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UBC Theses and Dissertations

Recirculating fluidized bed process for the roasting of molybdenite concentrates Wilkomirsky, Igor A. E. 1974

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RECIRCULATING FLUIDIZED BED PROCESS FOR THE ROASTING OF MOLYBDENITE CONCENTRATES by IGOR A.E. WILKOMIRSKY B . S c , U n i v e r s i t y o f C o n c e p c i o n , C h i l e , 1962 M.Sc., C o l o r a d o S c h o o l o f M i n e s , 1966 A THESIS SUBMITTED IN PARTIAL FULFILMENT OF THE REQUIREMENTS FOR THE DEGREE OF DOCTOR OF PHILOSOPHY i n t h e D e p a r t m e n t o f METALLURGY We a c c e p t t h i s t h e s i s as c o n f o r m i n g t o t h e req u i r e d s t a n d a r d THE UNIVERSITY OF BRITISH COLUMBIA November, 1 9 7** In presenting this thesis in partial fulfi lment of the requirements for an advanced degree at The University of Brit ish Columbia, I agree that the Library shall make i t freely available for reference and study. I further agree that permission for extensive copying of this thesis for scholarly purposes may be granted by the Head of my Department or by his representatives. It is understood that copying or publication of this thesis for f inancial gain shall not be allowed without my written permission. IGOR A .E . WILKOMIRSKY Department of Metallurgy The University of Brit ish Columbia Vancouver, Canada Date November 1974 A B S T R A C T The development of a new, recirculating f luidized bed process for the roasting of molybdenite concentrates has been successfully completed in a bench scale pi lot plant. The process employs high-efficiency cyclones with a novel pneumatic injection system for continuously recirculating the calcines and feeding the molybdenite concentrates into the reactor. The f luidized bed consists of a mixture of calcines and coarse sand with a wide size range. The latter provides smooth f lu idizat ion behaviour and an at t r i t ion effect to prevent calcine particles from agglomerating. The fine calcines (-325 mesh) are continuously elutriated from the bed and recycled to the reactor, while the coarser sand particles remain in the bed. A rotary mechanical scraper inside the f luidized bed prevents build-up of material along the reactor walls. The reactor design was based on studies of f luidizat ion character ist ics, gas and solid mixing, and part icle st rat i f icat ion using a two-dimension f luidized bed operated at room temperature. i i The performance of the process on molybdenite con-centrates from four different sources was evaluated. Con-centrates containing less than 0.04% calcium were roasted to molybdenum trioxide with sulphur levels below 0.15%, which is suitable for metallurgical uses. An economic comparison showed that the f luidized bed process is able to compete favourably with the existing multiple hearth process. The f luidized bed process has an output that is 30 to 50 times larger per total area of furnace than the multiple hearth roaster (3 to 5 times larger per unit f loor area), and a capital cost that is lower by 50%. In addition,potentially lower operating costs and higher levels of S02 in the off gases may be realized by the f luidized bed process. Tests showed that direct slurry feeding of MoS2 concentrates is feasible. This would eliminate the need for f i l t ra t ion and drying steps prior to roasting. Batch kinetic studies in the f luidized bed, and observations using hot stage and scanning electron microscopes indicate that the transformation of MoS2 to Mo03 is a complex process which involves an in i t i a l fast oxidation step followed by a slower second stage. The f i r s t step appears to be controlled by the rate of the chemical reaction which is f i r s t order with respect to oxygen concentration and strongly temperature dependent. In the slow second stage i t is possible that solid state diffusion is the rate l imiting process. The volat i l i za t ion and subsequent condensation of Mo03 seems to play an important role during the transformation. Continuous operation in a 12.5 cm diameter reactor showed that the main variables that control the f inal residual sulphur in the calcines are the average residence time of molybdenite particles in the f luidized bed, and the calcium content of the concentrates. Residence times in excess of 20 hours are required to achieve sulphur levels below 0.15%. The optimum temperature range for roasting was found to be very narrow, 520-550°C; in this range the f inal composition is not strongly dependent on temperature. i v TABLE OF CONTENTS Page ABSTRACT i i LIST OF TABLES . . . . . . . . . . . . . . . ix LIST OF FIGURES. x i i ACKNOWLEDGEMENTS . .xv i i i Chapter 1 INTRODUCTION 1 1.1 Minerals of Molybdenum. . 1 1.2 Ore Dressing of Molybdenum Minerals . . . . 3 1.3 Uses of Molybdenum 4 1.4 Extractive Metallurgy of Molybdenum . . . . 6 2 LITERATURE REVIEW 13 2.1 The Roasting of Molybdenite 13 2.2 Fluidized Beds 25 3 SCOPE OF THIS RESEARCH PROGRAM 33 4 EXPERIMENTAL EQUIPMENT AND OPERATION PROCEDURES 38 4.1 Two- and Three-Dimensional Models 38 v Chapter Page 4.2 Pilot Plant Equipment 43 4.3 Experimental Techniques . . . 5 5 5 GAS BEHAVIOR IN TWO- AND THREE-DIMENSIONAL FLUIDIZED BEDS 61 5.1 Fluidization Properties of Sand-Calcines Mixtures 61 5.2 Gas Bubble Measurement in Two-Dimensional Fluidized Beds 7 4 6 PARTICLE BEHAVIOR IN TWO- AND THREE-DIMENSION FLUIDIZED BEDS 85 6.1 Particle Strat i f icat ion . 85 6.2 Concentration Profi le of Reaction Solids in the Fluidized Bed 97 6.3 Elutriation of Calcines from the Fluidized Bed 103 7 GAS AND SOLID DISTRIBUTION IN THE FLUIDIZED BED REACTOR 113 7.1 Gas Tracer Experiments 113 7.2 Solid Tracer Experiments 129 8 KINETICS AND MECHANISM OF MOLYBDENITE OXIDATION 139 8.1 Batch Kinetic Oxidation of Molybdenite in the Fluidized Bed Reactor 139 8.2 Mechanisms and Morphology of Molybdenite Oxidation 152 8.3 Temperature of Particles During the Init ial Stage of Transformation 163 8.4 Estimation of the Total Time of Transformation 167 v i Chapter Page 9 CONTINUOUS ROASTING OF MOLYBDENITE CONCENTRATES IN THE FLUIDIZED BED REACTOR. 170 9.1 Operating Conditions Studied. . . . . . . . . 171 9.2 Optimum Operating Conditions 191 9.3 Material Balance on the Process 192 9.4 Slurry Feed Injection 196 10 . INDUSTRIAL APPLICATION OF THE FLUIDIZED BED PROCESS FOR MOLYBDENITE ROASTING 198 10.1 Scale-up of Fluidized Bed Reactor 199 10.2 Fluidized Bed Plant for Molybdenite Roasting 212 11 SUMMARY AND CONCLUSIONS. 221 SUGGESTIONS FOR FURTHER RESEARCH . 224 NOMENCLATURE 225 REFERENCES 232 APPENDICES 1 . 241 2 . 243 3 246 4 . . 274 5 250 6 251 7 259 v i i A p p e n d i c e s Page 8 264 9 . 267 10 270 11 . . . • • • 277 12 . . . . . . . . . 278 13 280 14 . . . . . . . . . . . . . 285 15 , 287 16 . . 289 ^ • • v i i i LIST OF TABLES Tab!e Page 1.1 Principal Molybdenum Minerals 2 1.2 Molybdenum Usage in the Western Countries 5 1.3 Chemical Analysis of Technical Grade Molybdenum Trioxide and Oxide Briquettes 8 2.1 Physical Properties of a - MoS2, Mo02 and Mo03. . . 14 2.2 Thermodynamic Functions of Some Molybdenum Compounds 16 2.3 Kinetic Studies on Molybdenite Oxidation 23 5.1 Minimum Fluidization Experiments. 63 5.2 Density of Pure and Bulk S i l ica Sand A n d C a 1 c i n e s 63 5.3 Bubble Diameter Measurements in 12.5 cm Two-Dimensional Fluidized Bed 77 6.1 Effect of Particle Size Distribution Strat i f icat ion 86 6.2 Particle Size Strat i f icat ion as a Function of Time 90 6.3 Particle Size and Terminal Velocity of Molybdenite Concentrates 106 ix Table Page 6.4 Elutriation Tests Performed in the Fluidized Bed Reactor, 12.5 cm Diameter 107 6.5 Elutriat ion Constants of Eq. (6.9) I l l 7.1 Gas Tracer Experiments 119 7.2 Average Values of K/.v, and D, n and D„ n 125 3 (be)b a,g r,g 8.1 Chemical Composition of Molybdenite Concentrates 141 8.2 Average Particle Size and Surface Area of Molybdenite Concentrates 141 8.3 Batch Kinetic Experiments 142 8.4 Calculated Rate of Reactionand Reaction Rate Constant for Molybdenite Oxidation 150 8.5 Kinetic Studies on Molybdenite Oxidation 153 8.6 Calculated Values of the Temperature at the Reacting Surface of MoS2 Particles During the Init ial Oxidation 166 8.7 Measured Values of Time of Reaction in "Chemical" and "Diffusional " Regimes 168 9.1 Variables Investigated for the Roasting of Molybdenite in the 12.5 cm Reactor 171 9.2 Optimum Range of Operating Conditions for Fluidizing Bed Roasting of MoS2 191 x Ta ble Page 9.3 Material Balance for Fluidized Bed Roasting of MoS2 Concentrates in 12.5 cm Reactor. 192 10.1 Dimensions and Operating Performance of a 10 TPD Plant for the Roasting of MoS2 214 10.2 Calculated Output of Reactors 214 10.3 Capital Cost for a Molybdenite Roasting Plant using a Multiple Hearth Furnace . . 216 10.4 Capital Cost for a Molybdenite Roasting Plant using a Fluidized Bed Process 217 10.5 Summary of Operating Costs 218 xi LIST OF FIGURES Fi gure Page 2.1 Phase stabi l i ty diagram for the system S(g) " 0 " M o ( s ) a t 9 0 ° O | < 1 8 2.2 Gas bubble models in gas-solid f luidized beds 28 4.1 Two-dimensional f luidized bed model and tracer injection device 39 4.2 7.5 cm diameter, three-dimensional ( left) and two-dimensional (right) f luidized bed pi exiglas models 41 4.3 7.5 cm diameter plexiglas f luidized bed model and recirculating system operating in closed c i rcui t 42 4.4 Schematic diagram of the 12.5 cm diameter f luidized bed reactor and continuous recirculating system 44 4.5 Cyclone system used in the f luidized bed reactor 46 4.6 Cyclone discharge rotary valve 48 4.7 Exploded view of recirculation and discharge system 49 4.8 Pneumatically operated venturi nozzle connected to the discharge valve (without insulation) 50 x i i Figure Page 4.9 Gas distribution grid and dispersion nozzle 52 4.10 12.5 cm diameter f luidized bed reactor and recirculating system during operation . . . . 54 4.11 General view of the f luidized bed pi lot plant for molybdenite roasting 56 5.1 Factor of volumetric increase of a static bed of sand as a function of the wt-% of calcines in the mixture 62 5.2 Pressure drop through the bed as a function of the superficial gas velocity 65 5.3 Minimum f lu idiz ing velocity for mixtures of sand and calcines 67 5.4 Fluidized bed height as a function of the superficial gas velocity 70 5.5 Bed expansion factor for the f luidized bed as a function of the superficial gas velocity 71 5.6 Pressure drop through the f luidized bed in the pi lot reactor as a function of temperature 73 5.7 Typical high speed pictures of two-dimensional f luidized bed model 75 5.8 Typical bubble measurements made on an "equivalent bubble diameter" 76 5.9 Bubble diameters measured in the two-dimensional f luidized bed model 78 5.10 Experimental and calculated values for the bubble diameter in the f luidized bed 81 x i i i Figure Page 5.11 Bubble diameter calculated using different empirical expressions 84 6.1 Particle size segregation in the f luidized bed as a function of the particle size distribution of the sand. 87 6.2 Particle size segregation in the f luidized bed as a function of the f luidizat ion time. . . . 89 6.3 Particle size segregation in the f luidized bed as a function of the in i t i a l composition of the sand-calcines mixtures 93 6.4 Particle size segregation in the f luidized bed as a function of the superficial gas velocity 95 6.5 Typical tracer test in the two-dimensional f luidized bed after an input of MoS2 tracer (black) into the mixture of Mo03 and sand (white) 99 6.6 Axial concentration profi le after a MoS2 tracer input. 