistent with the higher potassium concentration in the SG ash compared with SG char. The potassium-to-alumina molar ratio in the SG ash/coal sample was greater than 1 (K/Al = 1.397, Table  2-6). Thus, not all potassium could bind to Al and Si and form KAlSi3O8, and some of the peaks, 2θ = 22.3, 28.5, 29.1, and 43.5°, could correspond to excess potassium and/or calcium (Figure  2-24, pattern c). These peaks were not in the XRD profile for coal, but were visible in the profile for 100% SG ash (Figure  2-24, pattern d), which contained ~ 16 wt % K2O and 5 wt % CaO [150]. Co-gasification of acid-washed coal and switchgrass It is shown above that the ash content of coal has a negative effect on co-gasification of biomass and coal due to deactivation of the potassium originally presented in the biomass ash. Therefore the ash content of coal was reduced by washing the coal by agitation with five different types of solutions - distilled water, dilute hydrochloric acid, dilute acetic acid, dilute nitric acid, and dilute sulfuric acid (25 ml each). To remove inorganic components from coal, the sample was soaked and submerged in water or acids under different leaching conditions. After leaching, the   solutionsair at 105 FiguconversioPow Water orTemperaLeachingAcid con were filtere°C over 24 re  2-24. X-n) (a) Coal,der diffrac solution to ture, °C   time, h centration, d and washh. Table  2-7ray diffracti (b) SG/Coation files (PrespeTablcoal mass raM ed with dis summarizeon patterns ol 50:50 mixDF) for SiOctively. Pere  2-7. Coal Watertio 2020720 tilled water,s the experimf partially gture, (c) SG2 and KAlSiformed by Hash leaching Hydrocaci2020480.5 and the leaental condasified (wit ash/Coal 503O8 are nos.abibi [150] conditionshloric d A    ched samplitions [214]h CO2 at 75:50 mixture 85-0695 an. . cetic acid Su20 20 48 0.5 es were drie.  0°C to 50%, and (d) SGd 72-0077, lfuric acid Na20 20 48 40.5 069 d by   ash. itric cid 20 20 8 .5  70  After the leaching of the coal sample with the mentioned solutions, the coal ash content was reduced from 30.5 wt% (Table  2-1) to 20.8 wt%. Table  2-8 shows the ash elemental analysis after the leaching. It is clear that the treatment significantly reduced the total amount of minerals in the ash.    Table  2-8. Ash elemental analysis after ash removal treatment (all in wt%).  Si Al Ti Fe Ca Mg Na K P S Undetermined*Before leaching 26.9 12.5 0.32 2.0 4.0 0.8 1.9 0.6 0.04 0.9 50.0 After leaching 26.0 5.6 0.2 1.1 0.7 0.0 0.0 0.0 0.1 1 65.3 *Mainly oxygen, since the inorganic elements are present as oxides (e.g., Al2O3, SiO2, K2O).    Removal of minerals from the coal fractured the surface of particles and increased the total surface area of the fuel. SEM images of washed coal shown in Figure  2-25 confirm that the rock-like untreated coal (Figure  2-7) was converted to a more porous surface structure after the leaching. The BET and DRK surface areas of the treated coal increased from 12 and 151 m2 g-1 (Table  2-2) to 23 and 487 m2 g-1, respectively. Observation of mineral matter effects in gasification suggests that they are basically two-fold. On the one hand, of minerals catalyze the gasification process, while, on the other hand, removal of minerals enhances reactivity by opening up the pore structure and allowing greater accessibility of gasifying agents [17]. Figure  2-26 shows the conversion and rate plots for pure treated coal, 50:50 wt% treated coal:SG, and corresponding untreated samples at 800°C and 1 atm. Since the coal minerals were reduced significantly, the treated coal sample is less reactive and is converted much more slowly than untreated samples. This reveals that mineral removal from the coal had a much stronger adverse effect on coal reactivity than the surface area enhancement due to treatment. However, by reducing the Al and Si content of the coal ash, the K/Al and K/Si ratios increased so that switchgrass potassium could act as a catalyst. Therefore, the treated mixture sample showed faster conversion than treated 100% coal (showing synergistic effect), whereas the untreated   biomass/fully conConsistenreactivitiDeminercoals.   coal samplefirm the intet with the es in air oalisation wi showed inhractions betabove resulf chars preth HF resulFigure  2-25ibition, as ween biomats, Jenkins pared fromted in a fur. SEM imagpresented inss and fossiet al. [215] some lowther decreases of fresh  Figure  2-2l fuel ash du showed tha rank coale in char rcoal after le0. Our expring co-gasit HCl treats to decreaeactivity fo  aching.  erimental refication.   ment causese considerr the lower 71 sults d the ably. rank  72  (a)  (b)  Figure  2-26. (a) CO2 Gasification conversion vs. time and (b) rate vs. conversion for treated and untreated coal, and coal/SG mixtures at 800°C and 1 atm. All proportions are in wt%. Time, t, min0 1000 2000 3000 4000 5000 6000Conversion, α 0.00.20.40.60.81.0Coal:SG = 100:0 (untreated)Coal:SG = 50:50 (untreated)Coal:SG = 100:0 (treated)Coal:SG = 50:50 (treated)Conversion, α 0.0 0.2 0.4 0.6 0.8 1.0Rate, dα/dt, min-10.0000.0020.0040.0060.008 73  2.2.5.2  Fluid coke-switchgrass co-gasification For further study of the effect of fossil fuel and biomass ash during co-gasification, CO2 gasification of fluid coke (FC) and switchgrass was investigated next. The operating conditions were the same as in the previous section during gasification (800°C and 1 atm). As shown in Figure  2-27, the FC is not a reactive fuel. Hence it took about 6 days to achieve complete conversion. This is due to the low mineral content and non-porous morphology of fresh FC, as presented in Table  2-1, Table  2-2 and Figure  2-9. The blended sample gasification behaviour differed significantly from that of the coal/SG. Both 50:50 and 75:25 FC/SG mixtures had conversions between those of the pure biomass and pure coke. Comparison of the mixture curves with the corresponding calculated (additive) results using equation ( 2-11) reveals that when the samples contained SG, no catalytic effect could be observed, and the co-gasification kinetics were very similar to the calculated data based on independent biomass and coke. However when the biomass reached its full conversion (at t≈200 min) and only the SG alkali-metal-rich ash remained in the reactor basket with the non-gasified FC, significant improvement in the FC gasification can be observed, presumably due to catalytic effect. Table  2-9 shows the elemental amounts of C, K, Si, and Al produced at 750, 800, and 900°C (in mmol.g-1) for SG char and FC char. Based on Table  2-9, the potassium/carbon (K/C), potassium/aluminum (K/Al), and potassium/silicon (K/Si) initial molar ratios of different SG and FC samples at 750, 800, and 900°C were calculated. These are presented in Table  2-10. For the mixture samples, since FC has a low ash content (2 wt%), the K/Al and K/Si ratios are much higher than for the corresponding coal/SG samples presented in Table  2-6. Therefore, enough potassium was available to satisfy the aluminum and silicon demand, while still leaving enough to catalyze the FC gasification reactions.          74  (a)  (b)  Figure  2-27. (a) CO2 gasification conversion vs. time, and (b) rate vs. conversion for FC, SG and FC/SG mixtures at 800°C and 1 atm. All proportions are in wt%. Time, t, min0 1000 2000 3000 4000 5000 6000Conversion, α 0.00.20.40.60.81.0FC:SG = 100:0FC:SG = 0:100FC:SG = 50:50FC:SG = 75:25FC:SG = 50:50 (calculated)FC:SG = 75:25 (calculated)Conversion, α 0.0 0.2 0.4 0.6 0.8 1.0Rate, dα/dt, min-10.0000.0020.0040.0060.0080.0100.0120.014 75   Table  2-9. Elemental amounts of C, K, Si, and Al produced by pyrolysis at 750, 800 and 900°C (in mmol.g-1) for SG char and FC char.  SG FC Element  750°C 800°C 900°C 750°C 800°C 900°C C  58.84 57.11 54.55 50.95 47.60 45.54 K  0.98 0.90 0.78 0.53 0.49 0.42 Si  2.40 2.40 2.40 3.48 3.48 3.48 Al  0.11 0.11 0.11 1.16 1.16 1.16    Table  2-10. Potassium/carbon, potassium/aluminum, and potassium/silicon molar ratios for different SG and FC samples at 750, 800 and 900°C (calculated from Table  2-9).   Sample K/C K/Al K/Si 750°C 800°C 900°C 750°C 800°C 900°C 750°C 800°C 900°C FC:SG = 0:100 0.017 0.016 0.014 8.782 8.092 6.963 0.409 0.377 0.324 FC:SG = 50:50 0.008 0.007 0.006 2.629 2.424 2.089 0.367 0.338 0.292 FC:SG = 75:25 0.004 0.004 0.003 1.124 1.037 0.897 0.307 0.338 0.292 FC:SG = 100:0 0.000 0.000 0.000 0.058 0.055 0.052 0.049 0.047 0.044       The XRD 2-28) doethe molarrelated toSiO2 (2θ gasificatifor the e26). Thisall the po Figure  2at 800 2.2.5.3 Two-stagthe same profiles os not conta K/Al ratio  carbon in t= 22.2° andon process anhanced gas componenttassium was-28. X-ray d°C. Powder Coal-sawdue gasificatio operating f partially gin any peakexceeded 1 he fluid cok 36.4°) and s potassiumification rat was not fou saturated wiffraction pdiffraction firespest co-gasificn of coal aconditions asified (20%s related to (see Table  2e char. The most likely  reacts withe observed nd in the coith aluminoatterns of FCles (PDF) foctively. Peration nd sawdust as in the p gasified)the formatio-10). The twprofile for Sto K2CO3 (2 CO2 [216].in the co-fe-gasificationsilicate.  parent, 20%r SiO2 and formed by Hwas investigrevious sectFC and SGn of potassio broad peG/FC had sθ = 31.2°). The K2CO3eding exper experimen gasified FK2CO3 are nabibi [150]ated in the ions for th/FC (50% um aluminoaks at 2θ = 2mall peaks The latter f is believed iments (Figts of SG/CoC, and 50%os. 27-0605. final TGA e pyrolysis gasified) (Fsilicate, bec6° and 43.6correspondiormed durinto be responures 2-27 anal, likely bec  gasified SC and 01-100experiment,and gasific76 igure ause ° are ng to g the sible d 2-ause /FC 1,  with ation  77  stages. Figure  2-29 plots the change in conversion vs. time and rate vs. conversion of different samples for gasification at 800°C and 1 atm. As expected, coal, which contains less volatiles, reacted more slowly than the sawdust. However, the sawdust rate of reaction was much less than for switchgrass (Figure  2-20), due to its lower mineral contents (Table  2-1).  Figure  2-29 indicates that during the gasification stage, addition of biomass char had no positive catalytic effect on coal char gasification. Instead, minor inhibition can be observed. This is likely due to chemical reaction between biomass metals and coal minerals. It has been shown in the literature [36] that calcium can catalyze gasification, but not as well as potassium. It is possible that calcium compounds of the sawdust ash deactivated due to sintering via crystallite growth, as discussed by Radovic et al. [203]. Calcium can also be deactivated by coal Al and Si to form gehlenite (2CaO.Al2O3.SiO2) [210]. Note that Spiro et al. [45] observed that activation energies of alkaline earth salts are higher than alkali catalyzed and even un-catalyzed reactions for coal char gasification. The use of sawdust as a natural catalyst for coal gasification seems to be impractical, as it contains very little ash (0.4 wt%) and is limited in quantity.   2.2.6 Effect of operating conditions A series of experiments was performed on coal/SG gasification to ensure that the conclusions of the previous experiments are valid for different operating conditions, and to verify that external heat and mass transfer resistances did not affect the general behaviour of the co-gasification experiments. These experiments were conducted at different CO2 flow rates, sample loadings, basket heights, and CO2 concentrations. Figure  2-30(a) shows the effect of CO2 flow rate (250, 500 and 750 mlN min-1) on gasification of pure SG at 800°C and 1 atm. No significant influence can be observed. Figure  2-30(b) shows that the initial sample load (5, 10 and 15 mg) affected the gasification rate of pure coal and 75:25 coal:SG mixture samples. However, results of all sets of runs for different sample loads showed similar behaviour, and the inhibition effect can be observed in all mixture runs. Figure  2-30(c) assesses the effect of basket size with a basket of the same diameter (17 mm ID), but different heights (20 and 6 mm). The effect of basket height is not significant, and both sets of runs with different baskets showed inhibition effect during co-gasification. Finally, the reaction gas, CO2,     78  (a)  (b)  Figure  2-29. (a) CO2 gasification conversion vs. time and (b) rate vs. conversion for coal, SD and coal/SD mixtures at 800°C and 1 atm. All proportions are in wt%. Time, t, min0 200 400 600 800Conversion, α 0.00.20.40.60.81.0Coal:SD = 100:0Coal:SD = 0:100Coal:SD = 50:50Coal:SD = 75:25Coal:SD = 25:75Conversion, α 0.0 0.2 0.4 0.6 0.8 1.0Rate, dα/dt, min-10.0000.0020.0040.0060.0080.010 79  was diluted with N2 to study the effect of CO2 concentration (100, 60, and 40 vol%) on co-gasification of coal and SG. Again, the conversion results are consistent and in agreement with the general conclusions. Gomez-Barea et al. [217] also concluded that the CO2 partial pressure did not have a significant role on mass transfer limitation effects. From Figure  2-30(d), it can be seen that, unexpectedly, increasing the CO2 partial pressure did not increase the rate of reaction for the entire conversion range. Gomez-Barea et al. [217] showed that the relationship between conversion rate and CO2 partial pressure is not linear. This is likely due to changes in CO concentration inside the char particles which can inhibit the gasification, in agreement with the analysis of CO inhibition found in other kinetic studies [217–220]. Consistent with the mass transfer calculations presented in Appendix A, Gomez-Barea et al. [217] showed that for particles smaller than 0.8 mm, the overall effectiveness factor1 > 95%, and external mass transfer limitations are negligible. Hawley et al. [221] analyzed intraparticle mass resistance with a simple catalytic model of 1-2 mm wood char particles. They concluded that, up to 5 mm, the particle size did not alter the resulting kinetic expression. Thurner et al. [173] investigated the kinetics of wood pyrolysis and reported that for wood particles smaller than 2 mm and nitrogen flow rates higher than 2 mL/s, the external heat and mass transfer limitations are insignificant.  The findings from the results presented in this chapter are in good agreement with those of Habibi [179], who studied the same fuels, but with different charring procedure, particle sizes, heating rates and basket size.                                                   1 Ratio of the actual conversion at any instant to the conversion that would be observed at that instant if there were neither external nor intraparticle gradients.  80    Figure  2-30. Effect on conversion vs. time profiles of (a) CO2 flow rate (b) initial sample load (c) basket size (d) CO2 concentration, for CO2 gasification of coal, SG, and coal/SG mixtures at 800°C and 1 atm. All proportions are in wt%. Time, t, min0 100 200 300Conversion, α 0.00.20.40.60.81.0500 ml/min250 ml/min750 ml/minTime, t, min0 200 400 600 800 1000 1200Conversion, α 0.00.20.40.60.81.0Coal:SG = 100:0 - 15mgCoal:SG = 75:25 - 15mgCoal:SG = 100:0 - 10mgCoal:SG = 75:25 - 10mgCoal:SG = 100:0 - 5mgCoal:SG = 75:25 - 5mgTime, t, min0 200 400 600 800 1000 1200Conversion, α 0.00.20.40.60.81.0Coal:SG = 100:0 - large basket Coal:SG = 75:25 - large basketCoal:SG = 100:0 - small basketCoal:SG = 75:25 - small basketTime, t, min0 200 400 600 800 1000 1200Conversion, α 0.00.20.40.60.81.0Coal:SG = 100:0 - PCO2 = 1 atmCoal:SG =75:25 - PCO2 = 1 atmCoal:SG = 100:0 - PCO2 = 0.6 atmCoal:SG = 75:25 - PCO2 = 0.6 atmCoal:SG = 100:0 - PCO2 = 0.4 atmCoal:SG = 75:25 - P CO2 = 0.4 atm(a) (b) (c) (d)  81  2.3  SUMMARY (1) Characterization of biomass and non-biomass fuels demonstrated that the switchgrass was rich in alkali metals (particularly potassium), sawdust contained the highest amount of calcium, fluid coke had high fixed carbon, and the sub-bituminous coal had a high percentage of ash. Although the pyrolysis temperature affected the macro, meso, and micro surface areas of the fuel samples, it had no significant influence on char yields. Increasing the temperature led to evaporation of potassium. (2) TGA weight loss data showed three stages during pyrolysis of biomass and fossil fuels. No significant difference was found between the average experimental devolatilization rate data and the corresponding calculated ones if separate fuels acted independently. The pyrolysis rate data indicated that biomass and fossil fuel reacted independently in the blended samples. Switchgrass and sawdust had no significant effect on coal and fluid coke pyrolysis weight loss patterns.  (3) The Coats–Redfern method was used to analyze the general kinetic characteristics of biomass and fossil fuel pyrolysis. In stage I, pyrolysis rates of all samples were controlled by diffusion. In the second stage with temperatures from 200°C to the peak temperature (Tpeak), pyrolysis of the samples containing biomass exhibited approximately first-order reaction kinetic control, while pyrolysis of the stand-alone coal and fluid coke followed diffusion and random nucleation mechanisms, respectively. When the reaction temperature exceeded Tpeak, where nearly all the decomposition reactions take place, all samples appeared to follow a diffusion-controlled mechanism. (4) The nature of biomass and non-biomass samples and their ash content and composition play important roles in determining the co-gasification reaction rate of mixed fuels. (5) The alkali and alkaline earth metal constituents of biomass can react with minerals in fossil fuel to form metal aluminasilicate compounds (e.g. KAlSi3O8) which inhibit gasification. This was observed in co-gasification of switchgrass and sawdust with coal. (6) On the other hand, for high ratios of both K/Al and K/Si, biomass potassium can catalyze fossil fuel gasification. This was shown in co-gasification of switchgrass with fluid coke and water/acid-washed coal.     82  Chapter 3. STEAM GASIFICATION OF COAL AND SWITCHGRASS IN A BUBBLING FLUIDIZED BED REACTOR  Biomass, originating from crops and trees, has attracted increasing interest as the full cycle of biomass growth and energy utilization through gasification is considered to be largely carbon neutral. However, a major challenge for commercialization of biomass gasifiers is the low energy density and scattered distribution of the biomass, as well as high handling and transportation costs. This has hindered the potential for biomass utilization for energy and fuels [91, 222]. On the other hand, coal gasification is a well-developed and economically feasible conventional technology for energy utilization due to coal’s high energy density and concentration at source. Greenhouse gas emissions, however, constitute a major issue related to coal utilization for energy. Co-gasification of biomass and coal seems to be a feasible solution, reducing the GHG footprint of fossil fuels [223]. Co-feed gasification reduces production costs (compared to biomass gasification), and makes the process more sustainable (compared to coal gasification). Co-gasification of coal and biomass can also reduce problems associated with high ash and sulfur contents of the coal [224]. Also, as discussed in the previous chapter, some biomass minerals catalyze coal gasification reactions, enhancing the overall gas production efficiency, as well as reducing the product tar content through catalytic tar cracking. Fluidized bed reactors are among the most promising types of gasifier, with excellent fuel mixing, carbon conversion and thermal efficiency. They are able to operate with flexible feed specification and size, and are potentially suitable for scale-up. Therefore, it was decided to study gasification of single and mixture fuels in a pilot-scale fluidized bed reactor, with steam as the gasifying agent. Steam helps to produce high-quality product gas with good heating value rich in hydrogen. As noted in  Chapter 2 and by Pullen [17], effectiveness of particular catalysts for steam gasification has been reported to be generally similar to that for CO2 gasification. This chapter presents and discusses results of single feedstocks (coal and switchgrass) steam gasification.  Chapter 4 investigates steam co-gasification of coal/switchgrass mixtures and compares single-fuel and co-gasification results.     