101 6.7 Influence of the position of the dispersion nozzle on the concentration prof i le of MoS2 in the f luidized bed 102 6.8 Cumulative percent of MoS2 transferred along the f luidized bed 104 6.9 Elutriation rate and elutriation flux as a function of the superficial gas velocity 108 6.10 Semi-1ogarithmic plot of the rate of elutriation as a function of the superficial gas velocity 110 7.1 Response signal from S0 2 pulse input recorded by the infrared analyzer . . . . . . . . 115 xi v Figure Page 7.2 Exit age distribution function of tracer gas in the f luidized bed as a function of dimensionless time 117 7.3 Reactor dispersion number of gas in the 7.5 and 12.5 cm diameter f luidized bed reactors as a function of the superficial gas vel oci ty. 120 7.4 Calculated overall gas transfer coeff icient as a function of the f luidized bed height . . . . 124 7.5 Calculated profi le of the axial dispersion c o e f f i c i e n t o f g a s . . . . . . 126 7.6 Calculated profi le for the radial , dispersion coefficient of gas along the f 1 uidi zed bed . . . . . 1 26 7.7 Overall axial dispersion coefficient of gas in the f luidized bed as a function of the superficial gas velocity . . . . . . . . . 128 7.8 Response signal of the discharge point after a pulse input of MoS2 tracer 131 7.9 Dependence of solid residence time d i s t r i -bution function on dimensionless time . . . . . . 133 7.10 Calculated axial dispersion coefficient of solid as a function of the f luidized bed height 136 7.11 Calculated radial dispersion coefficient of solid as a function of the f luidized bed height . .136 7.12 Average axial dispersion coefficient of solids in the f luidized bed as a function of the superficial gas velocity 138 xv Figure Page 8.1 Mole fraction of molybdenite oxidized as a function of time for batch tests 143 8.2 Computed values of the rate of reaction as a function of the fractional conversion of MoS2 145 8.3 Rate of reaction as a function of temperature in the three stages of transformation 147 8.4 Influence of the oxygen partial pressure on the in i t i a l rate of reaction . . . . . . . . . 149 8.5 Arrhenius plot of the rate constant 151 8.6 Scanning electron micrographs of molybdenite samples oxidized in the hot stage microscope and in the f luidized bed reactor. . 156 8.7 Hypothetical view of the stages of oxida-tion of an MoS2 particle 162 8.8 Influence of temperature on the total time required for transformation of MoS2 to Mo03 in the f luidized bed reactor 169 9.1 Influence of the superficial gas velocity on the sulphur content in the calcines 173 9.2 Sulphur content in the calcines as a function of the temperature of roasting 175 9.3 Influence of the average residence time of solids in the 7.5 and 12.5 cm diameter f luidized bed reactors. . . . . 177 9.4 Sulphur content of the calcines as a function of the residence time of solids for different temperatures of roasting 179 9.5 Influence of the oxygen partial pressure on the sulphur content of calcines. . . . . . . . 181 xv i Figure Page 9.6 Influence of the temperature on the sulphur level of calcines fpr the Kennecott concentrates. 183 9.7 Sulphur content of calcines as a function of the average residence time of reaction for the Kennecott concentrates. 184 9.8 Sulphur level of calcines as a function of the residence time for low calcium concentrates. . 186 9.9 Sulphur content in calcines as a function of the roasting temperature for low calcium concentrates 188 9.10 Apparent rate of reaction of molybdenite in the f luidized bed reactor as a function of the fraction of sulphur oxidized 190 9.11 S0 2 in the off gases as a function of the MoS2 feed rate. . 195 10.1 Experimental and calculated residual sulphur in the calcines as a function of the average retention time and temperature . . . . . . . . . . . . . 205 10.2 Calculated reactor dimensions as a function of the sulphur level in calcines, for different output capacities 209 10.3 Calculated level of S0 2 in the off gases as a function of the reactor capacity . . . . . . 211 10.4 Layout of a 10 TPD fluidized bed plant for roasting molybdenite concentrates 213 10.5 Thermal balance for the f luidized bed reactor using slurry feed . . . 220 xv i i ACKNOWLEDGEMENTS I should like to extend my sincere thanks and appreciation to both Professors Keith Brimacombe and Paul Watkinson for their friendly and invaluable help, advice and continuous assistance in the form of discussions and original ideas through al l the steps of this research program. Thanks are also extended to the Professors of the Departments of Metallurgy and Chemical Engineering and fellow graduate students for helpful discussions. In particular Mr. Stewart Ballentyne, Mr. B.S. Prabhakar and Dr. Michael Fraser, were most generous with their time. Thanks also are due to Mr. Horst Tump and Mr. John Baranowski who sk i l fu l l y and patiently - most of the time - constructed and modified many times the different pieces of equipment required to complete this work; to Mr. Jim Brezden who ef f ic ient ly traced the figures of this thesis; and to other technicians who helped me in the several stages of this research. I am grateful to The University of Brit ish Columbia, and through i t to the people of Canada, for their generous assistance in the form of a graduate scholarship, to the xv i i i University of Concepcion, Chi le, for granting me a leave of absence and to the O.A.S. for further financial assistance. Very special thanks are given to my wife and our l i t t l e daughters for their love, patience and understanding. xix . . . and after nine hundred days of noise, moly -and sulphur dioxide Dear Lord, .1 pray now t i r e d and weary l e t my l a s t result f i t the bloody theory. xx Chapter 1 IN T R 0 D U C T I 0 N Molybdenum, discovered in 1778 by C.W. Seheele, found at the end of the nineteenth century i t s f i r s t p rac t i ca l appl icat ion as an a l loy ing element in s t e e l . Production of molybdenum f i r s t began i n d u s t r i a l l y in 1910, as a resu l t of i t s extensive usage in meta l lurg ica l and e l e c t r i c a l technology. 1 .1 Minerals of Molybdenum Molybdenum is one of the less common elements, with an average concentration in the earth 's crust of 0.001% [1] . Molybdenum-bearing deposits of indus t r i a l in te res t are of f i ve main types [2]: . ( a ) Q u a r t z v e i n s (b) Pegmat i t e s ( c ) H i g h t e m p e r a t u r e s k a r n b o d i e s (d) M e t a r p o r p h i c s c h i s t and g n e i s s e s ( e ) D i s s e m i n a t e d d e p o s i t s o f t h e p o r p h y r i c c o p p e r t y p e 1 Most of the commercially producing deposits are located along the cordi l lera systems of North and South America, but as an element, molybdenum is widely distributed over most of the earth's crust. The minerals of molybdenum of industrial interest are given in Table 1.1. Table 1.1 principal Molybdenum Minerals [3] Mineral • • j T — 1 : : . - ' • r : •• Composi tion Molybdenite MoS2 Wulfenite PbMo03 Molybdite Fe03 • 3Mo03 • nH20 Poweli te CaCMoWjO^ Ilsemanite Mo02 • 4Mo03 Bilonesite Mg Mo CU Pateraite Co Mo 0^  Of these, dply the f i r s t four are treated to recover the molybdenum, with molybdenite being the most important and most Common mineral. It alone accounts for over 90% of the.total world production of molybdenum. World production of molybdenum in 1971 was close to 200 mi 11ion pounds of molybdenum metal , with an annual 3 growth rate of about 9% [4]. The major producer of molybdenum is the United States, which generates about 60% of the total world production [5], followed by the USSR, Canada and Chile. Taken together, these countries account for over 90% of the total world production, as well as for the known reserves of ores. Other minor producing countries are Peru, Mexico, China, Norway, Finland, Greece and the Phil ippines. 1.2 Ore Dressing of Molybdenum Minerals Molybdenum-containing materials are concentrated almost exclusively by f lotat ion, which yields very high recoveries. Gravitational concentration and magnetic separa-tion are used in very limited cases, mainly to separate i ron. Molybdenite is easily f loated. Ores containing a few hundredths of 1% molybdenite yield concentrates containing 85-90% MoS2, with over 90% recovery [6]. Porphyric copper minerals, with 0.3 to 1.5% copper and 0.001 to 0.1% molybdenum are an increasingly important source of molybdenite. These minerals usually are treated by a collective f lotation of sulphides followed by a dif ferential f lotation in which the copper is depressed using sodium sulfide or another depressant. The molybdenite concentrate is then purified in three or more cleaning operations to obtain concentrates of 85 to 95% MoS2. An alternative practice is to depress the molybdenite with starch. The main impurities present in molybdenite concentrates usually are S i 0 2 , CaC0 3, A1 2 0 3 , CaSO^, Cu 2S, FeCuS 2, PbS, FeS2 and F e 2 0 3 , depending upon the source of the parent mineral. Due to the fine dissemination of molybdenite in the ores, the molybdenite is liberated by grinding the primary feed going to the f lotation c i rcui t down to -200 mesh, followed in some cases by a regrind of the primary molybdenite concentrates. As a result , the f inal molybdenite concentrates are extremely f ine, usually 100%-200 mesh, and in some cases, 100%-325 mesh. 1 • 3 Uses of Molybdenum Molybdenum finds its main use in ferrous metallurgy. Depending upon the country involved, from 75 to 90% of the total molybdenum produced is consumed directly by this industry. Table 1.2 gives a breakdown of usage of molybdenum in some of the western industrialized countries in 1971. With a solubi l i ty of about 8% in iron [7], molybdenum in steel exists mainly as a solid solution. However a fraction of the molybdenum exists in the iron as a complex iron-molybdenum carbide. The content of molybdenum in structural steels does not exceed 0.5% normally, while in high-speed tool steels i t reaches 7.5 to 8.5%, replacing tungsten. Molybdenum produces 5 Table 1.2 Molybdenum Usage in the Western Countries [8] Structural steels 45% Stainless steels 20 Tool apd high speed steels 11 Cast iron and steel mill ro l ls 6 Super al1oys 5 Molybdenum metal 4 Chemical and Lubricants 8 Mi seel 1aneous 1 TOTAL 100% several important modifications in steels: i t imparts a fine and uniform grain s ize; reduces the eutectoid decomposition temperature range for hardening and tempering; and improves the elast ic l imit of steels and the wear and impact resistance. When alloyed with chromium and nickel , molybdenum also eliminates the tempering brittleness of steels. In cast i ron, molybdenum reduces the grain s ize , thus improving the wear resistance and high temperature properties. Cast iron containing si l icon and molybdenum is employed as an acid-resistant material. Molybdenum is widely employed in 6 high temperature and ac id-proof s tee ls a l loyed with chromium, nickel and cpbal t . Molybdenum metal, due to i t s very high melting point of 2620°C [9 ] , second only to tungsten, is extensively used in e lec t ron ic and e l e c t r i c a l components, as f i lament supports and gr ids in e lec t ron ic tubes and as wire and ribbons in heaters for high temperature furances. Due to i t s low thermal neutron capture cross s e c t i o n , molybdenum metal is also used as s t ructura l material in atomic reac tors . Molybdenite is u t i l i z e d in a very pure form as a lubr icant due to the lamel lar nature of the hexagonal c rys ta l structure of MoS 2 . Sodium molybdate is used as pigment and dye, and molybdenum oxides Mo03 and Mo02 are increas ing ly employed as cata lysts in desulphurizat ion units at petroleum r e f i n e r i e s . 1 .4 Extract ive Metallurgy of Molybdenum Molybdenite concentrates are the s tar t ing material for the production of ferromolybdenum and several other ferromolybdenum a l l o y s , as well as molybdenum compounds such as molybdenum t r i o x i d e , ammonium paramolybdate, sodium molybdate and calcium molybdate. In almost a l l cases, the f i r s t step in processes which convert molybdenite to a usable form 7 involves the oxidation of MoS2 to Mo03. However, in some cases alternative hydrometallurgical routes may be used. 1.4.1 Pyrometallurgical Treatment of Molybdenite Concentrates .1 • The industrial practipe of oxidation (roasting) of the molybdenite concentrates is performed almost exclusively in the multiple hearth furnace. Rotary kilns and reverberatory furnaces which have also been used in the past for roasting, are either no longer iq use or are employed on a very limited scale. Fluidized bed furnaces have not been successfully ut i l ized although several attempts have been made to roast MoS2 in them on a pilot scale. Multiple hearth furnaces range in size from eight to sixteen hearths, and from three to six meters in diameter with an average throughput of 50 tq 1Q0 Kg/m2 x day of molybdenite concentrates, depending upon the size of the furnace and operating practice. The roasting temperature of gases along the hearths ranges from 400°C for the upper hearth to 650°C at the hottest middle hearth, decreasing to 25Q°C for the Iqwer hearth. Fuel is added to the upper and lower hearth levels to keep reaction at a reasonable rate. Gases usually are in countercurrent flow to the descending 8 sol ids, but provision of separate exit ports from each hearth for the gases is also common. From the total world production of molybdenite, over 95% is tregted in multiple hearth roasters to produce technical grade molybdic oxide. From this product, about 70% is retreated to obtain purer pxide or other molybdenum compounds and ferro a l loys. Technical grade molybdic oxide is marketed in powder or briquette form. Chemical analyses of standard products are given in Table 1.3. Table 1.3 Chemical Analysis of Technical Grade Molybdenum Trioxide and Oxide Briquettes Tech. M 0 O 3 M 0 O 3 Briquettes Moo3 79-90% 70-75% Mo equivalent 53-60% 47-50% Cu (maximum) 0.50% 0. 50% S (maximum) . 0.25% 0.25% C (maximum) 12% Si 0 2 , Al 20 3 , others balance balance The sulphur content of both technical grade molybdic oxide and oxide briquettes has to be s t r ic t ly controlled for steel 9 alloying purposes, with a maximum allowable l imit of 0.20 to 0.25% S. Copper and lead contents also cannt exceed 0.5%. Technical grade molybdic oxide and oxide briquettes are used directly for ferroalloys as alloying addition in struc-tural steels. In the latter case the oxide is added directly to the ladle where i t is reduced rapidly by the carbon in the stee l . Pure molybdic oxide is obtained from technical grade material by volat i l i zat ion of the oxide above 600°C to yield a product with about 99.975% Mo03. The fine powder and acicular crystals formed during condensation of the Mo03 vapor are compacted into briquettes to reduce the volume. Calcium molybdate is prepared by roasting the calcines with high grade limestone; the resulting product can be used directly as an addition to the steel furnace. Molybdenum metal is produced by hydrogen reduction of pure Mo03 or ammonium paramolybdate at 950 to 1100°C. The product obtained is a fine powder of about 0.1 to 6 microns in size. Ferromolybdenum is produced by the thermite process using technical grade molybdic oxide. Standard ferromolybdenum grades contain 58 to 64% Mo with less than 0.1% C. Molybdenum s i l i c i d e is also prepared by the thermite process. The most common grade marketed contains 60% Mo and 34% S i . 10 1.4.2 Hydrometallurgical Treatment of Molybdenum  T r i o x i d e At the present time, no hydrometallurgical process is in industrial operation to extract molybdenum directly from molybdenite concentrates or ores. Al l the existing hydro-metallurgical processes are ut i l ized to obtain molybdenum compounds by treating technical grade molybdic oxide. However, hydrometallurgical processes are used to treat the oxidized ores of molybdenum: powellite, molybdite and wulfenite. Hydrometallurgically treated calcines end up normally as two products: ammonium molybdate by treatment of the calcines with ammonia followed by a subsequent pur i f icat ion, and sodium molybdate by treating the calcines with sodium hydroxide. A r£sumd of the main extractive metallurgical processes employed for molybdenum ores and concentrates is gi ven i n Fi gure 1.1 1.4.3 Alternative Processes for Treatment of Molybdenite Several alternative methods'of treating molybdenite concentrates have been proposed. Pyrometal1urgical processes include decomposition of MoS2 in a plasma at 5000/8000°C [74,75], aluminothermic reduction of M0S2 followed by fused salt electrorefining [76] and electrooxidation of MoS2-C anodes in a fused salt bath [76]. Hydrometallurgical processes include pressure leaching in KOH [77], sodium hypochlorite 11 leaching [78,79] and electrooxidation [80]. A comprehensive survey of oxidizing agents for leaching molybdenite was given by Bhappu et al. [81]. None of the above processes has yet been applied on an industrial scale. P Y R O M E T A L L U R G I C A L T R E A T M E N T H Y D R O M E T A L L U R G I C A L T R E A T M E N T M o l y b d e n i t e C o n c e n t r a t e s >8535 MoS 2 O x i d i z e d O r e s Ca MoO* - PbMoC, - F e 2 (MoO*)• , , • >0. U Mo r o a s t i ng T e c h n i c a l M o l y b d i c O x i d e >55% Mo, <0.25% S , <0.5% Ca B r i q u e t t i n g w i t h p i t c h Thermi te r e d u c t i o n O x i d e b r i q u e t t e s >52% Mo <12S C <0.5% Cu <0.25% S D i s t i l l a t i o n L i m e s t o n e r o a s t i n g F e r r o -molybdenum 58-645! Mo C a l c i u m m o l y b d a t e >40% Mo M e t a l l u r g i c a l u s e s , 85 -0$ L u b r i c a n t s 2 . 1 * p i g m e n t s , f e r t i l i z e r s , c a t a 1 y s i s t s , 5 . 8 * Mo m e t a l l i c powder >95% Mo <2.S% Fe <1.5% S1 <0.5S Cu m e t a l s h e e t , r o d , w i r e It. 7* F i g u r e 1.1 P r i n c i p a l e x t r a c t i v e m e t a l l u r g i c a l p r o c e s s to t r e a t o r e s and c o n c e n t r a t e s o f m o l y b d e n u m . (_% r e p r e s e n t t h e a v e r a g e u s e I n 1 9 6 6 i n USA.) Chapter 2 L I T E R A T U R E R E V I E W 2.1 The Roasting of Molybdenite 2.1.1 Physical and Thermodynamic Properties of  Molybdenite and Molybdenum oxides 2.1.1.1 Physical Properties Natural molybdenite ( a - MoS2) has an hexagonal crystal lat t ice with a lamellar structure, in which the molybdenum atoms l ie between two layers of sulphur atoms in the form of a trigonal prismatic co-ordination polyhedron [12]. A r t i f i c i a l l y prepared molybdenite ( 8 - MoS2) has a rhombodedral structure [12]. Molybdenum dioxide has a distorted rut i le structure [13], while molybdenum trioxide has an orthorhombic crysta l lo-graphic structure. The change from the hexagonal structure of a - MoS2 to the orthorhombic structure of Mo03 during the roasting of the molybdenite plays an important role in the transformation process, as will be shown in Chapter 8. Some physico-chemical properties of these compounds are given in Table 2.1. 13 14 Table 2.1 Physical Properties of a - MoS2 s Mo02 and Mo03 a - MoS2 Mo02 Mo03 Molecular weight Density, gr/cc 25°C [16] Melting point °C [17] Boiling point °C [14] Crystal 1ographi c structure 160.08 4.80/4.88 1650/1700 decomposes hexagonal 127.95 6.342 decomposes decomposes di st. ruti 1 e 143.95 4.694 795 1100 orthorhombi c Molybdenum trioxide has one of the highest vapour pressures of the metal oxides. At 600°C, for example, the vapour pressure is 10 - 2 mmHg, whereas at 850°C i t is 22 mmHg. Kubaschewski, Evans and Alcock [14] give the following expressions for the partial pressure of Mo03 as a function of temperature: (a) from 298 to 1068°K: log p = -15,230T _ 1 - 4.02 log T + 27.16 mmHg (2.1) (b) from 1068°K to 1373°K: log p = -12,480T _ 1 - 4.02 log T + 24.60 mmHg (2.2) 15 Results from different vapour pressure studies show some discrepancies, due probably to the polymerization of the Mo03 molecules to a trimer form [15] (Mo0 3) 3. Calculated values of the vapour pressure of Mo03 as a function of temperature are l isted in Appendix 1. 2.1.1.2 Thermodynamic Functions The specif ic heat (Cp)» enthalpy of formation (AH° T) and free energy of formation (AG° T) are w,ell established for molybdenum disulphide and molybdenum tr ioxide. Few data exist however for the less stable molybdenum dioxide. In Table 2.2 equations -fo'rAC , A H ° T and A G ° T for these compounds are given. Calculated expressions for the enthalpies and free energies of reaction of other possible reactions occuring during the roasting of molybdenite are provided in Appendix 2. For t h e oxidation of MoS2 to Mo03 following the reaction: MoS, + 7/2 0 2 Z Mo03 + 2S02 (.2.3) (s) (s) at 550°C, a usual temperature for roasting operations, calcu-lated values of enthalpy, free energy of reaction and equilibrium constant are, respectively i T a b l e 2 . 2 Thermodynamic Functions of Same Molybdenum Compounds S p e c i f i c Heat, cal/mol jTemperature Range, °K a - Mo S;. MoCJj McO = AC, = -19.7 + 3,15 • TO-3 T .[19] P AC - 3.85 + 3.10 • IO"3 T - 3.08 . 10 s T 2 [20] 273 -298 -729 1068 R i , a c t i o n Thermodynamic Fun c t i o n , cal/mol Temperature Range, °K M°ts) + i S U g ) t « o s l ( s ) A H ! = &&j = -57,640 + 6.85T - 5.60 • IO"3 T 2 - 0.503 • 10 5 T"1 -89,000 - 14.42TlogT - 0.2 • l'0~ 3 T 2 + 87.86 T [21] [21] 273 - 729 M o U ) + *S 2 ( s ) I M o S 2 ( s ) -57,640 - 15.78TlogT + 5.60 • IO"3 T 2 - 0.252/ 10 5 T-1 • + 49 24T [21] 273 - 729 M o Cs) + 0 2 Mo02^.sj A H800°X = - 1 3 ° . ° 0 0 C22]-AG° = -140,100 - 4.6TlogT • 55.8 T [235 298 - 1300 M0(s) + fo2 t Mo0 3 ( s )" AH° = AGy = = -182,600 + 3.85T + 1.55 • 10" 3 T 2 +3.08 • 10 s T"1 = -182,600 - 3.85TlnT - 155 • 10" 3 T2 + 1.54 • 10 s T" [20] 1 + 89 .77 [20] 298 - 1068 17 A H 823°K -295,920 cal/mol A G 823°K -207,210 cal/mol K 823°K 4.98 x 10 These values indicate that the reaction occurs irreversibly at this temperature, with a large amount of heat generated. Coudurier, Wilkomirsky and Morizot [18] have ca l -culated the phase stabi l i ty diagram (Kellogg diagram) for the system - 0 - Mo ^ at 900°K, as shown in Figure 2.1. At the operating conditions of normal roasters with 1 to 3% S0 2 in the reactor and in the presence of air the only stable compound is Mo03. These conditions are depicted by the shadowed area in Figure 2.1. 2.1.2 Processes for the Roasting of Molybdenite 2.1.2.1 Multiple Hearth Process The multiple hearth process mentioned in Chapter 1 has been used for over sixty years, but l i t t l e research has been done on the transformation of molybdenite during roasting. Plant operations have been described in the l i terature [24,25]. Butters [26] has given a detailed description of the problems of gas disposal and solid 18 Figure 2 . 1 . Phase st a b i l i t y diagram for the system S(g) " 0 " M o ( s ) a t 9 Q 0 ° K -19 entrainment at the Murex plant in Rainham. The entrained solids from the reactor are recovered in a system of cyclones and a dust baghouse prior to return to the roaster. Wright [27] has described the operations at Endako Mines where about five mill ion pounds of molybdenite concentrates are roasted yearly in a 16 f t . diameter, 10 hearth furnace. Lastovitskaya et al. [28] have studied the different molybdenum compounds formed during roasting in an industrial multiple-hearth furnace, and Coudurier, Wilkomirsky and Morizot [18] invest i -gated the roasting of molybdenite in a 4-hearth p i l o t , multiple hearth furnace. 