83  3.1  FUEL CHARACTERIZATION As discussed in  Chapter 2, biomass alkali and alkaline earth metals, in particular potassium, can enhance fossil fuel gasification. However, other fossil fuel ash constituents, in particular aluminum and silicon, can be a major obstacle to having catalytic effect between biomass and fossil fuel during co-gasification. The results discussed in the previous chapter led us to select a biomass with ash content rich in potassium, and a thermal coal with low ash content, poor in aluminum and silicon, as fuels for pilot scale gasification experiments. After searching for available local and national biomass and fossil fuel resources, switchgrass from Ontario (Don Nott farm) and thermal coal from Vancouver Island (Quinsam mine) were chosen for this study. Table  3-1 presents the ultimate, proximate and elemental ash analyses of the parent fuels. Consistent with the previous chapter (Table  2-1), the coal contained much higher carbon and therefore heating value (80.3 wt% and 28.4 MJ/kg) than the biomass samples. Also the Quinsam mine coal had much less ash and moisture contents (12.9 wt% and 4.25 wt%, respectively) than the Alberta sub-bituminous coal tested in  Chapter 2 (30.5 wt% and 17.5 wt%, see Table  2-1), appropriate for this research. The silicon content of the Quinsam mine coal (16.9 wt%) was also less than for the sub-bituminous coal (26.9 wt%). Switchgrass ash samples were rich in potassium, 10.76 wt% for spring harvest switchgrass (SP-SG) and 21.83 wt% for fall harvest switchgrass (F-SG). The fall harvest switchgrass was therefore expected to have the highest catalytic effect on coal gasification, as it contained the highest potassium in its ash. The Quinsam mine coal is classified as a thermal coal, with low crucible swelling index (0.5), suitable for fluidization applications.  All samples were crushed before the characterization analyses and gasification experiments. A C.S. Bell Co. hammer-mill with 1/16 inch (1.59 mm) screen openings and a FRITSCH disc-mill were used for biomass and fossil fuel crushing respectively. Both biomass and coal had bimodal particle size distributions, as shown in Figure  3-1 (using standard sieves). The coal, SP-SG and F-SG bulk densities were 740, 235 and 242 kg/m3, respectively.      84    Table  3-1. Ultimate, proximate, and ash analysis of fresh feedstocks for pilot-scale experiments.    Sample Quinsam Mine Coal (Vancouver Island) Spring Switchgrass (Ontario) (SP-SG) Fall Switchgrass (Ontario) (F-SG)  Ultimate analysis (wt%), daf* Carbon, C   80.3   49.7   47.7 Hydrogen, H 5.5 6.2 5.9 Nitrogen, N 0.9 0.9 1.0 Sulfur, S 0.7 0.1 0.1 Oxygen, O (diff**)   12.6   43.1   45.3  Proximate analysis (wt%) Moisture 4.25  9.26    11.63 Ash (db*)   12.90  3.07  3.80 Volatile (db)   38.01    79.50    79.47 Fixed Carbon (db)   49.09    17.43    16.73 Higher heating value (db) (MJ/kg)   28.40    19.38    19.23 Crucible Swelling Number          0.5 n.a. n.a.  Ash analysis (wt%) Si   16.87   20.14    14.60 Al   13.39 0.22   0.06 Ti 1.43 0.02   0.01 Fe 4.69 0.38   0.19 Ca   13.36 9.40   7.28 Mg 0.24 3.26   2.80 Na 0.16 0.09   0.02 K 0.06   10.76     21.83 P 0.24               2.2   3.13 S 2.30 2.15   1.35 Ba 0.01 0.02   0.01 Sr 0.07 0.01   0.01 Mn 0.05 0.05   0.03 Cr 0.00 0.01   0.01 Cu 0.03 0.01   0.01 Ni 0.02 0.00   0.01 V 0.05 0.01   0.00 Zn 0.01 0.02   0.03 Hg 0.00 0.04   0.02 Undetermined***   46.96   51.21   48.6 * daf = dry and ash free, db = dry basis; ** calculated by difference, *** predominantly oxygen, since the inorganic elements are present as oxides (e.g., Al2O3, SiO2, K2O).  85  (a)  (b)  Figure  3-1. Particle size distributions of (a) coal (b) spring and fall switchgrasses.  3.2  EXPERIMENTAL  3.2.1 Experimental apparatus and operation Steam co-gasification of switchgrass and coal was performed in the Highbury Biofuel Technologies Inc. (HBTI) pilot-scale bubbling fluidized bed (BFB) reactor. This unit mainly consists of a high-temperature reactor tube inside a 305 mm ID high-pressure reactor shell, a screw feeder, steam super-heaters, a double cyclone, gas coolers, and a baghouse filter, as shown Particle diameter (µm)45 55 65 75 85 106 180 250 355 425 500 595 710 850>1.168Mass fraction, %0510152025Particle diameter (µm)106 180 250 300 355 425 500 595 710 850 >1.168Mass fraction, %051015202530Spring SGFall SG 86  in Figure  3-2. The detailed flow diagram is shown in Figure  3-3. Appendix B provides the unit calibration curves.  Fuel stored in the biomass storage hopper under N2 was fed to the screw feeder’s hopper in a batch feeding manner through a 51 mm ball valve. The biomass in the feed hopper (Vibra Screw Inc.) was fed to the fluidized bed gasifier through a pneumatic conveyor through a 8.48 mm pipe using N2 as the conveying gas. A Vibra Screw Inc. electric vibrator attached to the feeder hopper helped propel fuels into the pipe.  The biomass steam gasification unit is located in the high-head laboratory of the UBC Pulp and Paper Centre. The gasifier consisted of a 102 mm ID, 1219 mm long, stainless steel (800H/HT, S40, SMLS) pipe as an inner reactor with a gas distributor installed above the steam entrance, as shown in Figure  3-4. An internal cyclone (63 mm ID) was installed at the top of the column, immediately upstream of where the producer gases exit, to separate entrained solids from the gas stream, and to reduce the tar content of the gases. Pfeifer et al. [225] showed that the hot surface of an internal cyclone can thermally crack tars. The reactor is heated by two semi-cylindrical (127 mm ID, 229 mm OD, 914 mm L) ceramic fibre electrical heaters (240 V 4200 W each) and one full cylindrical (127 mm ID, 152 mm L) ceramic fibre electrical heater (120 V 1400 W) on the top section. Building saturated steam (586 kPag, 160°C) is supplied to the gasifier through a steam flow meter to the steam super-heaters (12×240 V×1800 W), super-heated to about 800°C, before entering the gasifier. N2 is introduced for system purge and preheating. The produced syngas leaves the gasifier through double external cyclones to gas coolers (double-pipe type) cooled by water. The cooled syngas passes through an orifice plate or a bypass before entering a baghouse filter, and then continues to the rooftop burner. A compressed air supply system was installed for optional biomass combustion preheating or biomass air-steam gasification [226].  Silica sand (Lane Mountain Company, Washington) was chosen as the inert fluidization bed material for the process. The sand improves the agitation in the fluidized bed, while also providing high rates of heat and mass transfer. It also serves to decrease the tendency of the fuel particles to agglomerate at the operating conditions [81]. Olivine sand, dolomite, and lime can be also used. However the attrition of dolomite and lime is significant. The sand was sieved to a particle size of 300-355 µm (US mesh #45-50) (resulting in Geldart group B particles). The static bed height for all experiments was ~ 0.30 m. This height was thought to be optimal for   temperatuthe sand Table  3-propertie0.37 m/s. Figur  Silica sare distributis shown in3. The valus as a fluid.  e  3-2. SchemCaOnd 0.04ion in the co Table  3-2es were ca The bubbliatic of HigfluidizTable  3 MgO  0.02 lumn based. The particlculated bang bed washbury Biofued bed reac-2. FluidizaAl2O30.64 on previousles and theised on sup operated wel Technolotor (adaptedtion bed ma Fe2O0.15 HBTI exper hydrodyner-heated stith an inletgies Inc. (H from HBTIterial analys3 K2O 0.2riments. Thamic propeream (at 52superficial BTI) pilot-s).  is.  Na27 0.0e compositities are list5°C and 1 gas velocitycale bubblinO Si8 9887 on of ed in atm)  of ~  g O2 .14  88   Figure  3-3. Highbury Biofuel Technologies Inc. unit flow diagram (adapted from HBTI).    Figure  3-4. Detailed dimensions of bubbling fluidized bHBTI).   ed gasifier (not to scale) (adapted fr89  om  90  Table  3-3. Material and hydrodynamic properties (based on steam at 525°C and 1 atm). Material ρb (kg/m3) (bulk) ds (µm) Ut (m/s)** Umf (m/s)*** Silica Sand 1600 327.5 1.24 0.064 Coal 740 142* 0.44 0.007 SP-SG 235 264* 0.45 0.008 F-SG 242 274* 0.47 0.009 * Based on ݀௦ = 1 ∑ ( ௫೔ௗೞ೔)௜൘ , ** Calculated based on Haider and Levenspiel [227]  correlation, *** Calculated based on Grace [228] equation.    3.2.2 Product gas analysis The product gas composition was measured by an on-line Varian Inc. CP-4900 micro gas chromatograph (GC) every 4 minutes during operation, with a Molsieve column for H2, CO, N2, and CH4 detection. Carbon dioxide was measured using a CP-PoraPLOT U column. The GC was calibrated using Praxair standard gas before the experiments. The product gas moisture and tar were fully removed by a shell and tube condenser, a silica-gel column, and a Genie membrane separator before being introduced to the GC. Helium was the GC carrier gas.  3.2.3 Tar sampling A modified version of European protocol CEN/TS 15439 [229] (Figure  3-5) was used for tar sampling of the product gas. The tar sampling port was branched from the main exit pipe located at the top of the gasifier. The port was heated by a rope-heater and insulated to keep the stream temperature above 300°C in order to prevent tar condensation along the pipe. A XC-60 Apex source sampling console which includes a gas meter, sample pump, and temperature control probe was used to sample the product syngas during operation. The product gas first passed through a glass fiber filter to remove fine particles. Next, the filtered gas entered six impinger bottles in series, each containing 50 ml of iso-propanol solvent. All six impingers were kept at    –20°C in a salt ice bath. After the tar sampling, the bottles were rinsed with extra solvent to collect the entire tar from the bottles. Then, the captured tar and solvent mixture was re-filtered through a 0.02 μm filter to remove particulates. The collected solvent/tar mixture was thermally treated in a rotary evaporator at 55°C and 17 kPa until all the solvent was extracted, taken as once the drip rate in the evaporator had fallen to one every 4 s. In order to remove the remaining water from the tar, 50 ml ethanol were added and thermal treatment was continued. The   evaporativacuum the remaremovedSee Appresults ar 3.2.4 CAs showexchangecarbon (Tanalyzer purgeablthe acidicarbon don was conflask was thining water , allowed toendix 14 ofe presented Figureondensed wn in Figure r. This watOC) conten(Shimadzu e organic cafied sample ioxide throutinued untilen initiateddroplets [22 cool in a d Butler [17in  Chapter 4  3-5. Atmosater analys 3-2, the uner can then t of the watTOC-VCPHrbon (NPOCusing an inegh acidific a drip rate, with the va9]. Air flowesiccator an4] for the c. pheric and iis  reacted steabe treated aer sample o). The org) method, wrt gas (nitroation, and t of one evcuum held  continued d weighed tomplete tarso-kinetic sa m condensend reused af each run wanic carbonhich is basgen in this he remaininery 4 s wasbelow 68 kfor 20 min.o determine sampling pmpling traind downstres a gasificatas analyzed was meased on purginwork). Inorg carbon co reached. APa. This pro The vacuum the qualityrocedure. T for tar [23am in a sheion agent. T in triplicatured by meg volatile oganic carbontent is oxir flow intcedure rem flask was of producehe tar sam 0].  ll and tubehe total ore using the ans of the rganic carbn is convertidized to ca91 o the oved  then d tar. pling  heat ganic TOC non-on of ed to rbon  92  dioxide through high-temperature catalytic oxidation. The total suspended solids (TSS) of the water samples were measured using a Whatman 934-AH filter. See the standard TSS procedure in Appendix C. The TOC and TSS results are presented in  Chapter 4. 3.2.5 Operating procedure Here is a brief operating procedure followed during each run:  • Perform leak test on the BFB reactor and piping. • Install carbon monoxide indicator on each floor of the unit to alarm upon gas leakage. • Ensure adequate nitrogen using liquid N2 tank.  • Initiate fluidization using nitrogen at 0.002 m3/s. • Start heating the reactor by providing power to the electrical heaters. • Start heating the steam super heater by supplying power to the corresponding electrical heaters. • Pass water through heat exchanger tubes by regulating the water valve to 0.18 m3/h.   • After heating the system for ~ 40 min, purge the downstream piping using nitrogen in order to remove any remaining solid particulates from the line. • Turn on tar sampling heater to maintain the port temperature above 300°C. • Turn on the rooftop incinerator by opening the natural gas valve. • Open the steam valve and reduce the N2 flow rate when the desired reactor temperature has been reached. • Pressurize the feed hopper using nitrogen.  • Operate the online GC for product gas measurement. • Start solid feeding into the reactor. • Initiate tar sampling by turning on pump.       93  3.2.6 Gasification indices To assess the process technology, the following variables were defined and determined: Carbon conversion efficiency, ηC   4 2,*1000*[CH % CO% CO %]*12 / 22.4(1 )*C%Mass of C in the product gasgasCfuel dry ashMass of C in the fuelvm xη + += −	  ( 3-1) Hydrogen conversion efficiency, ηH  2 4, , 2*1000*[2H % 4CH ]/ 22.4( (1 )*H%) (2* *H O% /18.0) (2* /18.0)Massof H in the product gasgasHfuel dry ash fuel wet steamMass of H in the inletvm x m mη += − + +	  ( 3-2)  where CH4%, CO%, CO2% and H2 (vol%) are the gas concentrations, vgas (Nm3/h) is the dry gas flow rate, mfuel,dry and mfuel,wet are the dry and wet fuel feeding rates (g/h) respectively, xash is the ash content in the feed, C% and H% are the carbon and hydrogen contents (weight based) in the ultimate analysis of fuel, respectively, and msteam (g/h) is the steam mass flow rate. The steam produced was not factored into the hydrogen utilization equation, as it was too difficult to measure in the current reactor setup.  The cold gas efficiency, ηcg  ,**gas gascgfuel dry fuelm Qm Qη =  ( 3-3) where mgas (g/h) is the dry gas mas flow rate, and Qgas (MJ/g) and Qfuel (MJ/g) are the product gas and fuel heating values, respectively.    The dry gas higher heating value at the standard state of 101.3 kPa and 273 K can be estimated from the gas composition by 2 4HHV 12.7H % 12.6CO% 39.8CH %= + +  ( 3-4) where the species heats of combustion are in MJ/N m3. This equation is based on heat of combustion data [231], assuming ideal-gas behaviour for the gaseous species. The micro-GC was not able to measure the concentration of higher hydrocarbons (e.g. ethylene). They are often too  94  low to be detected. However, accurate measurements of them is crucial when the gasifier operates at temperatures below 700°C or at elevated pressure, as the gas heating value can be significantly altered by the presence of even small fractions of these hydrocarbons [20].  The mass-based steam-to-fuel ratio, SF, is defined by ,SF steamfuel drymm=  ( 3-5) In this chapter, the error bars are based on 95% confidence intervals for a population mean (assuming a normal distribution).  3.3  RESULTS AND DISCUSSION 3.3.1 100% Quinsam mine coal gasification To study the effect of biomass on coal gasification in a pilot-scale fluidized bed reactor, single feedstocks were first gasified. Figure  3-6 shows the results for a dry nitrogen-free product gas composition and bed temperature for a 100% coal run with different fuel feed rates. The bed thermocouple was at a height of ~ 0.2 m above the distributer. Time (t) 0 corresponds to the start of feeding fuel to the reactor. As steam gasification is an endothermic process, introducing coal to the reactor caused a sharp drop in the bed temperature. The temperature and gas concentration were relatively stable during the experiment. This shows good performance of reactor heaters to provide the required gasification heat, and to maintain the reactor at a desired temperature. The endothermic gasification reactions (e.g. methane reforming, reaction ( 1-11) and steam-carbon, reaction ( 1-7)) dominated, and the product gas is rich in hydrogen. It is evident from Figure  3-6 that increasing the coal feed rate, caused the bed temperature to drop slightly, as more heat was required at lower steam-to-coal ratios. Full measured temperatures and pressures drop profiles are presented in Appendix D.  Figure  3-7 shows the effect of steam-to-coal mass ratio (SF) on gas composition at different temperatures. Even in the narrow range of steam-to-coal ratios covered during the gasification run, there was a visible effect of SF on the product gas composition. The temperatures for different SF were very close to each other. Increasing the steam-to-coal ratio caused a moderate increase in the H2 and CO2 concentrations and decreases in the CO and CH4 concentrations. This   95   Figure  3-6. Gasification results for dry nitrogen-free product gas composition and local bed temperature for coal run with different coal dry basis (db) feed rates at atmospheric pressure.   Figure  3-7. Effect of steam-to-fuel (coal) ratio on gas composition at different temperatures at atmospheric pressure. Time, t, min0 25 50 75 100 125 150 175 200Product gas concentration (vol%)0102030405060708090100Temperature, oC600625650675700725750775800825850875900H2COCH4CO2Bed temp.0.79 kg/h(SF=3.82)1.18 kg/h(SF=2.54)1.38 kg/h(SF=2.16)0.99 kg/h(SF=3.02) 1.77 kg/h(SF=1.70)2.35 kg/h(SF=1.28)Steam-to-fuel mass ratio, SF1.0 1.5 2.0 2.5 3.0 3.5 4.0Average product gas concentration (vol%)010203040506070Temperature, oC845850855860865870875H2COCH4CO2Bed Temp. 96  can be explained by more steam reforming of CH4 and more water-gas shift reaction of CO taking place because of the increased steam quantity. The results are in good agreement with reported literature data [232 –235].   Figure  3-8 presents the carbon, hydrogen and cold gas efficiencies, as well as product gas higher heating value vs. time, for different coal feed rates (see Figure  3-6). Also Table  3-5 summarizes the gasification indices for different steam-to-coal ratios. It appears that higher feed rates (lower SF) resulted in higher HHV due to a higher CH4 concentration (see Figure  3-7). The carbon and cold gas efficiencies were less than 40% during the experiment, due to high fixed carbon (see Table  3-1) and low reactivity of the coal. The high fixed carbon of the solid particles captured by the external cyclones confirms this, as shown in Table  3-4. The un-reacted coal particles and chars caused pressure build-up and gradual blockage downstream as shown in Figure  3-9. Recirculating entrained material to the bed can enhance the gasifier efficiency. The hydrogen efficiency slightly increased with increasing fuel feed rate, presumably due to higher methane production at lower SF (see Figure  3-7).  Table  3-4. 100% coal proximate analysis of entrained particles captured by external cyclones.   Moisture Ash and sand (db*) Volatiles (db) Fixed Carbon (db) Proximate analysis 8.6 46.7 4.7 48.6 * Dry basis  97   Figure  3-8. Coal gasification carbon efficiency, hydrogen efficiency, cold gas efficiency, and product gas higher heating value vs. time for different SF (see Figure  3-6) in atmospheric pressure gasification.    Figure  3-9. Coal gasification reactor lower and upper pressure vs. time. Time, t, min0 25 50 75 100 125 150 175 200Efficiency, η 0.00.10.20.30.40.50.6HHV, MJ/m3024681012141618CarbonHydrogenCold gas HHVTime, t, min50 100 150 200 250Pressure, kPag0246810121416P2 - lower bed (0.2 m above the distributor)P3 - upper bed (1 m above the distributor)blockage in the downstream piping  98    Table  3-5. Summary of 100% coal gasification parameters at different steam-to-fuel ratios (SF) and at atmospheric pressure.    SF 1.28 1.70 2.16 2.54 3.02 3.81Temperature (°C)  858 852 858 867 859 870H2 concentration (vol%) 60.7 61.3 61.0 61.4 62.8 62.4CO concentration (vol%) 14.9 14.5 14.5 14.5 13.7 14.4CH4 concentration (vol%) 9.4 8.4 8.5 7.9 6.0 6.9CO2 concentration (vol%) 15.0 15.8 16.0 16.2 17.5 16.3Carbon conversion efficiency (ηC) 0.15 0.16 0.16 0.15 0.20 0.20Hydrogen conversion efficiency (ηH) 0.18 0.15 0.13 0.11 0.13 0.10Cold gas efficiency (ηcg) 0.25 0.26 0.26 0.26 0.33 0.33HHV (MJ/Nm3) 13.4 13.0 13.0 12.8 12.1 12.5Gas yield (m3/kg coal) (db and N2-free) 0.54 0.57 0.57 0.57 0.81 0.75 99  3.3.2 100% spring switchgrass (SP-SG) gasification  The gasification of biomass is quite different from the gasification of coal. Biomass fuels tend to have considerably more volatile matter than coal. Hence, char gasification plays a minor role in the gasification of biomass because char typically represents less than 30 wt% of the pyrolysis products. Thus, devolatilization and secondary reactions play critical roles in determining the product distribution during the gasification of biomass [236, 237]. Figure  3-10 plots the product gas composition and temperature vs. time for the 100% spring switchgrass experiment. When biomass was fed to the reactor, the bed temperature dropped slightly due to endothermic reactions. The reactor reached a steady-state condition at t≈60 min. Increasing the biomass feed rate caused a major drop in bed temperature from ~ 860°C to less than 800°C in the final measurements. The CO/H2 molar ratio for the SP-SG run was greater than for 100% coal due to biomass having higher volatile and oxygen contents than coal (see Table  3-1). Based on Le Chatelier’s principle, higher oxygen helps the oxidation and partial oxidation reactions. The CO2 produced from the oxidation reactions can then react via the Boudouard reaction, ( 1-6), producing CO. It was also shown in  Chapter 1 that potassium can catalyze Boudouard and steam-carbon reactions. From Figure  3-10, it can be seen that biomass steam gasification produces less CO2 than coal gasification (cf. Figure  3-6). Figure  3-11 shows the effect of the steam-to-biomass mass ratio (SF) on the product gas composition. The temperature corresponding to each concentration data point is also presented. As for the 100% coal run, increasing the SF increased the H2 yield and decreased the CH4 product gas concentration. The largest error bars are for SF=3.58, corresponding to the early period of gasification when the compositions had not yet reached steady-state. It is evident from Figure  3-11 and Table  3-6 that supplying extra steam as a gasifying agent increased the partial pressure of H2O inside the gasification chamber, favouring the steam-carbon, water-gas shift and methane reforming reactions (reactions ( 1-7), ( 1-10) and ( 1-11)), leading to increased H2 production [238]. However, the gasification temperature needs to be high enough (above 750–800°C) for the steam reforming and water gas reactions to be favourable [239–241]. The results of Lv et al. [232] on the effect of SF on biomass steam gasification in a fluidized bed are in good agreement with the data presented.  