2.1.2.2 Fluidized Bed Processes The possibi l i ty that a f luidized bed with its inherent high efficiency and simpl ic i ty , could be used to roast molybdenite concentrates has attracted several workers, but attempts to use the process so far have met with l i t t l e success. The f i r s t known attempt was by Deev and Smirnov [29] who used a small batch f luidized bed to roast molybdenite, but no information was given on the feas ib i l i ty of the process. In the U.S.S.R. there have been several pi lot plgnt studies of f luidized bed roasting. Zelikman et al. [30] attempted to roast granulated molybdenite concentrates 20 using bentonite as a binder in a 1 m2 f lu idized bed at 570-580°C. Small pellets of 0.25 to 2 mm diameter were employed to minimize the dust entrainment prior to charging molybdenite to the furnace. The feed rate was 1160 Kg/m2 . day. Dust formed by at tr i t ion of the small granules and entrained by the gases amounted to 39% of the feed. This dust was only part ial ly oxidized and was granulated again before returning i t to the reactor. Thus the problem of dust entrainment seems to persist even when granulated molybdenite concentrates are used. The discharge product from the furnace contained about 0.6% S. Kononov et al. [31] have attempted to solve the dust problem by superimposing an electr ical f ie ld on the particles inside the f luidized bed. They found that dust in the off gases dropped by a factor of 12, whereas the output of the furnace increased about 10%. Unfortunately the inten-sity of the voltage, 70 KV, creates a new danger in operating conditions and results in high power consumption, making the innovation unattractive. In other work, Sada and Kobayashi [32] tried to roast coarse molybdenite concentrates of 35 to 60 mesh in a batch f luidized bed reactor of 10 cm diameter by 3 m length. The temperature of reaction was maintained at 630 to 650°C, which is well above the sintering temperature of the Mo03. Dust entrained in the off gases amounted to as much as 56% of the charge and contained 2 to 4% S. The remaining calcines, 21 comprising 44% of the original charge, had a sulphur content of 0.3 to 0.5%. Their research apparently did not progress beyond these batch experiments. Golanit, Korneeva and Stepanov [33], and Gorin et al. [34] have published work on the temperature control of f luidized bed furnaces for molybdenite roasting. However, industrial application of the process in the U.S.S.R. does not appear to have been achieved. Indeed, Lastovitskaya et al. [35] report in 1970 that "the fluidized-bed furnace cannot be used since the calcines obtained would contain 1.5 to 2% S." In a recent attempt, Grigoriu and Balasanian [36] roasted coarse molybdenite concentrates of 0.1-0.2 mm diameter in a batch laboratory f luidized bed between 450 and 600°C, obtaining a product with up to 85% of the sulphur removed. No published information has been found on f luidized bed roasting of molybdenite concentrates in North America or Europe. However, i t appears certain that in the U.S.A. attempts have been made in the past to use this process. None of these attempts at f luidized bed roasting have produced calcines of low enough sulphur to be used in the steel industry (0.25% S maximum), nor has the problem of entrainment of part ial ly converted dust been sat isfactor i ly solved. 22 2.1.3 Kinetics and Mechanism of Molybdenite Oxidation Heterogeneous gas-solid reactions such as the roasting of molybdenite have usually been analyzed mathe-matically by the unreacted shrinking core model. It is assumed that reaction occurs topochemically, beginning on the surface and moving toward the centre of the part ic le . A part ia l ly reacted particle then consists of an unreacted inner core of sulphide covered by a layer of oxide product. In the ideal case the overall rate of reaction of the particle can be controlled by the following steps: (1) Mass t r a n s f e r o f o x y g e n t h r o u g h t h e h y p o t h e t i c a l gas f i l m s u r r o u n d i n g t h e p a r t i c Ie, (2) D i f f u s i o n o f o x y g e n t h r o u g h t h e f o r m e d o x i d e l a y e r , i f t h e o x i d e i s not vo I a t i I e , (3) The r a t e o f t h e c h e m i c a l r e a c t i o n a t t h e o x i d e / s u l p h i d e i n t e r f a c e , (4) D i f f u s i o n o f t h e g e n e r a t e d S 0 2 t h r o u g h t h e o x i d e l a y e r , and (5) Mass t r a n s f e r o f S 0 2 t h r o u g h t h e gas f i l m s u r r o u n d i n g t h e p a r t i c l e . In actual cases, one or a combination of these steps can influence the rate,and a transition from one rate determining step to another can occur as the transformation progresses. As will be seen later , this classical picture of gas-solid reaction kinetics is inadequate for the molybdenite-oxygen reaction. T a b l e 2 .3 K i n e t i c S t u d i e s on M o l y b d e n i t e O x i d a t i o n W o r k e r s C o n d i t i o n s I m p o r t a n t R e s u l t s R e f e r e n c e Ghen B u i k c o n c e n t r a t e O x i d a t i o n w i t h a i r and o x y g e n T e m p e r a t u r e 3 8 0 - 5 5 0 ° C Ra te p a r a b o l i c f rom 4 3 0 - 4 9 0 ° C Above 5 5 0 ° C c o n v e r s i o n l i n e a r M o 0 3 o n l y p r o d u c t f o u n d 37 G o d f r e y and N e l s o n B u i k c o n c e n t r a t e O x i d a t i o n i n a i r T e m p e r a t u r e 4 8 0 - 6 0 0 ° C O x i d a t i o n s t a r t a t 4 0 0 ° C O x i d a t i o n i s l i n e a r 38 C a l i s t r u et al. B u l k p a r t i c l e s o f - 0 . 2 mm O x i d a t i o n i n a i r T e m p e r a t u r e 5 0 0 - 6 5 0 ° C C h e m i c a l r e a c t i o n c o n t r o l s a t i n i t i a l s t a g e s S o l i d d i f f u s i o n c o n t r o l s l a t e t r a n s f o r m a t i o n 39 Ong P r e s s e d c o n c e n t r a t e d i s c s T e m p e r a t u r e 3 6 0 - 6 4 0 ° C O x i d a t i o n i n a i r Rate p a r a b o l i c f rom 380 t o 5 3 5 ° C O x i d a t i o n r a t e l i n e a r 510 to 5 4 0 ° C A c t i v a t i o n e n e r g y 1 6 . 8 K c a l / m o l 40 Z e l i k m a n & B e l a e v s k a y a P r e s s e d c o n c e n t r a t e d i s c s O x i d a t i o n i n a i r T e m p e r a t u r e 4 0 0 - 6 0 0 ° C D i f f u s i o n i n s o l i d s t a t e a t 4 0 0 ° C C h e m i c a l r e a c t i o n + D i f f u s i o n a l i n s o l i d S t a t e c o n t r o l a t 5 0 0 ° C C h e m i c a l r e a c t i o n a t 5 0 0 ° C 4 2 , 4 3 Ammann and L o o s e T h i n c o n c e n t r a t e l a y e r O x i d a t i o n i n a i r and o x y g e n T e m p e r a t u r e 5 2 5 - 6 3 5 ° C C h e m i c a l r e a c t i o n c o n t r o l s o v e r a l l r a t e up to 70-80% o f c o n v e r t i o n A c t i v a t i o n e n e r g y 3 5 . 3 K c a l / m o l e 44 L a s t o v i t s k a y a et al. C o n c e n t r a t e s I n d u s t r i a l m u l t i p l e h e a r t h f u r n a c e O x i d a t i o n w i t h a i r D e t e c t e d M o 0 2 , M c O n , M o 9 0 2 6 , M 0 O 3 a t d i f f e r e n t l e v e l s o f f u r n a c e 35 C o u d u r i e r , W i l k o m i r s k y and M o r i z o t C o n c e n t r a t e s P i l o t m u l t i p l e h e a r t h f u r n a c e O x i d a t i o n w i t h a i r R e a c t i o n s e q u e n c e M o S 2 - M o 0 2 - M o 0 3 . O x i d a t i o n c o n t r o l l e d by i n t r a - p a r t i c 1 e d i f f u s i o n and c h e m i c a l c o n t r o l a t t h e end .18 Deev and S m i r n o v C o n c e n t r a t e s B a t c h f l u i d i z e d bed O x i d a t i o n w i t h a i r Below 4 0 0 ° C r e a c t i o n i s c h e m i c a l l y c o n t r o l l e d ; a b o v e t r a n s p o r t c o n t r o l 29 G r i g o r i u & B a l a s a n i a n C o a r s e p a r t i c l e s 0 . 1 - 0 . 2 mm B a t c h f l u i d i z e d bed T e m p e r a t u r e 4 5 0 - 6 0 0 ° C O x i d a t i o n w i t h a i r C h e m i c a l r e a c t i o n c o n t r o l s r a t e a t b e g i n n i n g , f o l l o w e d by d i f f u s i o n i n l a t e r s t a g e s 36 CONTINUED T a b l e 2.3 Continued) W o r k e r s C o n d i t i o n s I m p o r t a n t R e s u l t s R e f e r e n c e Z e l i k m a n & Voldman C o n c e n t r a t e s C o n t i n u o u s p i l o t f l u i d i z e d bed Between 500 t o 6 5 0 ° C r a t e i s c o n t r o l l e d by c h e m i c a l r e a c t i o n 46 Z e l i k m a n B u i k c o n c e n t r a t e S o l i d s t a t e r e a c t i o n between MoS 2 and M 0 O 3 R a t e f a s t 7 0 0 ° C , d e c r e a s i n g s h a r p l y w i t h d e c r e a s e i n t e m p e r a t u r e 47 G a l a t e a n u B u l k c o n c e n t r a t e S o l i d s t a t e r e a c t i o n between MoS 2 and M 0 O 3 O v e r 7 0 0 ° C , g a s e o u s M 0 O 3 p r o b a b l y r e a c t w i t h MoS 2 48 C a r d o e n P r e s s e d c y l i n d e r s 492 t o 670°C 0.003 t o 0.83 atm o x y g e n C h e m i c a l r e a c t i o n c o n t r o l t h e t r a n s f o r m a t i o n A c t i v a t i o n e n e r g y 42.4 ± 1 K c a l / m o l e 41 25 Despite several investigations having been conducted, the kinetics and mechanism of molybdenite oxidation have only part ia l ly been c l a r i f i e d . Contradictory results have been reported by workers using different techniques to study the oxidation kinet ics. Findings of these researchers are summarized in Table 2.3. It is evident from these results that a clearly defined mechanism for the reaction has not yet been estab-l ished. Nevertheless, the studies of Cardoen, Ammann and Loose, Ca l is t ru , Zelikman and Belaevskaya and Grigoriu and Balasanian al l suggest that at conversions up to about 70%, the overall rate is controlled by the chemical reaction rate. At higher conversions a diffusional or other slow step appears to control the overall rate [39,44,45]. For an industrial process very high fractional conversions of sulphide ot oxide are required. In this region very l i t t l e guidance can be obtained from the above kinetic studies since even the existence of a dense or porous layer of Mo03 has not yet been proven. 2.2 Fluidized Beds Fluidization is an operation by which fine solids are suspended in a f lu id - l i ke state by passing a f lu id upwards through a bed or part ic les. 26 In order to scale-up results from laboratory f luidized bed roasting studies to fu l l size industrial reactors, in addition to obtaining information about the kinetics of the reaction, a basic knowledge of gas and solid mixing and mass and heat transfer in f luidized beds is required. When a gas flows upwards through a bed at low flow rates, the bed remains stat ic as the gas percolates through. Increasing the flow rate causes the bed to expand and produces a limited movement of particles at the top of the bed; this is the so-called "expanded bed." As the gas flow rate is increased further, the particles become suspended in the gas: this is the "minimum f luidizat ion condition." Any further gas flow rate increase over this minimum condition produces gas bubbles that ascend through the bed in a f lu id - l i ke manner. The gas bubbles ascending along the f luidized bed are the major mechanism of solid mixing in the bed and represent a separate "dilute phase" which is dist inct from the f luidized solid i t s e l f , often referred to as the "emulsion phase." Penetration of gas from the bubble into the particulate phase forms the so-called "cloud" of the bubble, which plays an important role in the gas interchange between both gas and emulsion phase. Mathematical modeling of f luidized beds is based to a large extent on considerations of mixing and transfer 27 between the dilute phase (bubble plus cloud) and the dense particulate phase (emulsion). Research, theory and applications of f luidized beds up tp 1970 are well summarized in the books of Zenz and Othmer [49], Leva [50], Zabrodsky [51], Kunii and Levenspiel [52] and Harrison and Davidson [53]. 2.2.1 Gas Bubbles The importance of gas bubbles in f luidized beds was f i r s t ppinted out by May [54]. An attempt to predict mathe-matically the behaviour of individual bubbles was made by Davidson and Harrison [55]. In Figure 2.1-a, the main features of the Davidson and Harrison model are depicted. The model has one main noticeable deviation from experimental findings: the bubbles are not spherical as assumed, but are nearly hemispherical due to the formation of a wake at the bottom of the bubble. The wake is formed by the pressure gradient that exists from the top of the bubble to the lower section. Jackson [56] and later Murray [57] modified the original Davidson and Harrison model to account for these effects. In part icular, the Murray model (Figure 2.1-b) appears to represent the experimental findings more closely. The existence of the cloud predicted by the Davidson and Harrison model as well as the formation of the wake were demonstrated experimentally by Rowe [58] using coloured NO gas tracer. ggs recirculation flow lines inside bubble and cloud (flow is symmetrical on both sides) (a) Davidson - Harrison model emulsion flow lines gas recirculation flow lines inside bubble and cloud (flow is symmetrical on both sides emulsion recirculation flow lines in wake penetration of gas cloud gas bubble solid -gas emulsion flow lines for downward motion of emulsion penetration of gas cloud gas bubble solid-gas emulsion wake (b) Murray model Figure 2.2. Gas bubble models in gas-solid f luidized beds. 29 All proposed bubble models however have the main shortcoming that they can be applied only for individual , isolated bubbles ascending in a f luidized bed. The fact that bubbles coalesce, spl i t and change in size and shape as they rise through a f luidized bed has not yet been described quantitatively. 2.2.2 Fluidized Bed Models Mathematical models of the f luidized bed as a whole, attempt to predict the size and distribution of bubbles across and along the bed. The earl ier assumption that the be,d was an homogeneous mixture of gas and so l ids , or that gas and solids form two completely differentiated phases [54], was found to introduce large errors in the prediction of fast chemical reactions in the bed. The "bubbling bed" model of Kunii and Levenspiel [59] permits the calculation in a simple and, depending on the conditions, accurate form, of the mass and heat transfer amongst bubble, cloud and the particulate (emulsion) phase. This model considers a two-step interchange of mass and energy between the dilute phase (bubbles) and the particulate phase (emulsion): a transfer of gas between the bubble and cloud, and a subsequent transport from the cloud to the 30 emulsion. The model on based in an a r t i f i c i a l "equivalent bubble s ize ," that is assumed to be constant along the f luidized bed. However, this simplif ication introduces a substantial error. It has been shown that bubble size can vary by a factor of ten or more from the grid to the top of the f lu i di zed bed. Later, Kato and Wen [60] proposed a "bubble assemblage model" which accounts for the growth of the bubble along the bed. This model has been applied successfully to predict conversion of gas in catalyt ic f luidized bed reactors [60]. Yoshida and Wen [61] extended the application of the model tp the simulation of roasting of sphalerite. Calculated values agree .cfosely with some experimental data of Yagi et al. [62] but the fact that they calculated the bubble diampter using the Kobayashi et al. [63] expression, which has been proven to be the least accurate of a l l such equations, makes %he val idity of the results questionable. The fact that chemical reaction takes place in bubbles, clouds and wakes and in the emulsion at different rates requires that the bubble size and bubble population at any level of the bed be represented accurately. However, this apparently simple factor is most d i f f i c u l t to evaluate. That these models appear to be unable to account quantitatively in a rel iable manner for the many phenomena occurring inside the f luidized bed has precluded their application in the present work. 31 2.2-3 Elutriation of Solids Elutriation refers to the selective removal of fine particles from a f luidized bed reactor. This phenomena is particularly important for f luidized bed reactors that treat fine metallurgical f lotation concentrates, since much of the feed material may be carried over with the f lu id iz ing gases rather than being recovered as a bed overflow. Non-steady state studies on the elutr iat ion phenomenon have been made by Leva [64], Osberg and Charlesworth [65], Hanesian and Rankell [66] and others [67 ,68,69 ,70,71 ]. General relationships have been proposed by Wen and Hashinger [72] in the form •Vf7 = f Re. V U o "mfJ (2.4) which correlates the elutriation rate constant X* per unit area of the bed per unit time with the variables of the f luidized bed. Yagi and Aochi[73] proposed a similar expression in the form K*d = f 1 (Re )(Fr) u . I P (2.5) 32 A general expression that accurately describes the e lu t r ia -tibn rate for a wjde range of conditions remains to be developed. For design purposes, i t seems safer to use the actual data from the reactor to determine the value of the rate of e lutr ia t ion, at a given condition. Chapter 3 SCOPE OF THIS RESEARCH PROGRAM This research was undertaken with the primary goal of developing a f luidized bed process for the roasting of molybdenite concentrates. Normal constraints were placed on the direction of the development program, the most impor-tant being that the process devised had to be demonstrably competitive with the multiple hearth furnace, which currently enjoys widespread use for roasting molybdenite. A f luidized bed process was chosen as a probable alternative to the multiple hearth furnace since, as has been shown for the roasting of other metal sulphides, i t has several inherent advantages: (1) h i g h e r o u t p u t p e r u n i t a r e a o f h e a r t h (2) s i m p l e r p r o c e s s e q u i p m e n t w i t h p o s s i b l e r e d u c t i o n i n c a p i t a l c o s t s (3) c l q s e r c o n t r o l o f o p e r a t i o n (4) p o s s i b l y l e s s m a i n t e n a n c e (5) h i g h e r t h e r m a l e f f i c i e n c y 33 34 . (6) g r e a t e r c o n t r o l o f gas e m i s s i o n s (7) h i g h e r c o n c e n t r a t i o n o f SO2 i n • t h e r o a s t e r g a s e s As discussed i'n Chapter 2, f luidized beds have not yet been used on an industrial scale to roast molybdenite concentrates, unlike other metal sulphides (copper, z inc, nickel and iron)- The reasons for this lack of application have also been outlined in Chapter 2. Basical ly , attempts at f luidized bed roasting molybdenite have foundered on one or more Qf the d i f f i cu l t ies peculiar to molybdenite con-centrates. F i r s t l y , the concentrates are extremely f ine , haying an average particle size of about 10 microns, and are d i f f i cu l t to handle. Previous workers [82] have attempted to sol ve the problem of feeding such fine materials into a f luidized bed by granulating the MoS2. However these granules break down again due to at tr i t ion inside the reactor and the fines, which result , are blown out in the f lu idiz ing gas. In order to prevent excessive losses of the molybdenum, expensive dust collection systems are needed, with the dust being returned in granule form back to the f luidized bed. Secondly, $he vapour pressure of M 0 O 3 , the product; of the molybdenite roasting reaction, is very high evefi at relatively modest temperatures such as 600°C. This imposes an upper l imit on the temperature at which roasting can be conducted, since as the temperature increases above 35 600°C, the problem of material sintering inside the bed becomes severe. That the roasting reaction has to be carried out at such low temperatures results in low rates of oxida-t ion, further necessitating long retention times to achieve low levels of sulphur in the calcines. Thirdly, the sulphur content of the calcines must be less than 0.25% and indeed closer to 0.1%, the level routinely achieved in the multiple hearth furnace, i f the calcines are to be sold for metallurgical uses. The f luidized bed process developed in this work has been designed speci f ica l ly to overcome these major d i f -f icu l t ies that have plagued the earl ier workers. Problems associated with feeding the fine MoS2 particles and recovering the equally fine calcines from the bed have been solved by continuously recirculating the so l ids , and feeding them back into the reactor along with fresh MoS2 employing a specially designed pneumatic injector. In this system the fines e l u t r i -ated from the f luidized bed are collected in high efficiency cyclones and gravity fed back to the injector with a fraction being taken as discharge. Sintering of calcines inside the reactor has been minimized by using f luidized beds consisting of coarse sand and Calcines. The sand, which is not elutr iated, prpyides the necessary at tr i t ion to keep calcine particles separated. In addition, build-up of sinter along reactor walls has been avoided by employing a rotary scraper. F ina l ly , 36 low 'sulphur levels in the product calcines have been achieved by optimizing the roasting temperature and average part icle retention time. Long retention times are, of course, possible with [the solids recirculating system incorporated in this process. The design of the process and attainment of reason-able operating conditions have not been reached arb i t ra r i ly . Considerable research effort was also expended to gain knowledge of the micro- and macro-processes occurring inside the f luidized bed which affect i ts operation. Separate experiments were conducted in two- and three-dimensional f luidized beds, contained in plexiglas, to study the f lu id dynamics of the gas (both in the bubbles and in the emulsion) and the mixing behavior of the sol ids. The results of these experiments are presented and discussed in Chapters 5, 6 and 7. To study the oxidation kinetics under f lu id iz ing conditions, batch experiments were conducted. In addition, a qualitative kinetic study was carried out to determine the microscopic phenomena which proceed at the surface of a molybdenite particl during roasting. This work involved observations with a hot-stage- and a scanning electron-microscope, the results of wliich-are described in Chapter 8. F inal ly , the f luidized bed process was tested and proved using a 12.5 cm diameter reactor in a fu l ly continuous, bench scale pi lot plant. The results of these tests, in which four different commercial concentrates were roasted, are given in Chapter 9. The abi l i ty of this f luidized bed 37 p r o c e s s t o c o m p e t e w i t h t h e m u l t i p l e h e a r t h f u r n a c e i s d i s -c u s s e d i n C h a p t e r 1 0 . R e s u l t s o b t a i n e d i n t h e c o n t i n u o u s r o a s t i n g o p e r a t i o n h a v e b e e n s c a l e d u p t o m a k e t h i s c o m p a r i s o n . Chapter 4 EXPERIMENTAL EQUIPMENT AND OPERATIONAL PROCEDURES In order to obtain preliminary information on the gross behaviour of the f luidizat ion of molybdenite concentrates, two and three-dimensional plexiglas models were f i r s t con-structed. Then, based on the results obtained at ambient temperature with the models, 7.5 cm* and 12.5 cm* diameter stainless steel reactors were tested in a continuously operating pi lot plant. 