100   Figure  3-10. Gasification results for a dry nitrogen-free product gas composition and local bed temperature of 100% SP-SG run with different biomass feed rates (db) at atmospheric pressure.  Figure  3-11. Effect of steam-to-fuel (biomass) ratio on steady-state dry nitrogen-free gas composition at different temperatures and atmospheric pressure.  Time, t, min25 50 75 100 125 150 175 200 225Product gas concentration (vol%)0102030405060708090100Temperature, oC600625650675700725750775800825850875900H2COCH4CO2Bed Temp.0.88 kg/h(SF=3.58)1.05 kg/h(SF=2.97)1.23 kg/h(SF=2.52)1.40 kg/h(SF=2.18)1.66 kg/h(SF=1.87)Steam to biomass ratio, SF2.0 2.5 3.0 3.5Average product gas concentration (vol%)0102030405060Temperature, oC780800820840860880H2COCH4CO2Bed Temp. 101  Figure  3-12 and Table  3-6 show that the biomass gasification led to much higher carbon, hydrogen, and cold gas efficiency than coal gasification. As presented in  Chapter 2 (e.g. Figure  2-4), this is due to biomass being much more reactive than coal. The biomass ash rich in potassium (see Table  3-1) also enhanced the reactivity of the biomass. Also fewer solid particles were separated and captured by the external cyclone for biomass than for the coal run, indicating that most of the fuel was converted to gaseous products. The proximate analysis of the separated solids is provided in Table  3-7. No blockage or pressure build-up was observed during this biomass gasification experiment. The higher heating value of product gas did not change significantly (Figure  3-12), as the gas composition followed a very similar trend for different biomass feed rates. It is evident from Table  3-6 that for SF increasing from 1.87 to 3.58, the gas yield, HHV, carbon conversion efficiency, and cold gas efficiency all exhibited decreasing trends, which can be explained by the excessive quantity of low temperature steam lowering reaction temperature, decreasing reaction rates and degrading the gas quality [232].   Figure  3-12. 100% SP-SG carbon efficiency, hydrogen efficiency, cold gas efficiency and product gas higher heating value vs. time for different SF (see Figure  3-10) at atmospheric pressure. Time, t, min50 75 100 125 150 175 200 225Efficiency0.00.20.40.60.81.0HHV, MJ/m3024681012141618CarbonHydrogenCold gas HHV 102    Table  3-6. Summary of 100% SP-SG gasification parameters at different steam-to-biomass ratios at atmospheric pressure.   SF 1.87 2.18 2.52 2.97 3.58 Temperature (°C)  807 855 862 865 869 H2 concentration (vol%) 31.5 31.0 30.9 32.1 35.2 CO concentration (vol%) 44.8 45.8 46.2 46.1 45.5 CH4 concentration (vol%) 12.7 12.3 12.3 11.6 11.0 CO2 concentration (vol%) 11.1 10.8 10.7 10.2 8.3 Carbon conversion efficiency (ηC) 0.87 0.75 0.68 0.51 0.40 Hydrogen conversion efficiency (ηH) 0.22 0.16 0.13 0.09 0.06 Cold gas efficiency (ηcg) 0.95 0.81 0.73 0.56 0.45 HHV (MJ/Nm3) 14.72 14.66 14.66 14.54 14.62 Gas yield (m3/kg biomass (db and N2-free) 1.25 1.09 0.98 0.74 0.59   The firstcyclone sthe biom Table  3ProximateFi 3.3.3 1The fall pressure concentrapresentedswitchgrathe F-SG external cyeparated smass char part-7. Proxima  analysis gure  3-13. S00% fall swswitchgrass with a steamtion increas in Figure ss samples, gasificationclone captaller particuicles retainete analysis oMoistu5.EM image itchgrass (Fsteam gasif-to-biomased and the C 3-10. This  affecting th resulted inured larger lates (e.g. bd by the firsf particles cre Ash3 of biomass c-SG) gasifiication expes ratio of SO concentrmight be due catalysis  higher carbparticles (eiomass asht cyclone. aptured by eSG.   and sand (db62.4  har particlecation  riment wasF=2.4. Figuation decreae to differeof gasificaton, hydrog.g. biomass). Figure  3-xternal cyc) Volati5s caught by  performed re  3-14 andsed relativent ash comion reactionen and cold2 mm  char), whe13 shows thlones 1 and le (db) F.6  first externaat ~ 862°C  Table  3-8  to the correpositions os. As shown gas efficienreas the see SEM ima2 for 100% ixed Carbon32 l cyclone. and atmospshow that thsponding ref spring and in Figure cies than SP103 cond ge of SP- (db) heric e H2 sults  fall  3-15, -SG  104    Table  3-8. Summary of 100% F-SG gasification parameters at different bed temperatures at atmospheric pressure. Average bed temperature, °C  703 728 755 793 807 863SF 2.35 2.38 3.29 2.41 2.35 2.39H2 concentration (vol%) 23.6 25.9 27.5 29.3 32.7 36.3CO concentration (vol%) 45.8 46.0 45.2 44.3 42.0 39.8CH4 concentration (vol%) 15.2 13.9 14.1 13.8 13.1 12.1CO2 concentration (vol%) 15.4 14.2 13.3 12.6 12.2 11.8Carbon conversion efficiency (ηC) 0.69 0.64 0.76 0.78 0.78 0.75Hydrogen conversion efficiency (ηH) 0.12 0.11 0.12 0.15 0.17 0.17Cold gas efficiency (ηcg) 0.65 0.62 0.76 0.79 0.83 0.82HHV (MJ/Nm3) 14.84 14.64 14.81 14.84 14.69 14.38Gas yield (m3/kg fuel) (db and N2-free) 0.85 0.81 0.98 1.03 1.09 1.10     105   Figure  3-14. Gasification results for a dry nitrogen-free product gas composition and bed temperature of 100% F-SG run with SF=2.4 at atmospheric pressure.   Figure  3-15. 100% F-SG gasification carbon efficiency, hydrogen efficiency, cold gas efficiency, and product gas higher heating value vs. time for SF=2.4 at atmospheric pressure. Time, t, min0 10 20 30 40 50 60Product gas concentration (vol%)0102030405060708090100Temperature, oC600625650675700725750775800825850875900H2COCH4CO2Bed Temp.1.26 kg/h(SF=2.4)Time, t, min0 10 20 30 40 50 60Efficiency0.00.20.40.60.81.0HHV, MJ/m30510152025CarbonHydrogenCold gas HHV 106  gasification, probably due to higher potassium concentration of the F-SG. The N2-free gas yield also increased from 0.98 m3/kg (see Table  3-7) for SP-SG to 1.10 m3/kg biomass (db) for F-SG.     Effect of bed temperature Temperature is critical for biomass gasification. The effect of bed temperature on the gas product composition and HHV for F-SG steam gasification at almost constant steam-to-biomass ratio (SF≈2.4) is presented in Figure  3-16. It is evident that the H2 concentration increased with temperature, whereas the CH4 concentration showed an opposite trend. According to Le Chatelier’s principle, higher temperatures favour the reactants in exothermic reactions and the products in endothermic reactions. Therefore, the endothermic reactions (e.g. steam-methane reforming, steam-carbon reaction) were important at higher temperature, resulting in an increase of H2 concentration and a drop in the CH4 concentration. Hence raising the temperature was more favourable for hydrogen yield. At temperatures above 670°C, the water-gas shift reaction would dominate the process, resulting in decrease in CO and increase in H2 concentrations. The contents of CO and CO2 also depend on partial/complete combustion reactions which are exothermic [232]. Thus, higher temperature was not favourable for CO and CO2 production. At temperatures close to 900°C, both steam reforming and the Boudouard reactions dominate, resulting in CO production and CO2 consumption [238]. The concentrations of higher hydrocarbons (e.g. C2H2) were not measured in our experiments. It is likely that they show a downward trend with increasing temperature due to thermal cracking and steam reforming [232, 242]. Increasing temperature slightly reduced the HHV of the product gas from 14.84 MJ/m3 at ~ 700°C to 14.50 MJ/m3 at ~ 860°C, due to the decrease in CH4 and CO concentrations. Gupta and Cichonski [243] observed significant increases in the H2 concentration above 800°C for SF between 0.5 and 1.08. The maximum H2 yield was obtained at 1000°C for a feedstock consisting of paper, and at 900°C feedstocks consisting of cardboard and wood pellets. González et al. [244] observed that in air gasification, the H2 and CO concentrations increased from 700 to 900°C, whereas those of CH4 and CO2 decreased. Turn et al. [242] found that for a temperature increase from 750 to 950°C, the H2 concentration increased from 31 to 45%, while CH4 and CO remained nearly constant, CO2 decreased and the gas yield increased. These results are in good agreement with the data reported by Lv et al. [232].  107    Figure  3-16. Effect of temperature on a dry nitrogen-free gas composition of 100% F-SG gasification for SF≈2.4 at atmospheric pressure.  The effect of temperature on hydrogen efficiency, carbon efficiency, cold gas efficiency, and gas yield of fall switchgrass gasification for SF≈2.4 is shown in Figure  3-17. All gasification performance parameters increased with increasing temperature for the range covered. Higher temperature resulted in increased gas yield because of higher conversion efficiencies. Large error bars in Figure  3-17 might be related to limited measured data points at different temperatures.   It has been observed in the literature [186, 187] that higher temperatures (700 to 950°C) increased the gas yield and overall energy content of the gas. Kumar [238] reported that an increase in temperature (furnace set point from 750 to 850°C) led to increases in energy and carbon conversion efficiencies and percent gas compositions of H2. Boateng et al. [245] reported that an increase in gasification temperature from 700 to 800°C caused the gas yield, energy efficiency, carbon conversion efficiency and H2 content to increase, whereas the CH4, CO and CO2 contents all decreased. The decrease in CO content may have been due to the comparatively low temperature (relative to 850–900°C) for the Boudouard reaction to predominate [238]. Lv et Temperature, oC700 725 750 775 800 825 850 875Average product gas concentration (vol%)5101520253035404550HHV, MJ/m305101520H2COCH4CO2HHV 108  al. [232] also observed that the gas yield and carbon conversion efficiency generally increased with increasing temperature.           Figure  3-17. Effect of temperature on hydrogen efficiency, carbon efficiency, cold gas efficiency, and gas yield for 100% F-SG gasification with SF≈2.4 at atmospheric pressure.   3.4  SUMMARY (1) Single fuels (coal, spring switchgrass and fall switchgrass) were steam gasified in a pilot scale bubbling fluidized bed reactor at ~ 860°C and 1 atm. The gasifier temperature and product gas concentrations were relatively stable during the experiments. The bed temperature dropped slightly when the fuel feed rate was increased due to higher heat demand at lower steam-to-fuel ratios. (2) Steam gasification of Vancouver Island coal resulted in higher H2/CO molar ratio (~4.21) and CO2 concentration and lower CH4 production than for gasification of spring and fall switchgrass samples, due to the higher carbon content of the coal and therefore more steam-carbon and combustion reactions.  Temperature, oC700 725 750 775 800 825 850 875Efficiency0.00.20.40.60.81.0N2-free gas yield (m3 /kg db fuel) 0.70.80.91.01.11.21.3Carbon eff.Hydrogen eff.Cold gas eff.Gas yield 109  (3) Steam gasification of biomasses rich in potassium resulted in H2/CO molar ratios<1 presumably due to dominance of the Boudouard and steam-carbon endothermic reactions. As shown in  Chapter 1, potassium can catalyze these two reactions.    (4) Steam gasification of 100% coal resulted in lower hydrogen, carbon and cold gas efficiencies and gas yield than switchgrass gasification due to lower reactivity and higher required reaction residence time of coal particles.   (5) F-SG gasification resulted in higher carbon, hydrogen and cold gas efficiencies than for SP-SG gasification, possibly due to higher potassium concentration and hence, greater reactivity of the F-SG. (6) For both coal and switchgrass gasification experiments, increasing the steam-to-fuel ratio (SF) increased the H2 concentration and decreased the CH4 concentration, due to more steam-methane reforming.     (7) Increasing the gasifier bed temperature increased the H2 concentration, gas efficiencies and yield, whereas CO, CH4 and CO2 concentrations tended to decrease slowly. To enhance the syngas HHV, methane production can be increased by lowering the reactor temperature.                  110  Chapter 4. STEAM CO-GASIFICATION OF COAL AND SWITCHGRASS MIXTURES IN A BUBBLING FLUIDIZED BED REACTOR  It was shown in  Chapter 2 that for higher biomass-to-fossil fuel ratios (resulting in higher ratios of potassium to aluminum and silicon), potassium acted as a catalyst enhancing fossil fuel gasification. On the other hand, as mentioned in  Chapter 1, one of the main issues related to biomass gasification is the lack of security of feed supply. Therefore in practice, even in biomass/fossil fuel co-feed gasification, much less biomass than fossil fuel on a weight basis is typically fed to the gasifier. In view of the above considerations, it was planned to conduct steam co-gasification of 50:50 wt% coal:switchgrass in a bubbling fluidized bed. Kumabe et al. [82] observed that by varying the ratio of coal to biomass for gasification, the extent of the water-gas shift reaction reached a maximum at the ratio of 0.5. They attributed this to synergy between the coal and biomass. This chapter presents coal/switchgrass steam co-gasification and compares the single-fuel ( Chapter 3) and co-gasification results. As for the single-fuel experiments ( Chapter 3), the Highbury Biofuel Technologies Inc. (HBTI) pilot-scale bubbling fluidized bed (BFB) reactor was used for the co-gasification experiments covered in this chapter. See section  3.2 of  Chapter 3 for the experimental details. 4.1  RESULTS AND DISCUSSION 4.1.1 50:50 coal:spring switchgrass (SP-SG) co-gasification  In our tests, the coal and biomass were well-mixed by a concrete mixer prior to being fed to the hopper. In an effort to prevent segregation inside the feed hopper caused by vibrator-induced vibration of the hopper, the density of some feed samples at different feed rates and durations were analyzed. No significant segregation was observed. The co-feeding to the reactor was performed smoothly without any feeding difficulty. See Appendix B for feeder calibration curves.  111  Figures 4-1 to 4-3 and Table  4-2 present the results for steam co-gasification of a coal:SP-SG 50:50 wt% mixture. Figure  4-1 shows the dry product gas composition vs. time for different feed rates. The reactor reached steady-state after ~ 60 min. Similar to Chapter 3 results,  increasing the feed rate caused the bed temperature to drop sharply due to higher heat demand of endothermic reactions. All dry gas compositions were between the values for the 100% coal and SP-SG experiments (see Figures 3-6 and 3-10 of  Chapter 3). Consistent with Figure  3-7, by increasing the steam-to-fuel ratio, the H2 and CO concentrations increased, whereas the CO2 and CH4 concentrations decreased, as shown in Figure  4-2. Figure  4-3 and Table  4-2 show that the product gas HHV was increased by increasing the feed rate, due to an increase in CH4 concentration. The carbon and cold gas efficiencies and proximate analysis (average) of the entrained solids (see Table  4-1) also lay between the corresponding values for 100% coal and 100% SP-SG. This shows that most of the biomass portion of the mixture, which was more reactive than coal, was fully gasified, whereas the coal portion was only partially converted to product gas. Table  4-1 confirms that the first cyclone captured larger particles (e.g. char) with higher fixed carbon, whereas the second cyclone separated finer solids (e.g. ash) with higher ash/sand percentage. Consistent with Table  4-1, Figures 4-4 and 4-5 present SEM images of particles separated by the first and second cyclones. Large un-reacted coal char particles can be observed in Figure  4-4(a). Comparing Figure  4-4(b) and (c) with Figure  4-5(b) and (c), the coal and biomass chars in the second cyclone were covered with fine particulates, likely biomass and coal ash. After each co-gasification experiment, the reactor cooled down, and the downstream condenser tubes were opened for cleaning. As shown in Figure  4-6, large rigid rounded solid particles were found at the bottom of the first condenser, close to the TDV1 valve (see Figure  3-3). The particles were well rounded. It has been reported [246] that when the gas is cooled in heat exchanger, potassium condenses mainly as sulfate while ash deposits on the heat transfer surfaces. The lower melting point for some potassium species (e.g., KCl melts at ~ 770°C) leads to high risk of formation of hard deposits on furnace walls and convection tubes. Due to the interaction of chlorine and sulfur in ash, complex metal/ash reactions occur, resulting in severe corrosion of the heat transfer tubes [247–252]. Due to the great difference of the elemental composition of coal and biomass, the behaviour of alkali metals during co-combustion differs from that of pure coal or biomass combustion/gasification because of complex interactions  112  between the volatile elements (e.g. K and Na) and other mineral elements [247]. It was found that the major ash-forming elements (Al and Si) significantly influenced this behaviour [253], [254]. Consistent with the results in  Chapter 2, the experimental results for co-combustion showed that condensed alkali species were also significantly more abundant with increasing silicon and aluminum contents of blended fuels, usually in the form of Sanidine (K2O.Al2O3.6SiO2) and Albite (Na2O.Al2O3.6SiO2) [255].      Table  4-1. Coal:SP-SG 50:50 wt% run proximate analyses of particles captured by external cyclones.   Moisture Ash and sand (db) Volatile (db) Fixed Carbon (db) First cyclone  3.5 45.6 1.3 53.1 Second cyclone 7.5 58.7 6.2 35.1 Average 5.5 52.2 3.8 44.1   Figure  4-1. Co-gasification results for dry nitrogen-free product gas composition and bed temperature of coal:SP-SG 50:50 wt% run with different fuel feed rates (db) at atmospheric pressure. Time, t, min0 25 50 75 100 125 150 175Product gas concentration (vol%)0102030405060708090100Temperature, oC600625650675700725750775800825850875900H2COCH4CO2Bed temp.1.18 kg/h(SF=2.58)1.75 kg/h(SF=1.73)1.37 kg/h(SF=2.21)2.33 kg/h(SF=1.30) 113    Figure  4-2. Effect of steam-to-fuel (coal:SP-SG 50:50 wt%) ratio on gas composition at different temperatures and atmospheric pressure.  Figure  4-3. Coal:SP-SG 50:50 wt% co-gasification carbon efficiency, hydrogen efficiency, cold gas efficiency, and product gas higher heating value vs. time for different feed rates (see Figure  4-1) at atmospheric pressure.Steam-to-fuel ratio, SF1.2 1.4 1.6 1.8 2.0 2.2 2.4 2.6 2.8Average product gas concentration (vol%)01020304050Temperature, oC830840850860870880H2COCH4CO2Bed Temp.Time, t, min50 75 100 125 150 175Efficiency0.00.10.20.30.40.50.60.7HHV, MJ/m305101520CarbonHydrogenCold gas HHV   Figure  4-4. SEFigure  4-5. SE(a)M images of pa(a) M images of pa         rticles capturedrticles captured1 mm 2 mm  by first externa by second exteS(b)l cyclone for 50SG char.  (b) rnal cyclone foP-SG char.      200:50 coal:SP-SGr 50:50 coal:SP mm  300 mm  (a) captured so-SG (a) capture(c)lids (b) coal ch(c) d solids (b) coa114 ar (c) SP-l char (c)  200 mm  300 mm  115    Table  4-2. Summary of coal:SP-SG 50:50 wt% co-gasification parameters for different steam-to-coal ratios and atmospheric pressure.  SF 1.30 1.73 2.21 2.58 Temperature (°C)  836 856 857 867 H2 concentration (vol%) 39.4 43.6 46.1 44.4 CO concentration (vol%) 35.0 32.4 29.7 32.2 CH4 concentration (vol%) 13.4 11.8 10.5 11.4 CO2 concentration (vol%) 12.14 12.21 13.64 11.96 Carbon conversion efficiency (ηC) 0.35 0.37 0.42 0.45 Hydrogen conversion efficiency (ηH) 0.20 0.18 0.17 0.16 Cold gas efficiency (ηcg) 0.44 0.48 0.55 0.58 HHV (MJ/Nm3) 14.80 14.35 13.83 14.27 Gas yield (m3/kg fuel) (db and N2-free) 0.71 0.80 0.95 0.98   Figure  4.1.2 5Steam cobubblingproduct dof 2.41. Aspring swtemperatuthe correhigher hySP-SG (spotassiumproduct gconcentraEffect of Reactor tof coal aFigure  4- 4-6. Large r0:50 coal:fa-gasificatio fluidized bery N2-free glmost the sitchgrass cre. Similar sponding vadrogen, caree Figure  4 contents oas HHV wation of the pbed temperaemperature nd biomass9 displays thounded aggll switchgran of a 50:5d at ~ 800°as composiame CO2, lo-gasificatito the previlues of 100bon, and co-3), as showf F-SG coms also highroduct gas.ture is one of th steam gasie effect of tlomerates fo(see ss (F-SG) c0 wt% blenC and atmotion and bedess H2, but on test (comous co-gasif% coal andld gas efficn in Figurpared with er than for a        e most impfication, sinemperature und at bottoFigure  3-3)o-gasificatid of coal aspheric pres temperaturmore CO anpare Figuication run,  F-SG gasiiencies were  4-8. This SP-SG, as ll previous ortant operace the mainon the dry pm of first co. on  nd fall switsure. Figuree vs. time fod CH4 werere  4-1) dueH2 and CO fication expe obtained tis seeminglreported in experimentting variabl gasificationroduct gas cndenser nechgrass was  4-7 and Tar a steam-to produced t to gasificaconcentratioeriments ( Chan for co-y due to thTable  3-1 os, due to thees affecting  reactions omposition ar TDV1 va performedble  4-3 show-fuel molarhan for the tion at a lns were bethapter 3). Mgasificatione higher ashf  Chapter 3 higher metthe performare endothe during the 116 lve  in a  the  ratio coal-ower ween uch  with  and . The hane ance rmic.  117    Table  4-3. Summary of coal:F-SG 50:50 wt% co-gasification parameters at different bed temperatures at atmospheric pressure. Average bed temperature, °C  659 695 729 750 798 SF 2.61 2.26 2.30 2.31 2.41 H2 concentration (vol%) 32.7 37.7 39.3 42.8 42.5 CO concentration (vol%) 38.8 36.3 35.1 32.3 33.4 CH4 concentration (vol%) 14.9 13.9 13.8 12.8 13.2 CO2 concentration (vol%) 13.6 12.1 11.8 12.1 10.8 Carbon conversion efficiency (ηC) 0.36 0.43 0.51 0.65 0.68 Hydrogen conversion efficiency (ηH) 0.10 0.15 0.18 0.25 0.26 Cold gas efficiency (ηcg) 0.40 0.52 0.63 0.84 0.89 HHV (MJ/Nm3) 14.99 14.94 14.93 14.65 14.70 Gas yield (m3/kg fuel) (db and N2-free) 0.64 0.83 1.01 1.37 1.43     118   Figure  4-7. Co-gasification results for dry nitrogen-free product gas composition and bed temperature of coal:F-SG 50:50 wt% with SF=2.41 at atmospheric pressure.  Figure  4-8. Coal:F-SG 50:50 wt% co-gasification carbon efficiency, hydrogen efficiency, cold gas efficiency, and product gas higher heating value vs. time for SF=2.41 at atmospheric pressure. Bed temperature is presented in Figure  4-7. Time, t, min30 40 50 60 70 80 90 100Product gas concentration (vol%)0102030405060708090100Temperature, oC600625650675700725750775800825850875900H2COCH4CO2Bed Temp.1.30 kg/h(SF=2.41)Time, t, min30 40 50 60 70 80 90Efficiency0.00.20.40.60.81.0HHV, MJ/m305101520CarbonHydrogenCold gas HHV 119  steam co-gasification. Similar to Figure  3-16, with increasing the temperature the H2 concentration increased, whereas the CO, CH4, and CO2 concentrations fell slightly due to the dominant endothermic reactions (methane reforming and steam-carbon) resulting in more H2 production. The slight decrease in CO2 concentration with rising temperature may be due to the consumption of CO2 by dry reforming of CH4, and also due to the increase of light hydrocarbons, tars and biomass/coal dry reforming reactions, with temperature in this range [84, 256]. The water–gas shift reaction produces CO2, but results obtained by Gil et al. [257] indicate that when the temperature increases, CO2-consuming reactions would be more important than the shift reaction. The HHV remained almost constant over this range of temperature. The effect of temperature on gasification products compositions has been studied elsewhere (e.g. [84, 232, 256–258]), in good agreement with our data.  Figure  4-9. Effect of temperature on dry nitrogen-free gas composition of coal:F-SG 50:50 wt% co-gasification for SF≈2.4 at atmospheric pressure.  Figure  4-10 portrays the effects of temperature on hydrogen efficiency, carbon efficiency, cold gas efficiency, and gas yield for coal:fall switchgrass 50:50 wt% co-gasification with SF≈2.4 and atmospheric pressure. The trends are similar to those in Figure  3-17 ( Chapter 3) for fall Temperature, oC650 675 700 725 750 775 800Average product gas concentration (vol%)01020304050HHV, MJ/m30510152025H2COCH4CO2HHV 120  switchgrass without any coal. Increasing gas yield with increasing temperature is attributed to: (i) increased production of gas during the initial pyrolysis stage; (ii) steam cracking and reforming of the heavier hydrocarbons and tars; and (iii) enhanced gas production due to endothermic char gasification reactions [19]. The relationship between reaction temperature and gas yield is well supported by the literature [233, 236, 237, 259].  Figure  4-10. Effect of temperature on hydrogen efficiency, carbon efficiency, cold gas efficiency and gas yield of coal:F-SG 50:50 wt% co-gasification for SF≈2.4 at atmospheric pressure.  4.1.3 Comparison of single-fuel and co-gasification   Table  4-4 provides a comprehensive summary of the single-fuel ( Chapter 3) and coal/switchgrass mixtures steam gasification experimental data. Figure  4-11 compares the steam gasification results for 100% coal, SP-SG, F-SG, 50:50 coal:SP-SG, and 50:50 coal:F-SG mixtures in the atmospheric bubbling fluidized bed for a steam-to-fuel mass ratio of ~ 2.4 and (except for one case) similar temperatures. The average dry product gas compositions of coal/SP-SG co-feed gasification were between the corresponding values of the pure coal and SP-SG runs. For the coal/F-SG run, H2 and CO followed the same trend. However, co-gasification resulted in more  Temperature, oC640 660 680 700 720 740 760 780 800 820Efficiency0.00.20.40.60.81.01.2N2-free gas yield (m3 /kg db fuel) 0.40.60.81.01.21.41.61.8Carbon eff.Hydrogen eff.Cold gas eff.Gas yield 121    Table  4-4. Summary of steam gasification data for 100% coal, 100% SP-SG, 100% F-SG, 50:50 coal:SP-SG, and 50:50 coal:F-SG mixtures in the atmospheric pressure bubbling fluidized bed for SF≈2.4.    Experiment Coal (100%) SP-SG (100%) F-SG (100%) Coal:SP-SG (50:50) Coal:F-SG (50:50) Bed temperature (°C) (ave.)  863 858 863 862 798 H2 concentration (vol%) (ave.) 61.2 31.0 36.3 45.3 42.5 CO concentration (vol%) (ave.) 14.5 46.0 39.8 31.0 33.4 CH4 concentration (vol%) (ave.) 8.2 12.3 12.1 11.0 13.2 CO2 concentration (vol%) (ave.) 16.1 10.7 11.8 12.8 10.8 Carbon conversion efficiency (ηC%) (ave.) 15.6 71.5 74.9 43.3 68.0 Hydrogen conversion efficiency (ηH%) (ave.) 11.1 14.7 16.1 16.2 25.5 Cold gas efficiency (ηcg%) (ave.) 26.0 76.9 82.1 56.7 88.5 HHV (MJ/Nm3) (ave.) 12.90 14.66 14.38 14.05 14.70 Gas yield (m3/kg db biomass) (N2-free) (ave.) 0.57 1.03 1.10 0.96 1.43 Tar content (g/m3 dry gas) (N2 included) 6.59 9.61 7.20 5.82 5.54 Tar content (g/m3 dry gas) (N2-free) 21.98 21.28 15.94 13.63 12.96 TOC (ppm) 1514 2014 1998 1545 1810 TSS (ppm) 67.5 57.5 56.8 72.5 70.8     122   Figure  4-11. Comparison of steam gasification results for 100% coal, 100% SP-SG, 100% F-SG, 50:50 coal:SP-SG and 50:50 coal:F-SG mixtures in the atmospheric bubbling fluidized bed for SF≈2.4.H 2 C O C H4C O2H  e f f.C  e f f.C ol d  g as  ef f .G as  yi e ldH HVB ed  te mp .Product gas concentration (vol%) Utilization efficiency (%)N2-free gas yield (10*m3/kg db fuel) HHV (MJ/Nm3)020406080100Temperature, °C 600650700750800850900Coal (100%)SP-SG (100%)F-SG (100%) Coal:SP-SG (50:50)Coal:F-SG (50:50) 123  CH4 and less CO2 than in the corresponding pure fuel runs, due to gasification at a lower temperature (~ 800°C) than in the other runs (~ 860°C). It has been widely reported in the literature that raising the reactor temperature increases the gasification efficiency and energy content of products [238, 260, 261]. However, it is evident from Figure  4-11 that, although the steam co-gasification of coal with F-SG rich in potassium (21.83 wt% in ash) was performed at lower temperature, its hydrogen, carbon, and cold gas efficiencies were enhanced significantly. Its hydrogen and cold gas efficiencies, HHV, and gas yield were even higher than for single fuel gasification. The biomass alkali metals are considered as effective catalysts for H2O and CO2 gasification of carbon [262]. The co-gasification of coal and SP-SG with less potassium (10.76 wt% in ash) showed better performance than expected based on a linear combination of results for gasification of single fuels (similar approach as equation ( 2-4)). The hydrogen efficiency and gas yield were even higher than for the 100% coal and 100% SP-SG gasification tests, consistent with the Kumabe et al. [82] conclusion that the extent of the water-gas shift reaction is maximized at a SF ratio of 0.5 which they attributed to synergy between the coal and biomass. The 100% coal steam gasification experiment showed low carbon and cold gas efficiencies (15.6% and 26.0%, respectively). This can be enhanced by recirculation of entrained coal chars into the fluidized bed. Recirculation can also upgrade the co-gasification runs due to return of biomass ash fine particles into the bed, increasing the carbon-alkali metals solid-solid contact.     Tar and water condensate Figure  4-12 presents the product gas tar yield (with and without N2) for different steam gasification runs, based on the tar sampling standard method explained in section  3.2.3. The average bed temperature of different runs during the tar sampling period is indicated by means of a secondary Y axis. F-SG gasification resulted in lower tar yield due to higher ash and alkali metals content than for SP-SG gasification. Co-gasifying SP-SG with coal and F-SG with coal reduced the tar yields by ~ 28% and ~ 17%, respectively, compared with what would be expected based on a linear combination of the separate fuels data. Although higher bed temperature enhances thermal tar cracking, resulting in lower tar yield, the 50:50 coal:F-SG co-gasification with the lowest average bed temperature yielded the lowest tar due to greater presence of alkali and alkaline earth metals in the fall switchgrass ash. This shows that reducing the gasifier temperature without experiencing higher tar yield can be achieved by the presence of  124  natural catalyst (alkali metals in the switchgrass). Water-gas shift and methanation reactions begin to take place appreciably at about 425°C [17]. Therefore operation of the gasifier at lower temperatures and heat integration between the gasification and shift-methanation are desirable economically, favouring methane formation, improving thermal efficiency and increasing the heating value of the product gas [263].     Figure  4-12. Product gas tar yield (N2 included and N2-free) and average bed temperature for steam gasification of 100% coal, 100% SP-SG, 100% F-SG, 50:50 coal:SP-SG, and 50:50 coal:F-SG mixtures in a bubbling fluidized bed (atmospheric pressure).  Many previous authors have found that alkali metal catalysts are effective in reforming tars [264–268]. For example, McKee [267] successfully demonstrated that carbonates, oxides and hydroxides of alkali metals can effectively decompose tar during catalytic gasification. Simell and Leppälahti [129] related the catalytic activity for tar elimination of calcined rocks to several factors, such as large pore size and surface area of corresponding calcinates and relatively high alkaline (K, Na) content. Alkaline metals could act as promoters in commercial steam-reforming Coal (100%)SP-SG (100%)F-SG (100%)Coal:SP-SG (50:50)Coal:F-SG (50:50)g/N m3  dry gas0510152025Temperature, °C 700725750775800825850875900N2 includedN2 freeBed temp. 125  catalysts by enhancing the gasification of carbon intermediates deposited on the catalyst surface. Elliott and Baker [269] used 8 wt% potassium carbonate as a bed additive impregnated on wood for fluidized bed steam gasification at 750°C. They observed a reduction of phenolic tar compounds by a factor of 5 and in polycyclic aromatic hydrocarbons (PAH) by a factor of ~ 10. It was observed that alkali metals, especially potassium, acted as a promoter in unzipping cellulose chains during thermal decomposition of woody biomass, thus affecting the product gas distribution [270].  The allowable tar content for internal combustion (IC) engines for power generation is <100 mg/Nm3. Therefore, extra thermal or chemical tar cracking treatments are required to enhance the quality of the product gas before introducing it to an engine. Condensed water samples from the steam gasification runs were analyzed by the standard measurement procedures briefly explained in section  3.2.4 to determine the total organic carbon and total suspended solids, with the results shown in Figure  4-13. As the gas yields of the pure biomass gasification runs were higher than for the pure coal, the TOCs for the pure biomass experiments were greater than for the coal steam gasification. The TOCs of the coal and switchgrass blend samples were between those for pure coal and pure biomass. The coal/F-SG co-gasification resulted in higher TOC, seemingly due to higher gas yield, than for coal/SP-SG gasification. Typically, tars containing heterocyclic aromatics with high polarity are soluble in water (e.g. pyridine, phenols and cresols).    As a side-project, the coal:SP-SG 50:50% water sample was treated to reduce its organic carbon content using ion exchange resin. Appendix E provides details.  Since Quinsam mine coal was crushed to finer particle sizes than the biomasses (see Figure  3-1), a greater amount of suspended solid particulates was found in the condensed water from 100% coal gasification than 100% biomass experiments.    126   Figure  4-13. Condensed water TOC and TSS for steam gasification of 100% coal, 100% SP-SG, 100% F-SG, 50:50 coal:SP-SG, and 50:50 coal:F-SG mixtures in atmospheric bubbling fluidized bed. Temperatures are given in Figure  4-12.  4.2  SUMMARY (1) Blends of 50:50 wt% coal:SP-SG and 50:50 wt% coal:F-SG were steam gasified in a pilot-scale bubbling fluidized bed reactor. The effect on product gas efficiencies, yield and heating value of co-gasificating coal with switchgrass rich in potassium were investigated.  (2) All co-feed dry gas compositions were between the values for the 100% coal and switchgrass experiments. (3) The hydrogen, carbon and cold gas efficiencies, gas yield and HHV were enhanced significantly during co-gasification of coal and fall switchgrass. (4) The gas tar yield was found to be reduced by ~ 28% and ~ 17% (compared with what would be expected based on a linear calculation of the single fuels data) when SP-SG or F-SG, respectively, were co-fed with coal to the gasifier. The results indicate that Coal (100%)SP-SG (100%)F-SG (100%)Coal:SP-SG (50:50)Coal:F-SG (50:50)ppm(mg/L)05001000150020002500TOCTSS (data multiplied by 10) 127  biomass ash alkali metals can act as inexpensive natural catalysts for steam gasification and tar cracking. Lower tar yield due to catalytic cracking facilitates operation at lower reactor temperatures for higher thermal efficiency.    (5) Increasing the gasifier bed temperature increased the H2 concentration, gas efficiencies and yield, whereas the CO, CH4 and CO2 concentrations tended to decrease. To enhance the product gas heating value, CH4 production can be increased by lowering the reactor temperature. Switchgrass alkali metals can assist in operating at lower temperatures without being penalized by a rise in product gas tar content. The hydrogen concentration can also be increased by augmenting the steam-to-fuel ratio.                    128  Chapter 5. LIME ENHANCED STEAM CO-GASIFICATION OF COAL AND SWITCHGRASS IN A BUBBLING FLUIDIZED BED REACTOR  It was shown in  Chapter 4 that adding switchgrass to coal can enhance steam gasification by reducing the product gas tar and augmenting the gasification efficiency, product gas heating value and yield of useful products. Biomass/coal co-feed gasification could become a more sustainable process if most or all of the CO2 produced could be captured from the system. This could be achieved by capturing the CO2 in-situ, in the same bed where gasification is carried out [24]. Integration of CO2 capture and storage technologies with biomass gasification could even result in net removal of CO2 from the atmosphere [271], often referred to as “negative emissions”.  CO2 capture during gasification via CaO(s)+CO2(g)↔CaCO3(s) can also shift several reactions (e.g. water-gas shift, steam-carbon and steam reforming) towards more H2 production, thereby improving product quality. The lime-enhanced gasification (LEG) process is an integrated CO2 capture process in which CaO is employed as a high-temperature CO2 carrier between two environments: a steam gasifier and a regenerator (calciner). Adsorbed CO2 captured by lime is transferred to the regenerator as CaCO3 where the sorbent is calcined, generating a CO2-rich stream suitable for storage. Furthermore, the rejected CaO/ash mixture could be a valuable by-product for the cement industry as a pre-calcined raw meal, offering large energy savings [114]. Co-gasification coupled with in-situ CO2 capture using CaO also exploits heat generated by the exothermic carbonation reaction to provide the heat required for the endothermic gasification reactions. The energy requirements associated with the regeneration of the sorbent must, of course, be taken into account when estimating the overall process efficiency. In this regard, the transfer of heat between the gasifier and sorbent regenerator is critical to maximizing overall process efficiency [19]. Gasifiers, especially coal gasifiers, operating at lower temperatures have incomplete carbon conversion. An additional advantage of the LEG process is the possibility to use the remaining char from the gasifier for supplying part of the regeneration heat demand. Therefore, high energetic efficiencies for the overall cycle can be achieved without high carbon  129  conversion in the gasifier [114]. Alternatively a portion of the H2 generated may be oxidized to generate the heat needed for the regeneration [272]. Limestone has some significant advantages as a solid sorbent due to its abundance, low cost, high reactivity as lime in the temperature range 500-700°C, and low vulnerability to chemical poisoning. CaO is also effective for sulfur, chlorine and mercury capture and can catalyze tar cracking and ammonia oxidation reactions.    Given their ability to exchange particles, fluidized bed reactors are likely to be most suitable for the proposed process. The rapid solid motion and mixing in fluidized beds provide good heat and mass transfer and superior temperature uniformity compared to fixed beds [16, 273]. These characteristics promote high fuel conversion and high sorbent utilization for CO2 capture. Hence, in this thesis project, steam co-gasification of switchgrass and coal integrated with CO2 capture was investigated in a bubbling fluidized bed (BFB) reactor with limestone as the bed material.  5.1  EXPERIMENTAL  5.1.1 Experimental apparatus and operation A blend of 50:50 wt% coal and spring switchgrass (SP-SG) was chosen as the fuel, as fall switchgrass (F-SG), richer in potassium than SP-SG was not available at the time of the experiments. Figure  3-1 and Table  3-1 of  Chapter 3 provide the particle size distribution and ultimate, proximate and ash analyses of both fuels. In this study, gasification was integrated with both carbonation (CO2 capture) and calcination in a single reactor by conducting experiments in a semi-batch cyclical manner. The blended fuel was fed to a fluidized bed loaded with solid sorbent particles as bed material in the gasification/carbonation stage. When the CaO had been fully utilized and the product gas compositions became relatively stable, the process was then switched to calcination by stopping feeding the fuel and steam to the reactor, and raising the reactor temperature. A portion of the heat needed for calcination was supplied by introducing air to fully combust unreacted char remaining in the bed from the gasification stage. The reactor temperature was controlled by the N2-to-air ratio during this period. Absence of CO2 in the product gas indicated that calcination of bed material was completed. The system was then again switched to the gasification/carbonation  130  mode for the next cycle. Using a bubbling fluidized bed reduced particle attrition and entrainment due to operation at modest gas velocities relative to circulating fluidized beds.    A lime-enhanced gasification experiment was first conducted in the same HBTI bubbling fluidized bed unit as the experiments covered in  Chapter 3 (see Figure  3-3) using 300-355 µm (US mesh #45-50) Strasburg limestone as the bed material. However, this run was unsuccessful due to blockage in the downstream piping due to high lime attrition and elutriation. The HBTI system is large enough that it was practically impossible to un-plug the piping during the experiment. Therefore, it was decided to move to a smaller bubbling fluidized bed set-up initially constructed by Sakaguchi [274, 275] located in the UBC Clean Energy Research Centre (CERC).  