4.1 Two and Three-Dimensional Models To study the particle transfer and bubble d i s t r i -bution, as well, as particle segregation in the f luidized bed, two-dimensional reactors were i n i t i a l l y fabricated, one 7.5 cm* wide by 1 m high by 1.25 cm thick, and a second .1 2.5 cm* wi de by 1 m high and 1.25 cm thick. These were constructed to represent a central section of the three-dimensional cyl indrical reactors. The 12.5cm wide, Reactors were made of 3-in and 5-in diameter pipe. 38 tluid bed / light source / / recycling nozzle distributor ceramic balls high speed movie camera tracer injection • device V • flowmeters 11 fluidizing air recycling air 3-way valve Figure 4.1. Two-dimensional fluidi;:ed bed model and tracer injection device. CO 40 two-dimensional reactor is shown in Figure 4.1. A general view of the equipment used in the f luidizat ion studies is shown in Figure 4.2. To determine the main f lu idiz ing characteristics of t h e material used, a. 7.5 cm diameter by 1 m model reactor, shown in Figure 4.2, was employed. The gas distributor in the model has 42 holes, each with a diameter pf j mm. To separate the entrained material from the f lu idiz ing bed gases, a cyclone system was also designed and constructed [83,84]. The basic principle of the f luidized bed roaster used in this work was the elimination of losses of material from the entrained so l ids , which is a serious problem i'n multiple hearth roasting, by continuously re-circulating these fine solids to the reactor. For this purpose a venturi-type injector was designed to feed the recirculated calcines from the cyclones back into the bed continuously, as shown in Figure 4.3. Using a single cyclone, the collection of solids was found always to be less than 80 to 90% of the entrained material. Therefore, to increase the collection efficiency a second cyclone was designed to operate at about 1.5 times th,e tangential velocity of the main cyclone. Figure 4 . 2 . 7.5 cm diameter, three-dimensional ( left) and two-dimension (right) f luidized bed plexiglas models . F i g u r e 4 . 3 . 7 . 5 cm diameter p l e x i g l a s f l u i d i z e d bed model and r e c i r c u l a t i n g system o p e r a t i n g in c l o s e d c i r c u i t . 43 4.2 Pi lot Plant Equipment Based on the information obtained from the plexiglas models, a 7.5 cm diameter by 1 mm high reactor was f i r s t constructed and tested. Subsequently, due to the problems associated with the size of this reactor, a larger, 12.5 cm diameter reactor was constructed. The major part of the study was then conducted on the larger reactor, schematically Shown in Figure 4.4. The reactor was made of 316 stainless steel tube, and was externally heated in the lower 25 cm section to prevent excessive heat losses. The external heaters were controlled manually with two separate variac control lers. The reactor was ipsulated with a 10-cm thick fiberglass packing. Two chrome!-alumel thermocouples, positioned at 10 and 15 cm above the gas distributor gr id , measured the average bed temperature. Another thermocouple at the top of the reactor monitored the temperature of the gases at the outlet. To avoid build-up of material along the walls of the reactor, a 6 rpm rotating arm scraper was insta l led. It vya^  driven from the top of the reactor by a 1/24 HP high torque motor. Air for f lu idiz ing the bed was preheated in an electr ical preheater, which was controlled automatically to within ± 1°C by a Honeywell feedback controller coupled with a chrome!-alumel thermocouple. The preheater was 33 4a. F i g u r e 4.4. S c h e m a t i c D i a g r a m o f t h e 12.5 cm d i a m e t e r f l u i d i z e d bed r e a c t o r and c o n t i n u o u s r e c i r : u l a t i n g s y s t e m . 45 1. Recycling Air Flowmeter 2. Distributor Air Flowmeter 3. Air Blowers 4. Compressed Oxygen, Sulphur Dioxide and Nitrogen 5. 12.5 cm Diameter x 1 m length Reactor 6. 6 rpm Scraper Device 7. Distributor Plate, 54 holes, 1 mm Diameter 8. Recycling Nozzle, 3 holes, 2.5 mm Diameter 9. Gas Calming Section of 1.25 cm ceramic beads 10. Fluidizng Air Preheater, 7.5 cm Diameter x 45 cm length 11. Recycling Air Preheater, 3.8 cm Diameter x 30 cm length 12. Modified Venturi Nozzle for recycling and feeding the solids 13. Discharge Flap Valve for Calcines 14. Solenoid for operating the Flap Valve 15. Automatic Adjustable Time for Solenoid 16. Bin and 2.50 Diameter Screw Feeder for Molybdenite 17. Adjustable Speed Motoreducer, 0.25 to 25 rpm 18. Product Discharge Bin 19. Compensating Pressure Chamber, 12.5 cm Diameter x 14 cm height 20. Rotary Star Valves, 44 mm Diameter and 38 mm Diameter 21 j Main Cyclone, 4 cm Diameter 22. Secondary Cyclone, 3.2 cm Diameter 23. Coil Heated Pipes and Cyclones 24. Gas Sampling Cooler for S0 2 Analyzer 25. Sulphur Dioxide Infrared Analyzer 26. Water Ejector Scrubber, 2.5 cm Diameter x 20 cm height 27 Water Tank, 25-1 28. High Pressure Centrifugal Pump 29. Exhaust Gases 30. Chromel-Alumel Thermocouples 31. Water Manometer 32. Electromagnetic Vibrators 33. S0 2 Analysis Recorder 34. Slurry Tank, 1-1 3,5. Agitator, 200 rpm 36. High Pressure Venturi for Slurry Feeding Figure 4.4. Legend. 46 47 f i l l e d with ceramic beads. The gas distributor was made of 316 stainless s tee l , 1/4 inch thick, with 54 holes, 8 mm in diameter (Figure 4.9). The solids separated in the cyclone system, of 4 and 3.2 cm diameter respectively (Figure 4.5), were dis-charged through an expanded cone at the apex of each cyclone, to decrease re-entrainment of sol ids . To avoid back flow of gases and solid from the recirculating system, a rotary valve system was designed, consisting of a stainless steel rotor rotating against a Teflon lining (Figure 4.6). Air cooling proved to be suff icient to avoid softening of the Teflon l ining even at very high discharges rates, resulting in reasonably long l i f e . The valves were rotated at 6 rpm (main cyclone) and 3 rpm (secondary cyclone), by means of a high-torque, 1/24 HP motors. Sol ids, discharged from the cyclones, were collected together through a 1 inch pipe and passed continuously to a distribution valve. This latter valve had an internal flapper connected to a solenoid so that the solid material could be directed either toward the product bin or the recycling system. An adjustable timer, with ranges of 0.6 %o 60 seconds allowed the proportions of calcines going to product or recycle to be varied. To minimize the fluctuations in pressure due to bubbling inside the bed, a pressure-absorbing chamber was connected to the upper section of the 48 Figure 4.6. Cyclone discharge rotary valve. discharge valve. The system is shown in Figure 4.7 before assembly. Al l components were made of 316 stainless s tee l , and insulated externally with 0.6 cm ceramic wool to avoid heat losses. The calcines to be recycled to the reactor were discharged by the flap valve to a specially designed, modified venturi nozzle (see Figures 4.7, 4.8 and AppendixiA). In order to maintain the operating drag at the venturi throat, a secondary gas-loop was added. The nozzle was externally insulated with ceramic wool. The venturi nozzle system operated with air which was preheated in a separate electr ic preheater similar to that used for the f lu id iz ing Figure 4 . 7 . Exploded view of recirculation and discharge system. 1 . D i s c h a rge b e l l 2. P r e s s u r e c h a m b e r 3 . D i s t r i b u t i o n v a l v e 4 . D i v i d e d d i s c h a r g e c o n e 5 . V e n t u r i n o z z l e 6. S c r e w f e e d e r i n l e t 7. F l a p p e r v a l v e arm 50 Figure 4.8. Pneumatically operated venturi nozzle connected to the discharge valve (without i n s u l a t i o n ) . 1. V e n t u r i n o z z l e 2. S e c o n d a r y a i r l o o p 3. A i r p r e h e a t e r k. R e c y c l e d c a l c i n e s p i p e 5. D i s c h a r g e d c a l c i n e s p i p e 6 . D i v i d e d d i s c h a r g e c o n e 51 a i r . Control of the preheat temperature, which was measured with a chromel-alumel thermocouple, was maintained manually by means of a variac. The gas and solids from the venturi nozzle were then injected pneumatically back into the reactor through a 0.5 cm diameter stainless steel tube. This tube crosses the distributor and bends upwards at the center of the reactor, ending 15 cm above the distr ibutor. In this way, a high pressure drop at the lower level of the bed was avoided. A dispersion nozzle, with three equiaxially located holes of 2.5 mm diameter through which the solids and gas f ina l ly discharged into the f luidized bed, was attached to the upper end of the stainless steel tube (Figure 4.9). Fresh molybdenite concentrates were fed by means of a 1-inch screw feeder connected to a variable speed motoreducer (Figures 4.4 and 4.7). To avoid agglomeration and compaction of the molybdenite concentrates inside the bin, a ver t ica l , 4 rpm rotary arch-breaker was insta l led. The molybdenite from the screw feeder was discharged below the flap valve into the divided cone (Figure 4.7). In this way, a continuous stream of recycled solids and fresh molybdenite concentrate entered the venturi nozzle to be recirculated to the bed. A 300 watt electromagnetic vibrator was clamped to the discharge system, providing a smooth, continuous flow of solids through the recirculation system. u r e 4 . 9 . Gas d i s t r i b u t i o n g r i d and d i s p e r s i o n n o z z l 1. Gas d i s t r i b u t o r 2. D i s p e r s i o n n o z z l e 3. T h e r m o c o u p l e s 53 Two rotary vane blowers provided air to the f luidized bed and the recycling system. Air flows were measured with two flowmeters. Compressed oxygen, sulphur dioxide and nitrogen, when required, were connected to the recirculation l ine , and metered through gas flowmeters. To avoid condensation of gaseous lower oxides of molybdenum, al l lines as well as the cyclones were kept above 250°C by means of heating tapes. A chromel-alumel thermocouple located outside the main cyclone measured the temperature of this system. Al l temperature-measuring thermocouples were connected to a Texas Instrument multipoint recorder. The particles that were too small to be recovered in the cyclone system were collected in a water scrubber. The scrubber operated in a closed loop with a high pressure centrifugal' pump, connected to a 25-T water tank. The water ejector had a modified throat and water entrance to avoid condensation and deposition of material prior to the water and gases being contacted at the ejector constriction (Figure 4.4). A Beckman infrared analyzer was connected to the off gases from the cyclones to monitor continuously the S0 2 content in the gases. From the water scrubber the gases Were f ina l ly scrubbed with a solution of sodium carbonate to absorb the S0 2 . The aqueous solution was continuously F i g u r e 4.10. 12.5 cm d i a m e t e r f l u i d i z e d bed r e a c t o r and r e c i r c u l a t i n g s y s t e m d u r i n g o p e r a t i o n . 55 recirculated by means of a small centrifugal pump connected to a reservoir tank containing 150 1 of solution. A view -of the 12.5cm reactor and recirculating system is shown in Figure 4.10 and an overall view of the pi lot plant in Figure 4.11. 4 . 