Prior to performing the experiments, different sections of the unit (e.g. after burner heater, screw feeder piping, tar sampling port and downstream filter) were modified/repaired. As shown in Figure  5-1, this reactor was constructed of 310 stainless steel, and comprised of a 76 mm ID, 707 mm tall, atmospheric-pressure fluidized bed reactor, with a distributor plate and internal cyclone of 50 mm ID. A dip leg returned solids to the lower bed. The reactor tube was electrically heated by four radiative clam-shell heaters, two heating the pre-heater and lower bed, and the other two heating the upper half of the fluidized bed. The distributor was a perforated plate of 3.2 mm holes, 6.4 mm apart, 95 in total. This plate was covered with a single layer of stainless steel screen, 400 mesh, to prevent solids from dropping into the pre-heater [174]. Appendix C of Sakaguchi [275] provides detail dimensions and specifications of this facility. The blended coal-switchgrass mixture was fed from the side of the fluidized bed at a height of 100 mm above the distributor through a screw feeder, which was fully surrounded by a water cooling jacket to keep the feedstock temperature below 80°C to avoid plugging. The screw was 12.7 mm in diameter, 600 mm in length and drew the fuel from a Schenck Accu-Rate Tuf-Flex 300 feed hopper which was slightly pressurized with a low flow of nitrogen (~ 0.12 L/min) to prevent back-flow. All gas flows were controlled with the aid of rotameters. The feedstock was gasified using steam, generated by pumping de-ionized water through a steam generator where it was heated to ~ 250°C, then into the pre-heater, which was held near the bed temperature. Nitrogen was the fluidizing gas, in addition to steam, to ensure consistent fluidization on the small scale. Nitrogen and steam were mixed in the pre-heater before entering the bed.    Figure  5 Thirteen the axis condensethe reactothe condthermocoafter-burnAgilent 4the operaprocedurThe bed heat requ-1. SchematK-type therof the reacr. Fine charr, and a SHenser. Produuple and a er (~ 600°C900 gas chrting condite and calibratemperatureired for calcic of CERC mocouples wtor. Excess  particles gELCO FOSct gas flowpressure tra). The comomatographions. Appention curves was held aination wasfluidized beere deploysteam in thenerated in BN-78 sing was measnsducer. Thposition of  (GC) withdices 14 an, respectivels close as p supplied byd experimen[275]).  ed throughoe product gthe reactor le-cartridgeured by a ve producedthe productargon and hd 17 of Buy.  ossible to 70 feeding airtal apparatuut the reactas was conwere captur filter (10 molumetric fgas was the gas was melium as cartler [174] p0°C during (see Tables (modifiedor system, idensed anded by an inicron fabriclow meter cn combusteeasured usinrier gases. Srovide the  the gasific  5-1) into th from Sakagncluding fiv separated ternal cyclo) downstreaombined wd with air ig a four-coee Table  5-system operation stagese bed in ord131  uchi e on by a ne in m of ith a n the lumn 1 for ating . The er to  132  combust the residual char. The bed pressure was close to atmospheric, in the range of 10 - 30 kPag measured at three positions: pre-heater, lower bed (50 mm above the distributor), and freeboard (400 mm above the distributor). The pressure was limited to below 50 kPag, due to the low pressure feed hopper.  Table  5-1. Operating conditions in CERC gasifier. The hydrodynamic parameters based on steam at 525°C and 1 atm. Bed material (limestone) hydrodynamic parameters ρb (kg/m3) (bulk)  1350 ds (µm) 563 Ut (m/s)* 1.98 Umf (m/s)** 0.21 Fluidization parameters Operating gas velocity (m/s) 0.29 Steam flow rate (m3/h) 3.91 N2 flow rate (fluidization and hopper) (m3/h at ambient) 0.87 Gasification/Carbonation conditions Fuel feed rate (kg/h) 0.45 Steam-to-fuel mass ratio (SF) 2.56 Calcination conditions Air flow rate (m3/h at ambient)*** 0.54 * Calculated based on Haider and Levenspiel [227] correlation, ** Calculated based on Grace [228] equation, *** N2 was also entered to control the reactor temperature during char combustion.   The static bed height at the beginning of each run was ~ 30 cm, calculated based on the mass of solids added. Strasburg limestone was the solid sorbent and bed material. See Table  5-2 for its composition. As the internal cyclone was not efficient enough for complete solids separation, to reduce particle elutriation relatively large lime particles (563 µm volumetric weighted average) were used for the CERC experiments, coarser than the sand displayed in the HBTI experiments described in  Chapter 3. The particle size distribution was analyzed using a HYDRO Mastersizer 2000, with nitrogen as the carrier gas. Results are shown in Figure  5-2. Prior to the experiments, the limestone was calcined at 900°C for 15 h in an oven with 100 mL/min N2. The particles were then promptly loaded into the reactor to prevent hydration.   Measurements with nitrogen were performed for fresh limestone samples, as well as for calcined samples from each cycle, to determine surface areas using a Micromeritics ASAP 2020 surface  133  area and porosity analyzer. The main source of adsorption measurements errors came from sample weight measurements, as well as from electrostatic charges acquired during loading of the samples into the measurement tubes. Prior to the adsorption measurements, the samples were degassed at 105°C and 0.15 mbara for 24 h.  Table  5-2. Strasburg limestone material properties (wt%).  CaO MgO Al2O3 Fe2O3 K2O Na2O SiO2 LOI* Limestone 53.7 1.25 0.19 0.94 0.08 0.02 0.94 42.9 * Loss on ignition   Figure  5-2. Limestone particle size distribution, determined by Mastersizer 2000.  Tar sampling was conducted during three cycles of gasification/calcination, following the same procedures as described in section  3.2.3.  Particle diameter, μm1 10 100 1000Mass fraction, %02468 134  5.1.2 LEG indices The hydrogen, carbon, and cold gas efficiencies were calculated in similar manner to  Chapter 3 (see section  3.2.6). Reported error bars are based on 95% confidence intervals for population means (assuming normal distributions). Utilizing four columns, the GC was able to measure the composition of C2+ hydrocarbons. The heating value of the product gas was then estimated based on a modified version of equation ( 3-4):    2 42 6 2 4 3 6 3 8 4 8HHV 12.7H % 12.6CO% 39.8CH %70C H % 63C H % 92C H % 100C H % 123C H %= + + ++ + + +  ( 5-1) where the species heats of combustion are in MJ/Nm3. This equation is based on heat of combustion data [231], with ideal-gas behaviour assumed for the gaseous species. The calcium utilization is defined as the moles of CO2 captured over the moles of CaO (or Ca) available, given [102] by the equation:  23 0,0 0( )*CaCO t CaOCaO CaO COn m m MXn m f M−= =  ( 5-2) where fCaO is the mass fraction of CaO in the sorbent, m is the mass of the sorbent at time 0 or t, and M is the molecular weight of the different species. To calculate the calcium utilization, particulate samples were collected from the bed during each stage of cycling and analyzed using the Thermax500 TGA.  5.2  RESULTS AND DISCUSSION Limestone was calcined twice, once in an oven at 900°C as explained above, and once after being loaded into the CERC BFB reactor. It was found that even after calcination in the oven for 15 h, some particles still contained CO2, as shown in Figure  5-3. This figure shows that when the equilibrium CO2 partial pressure exceeded the actual CO2 partial pressure, calcination occurred.   135   Figure  5-3. CO2 concentration, difference between equilibrium and actual CO2 partial pressure and temperatures during calcination of sorbent material in the CERC BFB.   In total, five cycles of gasification/carbonation and calcination were performed in the CERC BFB. As explained above, gasification cycling was accomplished by feeding the 50:50 coal:SP-SG mixture until the end of the breakthrough period, when the syngas composition became almost steady. The reactor was kept as close as possible to 700°C during each gasification/carbonation stage. Figure  5-4 presents the evolution with time of the product gas composition and reactor temperatures for the gasification/carbonation stages for cycles 1 and 2. The measured C2+ included ethylene (C2H4), ethane (C2H6), acetylene (C2H2), propylene (C3H6), propane (C3H8), 1-butene (C4H8), i-butene (C4H8) and butane (C4H10). Periods of high H2 production and decline in CO2 concentration until t≈50 min for cycle 1 and up to t≈45 min for cycle 2 signify in-situ CO2 capture during co-gasification of coal and switchgrass. During this period, the exothermic carbonation reaction assisted in providing heat needed for the endothermic gasification reactions, and therefore the lower bed temperature (T8 in Figure  5-1) was relatively stable. However, when the sorbent particles utilized for CO2 capture approached their carbonation limit, the lower bed temperature started to fluctuate.  Time, t, min0 25 50 75 100 125 150 175Temperature, oC550600650700750800850900CO2 gas concentration (vol%)Peq,CO2-PCO2, kPa0510152025303540T8 - Lower Bed (150 mm)T7 - Mid Bed (200 mm)T6 - Upper Bed (250 mm)CO2 concentrationPeq,CO2-PCO2 136    Figure  5-4. Dry N2-free product gas compositions and temperatures during gasification/carbonation in the CERC BFB at 1 atm: (a) Cycle 1 (b) Cycle 2. For cycles 3 to 5, see Appendix F.  Time, t, min0 5 10 15 20 25 60 75 90 105 120Product gas concentration (vol%)0102030405060708090100Temperature, oC0100200300400500600700800H2CH4COCO2C2+T8 - Lower Bed (150 mm)T7 - Mid Bed (200 mm)T6 - Upper Bed (250 mm)(a) Cycle 1Time, t, min0 10 20 30 60 80 100 120Product gas concentration (vol%)0102030405060708090100Temperature, oC0100200300400500600700800(b) Cycle 2 137  The T8 thermocouple is located directly above the feed port where cold coal and biomass were introduced to the reactor, causing a localized decrease and fluctuation in the lower bed temperature. Thermocouples at 200 and 250 mm did not experience the same degrees of temperature drop and fluctuation. After the sorbent solids were mostly saturated with CO2, the H2 concentration started to decline, while the CO2 concentration began to rise. Other authors (e.g. [174, 259]) reported similar behaviour during batchwise gasification/carbonation in a fluidized bed. Appendices F and G provide complete experimental results for all five cycles. While performing all five cycles of the experiments, plugging occurred in the condenser tube. The condenser design resulted in periodic increases in bed pressure across the bed, as shown in Figure  5-5, requiring interrupting the experiment (i.e. stopping the feeding of fuel and steam to the reactor) and cleaning the condenser tube and downstream filter. Figure  5-6 shows char and tar accumulation at the entrance of the downstream filter during the blockage period of Cycle 1 gasification/carbonation. The “breaks” in the abscissa-axes of Figure  5-4 correspond to these plugging periods. After the condenser was un-plugged, the experiment was resumed by re-introducing steam and fuel to the reactor. The product gas compositions fluctuated until steady-state was achieved and maintained until the capacity of the sorbent for CO2 capture was reached. A new condenser was designed and will be used for future experiments (see Appendix 20 of Butler [174] for more details regarding the condenser design). Another factor that likely affected the results (e.g. the fluctuations in gas compositions) was segregation of the bed materials due to their high density difference. According to Rowe et al. [276], the dependence for segregation tendency is  0.22.5Segregation tendency ( ( ) ) J Bmf FF SdU Udρρ−− ⎛ ⎞⎛ ⎞∝ − ⎜ ⎟⎜ ⎟⎝ ⎠ ⎝ ⎠  ( 5-3) where U and (Umf)F are operating and flotsam minimum fluidization velocities, ρJ and ρF are jetsam and flotsam particle densities, and dB and dS are particle sizes of the big and small particles, respectively1. According to equation ( 5-3), the effect of particle density difference is                                                  1 The flotsam is the lighter or smaller component which tends to remain in the top layer of the bed upon fluidization, while the jetsam is those heavier or larger components, which tends to stay in the bottom part of the fluidized bed.    considerasize is to Figure  Figur bly more imalter the Um 5-5. Initial ae  5-6. Char Reactor pressure drop, kPa012345portant thaf of the mixtnd final reaand tar accuCycle 1InitialFinaln that of parure.   ctor pressurfive mulation atGasificatioCycle 2ticle size die drops duriLEG cycles entrance ofn/CarbonationCycle 3 Cfference. Thng gasificati.    downstream periodycle 4 Cyce main effeon/carbonat  filter causile 5ct of the pa ion periodsng blockage138 rticle for .    139  In the course of the experiment, the CaO particles gradually carbonated to denser CaCO3 particles. Hence, segregation could be worse as time advanced. The mixing index was calculated based on Nienow et al. [277] correlations for binary systems assuming the limestone as jetsam and the switchgrass as flotsam solids. The estimated mixing index, M, of the LEG experiments was ~ 0.23, close to complete segregated bed (M=0) and far below the state of perfect mixing (M=1).      The hydrogen and cold gas utilization efficiencies decreased as the bed deactivated, resulting from the reversal of the water-gas shift (reaction ( 1-9)), steam-carbon (reaction ( 1-7)) and steam reforming (reaction ( 1-11)) reversible reactions, lowering the conversion of H2O, see Figure  5-7. However, the carbon utilization efficiency remained high during the experiment. As there was no way to measure the instantaneous uptake of CO2 on CaO, it was averaged over the entire cycle. The carbon and cold gas efficiencies were less than for the corresponding co-gasification results (without carbon capture) presented in  Chapter 4 (compare Figure  4-11) due to higher particle entrainment because of poor cyclone performance and shorter reactor height than for the HBTI facility. Because of the decrease in H2 and increase in CO2 concentrations in the late period of co-gasification, the HHV gradually declined. During gasification, char accumulated in the bed and on the bed material. Figure  5-8 shows a photograph of large char/lime/ash agglomerates produced during limestone enhanced co-gasification of coal and biomass. The catalytic effect of CaO on tar cracking led to accumulation of char on the particle surfaces. This adhering char would be of benefit to a dual-bed CO2 capture system, as more char would be transported with the bed material to the combustion reactor [174]. The formation of melts involving mixtures of CaO, CaCO3 and Ca(OH)2, at temperatures significantly lower than the melting temperatures for the pure components were reported in  Chapter 1 (section  1.5). The formation of low-temperature melts hampered operations by causing blockages in equipment, whilst agglomeration of sorbent and fuel particles is expected to limit gas-solid interactions, hence reducing CO2 capture and fuel conversion efficiencies. Thus, melt formation occurs imposes additional constraints on the LEG system [19].  After gasification, fuel feeding was stopped, the steam feed was replaced by air, and the reactor temperature increased to calcine and regenerate the sorbent material in the bed, as shown in Figure  5-9 for cycles 1 and 2. The bed temperature increased rapidly as a result of combustion of   residual flows of completewhich hastream, thFigure  5efficiencFigure char inside air and nitroly combustd been capis indicated-7. Cycle 1 y, and HHV 5-8. Agglom0Efficiency0.00.20.40.60.81.0the reactor. gen. A droped. Then, ctured during that calcinaco-gasificati vs. time. Teration of cHydCaCoHHThe temper in CO concalcination o gasificatiotion of the bon/carbonatemperaturesin har, lime ancaT20rogenrbonld gasVature was thentration inf the bed sn. When theed materialion hydroge are shown iTable  5-1. d ash after 5lcination.  ime, t, min60en controlldicated thatorbents prore was no  had been con efficiencyn Figure  5-4 cycles of g80ed by regul the residualceeded to more CO2 impleted.   , carbon effi, other operasification/c100 120ating the rel char was alremove the n the outpu ciency, coldating condit arbonation HHV, MJ/m305101520140 ative most CO2 t gas  gas ions and  141      Figure  5-9. Dry product gas compositions, difference between equilibrium and actual CO2 partial pressure and temperatures during calcination in CERC BFB: (a) Cycle 1 (b) Cycle 2. For cycles 3 to 5, see Appendix G.   Time, t, min0 20 40 60 80 100 120Product gas concentration (vol%)Peq,CO2-PCO2, kPa0102030405060708090100Temperature, oC01002003004005006007008009001000(a) Cycle 1Time, t, min0 25 50 75 100 125 150 175Product gas concentration (vol%)Peq,CO2-PCO2, kPa020406080100120140160Temperature, oC1002003004005006007008009001000(b) Cycle 2H2CH4COCO2O2N2T8 - Lower Bed (150 mm)T7 - Mid Bed (200 mm)T6- Upper Bed (250 mm)Peq,CO2-PCO2 142  Although LEG was performed at much lower bed temperature (~ 586°C on average for the five cycles) than for the corresponding sorbent-free steam co-gasification test in the HBTI unit (sand as the bed material) presented in  Chapter 4 (~ 862°C), the hydrogen production (averaged over the entire period of gasification) was enhanced by ~ 22%. Figure  5-10 summarizes the lime-enhanced 50:50 coal:SP-SG co-gasification for 5 cycles. It is evident that H2 concentration and hydrogen utilization efficiency declined due to the decrease in calcium utilization over the course of cycling, whereas the CO, CO2, CH4 and C2+ concentrations rose. The lower calcium utilization efficiencies in cycles 1 and 5 were due to the decrease in co-gasification duration in those cycles of ~ 90 and 86 min, respectively, compared with gasification/carbonation over the periods of 109-118 min for cycles 2 to 4. The highest calcium utilization efficiency was for the 3rd cycle (~ 65.6%), probably due to carbonation at the lowest bed temperature (~ 568°C) compared to the other cycles. The carbon utilization efficiency varied between 18 and 23%, showing no clear trend with cycle number, consistent with data reported by Butler [174]. The product gas higher heating value was sensitive to CH4 and C2+ concentrations. The peak syngas concentrations for the five cycles, see Figure  5-11, had a similar trend as the time-averaged data presented in Figure  5-10. The H2 concentration decreased by a small percentage between the 1st and 5th cycles. This slight decrease can be attributed to the reduction in CO2 capture effectiveness of the bed. The maximum HHV increased in Cycle 5 relative to Cycle 1 due to higher peak CH4, CO and C2+ concentrations.  Lime surface area analysis The surface areas of the original Strasburg limestone, oven-calcined lime and bed samples after completion of each cycle were measured using BET nitrogen adsorption, as shown in Figure  5-12. The nitrogen isotherms of the surface area analysis are provided in Appendix H. It is evident that the sorbent specific surface area increased due to calcination in the oven. The surface area was enhanced more by calcination inside the reactor before beginning the cycling, showing that 15 h oven heat treatment calcined the limestone particles partially, consistent with Figure  5-3. The data indicated a general drop in surface area during the course of cycling. The loss of sorbent activity during cycling is attributed to sintering causing surface area reduction during high-temperature regeneration. The cumulative volume of pores for the 5th cycle sample was close to that of the original limestone, but pores became larger, resulting in a decay of the   143    Figure  5-10. Comparison of experimental concentrations and performance indices for LEG cycles with 50:50 coal:SP-SG mixture in the CERC atmospheric fluidized bed for SF≈2.6. Averages were calculated over entire period of co-gasificaiton.  H 2 C H4C O C O2C 2+H  e f f.C  e f f.C ol d  g as  ef f .C al c iu m u ti l i za t io n e ff .G as  yi e ldH HVT 8 -  L ow er  Be d ( 15 0 mm )Time averaged product gas concentration (vol%) Utilization efficiency (%)N2-free gas yield (100*m3/kg db fuel) HHV (MJ/Nm3)010203040506070Temperature, °C 0100200300400500600700Cycle 1Cycle 2Cycle 3Cycle 4Cycle 5 144   Figure  5-11. Effect of sorbent cycling on peak product gas composition (dry, N2-free) and HHV of 50:50 coal:SP-SG steam co-gasification at atmospheric pressure.    CaO surface area in the course of cycling due to the weak mechanical structure of lime particles. A number of researchers have reported a decay in sorbent reactivity, accompanied by a significant decrease in pore volume and surface area [97, 98, 278–281]. It is well known that sintering of CaO during the sorbent cycling process (calcination/carbonation) leads to a decrease in surface area and porosity and an increase in average crystal size [120, 121].  Pore size distributions, calculated by the BJH method [282] based on N2 adsorption, are plotted in Figure  5-13. With the exception of the original limestone, there is meso porosity in the samples. The volume of the pores decreased during the course of cycling (except for Cycle 4), consistent with the decrease in N2 uptake, see Figure  5-12. Consistent with the BET surface area results, Figure  5-14 shows that calcination at high temperatures sintered the sorbent particles after five cycles. Following Bhatia and Perlmutter [283], Sun et al. [99] and Abanades and Alvarez [97] concluded that the decay in reactivity over multiple cycles was due to a decrease in micro-porosity and an increase in meso-porosity. Their argument was supported by a qualitative Cycle 1 Cycle 2 Cycle 3 Cycle 4 Cycle 5Product gas concentration (vol%)0102030405060708090Temperature, oC610620630640650660670680HHV, MJ/m3121314151617181920H2COCH4CO2C2+Bed temp. (150 mm)HHV 145  SEM study, consistent with the photographic images. SEM images of all five cycles are provided in Appendix I.   Figure  5-12. BET surface areas and BJH cumulative pore volumes of sorbent samples.   Figure  5-13. BJH N2 adsorption pore volume distribution of sorbent samples.  Original limeAfter ovenCalcined in the reactorCycle 1Cycle 2Cycle 3Cycle 4Cycle 5BET surface area, m2 /g 0.00.51.01.52.02.53.03.5BJH adsorption cumulative volume of pores, cm3 /g 0.0000.0020.0040.0060.0080.0100.0120.0140.0160.018BET surface areaBJH pore volumePore width, Å0 50 100 150 200 250 300 350 400 450 500 550 600Pore volume, cm³/g·Å 01e-52e-53e-54e-55e-56e-5Original limeAfter ovenCalcined in the reactorCycle 1 Cycle 2Cycle 3Cycle 4Cycle 5   Cycle 1 Cycle 5  Tar measCycles 1Figure  5presumaband loss activity ocorresponmuch higis well k[122–124for tar crFigure  5-14urement  , 2 and 5 ta-15. The taly because of fine sorf calcium wding tar yieher temperanown to hav, 127–129].acking has b5k X magn. SEM imar yields of tr yield incof a decreasbent particlas found told of the satures (~ 86e a catalyti The use ofeen discusseification ges of calcinhe product reased frome in bed temes from the decrease dund bed co-g2°C), indicac effect on  CaO or cald in detail e  ed bed sorbgas during g 4.4 g/m3perature le reactor in ring gasificasification pting that Cathe decompocined dolomlsewhere [110ents after firasification/in Cycle 1ading to lesthe course ation [50]. resented in O catalyzedsition of taite in-situ o25, 130, 13k X magnifist and fifth carbonation to 10.5 gs tar thermaof cycling.The results  Chapter 4 ( tar crackinr species dur downstrea1]. cation cycles.  are present/m3 in Cycl decompos In additionare similar tsee Table  4g reactions.ring gasificm from gas146   ed in le 5, ition, , the o the -4) at  CaO ation ifiers  147   Figure  5-15. Product gas tar yield and average bed temperature for 1st, 2nd and 5th cycles of lime enhanced steam co-gasification of 50:50 coal:SP-SG mixture in atmospheric CERC BFB.   Particle size distribution The particle size distribution of the calcined bed material after each cycle was analysed using the Mastersizer. Results are shown in Figure  5-16. A shift towards finer particles is evident due to attrition of solid sorbent during fluidization. It should be noted that these results were based only on the bed material samples, so that entrained fines are excluded. The volumetric weighted average particle size dropped by ~ 25% from 563 µm (original sample) to 427 µm at the end of the fifth cycle.  Cycle 1 Cycle 2 Cycle 5Tar yield, g/m3  dry gas024681012Temperature, oC540560580600620640Tar yieldBed temp. (15 cm) 148   Figure  5-16. Particle size distribution of original limestone and calcined bed material after 1st, 3rd and 5th cycles.   5.3  SUMMARY (1) Lime enhanced steam co-gasification of a 50:50 wt% coal/spring switchgrass mixture in a bubbling fluidized bed was studied for five cycles of gasification/carbonation and calcination/regeneration. Using CaO instead of silica sand as the bed material, the steam co-gasification hydrogen production increased by ~ 22%. The hydrogen utilization efficiency also increased as a result of the carbonation. The experimental data indicated that char combustion could supply the heat required for the endothermic CaO re-generation reaction. The calcium utilization efficiency of the sorbent decreased somewhat with calcination/carbonation cycling. (2) Blockage of the condenser tube due to high entrainment of sorbent and char fines was a major issue during the experiments, interrupting operation.  (3) The sorbent decayed during cycling due to sintering of CaO pores as a result of high-temperature calcination. This led to a decrease in the sorbent specific surface area and cumulative pore volume. SEM images confirmed sintering of sorbent particles. Particle diameter, μm1 10 100 1000Mass fraction, %02468Original lime - Vol. wei. ave. size: 563 μm   Cycle 1 - Vol. wei. ave. size: 487 μm   Cycle 3 - Vol. wei. ave. size: 457 μm   Cycle 5 - Vol. wei. ave. size: 427 μm    149  (4) Tar yield measurements showed that reducing the gasification temperature can be achieved without experiencing higher tar yield aided by the presence of CaO sorbent particles as the bed material, likely due to Ca catalysis of tar cracking reactions. (5) The weak mechanical structure of CaO led to attrition during fluidization of solid particles. The average particle size dropped by ~ 25% from 563 µm (for the original limestone) to 427 µm at the end of the fifth cycle. Particle attrition led to increased fines entrainment and loss of the CO2 sorbent during the course of cycling.                 150  Chapter 6. EQUILIBRIUM MODELING  At chemical equilibrium, a reacting system is at its most stable composition, a condition achieved when the entropy of the system is maximized, while its Gibbs free energy is minimized. Two approaches have been developed for equilibrium modeling: stoichiometric and non-stoichiometric [284]. The stoichiometric approach requires a clearly defined reaction mechanism incorporating all chemical reactions and species involved. In non-stoichiometric formulations, on the other hand, no particular reaction mechanisms or species are involved in the numerical solution. The only input needed to specify the feed is the elemental composition, which can be readily obtained from ultimate analysis data. It is particularly suitable for systems with unclear reaction mechanisms and feed streams like biomass whose precise chemical compositions are unknown [285]. Both methods were applied in this project, and similar results were obtained. 6.1  SINGLE-FUEL AND CO-GASIFICATION EQUILIBRIUM MODELING The Aspen Plus V7.2 process simulator (solids simulation class) was used for the equilibrium simulation. Figure  6-1 displays the process flowsheet. It considers three stages: reaction, cooling, and separation. The Peng-Robinson model [286] was used for the simulation. First, the ultimate and proximate analyses of the fuel (switchgrass, coal, or mixed) were specified in the FUEL stream. All species were defined as “conventional”, except for C as solid and Ash as a “non-conventional (NC)” component. The FUEL stream then entered the decomposition (yield) reactor where the fuel was converted into its constituent components - carbon, hydrogen, oxygen, sulphur, nitrogen, and ash - by specifying the yield distribution according to the fuel ultimate analysis. The heat generated from the yield reactor was transferred to the endothermic gasifier. The gasifier was then fed by the FEED and STEAM streams, the reactor temperature and pressure were determined, and the system Gibbs free energy was minimized to yield the products. The PRODUCT stream was next cooled to 25°C using a shell-tube heat exchanger. Ash particles were separated using a solid splitter (SSPLIT) unit. Free of H2O, N2 and O2 gas was finally obtained using a separator block by defining the desired split fractions. Using a secondary separator, the water content of the product was also measured to calculate the steam conversion efficiency. In this simulation, assumptions included steady state conditions,  151  isothermal reactor, SO2, H2S, CS2 and COS as the only products of S reactions [285, 287], NH3, NO and HCN as the only product of N reactions [285, 287], and instantaneous drying and pyrolysis in the gasifier [288].    Figure  6-1. Aspen Plus equilibrium simulation diagram for steam gasification process.   6.1.1 Modeling results There is a considerable advantage to gasifying under pressure, with the result that practically all modern processes are operated at pressures of at least 10 bar, and some as high as 100 bar [16]. Figure  6-2 shows the effect of temperature and pressure on the H2, CO, CH4 and CO2 equilibrium concentrations for the 50:50 coal:F-SG steam gasification with SF=2.41. The endothermic methane reforming, water gasification, and Boudouard reactions are enhanced by increasing the reactor temperature, resulting in an increase in H2 and CO concentrations, whereas at lower temperatures, forward oxidation and methanation reactions are more active, causing an R-DECOMPR-GIBBSSSPLITHEATXSEP1SEP2FUELFEEDDRY-GASWASTESIDE-PWATERC-WATERH-WATERQSTEAMPRODUCTCOLD-PCOLD-GASASH1-Reaction 2-Cooling3-Separation  152  increase in CH4 and CO2 concentrations. An increase in pressure produces an increase in CH4 and CO2, to the detriment of H2 and CO, as, in accordance with Le Chatelier’s principle, an increase in pressure shifts the equilibrium to the side with the fewer moles of gas (i.e. methanation, steam-methane reforming and oxidation reactions) [16, 289, 290].     Figure  6-2. Effect of temperature and pressure on H2, CO, CH4 and CO2 equilibrium compositions for 50:50 coal:F-SG steam gasification with SF=2.41. Colour bars are related to Y-axis.   010203040506070500600700800900100011000510152025H2 concentration (vol%)Temperature, o CPressure, bar10 20 30 40 50 60 70 051015202530500600700800900100011000510152025CO concentration (vol%)Temperature, o CPressure, bar5 10 15 20 25 30 01020304050500600700800900100011000510152025CH4  concentration (vol%)Temperature, oCPressure, bar10 20 30 40 50 01020304050500600700800900100011000510152025CO2 concentration (vol%)Temperature, oCPressure, bar10 20 30 40 50  153  The higher heating value of the product gas, calculated based on equation ( 3-4) is predicted to increase as a result of decreasing the temperature and increasing the pressure, due to enhanced methane production, as shown in Figure  6-3.  Figure  6-3. Effect of temperature and pressure on HHV of equilibrium product gas for 50:50 coal:F-SG steam gasification with SF=2.41. Colour bars are related to Y-axis.     The effects of temperature and pressure on SO2, H2S, NH3, CS2, NO, HCN, and COS equilibrium concentrations are shown in Figure  6-4. The most concentrated side-product is H2S where concentration was increased by lowering the temperature and raising the pressure. Sulfur compounds corrode metallic surfaces. Even small amount of sulfur can poison gasification catalysts. Sulfur removal is often required to avoid its detrimental effects. The molar fractions of HCN and NH3 are sufficient to trace the two most important reaction pathways for nitrogen chemistry in combustion and gasification. These two species are important final products under reducing conditions, as well as key intermediate species for NO and N2O formation in oxidizing atmospheres [291, 292]. Although most of the ammonia decomposes to N2 at typical gasification temperatures, its concentration should be controlled to avoid catalyst poisoning and extra NOx  1012141618205006007008009001000 1100051015202530HHV, MJ/N m3Temperature, oCPressure, bar12 14 16 18 20  154     Figure  6-4. Effect of temperature and pressure on SO2, H2S, NH3, CS2, NO, HCN, and COS equilibrium compositions for 50:50 coal:F-SG steam gasification with SF=2.41. Colour bars are related to Y-axis. 0.00.10.20.30.40.5500600700800900100011000510152025SO2 c on ce nt r at i on  ( vo l pp m)T emp er at ur e,  o CPressure, bar0.1 0.2 0.3 0.4 0.5 600800100012001400160018002000500600700800900100011000510152025H2S co nc en tr at i on  ( vo l pp m)T e m p e ra t u r e ,  o CPressure, bar800 1000 1200 1400 1600 1800 2000 050100150200250300500600700800900100011000510152025NH3 c on ce nt r at i on  ( vo l pp m)T e mp e r at u r e , o CPressure, bar50 100 150 200 250 300 0.05.0e-51.0e-41.5e-42.0e-42.5e-4500600700800900100011000510152025CS2 a nd  NO co nc en tr at i on  ( vo l pp m)T em pe r at u re ,  o CPressure, bar5e-51e-41.5e-42e-42.5e-4CS2NO0.000.050.100.150.200.250.30500600700800900100011000510152025HCN co nc en tr at i on  ( vo l pp m)T emp er at ur e,  o CPressure, bar0.05 0.10 0.15 0.20 0.25 0.30 024681012141618500600700800900100011000510152025COS co nc en tr at i on  ( vo l pp m)T e mp e r at u r e , o CPressure, bar2 4 6 8 10 12 14 16 18  155  production in gas turbines. The results are in agreement with equilibrium data reported by Li et al. [293]. Figure  6-5 presents the predicted equilibrium product dry gas composition and HHV of 100% SP-SG, 100% coal, and 50:50 coal:SP-SG mixture steam gasification vs. temperature for SF≈2.4 at atmospheric pressure. The predicted co-feed gas compositions lie between those predicted for corresponding single-fuel gasification. For the decomposition reactor, the simulation modeled the fuel as a simple combination of elements based on the ultimate analysis. Therefore, the model did not predict tar, char and heavier hydrocarbon elements. It also significantly under-estimated the methane concentration, reflecting the relative stability of CH4, retarding its breakdown into equilibrium products [285, 293].     Figure  6-5. Equilibrium product dry gas composition and HHV of 100% SP-SG (solid lines), 100% coal (long-dashed lines), and 50:50 coal:SP-SG mixture (short-dashed lines) steam gasification vs. temperature with SF≈2.4 and 1 atm.    Temperature, oC550 600 650 700 750 800 850 900 950 1000 1050 1100Product gas concentration (vol%)010203040506070HHV, MJ/N m368101214HHVH2CO2COCH4100% SP-SG100% Coal50:50 coal:SP-SG 156  6.1.2 Comparison of model predictions with experimental results As shown in Figure  6-6, the equilibrium model was able to predict with reasonable accuracy the gas composition, except for CH4, for coal steam gasification. CO is over-predicted, presumably because the Boudouard reaction did not reach equilibrium. The model, however, fits poorly for biomass gasification runs due to model limitations (e.g. not considering the pyrolysis of the biomass with high volatile content or the selective catalytic effects of switchgrass potassium on the Boudouard and steam-carbon gasification reactions, discussed in  Chapter 1). Previous studies [294, 295] have shown that high measured concentrations of methane from coal gasification result from incomplete conversion of pyrolysis products; equilibrium molar concentrations of methane in the off-gas are less than 0.1% for all runs, whereas actual methane concentrations are of the order of a few per cent. The high measured methane concentration in the product gas cannot be explained on an equilibrium or thermodynamic basis. The deviation must result from non-equilibrium factors, e.g. incomplete cracking of pyrolysis products. It is often taken for granted that the amount of each element participating in the chemical equilibrium is the same as for the feed. This is true when slow reaction kinetics and mass transfer processes do not impede the achievement of equilibrium. However, this assumption is not valid for real processes in which reactions (mostly heterogeneous) are influenced by kinetics and/or mass transfer limitations so that some elements and species never achieve equilibrium [285]. The model assumes that all reactions proceed to equilibrium, in which case the methane would mostly be reformed into H2, CO and CO2 at elevated temperatures and excess H2O. With the addition of forced production of CH4 and C2+, in particular ethylene, a better fit could be obtained, as suggested by Li et al. [285] who introduced a bypass for a portion of the fuel C and H to form CH4 in their model. For a steam-to-coal mass ratio range of 1-4, the equilibrium modeling product gas composition and HHV data followed the same trend as the experimental results, see Figure  6-7. Increasing steam shifted the equilibrium of product gas towards H2 and CO2, whereas CO and CH4 decreased. Again the calculated HHV of the product gas declined, mainly because of the drop in CH4 and CO. The model gave more accurate predictions at higher SF. Coal particles with high fixed carbon were not reactive enough so that increasing steam-to-coal ratio moved the gasification reactions further towards equilibrium.      157    Figure  6-6. Comparison of experimental data and equilibrium simulation predictions for SF≈2.4 and 1 atm (experimental data same as in Table  4-4 of  Chapter 4).    Figure  6-8 displays the effect of temperature on product composition and HHV of coal and 50:50 coal:SP-SG runs and compares the experimental and simulation data. The 100% coal run H2 concentration was perfectly predicted by the equilibrium model. However, the model under-predicted the HHV for both cases because of under-calculating the methane concentration.     CoalSP-SGF-SGCoal:SP-SG=50:50Coal:F-SG=50:50Product gas concentration (vol%)010203040506070Temperature, oC7508008509009501000H2 - Eq. ModelH2 - ExperimentalCO - Eq. ModelCO - ExperimentalCH4 - Eq. Model CH4 - ExperimentalCO2 - Eq. ModelCO2 - ExperimentalBed temp. 158   Figure  6-7. Comparison of experimental data and equilibrium modeling predictions: effect of steam-to-fuel mass ratio on dry gas composition and HHV for 100% coal run (same experimental data as in Figure  3-7 of  Chapter 3).          Steam-to-fuel ratio, SF1.0 1.5 2.0 2.5 3.0 3.5 4.0Time average product gas concentration (vol%)01020304050607080Calculated HHV, MJ/N m37891011121314H2 - ExperimentalCO - ExperimentalCH4 - ExperimentalCO2 - ExperimentalHHV - ExperimentalH2 - Eq. ModelCO - Eq. ModelCH4 - Eq. modelCO2 - Eq. modelHHV - Eq. Model 159    Figure  6-8. Comparison of experimental data and equilibrium modeling predictions: effect of temperature on dry gas composition and HHV for (a) 100% coal (experimental data same as in Figure  3-6 of  Chapter 3) (b) 50:50 SP-SG (experimental data same as in Figure  4-1 of  Chapter 4).  Temperature, oC845 850 855 860 865 870 875Product gas concentration (vol%)01020304050607080HHV, MJ/N m368101214Temperature, oC845 850 855 860 865 870 875Product gas concentration (vol%)010203040506070HHV, MJ/N m3678910111213141516H2 - ExperimentalCO - ExperimentalCH4 - ExperimentalCO2 - ExperimentalH2 - Eq. modelCO - Eq. modelCH4 - Eq. modelCO2 - Eq. modelHHV - ExperimentalHHV - Eq. Model(a) (b)  160  6.1.3 Modified equilibrium model As shown above, the high measured methane concentration in the product gas cannot be explained on an equilibrium or thermodynamic basis. The carbon conversion in a real gasifier depends on many factors: thermodynamics, chemical kinetics, hydrodynamics, heat and mass transfer, residence time and even particle size distribution. Equilibrium conversion provides an upper bound on the conversion efficiency, while the fractional achievement of equilibrium for a real system depends on such operating parameters as temperature, system pressure and steam-to-carbon ratio [293]. The discrepancy between equilibrium model and experimental data can be substantially reduced if the experimental data, such as the overall carbon, hydrogen and steam conversions, and the produced methane molar concentration, are incorporated in the equilibrium model. Figure  6-9 shows the modified version of the original Aspen Plus equilibrium simulation presented in Figure  6-1. Based on the experimental CH4 molar data, a fraction of carbon and hydrogen was split from the decomposition reactor output stream and was sent to a methanator stoichiometric reactor (R-CH4). The CH4 produced according to reaction ( 1-9) was then combined with the Gibbs reactor products downstream. In addition, fractions of the C(s), H2 and H2O were taken out (through separator blocks 2 and 3) from the decomposed feed and steam streams, based on the experimental carbon, hydrogen and water conversion efficiencies.  From the parity plot in Figure  6-10, it is evident that the modeling product gas composition predictions were enhanced significantly due to kinetic modifications based on the experimental data. Comparing with Figure  6-8(a), we find good agreement between the empirical and modified model data of the 100% coal steam gasification, as portrayed in Figure  6-11. Figure  6-12 compares the experimental and modified model product gas compositions and HHV vs. temperature for the 50:50 coal:F-SG steam co-gasification experiment. While the model predicted accurately at high temperatures, it started to diverge from the experimental data with a decrease in temperature, as the total water conversion was only measured in the period that the reactor temperature was close to 800°C while performing the experiment.           