3 Experimental Techniques 4.3.1 Plexiglas Models To study the bubble distribution in the two-dimen-sional plexiglas model, a mixture of s i l i c a sand and calcines of Mo03 was prepared at the desired composition prior to being charged into the bed. A high speed camera and a 35 mm camera with high speed f i lm, using angular illumination, was employed to record the instantaneous size and distribution o f b u b b l e s i n t h e b e d . For particle distribution measurements, a tracer of 2 to 5 gr of MoS2 was added to a tracer injection device (Figure 4.1). After a free bubbling period of the f luidized bed the recycling gas was diverted to the tracer injector for 3 seconds by means of a 3-way valve. Al l gas flows were then shut off simultaneously. After the test , samples of the sand-calcines mixture were taken every 2 cm from the distributor to the top of the bed, screened to separate the F i g u r e 4.11. G e n e r a l v i e w o f t h e f l u i d i z e d bed p i l o t p l a n t f o r m o l y b d e n i t e r o a s t i n g . 57 calcines from the sand, then weighed and further analyzed for sulphur. Results from the two-dimensional bed tracer studies were used to predict the solid distribution along the bed, as well as the particle size distr ibut ion. To determine the main f1u i di zing characteristics of the sand-calcines mixtures, a given amount of the mixture of the desired composition was added to the 7.5 cm three-dimensional reactor model. Gas flowmeters and water manometers permitted direct measurements of air flow and pressure drop through the f luidized bed. 4.3.2 Pilot Plant Tests Gas tracer studies were performed at 510-520°C to reproduce the actual conditions of operation. A delta input of S0 2 tracer was injected through the recycling nozzle, and the output signal recorded continuously with a recorder coupled to the infrared analyzer. To avoid overflow of gas, the tracer was added in 5 seconds at a flow rate of air identical to that blown under normal operating con-dit ions. The downstream detecting point was located at the top of the reactor. Glass wool f i l t e rs were used to prevent the entrance of solids to the analyzer during the tests. Solid residence time distr ibution tests were performed with, the aid of a special injector which could 58 feed 20 gr of tracer MoS2 directly to the discharge valve in about 3 seconds. The tests were conducted at 300°C to avoid agglomeration of solids in the bed at lower temperature or oxidation of the tracer at higher temperature. Samples were taken from the discharge system during the test while fresh calcines were fed at the same rate to maintain a constant total weight of material in the bed. Batch kinetic studies were performed by feeding 20 to 40 gr of molybdenite concentrates to the bed in two minutes. Samples of solid product were then taken periodically for analysis while the rest of the calcines were kept con-tinuously recirculat ing. Continuous experiments were also performed to test the operation and feas ib i l i ty of the pi lot plant process at steady state. Fpr these experiments, a charge consisting of coarse s i l i c a sand and previously roasted calcines, in the amount desired to obtain a given average residence time in the bed, were f i r s t charged to the reactor. The s i l i c a sand provided a thermal sink for the bed, an at tr i t ion effect overcoming possible sintering problems, and relat ively smooth f luidizat ion properties to the charge. The difference in size between the coarse sand (-40/+140 mesh) and the finer calcines (-325 mesh) ensured that at the operating superficial gas ve loc i t ies , the sand was not elutriated with the calcines. The reactor was then heated by means of the 59 external and gas heaters, to near the desired temperature, whereupon feeding at the fixed rate was commenced. The tempera-ture in the bed was then adjusted by decreasing the power input to the external heaters. After 1 to 1.5 hours of feeding, by which time steady state conditions were achieved, samples were taken every 15 minutes for 1 to 3 hours for analysis. During the tests , steady state conditions were checked con-tinuously by the temperature recording and by monitoring the S0 2 analysis on a recorder coupled to the S0 2 analyzer. At the completion of the test , the total calcines from the bed were discharged and weighed, as were the total calcines roasted during the test. The scrubber tank was also discharged and the solids weighed after f i l t e r ing and drying, Filtered l iquid was periodical ly analyzed for molybdenum. Elutr iat ion tests were conducted during or at the end of roasting experiments by taking samples from the d is -charge bin during 1 to 5 minutes. Slurry feeding to the reactor was investigated in a single test. A small , high pressure venturi nozzle and a pinch valve were used to control the drop-by-drop flow of slurry from an agitated tank to the bed. Compressed air operated the venturi as well as providing pressure to feed the slurry towards the venturi throat. The gas-slurry emulsion 60 was injected at the lower level of the bed, direct ly through a 3.5 mm stainless steel tube. Chapter 5 GAS BEHAVIOUR IN TWO AND THREE DIMENSIONAL FLUIDIZED BEDS 5.1 Fluidization Properties of Sand-Calcines Mixtures 5.1.1 Minimum Fluidization Conditions To determine the minimum f lu idiz ing conditions of sand-Mo03 calcine mixtures, a study was conducted in the three-dimensional, plexiglas f luidized bed. The mixtures used and conditions of the experiments are l isted in Tables 5.1 and 5.2 as well as in Appendix 5. The volume of the static beds containing mixtures of Mo03 calcines with 1 Kg of sand was measured. The beds were f i r s t packed with an electromagnetic vibrator until the volume reading became constant. The experimental results are plotted in Figure 5.1. The static bed height of the sand-calcine mixture is given by L = f. m i "m,sil (cm) (5.1) 61 62 Figure 5.1 Factor of volumetric increase of a s t a t i c bed of sand as a function of the wt-% of ca lc ines in the mixture. 63 Table 5.1 Minimum F Iu id iza t ion Experiments 3-inch diameter, three-dimensional f 1 uidized bed S i l i c a sand -40/+140 mesh; d si 1 = 0.028 cm Mo03 ca lc ines -325 mesh; * c a l = 0.001 cm wt-% calc ines in sand: 0 to 1 00 Un, s u p e r f i c i a l gas v e l o c i t y : 0 to 30 cm/sec FI ui di zi ng gas : A i r , 25°C, 1 atm Table 5.2 Density of Pure and Bulk S i l i c a Sand and Calcines p , Density (gr/cm 3 ) e 0 » Voidage ( f rac t i on) S i l i c a sand (quartz) 2.65 -S i l i c a sand, bulk, -40/+140 mesh 1 .58 0.404 M 0 O 3 4.70 -Calcines of Mo0 3 , bulk, -325 mesh 1 .24 0.724 64 where L . , is the height of the stat ic sand bed, and f. is the volumetric expansion factor. The following expressions were applied to calculate the average particle s ize , density and voidage: A = *s11 X s i l + *cal X cal ^ (5.2) Ps = Psil X s i l + Peal X cal ^ r ' c m ^ <5-3> £o = e 0 > s 1 l X S 1 1 + £ 0 j C a l X c a l (5.4) An alternative relationship often recomended for the calcu-lation of the average particle size based on the volume/ surface ratio [85], 's = X ... \ x , ( 5- 5) si 1 + cal d . • d , s i l cal could not be used since i t gave erroneous values of the minimum f luidizat ion velocity. For example, even at 20 wt-% of calcines, the calculated value for um^ was 0.035 cm/sec, compared with 5 cm/sec measured experimentally. Figure 5.2. Pressure drop through the bed as a function of the superficial gas velocity. 66 The c h a r a c t e r i s t i c curves of pressure drop through the f l u i d i z e d bed as a funct ion of the s u p e r f i c i a l gas v e l o c i t y , for the d i f fe ren t sand-calc ine mixtures tes ted , are given in Figure 5.2. In a l l cases, with the exception of the pure ca lc ines bed, a constant weight of sand ( 1 . 5 Kg) was used, and the weight of ca lc ines was var ied . Under these cond i t ions , a progressive increase of AP with increasing weight f rac t ion of ca lc ine can be seen. The unusual shape of the curve for pure ca lc ines is probably due to severe channeling which was observed even at very low s u p e r f i c i a l gas v e l o c i t i e s . P a r t i c l e s began to agglomerate af ter a few minutes of operat ion, apparently as a resu l t of the humidity of the a i r , and possib ly e l e c t r o s t a t i c e f f e c t s . The humidity problem does not e x i s t , of course, in the f l u i d i z e d bed reactor . The unusual curves were observed in a l l cases where the ca lc ine content exceeded 50 wt-%. As can be seen in Figure 5.2, at 50 and 60 wt-% calc ines the pressure drop decreases at high s u p e r f i c i a l gas v e l o c i t i e s due to the onset of agglomeration. Experimental values of the minimum f l u i d i z a t i o n v e l o c i t y , u ^ , were determined from Figure 5.2, and are plotted in Figure 5.3. Also presented in Figure 5.3 are calculated values of u ^ obtained from the expression given by Kunii and Levenspiel [85]: CD Wt. % calcines Figure 5.3. Minimum f luidiz ing velocity for mixtures of sand and calcines. 68 _ 2 d s P c - P, Umf = " 1650 u ' q ' "Re < 20 (5, The agreement between experimental and calculated values for u m f is only fa i r between 0 and roughly 60 wt-% calcines, and poor above this range. This discrepancy may result from the fact that the empirical equation does not take into consideration electrostatic attraction between small particles which as shown by Baerns [87], can be signif icant ly large. It is also possible that the interaction between particles in this system involving a bimodal particle size d istr ibu-tion was responsible for the observed difference between measured and calculated u ^.••The experimental results suggest that the smaller particles (Mo03) play a more important role in the f luidizat ion characteristics of the bed than the coarse particles (sand). From Figure 5.3, i t can be seen that the measured values of u m f drop rapidly up to about 50 wt-%; above this value the u m f decreases slowly with increasing calcine content. Due to the large discrepancy between calculated and experimental values for the u ^ , i t was decided to use the latter values for calculation purposes on the f luidized bed reactor. 69 5.1.2 F l u i d i z e d Bed Expansion Direct measurements were made of the height of the f l u i d i z e d bed as a funct ion of the s u p e r f i c i a l gas v e l o c i t y . The resul ts of these tests are given in Figure 5.4, where i t can be seen that above u 0 ~ 25 cm/sec, the expansion of the bed is no longer l i near with u 0 . This may be the consequence of slug formation. The measured values of the bed height at minimum f l u i d i z a t i o n from Figure 5.4 can be re lated to the s t a t i c bed height by the fol lowing r e l a t i o n s h i p : L m f - 1.048 L m (cm) (5 From Figure 5.4, the bed expansion factor, f = Lf/Ljpf for s u p e r f i c i a l gas v e l o c i t i e s greater than u m f was c a l c u l a t e d . Values are plotted in Figure 5.5. The bed height under f l u i d i z e d condit ions then can be computed by the re la t ionsh ip The voidage of the bed at minimum f l u i d i z a t i o n condit ions was ca lculated as fo l lows: 25 2 0 E o CD •o M 15 h Lm S 10 0, L m f A A - A - A _ A L m f 2 -A-A-A A. silica sand - 4 0 / + 1 4 0 mesh calcines - 3 2 5 mesh air, 2 5 ° C , I atm. 0 5 10 15 2 0 2 5 u 0 , Superficial ga:; velocity (cm/.sec) 3 0 o F i g u r e 5 .4 . F l u i d i z e d b e d h e i g h t a s a f u n c t i o n o f t h e s u p e r f i c i a l g a s v e l o c i t y —• 1—: - i — r r silica sand - 4 0 / 4 - 1 4 0 mesh calcines - 3 2 5 mesh air 2 5 °C , I atm. wt. % calc ines k o 10 u 0 , Superficial gas velocity Figure 5.5. Bed expansion factor for the f luidized bed as a function of the superficial gas velocity. 