161   Figure  6-9. Modified Aspen Plus equilibrium simulation diagram for steam gasification process.R-DECOMPR-GIBBSSSPLITHEATXSEP1SEP2FUELDRY-GASWASTESIDE-PWATERC-WATERH-WATERQCOLD-PCOLD-GASASHSEP1SEP2C+2H25R-CH4C,H2FEEDCH4-OUTMIXERR-G-OUTPRODUCTSEP3STEAM STEAMOUTSTEAM-IN1-Reaction2-Cooling 3-Separation 162      Figure  6-10. Parity plot comparing experimental and modified equilibrium model product gas compositions for 100% coal, SP-SG, F-SG, 50:50 coal:SP-SG and 50:50 coal:F-SG steam gasification experiments (experimental data same as in Figure  6-6). Product gas concentration (vol%) - Experimental 0 10 20 30 40 50 60 70Product gas concentration (vol%) - Modified equilibrium model010203040506070H2COCH4CO2 163   Figure  6-11. Comparison of experimental data and modified equilibrium modeling predictions: effect of temperature on dry gas composition and HHV for 100% coal (experimental data same as in Figure  3-6 of  Chapter 3).  Figure  6-12. Comparison of experimental data and modified equilibrium modeling predictions: effect of temperature on dry gas composition and HHV for 50:50 coal:F-SG mixtures.   Temperature, oC845 850 855 860 865 870 875HHV, MJ/N m367891011121314Product gas concentration (vol%)051015206080100HHV - ExperimentalH2 - ExperimentalCO - ExperimentalCH4 - ExperimentalCO2 - ExperimentalH2 - Eq. modelCO - Mod. Eq. modelCH4 - Mod. Eq. modelCO2 - Mod. Eq. modelHHV - Mod. Eq. modelTemperature, oC640 660 680 700 720 740 760 780 800 820HHV, MJ/N m31314151617Product gas concentration (vol%)05101520253035404550HHV - ExperimentalH2 - ExperimentalCO - ExperimentalCH4 - ExperimentalCO2 - ExperimentalH2 - Mod. Eq. modelCO - Mod. Eq. modelCH4 - Mod. Eq. modelCO2 - Mod. Eq. modelHHV - Mod. Eq. Model 164  6.2  LEG EQUILIBRIUM MODELING Similar model as in section  6.1 was applied in an effort to simulate the lime-enhanced steam co-gasification of 50:50 coal:SP-SG. Two extra solid components, CaO and CaCO3, were defined in the Aspen Plus V7.2 process simulator. In addition, an extra stream containing CaO as sorbent was introduced to the Gibbs reactor. The CaO mass flow rate was defined based on the limestone properties given in Table  5-2. The fuel properties were defined based on the 50:50 wt% coal:SP-SG mixture. Other properties were the same as in Table  5-1. During the LEG experiments, it was practically impossible to collect the condensate water in order to calculate the steam conversion efficiency. Therefore for the kinetically modified equilibrium model, only a methanator reactor was added to the process, and the steam flowrate to the Gibbs reactor was not manipulated.    6.2.1 Modeling results Figure  6-13 presents the predicted equilibrium product dry gas compositions, HHV and calcium utilization efficiency for 50:50 coal:SP-SG steam co-gasification vs. temperature with SF≈2.6 and atmospheric pressure. In the temperature range where CaO is active for CO2 capture (i.e. 500 to ~ 725°C), the H2 concentration increased to a peak of ~ 91%. Lowering the temperature after that peak caused partial consumption of H2 by methanation to produce CH4. Methanation was noticeably a factor at lower temperatures [296]. Raising the temperature beyond 675°C decreased the predicted calcium utilization efficiency and caused a slow decline in H2 concentration. The sorbent was completely inactive above ~ 725°C, and the predicted gas concentrations for co-gasification without any sorbent and sorbent-enhanced gasification converged with the LEG predictions.  Operating the LEG process at higher pressures would allow gasification/carbonation to be performed at higher temperatures, without losing calcium utilization efficiency, as predicted by the equilibrium model and as shown in Figure  6-14. Operating LEG at higher temperatures improves thermal tar cracking of product gases. Appendix J provides the predicted effects of temperature and pressure on the H2, CO, CH4, CO2, C2+, SO2, H2S, NH3, CS2, HCN and COS equilibrium concentrations and HHV for LEG of the 50:50 coal:SP-SG mixture.       165     Figure  6-13. Predicted equilibrium product dry gas composition, HHV and calcium utilization efficiency for 50:50 coal:SP-SG steam co-gasification vs. temperature with SF≈2.6 at 1 atm. HHV calculated based on equation ( 5-1).  Temperature, oC500 550 600 650 700 750 800 850 900 950 1000 1050 1100Product gas concentration (vol%)0102030405060708090100HHV, MJ/m39101112131415Calcium utilization efficiency, X0.00.20.40.60.81.0H2 - No sorbentH2 - CaO sorbentCO - No sorbentCO - CaO sorbentCH4 - CaO sorbentCH4 - No sorbentCO2 - No sorbentCO2 - CaO sorbentHHV - CaO sorbentHHV - No sorbentCa ut. eff., X 166   Figure  6-14. Calcium utilization efficiency vs. temperature for different pressures predicted by the equilibrium model for LEG of 50:50 coal:SP-SG mixture. See Table  5-1 for the operating conditions.    6.2.2 Comparison of model predictions with experimental results Figure  6-15 compares predictions of the proposed original and modified equilibrium models with experimental data presented in  Chapter 5. The experimental points represent the product gas compositions during the phase of maximum H2 concentration (maximum CO2 capture). Better agreement was observed between the LEG original model (without a methanator) and experimental data presented in Figure  6-15 than for the co-gasification results in section  6.1.2 (compare Figure  6-6), probably due to enhancement of gasification reaction kinetics by Ca catalysis of hydrogen production, as well as simultaneous capture of CO2 through carbonation, shifting the reactions closer to equilibrium. The H2 concentration is still over-predicted, though closer to model predictions than for the co-gasification without sorbent in a bed of silica sand (cf. Figure  6-6). The model predicts the CO2 and CO concentrations in the CaO bed with fair accuracy. The CH4 concentration is still predicted weakly by the original model, since methane is relatively stable. Adding the methanator to the original model only slightly enhanced the CH4 Temperature, oC650 675 700 725 750 775 800 825 850 875 900Calcium utilization efficiency, X0.00.10.20.30.40.50.60.70.80.91.01 atm10 atm20 atm30 atm5 atm 167  and H2 predictions.  The trends in syngas composition with increasing temperature are generally similar for the model and experimental data. It should be noted that zero CaO loss was assumed in the equilibrium model, whereas some sorbent fines were entrained from the reactor in the course of the experiments. Butler [174] reported ~ 37% sorbent loss by the end of 5 LEG cycles using the same CERC BFB facility with almost the same duration and superficial gas velocity. The best predictions in Figure  6-15 correspond to the first and second cycles. Therefore, incorporating the sorbent loss in the Aspen simulation would likely improve the model predictions for the later cycles.    Figure  6-15. Comparison of experimental data for five cycles and equilibrium modeling predictions: effect of temperature on dry gas composition of 50:50 coal:SP-SG LEG with SF≈2.6 and 1 atm.  6.3  SUMMARY (1) An equilibrium model was programmed in Aspen Plus, based on minimization of steam gasification Gibbs free energy. The original model was simple and excluded tar, char and heavy hydrocarbons. This model was unable to predict the methane composition properly due to the simplifying assumptions. The modeling predictions were in good agreement Temperature, oC680 690 700 710 720Product gas concentration (vol%)0102030405060708090100Y Axis 2H2 - ExperimentalCO - ExperimentalCH4 - ExperimentalCO2 - ExperimentalH2 - Eq. modelCO - Eq. modelCH4 - Eq. modelCO2 - Eq. modelH2 - Eq. model with methanatorCO - Eq. model with methanatorCH4 - Eq. model with methanatorCO2 - Eq. model with methanator 168  with the 100% coal experimental data, but in poor accord with biomass experimental data, in particular over-predicting the hydrogen concentration.  (2) The equilibrium model was useful in predicting what is thermodynamically attainable. The original equilibrium model was modified by empirical data to account for the deviation from experimental results due to partial carbon and steam conversion. The modified model predictions compared reasonably well with measured gas compositions for the single-fuel and co-gasification experiments. The modified equilibrium model has potential to be utilized for scale-up of steam gasification fed by single or combined feedstocks. (3) An Aspen equilibrium model worked better for limestone-enhanced gasification (LEG) than for steam gasification without sorbent. Again, this model was limited, as it did not predict condensable tars due to the simplified representation of the fuel by its constituent elements. Methane concentration was also under-predicted for all cycles due to kinetic limitations. The model accurately predicted the CO2 and CO concentrations for the early cycles, but slightly over-predicted H2 concentrations.  (4) Reasonable accuracy of the enhanced gasification equilibrium model justifies the use of such a basic model to provide best-case scenario predictions in the design of LEG systems. Empirical modifications to the model could be made to account for deviations from the experimental results.                      169  Chapter 7. CONCLUSIONS AND RECOMMENDATIONS  7.1  CONCLUSIONS  Alarming statistics about the dramatic rise in atmospheric GHG emissions and consequent rise in the global surface temperature in recent decades have obliged engineers to move towards sustainable technologies. Gasification offers a gaseous product (syngas) that can be converted to diverse end products such as hydrogen, methanol, ethanol, transportation fuel and power. Biomass is considered a carbon-neutral form of energy as the CO2 released during its utilization is equal to the CO2 absorbed from the atmosphere during its growth through photosynthesis. However, biomass gasification technologies are typically not feasible economically due to the low energy density of biomass and the high transportation cost of the feedstock. On the other hand, as a fossil fuel and one of the main sources of CO2 emissions, coal has a high energy density and coal-gasification is a well-developed technology. Therefore, co-feed gasification of biomass and fossil fuel has a promising prospect to accelerate the transition from energy production based on fossil fuels to renewable fuel technologies. The addition of CO2 capture and storage (CCS) technology to co-gasification could even result in net removal of CO2 from the atmosphere, often referred to as negative emissions. The main conclusions of this thesis are as follows: (1) In order to help understand the interactions between biomass and fossil fuels, the kinetic of two-stage CO2 gasification of biomass and fossil fuels was first studied in a thermogravimetric analyzer (TGA). Two types of Canadian biomass, Manitoba switchgrass (SG) and beetle-killed BC pine sawdust (SD) from British Columbia, and two types of Canadian fossil fuels, Alberta Genesee sub-bituminous coal and fluid coke (FC) from Syncrude Canada Ltd. in Fort McMurray, with widely differing ash compositions, were chosen. The switchgrass ash was rich in potassium (16.8 wt%), while the coal contained a considerable proportion of ash (30.5 wt%) rich in aluminum and silicon. The sawdust and fluid coke contained much less ash. The BET surface area of all samples increased by making char at 750, 800 and 900°C, consistent with SEM images.  170  The higher CO2 adsorption surface areas of the char samples (except for the FC) than the N2 adsorption surface areas suggest that the samples had microporous structures. The surface areas for both N2 and CO2 adsorption decreased with increasing charring end temperature from 750 to 900°C, possibly due to sintering. During the co-pyrolysis with N2 of the single and blended samples of biomass and fossil fuel from room temperature up to 800°C, no remarkable synergistic kinetic effect was observed. Adding 25 wt% and 50 wt% switchgrass to coal had an adverse effect on the CO2 gasification rate. On the other hand, when the biomass char was dominant in the mixture (75 wt % SG), the rate of conversion was faster than for coal alone. XRD analyses show that the inhibition effect is likely due to interaction of coal mineral matter with biomass potassium and formation of aluminosilicate crystals. Therefore in another TGA experiment, 50:50 wt% coal char:SG ash was gasified and the coal gasification rate was enhanced significantly. The switchgrass ash displayed impressive catalytic activity when mixed with the coal char. Washing the coal with water and different acids and reducing the coal minerals enhanced co-gasification of switchgrass and treated coal, with no inhibitory effect.  Since fluid coke had low Al and Si contents, significant catalytic enhancement was observed when co-gasifying switchgrass with fluid coke, with no formation of aluminosilicates. This was an encouraging result, showing that low reactivity of some fuels can be improved by the presence of biomass minerals, acting as inexpensive catalysts. (2) To demonstrate the key characteristics of co-gasification in a pilot plant scale, Ontario spring and fall harvest switchgrass as biomass and Vancouver Island coal as fossil fuel were chosen to perform the gasification of single and 50:50 wt% mixture fuels in an atmospheric pilot-scale fluidized bed reactor using steam as a gas agent, mostly at ~ 860°C, with silica sand as the bed material. For both mixed feeds, higher gas yields were obtained than from the single fuel experiments, presumably due to switchgrass potassium catalytic interactions with carbon of the coal enhancing the rates of gasification reactions. Fall switchgrass gasification resulted in lower tar yield due to higher ash and alkali metals content than for spring switchgrass. Although higher bed temperature enhances thermal tar cracking, reducing tar yield, the 50:50 coal:fall switchgrass co-gasification  171  with the lowest average bed temperature yielded the lowest tar due to greater presence of alkali and alkaline earth metals in the fall switchgrass ash. Increasing the steam-to-fuel ratio for coal and coal:switchgrass mixtures caused a moderate increase in the H2 and CO2 concentrations and decreases in the CO and CH4 concentrations, due to more steam-CH4 reforming and water-gasification reaction of CO. With increasing the reactor temperature, the H2 concentration increased, whereas the CO, CH4, and CO2 concentrations fell slightly. (3) In the final experimental phase of this work, in-situ CO2 capture was integrated with gasification to study lime-enhanced steam co-gasification of 50:50 coal:spring switchgrass in an atmospheric pressure bubbling fluidized bed with limestone as the bed material. Five cycles of gasification and carbonation were performed at <700°C, while the sorbents were regenerated by calcination at >850°C. In-situ CO2 removal during steam co-gasification increased the hydrogen production as a result of the shift in equilibrium. Hydrogen concentration in the outlet gas increased to over 70% in a bed of CaO. The exothermic carbonation reaction improved the bed temperature stability until the sorbents reached their utilization limit. However, the density differences among the biomass, coal and CaO particles resulted in segregation of the solids inside the reactor. Also formation of low-temperature melts caused unwelcome blockages, whilst the agglomeration of sorbent and fuel particles may reduce CO2 capture and fuel conversion efficiencies. In the course of cycling, H2 concentration, hydrogen utilization efficiency decreased due to lower calcium utilization, whereas CO, CO2, CH4 and C2+ concentrations increased. BET analyses showed a general drop in surface area during cycling. The SEM images after the 1st and the 5th cycles revealed particle sintering during high-temperature regeneration of CaO, causing loss of sorbent activity and total surface area. Product gas tar yields were close to the corresponding tar yield for sand bed co-gasification at a much higher temperature (~ 862°C), indicating that CaO catalyzed tar cracking reactions. The volumetric weighted average particle size of the calcined sorbent material dropped by ~ 25% from 563 µm (original sample) to 427 µm at the end of the fifth cycle as a result of particle attrition.   172  (4) To examine the reliability of equilibrium modeling to predict the product gas compositions, the steam gasification equilibrium reactor was simulated by minimizing the Gibbs free energy with Aspen Plus software. The original equilibrium model was only useful in predicting what is thermodynamically achievable. It over-predicted H2 while significantly under-predicting CH4 and was unable to predict heavier hydrocarbons (C2+) or tars due to chemical kinetic constraints. The model was able to accurately predict the product gas compositions after adding an extra methanator stoichiometric reactor to produce methane based on the empirical CH4 concentration, and removal of part of the carbon, hydrogen and steam before introducing the feed and gas agent streams to the Gibbs reactor based on experimental carbon, hydrogen, and steam efficiencies. The original equilibrium model was modified by adding CaO in order to simulate the lime-enhanced system. The model was able to predict the product gas composition of the early cycles with reasonable accuracy.     Research Significance  The catalytic effect of synthetic alkali and alkaline earth metals on carbon gasification has been well investigated in the literature. However, the effect of natural minerals as inexpensive catalysts (e.g. biomass ash metals) on coal gasification has received much less attention. In this thesis, a systematic approach was adopted towards studying the thermo-chemical interactions between biomass and fossil fuel elements during co-gasification. The study tried to address several confusions and misinterpretations in the literature about whether co-feeding biomass with fossil fuels can enhance the gasification rate. The thermogravimetric and large scale co-gasification analyses can be a useful guideline to select proper biomass and fossil fuels to improve the co-gasification performance and minimize inhibitory effects. Biomass/coal co-gasification can be considered as an environmentally-friendly technology, if a portion of CO2 in the product gas originated from coal can be captured and stored through CCS technology. The lime enhanced co-gasification study was an attempt to integrate gasification with CO2 capture to introduce a more sustainable process and to investigate the operational challenges of having three different types of particles (coal, switchgrass and lime) in a single environment.  173  7.2  RECOMMENDATIONS FOR FUTURE WORK (1) It is recommended to extend the thermogravimetric analyses presented in this work by investigating steam (instead of CO2) co-gasification of biomass and fossil fuels and comparing the kinetic results with the data presented in  Chapter 2. Having a gas chromatograph capable of detecting very low syngas compositions would also allow the pyrolysis/gasification product gas composition of single and blended feedstocks to be investigated. (2) For pilot scale future study, dual fluidized bed reactors seem to be suitable to address some major issues which arose during the co-gasification experiments in bubbling fluidized bed and in a lime-enhanced reactor. Low carbon conversion efficiency with coal (by itself or comprising 50% by mass of the fuel) can be improved by circulation of the char between the two environments of a dual-bed, and increasing the residence time. Solids circulation could also enhance the catalytic interaction between biomass ash minerals and fossil fuel carbon in co-gasification experiments. Dual fluidized beds can ideally be operated as autothermal reactors for lime-enhanced gasification by providing two environments, one suitable for gasification and simultaneous CaO carbonation at low temperatures, and the other appropriate for char combustion and regeneration of the sorbent bed materials at higher temperatures. A portion of heat needed for endothermic gasification reactions can be supplied by returning the hot solids to the gasifier. Dual fluidized beds could also improve the mixing and reduce particle segregation. 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Assumptions: • The sample bed and the basket are at same temperature. • Isothermal bed with no temperature gradient along the bed.  • No gas flow through the bed. • Lumped capacitance method for heat transfer. • No radiation. • Full sphere basket for estimating Nusselt number. • Only Boudouard reaction takes place. Formulation: Energy balance over the basket: 0 0/( )/( ) /1 1( )bpreaction convection preaction pT treaction pT Treaction reactionb Aht mcEnergy Energy mc dT dtEnergy Ah T T mc dT dtdT dtEnergy Ah T T mcEnergy EnergyT TAh Ahe∞∞∞=∞− =− − =⎛ ⎞ =⎜ ⎟− −⎝ ⎠= + −∫ ∫ Where h is a convective heat transfer coefficient of the fluid:  199  2 *CObk Nuhd=  kCO2 is a thermal conductivity of the fluid and db is the basket diameter.  For flow over a sphere, Whitaker [297] proposed:  0.5 041.67 0.4 0.253.5<Re<7.6*10 ,0.7<Pr<380, and 1< / <3.22 (0.4Re 0.06Re )P ):r (wwif µ µNuμμ= + + Alternatively, Churchill [298] introduced: 0.259/16 4/91120.59Re2(1 (0.47 /Re<10 ,  Pr>Pr) )0.5:if anyNu⎛= ⎞+ ⎜ ⎟+⎝ ⎠ where  222RePrR bCOCOCOU dννα== UR is the reactor gas velocity, νCO2 is the gas kinematic viscosity and αCO2 is the gas thermal diffusivity.   200  Calculations: Parameter Symbol Equation Value Units NotesSample mass ms   15.00 mg         1.50E-05 kg   Sample bed thickness ts  1.00E-3 m from experimental data Bed voidage ε  0.4  from experimental data TGA temperature T∞   800 °C   Diameter of the basket db   1.70 cm         0.017 m   Area of the basket Ab π*db2/4 2.27E-04 m2   Basket thickness tb  7.50E-04 m  Basket mass mb  2.91E-03 kg  Thermal PropertiesHeat capacity of switcghgrass cp,SG   5084.32 kJ/kg*K from Dupont [299] Heat capacity of quartz glass cp,b   1500.00 kJ/kg*K from Kelley [300]  Weighted ave. heat capacity  cp,w xSG*cp,SG+xb*cp,b 1205.18 kJ/kg*K  Switchgrass thermal conductivity kSG 4.429*10-2+1.477*10-4 T∞ 0.20 W/m*K from Hankalin [301] Sample thermal conductivity ks 0.9065/(0.667/kSG+0.13/kCO2) 0.18 W/m*K from Gabor [302] Quartz glass thermal conductivity kb  2.00 W/m*K from Sergeev [303] CO2 thermal conductivity kCO2   7.53E-02 W/m*K at 800°C and 1 atm CO2 kinematic viscosity νCO2   8.79E-05 m2/s at 800°C and 1 atm CO2 thermal diffusivity αCO2   1.22E-04 m2/s at 800°C and 1 atm       Prandtl number Pr νCO2 /αCO2 0.72                 Heat of reaction QR''   172.60 kJ/mol Boudouard reaction   QR' QR''*MC 14383.33 kJ/kg               Conversion α   0.50     Time to reach α=50%     2682.00 s   Gasification rate rgas -1.63E-03 mg/s from TGA experimental data Gasification energy ER rgas*QR' -2.34E-05 kW    201  Reactor Hydrodynamics Gas flow in reactor q   0.50 LPM at ambient     T∞/Tamb * q 1.83 LPM at reactor temperature       3.051E-05 m3/s   Reactor diameter dR   1.25 in         0.03 m   Area of reactor AR' π* dR2/4 7.92E-04 m2 total cross section   AR AR' - Ab 5.65E-04 m2 cross section minus basket Gas velocity in reactor UR q/AR 5.40E-02 m/s   Reynolds number Re UR* db /νCO2 10.45   vortex creation (better heat/mass transfer)  Convective Heat Transfer (Lumped Capacitance Method) Nusselt number  Nu1 see formulation 3.39   from Whitaker [297] Nusselt number  Nu2 see formulation 2.82   from Churchill [298] Heat transfer coefficient h1 kCO2Nu1/db 15.00 W/m2*K         1.50E-02 kW/m2*K               Heat transfer coefficient h2 kCO2Nu2/db  12.49 W/m2*K         1.25E-02 kW/m2*K               Basket Biot number Bib1 h1tb/kb 5.62E-03  lumped capacitance is valid for Bi<0.1  Bib2 h2tb/kb 4.68E-03  lumped capacitance is valid for Bi<0.1       Sample Biot number  Bis1 h1ts/ks 0.08  lumped capacitance is valid for Bi<0.1  Bis2 h2ts/ks 0.07  lumped capacitance is valid for Bi<0.1       Combined Biot number Bic1 Bib1+Bis1 0.09  lumped capacitance is valid for Bi<0.1  Bic2 Bib2+ Bis2 0.07  lumped capacitance is valid for Bi<0.1       Particle bed temperature Tb1 T∞+ER/(Ah1) 799.98 °C from Whitaker [297] Particle bed temperature Tb2 T∞+ER/(Ah2) 799.98 °C from Churchill [298]     202  A.2 Mass Transfer Calculations Here is a simple mass transfer calculation to estimate the TGA internal and overall effectiveness factors for the 100% coal CO2 gasification experiment at 800°C and 1 atm. Assumptions: • Flat plate sample bed. • 1st order reaction. • Full sphere basket for estimating Sherwood number. • Only Boudouard reaction takes place. Formulation: For the internal effectiveness factor of a flat plate [304, 305]: inttanhernalφη φ=   where ϕ is the Thiele modulus and can be defined as   22 2,@8001.751 22 2 1 12 1/21( , ) ( , )eCO CO C ceCO CO CO COL k Dplate thicknessLkkDDP TD T P D T PP Tφεεστ− °− −′′′==′′′ = −=⎛ ⎞⎛ ⎞= ⎜ ⎟⎜ ⎟⎝ ⎠ ⎝ ⎠ De is the effective diffusivity. Typical values of the constriction factor, the tortuosity and the sample porosity are as follows, respectively [306]:  0.8, 3.0, 0.4cσ τ ε= = =  203  The overall effectiveness factor for a 1st order reaction is:  intint1 /ernaloverallernal a b c ck S k aηη η ρ= ′′+  where  (1 )1p ackkSaplate thicknessρ ε′′ = −=  Sa is the surface area per unit mass, k´´ is the specific reaction rate and ρb is the bed bulk density.  The mass transfer coefficient is  ShecbDkd= For flow around a sphere, the Sherwood number was estimated using Frössling [307] correlation:  21/2 1/3Sh 2 0.6 Re ScSc COeDν= +=          204  Calculations: Parameter Symbol Equation Value Units NotesSample bed thickness ts  1.00E-3 m from experimental data Characteristic length L 0.5ts 5.00E-04 m  Bed voidage ε  0.4  from experimental data TGA temperature T∞   800 °C   Diameter of the basket db   1.70 cm         0.017 m   Internal Mass Transfer Limitation CO2-CO diffusivity  DCO2-CO   1.52E-05 m2/s at 25°C and 1 atm from Fuller at al. [308] CO2-CO diffusivity  DCO2-CO,800°C DCO2-CO*(1073/298)1.75 1.45E-04 m2/s at 800°C and 1 atm       Effective diffusivity De 0.11DCO2-CO,800°C 1.54E-05 m2/s at 800°C and 1 atm Specific reaction rate k   6.38E-03 1/s from experimental data for 100% coal  Specific reaction rate k´´´ k/(1- ε) 1.06E-02 1/s   Thiele modulus ϕ L*( k´´´/ De)0.5 1.31E-02           Internal effectiveness factor ƞinternal see the formulation 0.99  Negligible internal resistance to mass transfer       Overall Mass Transfer Limitations      Surface area/particle mass Sa  243 m2/g see Table  2-2 of  Chapter 2 External surface area/particle volume ac 1/ts 1000 1/m for a flat plate Coal bulk density ρb  642.70 m3/kg from experimental data Specific reaction rate k´´ k/(ρb*Sa) 4.09E-11 m/s        Schmidt number  Sc νCO2 /De 5.69   Sherwood number Sh 2+0.6Re0.5Sc0.33 5.46  from Frössling [307] Mass transfer coefficient ke De*Sh/db 4.96E-03 m/s        Overall effectiveness factor ƞoverall see the formulation 0.99  Negligible overall resistances to mass transfer   205  Sensitivity analysis:  Figure A1. Effect of TGA sample bed thickness on the internal and overall effectiveness factors.        Sample bed thickness, ts, mm0 2 4 6 8 10 12 14 16 18 20Effectiveness factor, η0.950.960.970.980.991.00InternalOverall 206  APPENDIX B. HBTI1 BUBBLING FLUIDIZED BED CALIBRATION CURVES  Feeder calibration curves of single and mixed fuels:    Figure B1. Feeder calibration curve of 100% coal.                                                   1 Highbury Biofuel Technology Inc. Controller setting, %12 14 16 18 20 22 24 26 28 30 32Feed rate, kg/h0.60.81.01.21.41.61.82.02.22.42.62.8y=0.0105x-0.0624 207   Figure B2. Feeder calibration curve of 100% SP-SG.   Figure B3. Feeder calibration curve of 100% F-SG. Controller setting, %10 15 20 25 30 35 40 45Feed rate, kg/h0.40.60.81.01.21.41.61.82.0y=0.047x-0.078Controller setting, %14 16 18 20 22 24 26 28 30 32Feed rate, kg/h0.60.81.01.21.41.61.82.02.2y = 0.083x - 0.462 208   Figure B4. Feeder calibration curve of 50:50 coal:SP-SG.   Figure B5. Feeder calibration curve of 50:50 coal:F-SG. Controller setting, %14 16 18 20 22 24 26 28 30 32Feed rate, kg/h0.81.01.21.41.61.82.02.22.42.6y=0.0103x-0.584Controller setting, %10 12 14 16 18 20 22 24 26Feed rate, kg/h0.60.81.01.21.41.61.82.02.22.4y = 0.121x - 0.648 209  Rotameters calibration curves:  Figure B6. BFB shell N2 rotameter calibration curve.   Figure B7. Fluidization N2 rotameter calibration curve. Rotameter reading0 2 4 6 8 10 12Actual flow rate (cfm)0.050.100.150.200.250.300.350.40y = 0.0319x + 0.0183Rotameter reading1 2 3 4 5 6 7Actual flow rate (cfm)1234567y = 0.9421x + 0.0975 210    Figure B8. Air rotameter calibration curve.   Figure B9. Conveying N2 calibration curve.  Rotameter reading0.0 0.5 1.0 1.5 2.0 2.5 3.0 3.5 4.0 4.5Actual flow rate (cfm)012345y = 1.0326x - 0.145Rotameter reading0.0 0.5 1.0 1.5 2.0 2.5 3.0 3.5 4.0 4.5Actual flow rate (cfm)012345y = 0.9545x + 0.0236 211    Figure B10. Hopper N2 calibration curve.            Rotameter reading3 4 5 6 7 8 9 10Actual flow rate (LPM)0.00.51.01.52.02.53.03.5y = 0.3936x - 0.8904 212  APPENDIX C. TSS STANDARD METHOD [309]  The total suspended solids (TSS) of the gasification condensate water samples was measured based on the following procedure:    • Filter Type: Whatman 934-AH (47 mm OD) • Initially each filter is rinsed 5 times (approximately 10 mLs of water each time) with distilled water using the vacuum filter cassette holder. Apply the vacuum to rinse the water through the filter. After washing, place each filter into separate aluminum weigh boats and place in 103°C oven for up to 2 h.  • After drying, weigh each filter in its aluminum boat and record this weight as the pre-weight.  • Using a 500 mL graduated plastic cylinder, pour out 500 mLs of the sample into the cylinder.  • Assemble the vacuum filter cassette apparatus.  • Remove the filter from the weigh boat and place it in the vacuum filter cassette.  • Pour the sample slowly through the filter as the vacuum is being applied. As the filter gets loaded with solids, the flow will slowly diminish.  • If the flow diminishes markedly and the sample becomes difficult to filter, then stop pouring and allow this volume of sample to complete filtering. This is taken as the final sampling volume (L) that is recorded.  • Remove the filter from the vacuum cassette apparatus and place it back into its original weigh boat.  • After completing a batch of filtering, place all the sample filters in the 1030 C oven for 2 to 3 hours.  • Perform a post weigh of all the dried filters and record this as the final post weight.  • Final results are calculated as:  TSS (mg/L) = [Post Weight (mg) - Pre Weight (mg)] / Volume of Sample (L)   213  APPENDIX D. TEMPERATURE, PRESSURE DROP, AND STEAM FLOW RATE PROFILES OF EXPERIMENTS IN HBTI FLUIDIZED BED GASIFIER  Full measured temperatures, pressures drop and steam flow rate profiles of performed experiments in HBTI BFB are presented below:   Figure D1. 100% coal run temperature variation with time. 100% coal temperatures profileTime, t, min0 50 100 150 200 250Temperature, oC02004006008001000T0 - Steam inT1 - Bottom bedT2a - BedT2b - Lower freeboardT4 - Cyclone inT5 - Tar samplingT6 - Before first condenserT7 - After first condenserT3a - Upper freeboardT9 - After second condenserT11 - Before second condenserT3b - Reactor topT13 - Freeboard wallT14 - Before baghouseT15 - Bed wall 214   Figure D2. 100% SP-SG run temperature variation with time.   Figure D3. 100% F-SG run temperature variation with time. 100% SP-SG temperatures profileTime, t, min0 50 100 150 200Temperature, oC020040060080010001200100% F-SG temperatures profileTime, t, min0 50 100 150 200Temperature, oC02004006008001000 215    Figure D4. 50:50 coal:SP-SG run temperature variation with time. .    Figure D5. 50:50 coal:F-SG run temperature variation with time. 50:50 coal:SP-SG temperatures profileTime, t, min0 50 100 150 200Temperature, oC0200400600800100050:50 coal:F-SG temperatures profileTime, t, min0 20 40 60 80 100 120 140 160 180Temperature, oC02004006008001000 216    Figure D6. Pressure drop variation with time of different experiments.    Figure D7. Steam flow rate variation with time of different experiments. Pressure drops profileTime, t, min0 50 100 150 200 250Bed pressure drop (kPa)0246810100% Coal100% SP-SG100% F-SG50:50 coal:SP-SG50:50 coal:F-SGSteam flow rate profileTime, t, min0 50 100 150 200 250Steam flow rate (kg/h)2.62.83.03.23.4100% Coal100% SP-SG100% F-SG50:50 coal:SP-SG50:50 coal:F-SG 217  APPENDIX E. WATER TREATMENT ON THE 50:50 COAL:SP-SG  GASIFICATION CONDENSED WATER1  DOWEX TAN-1 ion exchange resin was used to treat the condensed water. It is a macroporous strong base anion resin with an open pore structure allowing for a more effective binding and removal of total organics, while offering a more efficient regeneration due to faster kinetics. It is well suited to bind and remove many larger organics like tannic acids, which pass through simple UF, sand filtration and other first-stage processing steps. The DOWEX TAN-1 resin is certified under ANSI STD 61, making it a useful tool for removal of organics from surface water sources being targeted for potable water. The DOWEX TAN-1 resin also meets the 21CFR173.25 for food contact and can be applied to bind organics from process streams for flavor and color control.  Two water samples, 450 mL each, were diluted with 550 mL distilled water. Then, 25 g and 50 g resins were gravimetrically measured and added to the water samples, and mixing was conducted at room temperature and atmospheric pressure using a stirred jar at 150 rpm for both samples. The TOC of the samples was measure during the mixing period using a TOC analyzer (Shimadzu TOC-VCPH). Figure F shows that the major drop in TOC occurred during the first 10 min for both samples. The TOC of the sample with 50 g resin reduced by ~ 50% from 1511 to 767 ppm. The total carbon content of this sample was reduced to 457 ppm, see Figure F, by adding 100 g extra resin. Figure G shows the lighter color of treatment 3 than treatment 2 due its less TOC after treatment. More advanced industrial water treatment methods is required for better performance.                                                  1 This work was done in collaboration with Mohammad Mahdi Bazri, water treatment group, Chemical and Biological Engineering Department, University of British Columbia.    Figure  Figure E1. TOC vsE2. Pictures0TOC, ppm4006008001000120014001600. time result of (right jar50TTTs of 50:50 cinitial ): treatment3100 15reatment 1 - 25reatment 2 - 50reatment 3 - 10oal:SP-SG cresin loading 2 sample af60 min.   Time, t, min0 200 g resin g resin0 gr resin (perfoondensed ws.  ter 240 min250 30rmed on treatmater sample, (left jar): tr0 350ent 2 sample)  s with differ eatment 3 a218 ent fter  219  APPENDIX F. LIME ENHANCED GASIFICATION (LEG) GASIFICATION/CARBONATION RESULTS  Product gas compositions and temperature profiles of all five LEG gasification/carbonation experiments are presented below:   Figure F1. H2 product compositions vs. time during steam gasification/carbonation of 50:50 coal:SP-SG mixtures for five cycles in the CERC BFB at 1 atm. Time, t, min0 20 40 60 80 100 120H2 concentration (vol%)4050607080Cycle 1Cycle 2Cycle 3Cycle 4Cycle 5 220   Figure F2. CO2 product compositions vs. time during steam gasification/carbonation of 50:50 coal:SP-SG mixtures for five cycles in the CERC BFB at 1 atm.  Figure F3. CO product compositions vs. time during steam gasification/carbonation of 50:50 coal:SP-SG mixtures for five cycles in the CERC BFB at 1 atm. Time, t, min0 20 40 60 80 100 120CO2 concentration (vol%)010203040Cycle 1Cycle 2Cycle 3Cycle 4Cycle 5Time, t, min0 20 40 60 80 100 120CO concentration (vol%)0246810121416Cycle 1Cycle 2Cycle 3Cycle 4Cycle 5 221   Figure F4. CH4 product compositions vs. time during steam gasification/carbonation of 50:50 coal:SP-SG mixtures for five cycles in the CERC BFB at 1 atm.  Figure F5. C2+ product compositions vs. time during steam gasification/carbonation of 50:50 coal:SP-SG mixtures for five cycles in the CERC BFB at 1 atm. Time, t, min0 20 40 60 80 100 120CH4 concentration (vol%)024681012141618Cycle 1Cycle 2Cycle 3Cycle 4Cycle 5Time, t, min0 20 40 60 80 100 120C2+ concentration (vol%)024681012Cycle 1Cycle 2Cycle 3Cycle 4Cycle 5 222   Figure F6. Lower bed temperature (T8) vs. time during steam gasification/carbonation of 50:50 coal:SP-SG mixtures for five cycles in the CERC BFB at 1 atm.           Time, t, min0 20 40 60 80 100 120Temperature, oC400450500550600650700750800Cycle 1Cycle 2Cycle 3Cycle 4Cycle 5 223  APPENDIX G. LIME ENHANCED GASIFICATION (LEG) CALCINATION RESULTS  Product gas compositions, temperature profiles and CO2 partial pressures of all five LEG calcination experiments are presented below:   Figure G1. Cycle 1 dry product gas compositions, temperatures and pressures during calcination in CERC BFB.  Time, t, min0 20 40 60 80 100 120CO2 partial pressure (kPa)0102030405060Product gas concentration (vol%)0102030405060708090100Temperature, oC01002003004005006007008009001000Cycle 1CO2 eq. partial pressureCO2 partial pressureH2CH4COCO2O2N2T8 - Lower Bed (150 mm)T7 - Mid Bed (200 mm)T6- Upper Bed (250 mm) 224   Figure G2. Cycle 2 dry product gas compositions, temperatures and pressures during calcination in CERC BFB.   Figure G3. Cycle 3 dry product gas compositions, temperatures and pressures during calcination in CERC BFB. Time, t, min0 25 50 75 100 125 150 175CO2 partial pressure (kPa)020406080100120140160Product gas concentration (vol%)0102060708090100Temperature, oC1002003004005006007008009001000Cycle 2Time, t, min0 20 40 60 80 100 120 140 160CO2 partial pressure (kPa)01020304050Product gas concentration (vol%)01020708090100Temperature, oC01002003004005006007008009001000Cycle 3 225   Figure G4. Cycle 4 dry product gas compositions, temperatures and pressures during calcination in CERC BFB.  Figure G5. Cycle 5 dry product gas compositions, temperatures and pressures during calcination in CERC BFB. Time, t, min0 10 20 30 40 50 60 70 80 90 100 110CO2 partial pressure (kPa)020406080100Product gas concentration (vol%)01020708090100Temperature, oC01002003004005006007008009001000Cycle 4Time, t, min0 20 40 60 80 100 120 140CO2 partial pressure (kPa)020406080100Product gas concentration (vol%)0102030708090100Temperature, oC01002003004005006007008009001000Cycle 5 226  APPENDIX H. LIME ENHANCED GASIFICATION (LEG) NITROGEN ISOTHERMS OF DIFFERENT SAMPLES    Figure H. Nitrogen isotherms for different sorbent samples. P is the pressure of gas adsorption (atm) and P0 the saturation pressure of the analysis gas at the analysis temperature (atm).       Relative pressure, P/P00.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0N2 uptake, mmol/g0.00.10.20.30.40.5Original limeAfter ovenCalcined in the reactorCycle 1Cycle 2Cycle 3Cycle 4Cycle 5  APPENDSAMPLE Figur  IX I. LIMES e I1. SEM imFiguFigu ENHANCages of there I2. SEM re I3. SEM ED GASIFI CaO sorbestartinimages of thimages of thCATION (Lnt sample afg the cycline CaO sorbee CaO sorbeEG) SEMter calcinatig. nt sample ant sample a IMAGESon inside thfter Cycle 1fter Cycle 2 OF CALCIe BFB befor. . 227 NED e        FiguFiguFigure I4. SEM re I5. SEM re I6. SEM images of thimages of thimages of the CaO sorbee CaO sorbee CaO sorbent sample ant sample ant sample after Cycle 3fter Cycle 4fter Cycle 5. . . 228  229  APPENDIX J. LIME ENHANCED GASIFICATION (LEG) EQUILIBRIUM MODEL RESULTS    Figure J1. Effect of temperature and pressure on H2, CO, CH4 and CO2 equilibrium compositions predictions for 50:50 coal:SP-SG LEG with SF=2.6 using Aspen Plus. Colour bars are related to Y-axis.  020406080100500 600 700 800900 1000 11000510152025H2 concentration (vol%)Temperature, oCPressure, bar20 40 60 80 100 051015202530500600700800900100011000510152025CO concentration (vol%)Temperature, o CPressure, bar5 10 15 20 25 30 02468101214500600700800900100011000510152025CH4  concentration (vol%)Temperature, oCPressure, bar2 4 6 8 10 12 14 0510152025500600700800900100011000510152025CO2 concentration (vol%)Temperature, oCPressure, bar5 10 15 20 25  230    Figure J2. Effect of temperature and pressure on SO2, H2S, NH3, CS2, NO, HCN, and COS equilibrium compositions for 50:50 coal:SP-SG LEG with SF=2.6 using Aspen Plus. Colour bars are related to Y-axis.   0.00.10.20.30.40.5500600700800900100011000510152025SO2 c on ce nt r at i on  ( vo l pp m)T emp er at ur e,  o CPressure, bar0.1 0.2 0.3 0.4 0.5 600800100012001400500600700800900100011000510152025H2S co nc en tr at i on  ( vo l pp m)T e m p e ra t u r e ,  o CPressure, bar800 1000 1200 1400 050010001500200025003000500600700800900100011000510152025NH3 c on ce nt r at i on  ( vo l pp m)T e m p e ra t u r e ,  o CPressure, bar500 1000 1500 2000 2500 3000 0.05.0e-51.0e-41.5e-42.0e-42.5e-4500600700800900100011000510152025CS2 a nd  NO co nc en tr at i on  ( vo l pp m)T em pe r at u re ,  o CPressure, bar5.0e-5 1.0e-4 1.5e-4 2.0e-4 2.5e-4 0.000.050.100.150.200.25500600700800900100011000510152025HCN co nc en tr at i on  ( vo l pp m)T emp er at ur e,  o CPressure, bar0.05 0.10 0.15 0.20 0.25 0246810121416500600700800900100011000510152025COS co nc en tr at i on  ( vo l pp m)T e mp e r at u r e , o CPressure, bar2 4 6 8 10 12 14 16  231   Figure J3. Effect of temperature and pressure on C2+ equilibrium composition for 50:50 coal:SP-SG LEG with SF=2.6 using Aspen Plus. Colour bars are related to Y-axis.   Figure J4. Effect of temperature and pressure on HHV of equilibrium product gas for 50:50 coal:SP-SG LEG with SF=2.6 using Aspen Plus. Colour bars are related to Y-axis.    0.00.20.40.60.81.01.21.41.6500600700800900100011000510152025C2+ concentration (vol ppm)Temperature, o CPressure, bar0.2 0.4 0.6 0.8 1.0 1.2 1.4 1.6 10121416185006007008009001000 1100051015202530HHV, MJ/N m3Temperature, oCPressure, bar12 14 16 18   APPENDIX K. UBC DUAL FLUIDIZED BE FiguD REACTOR re K1. UBC duFLOW DIAGRal fluidized bedAM   reactor flow diagram. 232  

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