72 e m f = 1.048 e 0 (5.9) where, as shown in Eq. 5.7, the factor 1.048 was the volumetric ratio between the bed at minimum f lu idiz ing conditions and the stat ic bed. 5.1.3 Pressure Drop and Temperature At 25°C, the pressure drop through the f luidized bed (u 0 > u m f ) remained approximately constant as shown in Figure 5.2. However, a sharp decrease in the measured value of AP was always observed in the f luidized bed reactor as the temperature of the bed was increased, as depicted in Figure 5.6. This phenomenon, which seems to have received l i t t l e attention in the l i terature, is possibly caused by an increase in the "bed f lu id i ty . " This has been suggested by Mi i , Yoshida and Kuni [88] who found a decrease in the value of u m f with increases in the bed temperature. Channeling in the bed also may play an important ro le , but no direct evidence was obtained which could substantiate this poss ib i l i ty . From the experimental data obtained, between 150 and 550°C the pressure drop through the f luidized bed decreases, on the average, by 10 cm of water for every 100°C increase in temperature. 73 6 0 i O CM X e CD X ) o> N l CD t o CD O O t -5 0 40 2 3 0 j CL o 20 10 0 0 • A Of V O A • A T A ® ® ® A • ZS-zv cm/sec Test no. o A B 64 B 68 B 69 ^ B 70 V B 71 B 73 B 105 ® B 114 Q^, B 116 & B 119 • • JL _L 200 400 600 Temperature of bed (°C) Figure 5 . 6 . Pressure drop through the f luidized bed in the pilot reactor as a function of temperature. 74 5.2 Gas Bubble Measurements in Two-Dimensional Fluidized Beds In order to determine the bubble diameter and distribution in the f luidized bed, experiments were conducted in the 5-inch, two-dimensional model. Flash pictures were taken in each test (see Chapter 4, Figure 4.1) and the bubble diameter measured from the pictures. The conditions used in the experiments are given in Table 5.3. Typical pictures for different compositions of the bed and superficial gas velocit ies are shown in Figure 5.7. The adherence of the fine Mo03 calcines to the plastic cover of the f luidized bed made i t d i f f i cu l t to identify clearly the bubble diameter. However, an a r t i f i c i a l "bubble diameter" could be measured which approximates the diameter of the equivalent volume of gas for each bubble. This value was used to calculate the average diameter of the bubbles at a given level in the f luidized bed. Typical bubble measurements are shown in Figure 5.8. Average measured values of the bubble diameter are plotted in Figure 5.9 (a) to (d) and given in Appendix 6. From the experimental values of d b given in Figure 5.9, i t can be concluded that, despite the fa i r ly large scatter, the bubble diameter is not a function of the average size of the particles in the bed. This finding agrees with the results of Geldart [89]. When pure sand was used, a few smaller values of d h were measured probably due to Figure 5.7. Typical high speed pictures f luidized bed model . of two-dimensional 76 Figure 5.8. Typical bubble measurements made on an "equiva-lent bubble diameter." 77 Table 5.3 Bubble Diameter Measurements in 12.5 cm Two-Dimensional Fluidized Bed Gas: A i r , 25°C, 1 atm L : 28 ± 2 cm m High speed film ASA 400, f .1 .8 , 1/1000 sec Run Calci nes G f u 0 No. (wt-%) (1/min) (cm/sec) B . 1 0 6 - A . 1 0 20 20 A . 2 n 29 30 A . 3 II 38 40 B . 1 0 6 - B . 1 20 14 15 B . 2 II II 20 20 B . 3 •I 29 30 B . 4 II 38 40 B .106-C.1 40 14 15 C.2 II 20 20 C.3 M 29 30 C.4 38 40 B.106-D.1 60 . 14 15 D. 2 n 29 20 D.3 II 29 30 D.4 38 40 D.5 " 6 10 B . 1 0 6 - E . 1 80 6 10 E . 2 n 14 15 E . 3 20 20 E . 4 n 29 30 E . 5 n 38 40 78 £ o CO <L> E o — X2 o> O i_ CU > < l-o E o Cl> OJ t o "O Q> X) X) x> OJ o> O i_ CD > 5 0 4 3 21 u 0 = 10 cm/sec Run no. wt. % calcines A B-I06-D-5 60 • B-I06-E-1 80 u 0 = 20 cm/sec Run no. wt. % calcines o B-I06-AI A BI06-B-2 |._ ^ g. ir>^_r.o A B- I06-D2 • B- I06-E-3 0 20 40 0 (a ) (b) JL 0 10 20 30 40 l f , Fluidized bed height (cm) Figure 5.9. Bubble diameters measured f1ui di zed bed model. in the two-dimensional T T E E o 7 6 5 u 0 = 3 0 c m / s e c Run np_. wt. % colcines O B I 0 S - A - 3 0 A B I 0 6 - B - 3 2 0 © B - I 0 6 - C - 3 A B - I 0 6 - D - 3 • B 1 0 6 - E - 4 " i — — r ( c ) ( d ) Ay u 0 = 4 0 c m / s e c Run no. wt. % c a l c i n e s 0 O B - I 0 6 - A - 4 A B I 0 6 - B - 4 © B I 0 6 - C - 4 A B - I 0 6 - D - 4 • B - I 0 6 - E - 4 0 2 0 4 0 6 0 8 0 10 2 0 3 0 4 0 l f , F l u i d i z e d b e d he igh t (cm) 5 0 F i g u r e 5 . 9 ( C o n t i n u e d ) 80 more gas (from 50 to 25%)being required for f luidizat ion alone and proportionately less for bubble formation over the range of velocit ies studied (20 to 40 cm/sec). From the measured values of d b , a best - f i t line was traced for each case. No attempt was made however to treat the data obtained at u 0 = 15 cm/sec due to the large scatter. The data could be f i t ted by the following relat ion-ship. d b = 0.00265 l f 'mf (cm) (5.10) where d b is the diameter of the bubble in cm at the distance l.f (cm) from the distributor gr id; N^  is the number of holes per unit area in the distributor and u 0 and u m f are in cm/sec. This expression is valid for 1 ^ > 3-4 cm. Values of bubble diameter calculated using this equation are plotted in Figure 5.10 together with the experi-mental values of d^. It can be seen that agreement over the range of Uo from 18 to 45 cm/sec is reasonably good, considering the scatter of the experimental data, but is poor for lower values of u 0 . This discrepancy does not affect the usefulness of Eq. 5.10 in calculation for the f luidized bed reactor however since the range of superficial gas velocit ies used in the reactor was 17 to 35 cm/sec. In 81 Figure 5.10. Experimental and calculated values for the bubble diameter in the f luidized bed. 82 this range, the estimated error in d b using Eq. 5.10 is less than 10%. In applying the results of bubble diameter, measured using a two-dimensional f luidized bed model, to a three-dimensional reactor, i t must be assumed that wall effects in the model are not excessive. This assumption cannot be verif ied in this work but is necessary due to the d i f f icu l ty of measuring bubble diameters in a three-dimensional f luidized bed. Equation 5.10 can be compared with the empirical expressions found in the l i terature: (a) Kobayashi, Arai and Chiba [90]: d. = 1.4-1- d p b f s ys Uo ^ m f (5.11) (b)1 Geldart [91] d b = 1.43 ( > mf 0.4 1 I h J 19 J + 0.027 l f (u 0 - u m f ) (5.12) 83 (c) Chiba, Terashima and Kobayashi [92]: i 0 . 286 d b " d b 0 1 .25 + 1 (5. The calculated values of d b using these equations and the equation derived for the present work are compared in Figure 5.11. The best agreement is achieved with the equation of Chiba et al. , which they claim can be used for fine particles Poorer agreement was found with the relation of Geldart which holds for f luidized beds larger than about 1 m in diameter. The lack of agreement with the relation proposed by Kobayashi et al. suggest that the latter equation is in gross error. At only 9 . 6 cm from the distributor their equation predicts a bubble diameter that is equal to the width of the f luidized bed. Such conditions were never observed. 84 F i g u r e 5 . 1 1 . B u b b l e d i a m e t e r c a l c u l a t e d u s i n g d i f f e r e n t e m p i r i c a l e x p r e s s i o n s . Chapter 6 PARTICLE BEHAVIOR IN TWO- AND THREE-DIMENSIONAL FLUIDIZED BEDS 6 .1 Particle Strat i f icat ion From early tests performed in the three-dimensional f luidized bed consisting of mixtures of calcines and a narrow size range of sand, i t became clear that signif icant part icle size strat i f icat ion could occur during f lu id iza t ion . It was further recognized that this phenomenon could present problems in the f luidized bed reactor since the hoped for at t r i t ion effect of the sand, that i s , to prevent sticking of the calcines, and the smooth f lu idizat ion properties (expected with a homogeneous mixture of calcines and sand) might not be real ized. In order to determine the operating conditions under which strat i f icat ion could be avoided, a number of experiments were undertaken. An experimental approach was necessary due to the lack of information in the l i terature on the phenomenon: only two papers have been published 85 86 [93,94] neither of which applies to the present bimodal system. In the experimental work the effect of part icle size d istr ibut ion, bed composition and superficial gas velocity on the extent of strat i f icat ion has been assessed. The experiments were carried out in the two-dimensional f luidized bed. 6.1.1 Effect of Particle Size Distribution  S t ra t i f i cation Two ranges of size distribution of sand were tested: a narrow range and a wide range. Conditions of the tests are given in Table 6.1 and Appendix 7. Table 6.1 Size Distribution of Sand and Test Conditions Init ial composition of mixture = 20 wt-% calcines, -325 mesh Superficial gas velocity, u 0 =30 cm/sec Fluidization time, t = 3 min Test B.l Sand (-40/+70) mesh, 520 g Test B.95 Sand (-40/+140) mesh, 520 g Prior to each test, the bed was charged with a homogeneous mixture of sand and calcines and fluidized for three minutes. 87 l O O p 90 80 70 tn tu c 1 1 1 1 1 r = 30 c m / s e c initial : 3 min B l © B-95 O 20 wt. % ca lc ines f l u id i za t ion sand - 3 0 / + 7 0 mesh - 4 0 / + 140 " I top of bed 60 o o I •o c ° 501 a> 4 0 o o 5 30 20 I0h I p initial composition „ O ^ K o-=o—-—O—O _ j _ o - o - o - o - " i t o p o f b e d | ho o-J I L 10 20 30 \ w , Distance from the distributor (cm) Figure 6.1. Particle size segregation in the f luidized bed as a function of the particle size distribution of the sand. 88 The axial size distribution of so l ids , expressed in terms of weight per cent of calcines versus f luidized bed height, after three minutes of f lu id iza t ion , is shown in Figure 6.1 for the two size ranges of sand. It can be seen that for the narrow size distribution of sand, complete segregation between sand and calcines occurs. On the other hand, only limited segregation is found when a sand bed with a wide size range is used. In the latter case, no segregation exists in the upper two-thirds of the f luidized bed, whereas in the lower one-third of the bed, a gradient in the calcine concentration towards the distributor is evident. A small increase in the calcines fraction can also be seen adjacent to the distr ibutor. This is due to the fact that some material remains between the holes of the distributor where f luidizat ion is absent. The mechanism of strat i f icat ion for the case ofthe narrow size range of sand particles and calcines is discussed in detail in Section 6.1.3. Tests were also performed with the narrow range of sand (-40/+70 mesh) and calcines, -325 mesh, in which samples were f luidized for different periods of time and sub-sequently size-analyzed along the bed. Conditions of the tests and results are given in Table 6.2 and Figure 6.2. 89 100 D X E 03 c D U c o in m <u c o l m , Distance from the distributor (cm) Figure 6.2. Particle size segregation in the f luidized bed as a function of the f luidizat ion time. 90 Table 6.2 Particle Size Strat i f icat ion as a Function of Time Init ial composition of mixture = 20 wt-% calcines -325 mesh Superficial gas velocity, u 0 = 30 cm/sec Sand A -40/+70 mesh, 520 g Test Fluidization time t (sec) Total gas flow G f (1) Rate of Fines Transport No AM/At (g/sec) AM/AGf (g/cc gas) B.4 15 6.21 2.989 7.22 x IO"3 B.2 30 1 2.42 1 .416 3.42 x IO"3 B.l 180 74.52 0.472 1.14 x 10~3 The results in Table 6.2 and Figure 6.2 show clearly that the strat i f icat ion of fines in this bimodal system of particles with a large difference in particle size is a very rapid phenomenon, being vir tual ly completed in three minutes. The time-averaged rate of fines transport through the f luidized bed was measured from the gas distributor to the level of the bed where the solids composition equalled the i n i t i a l composition. Adopting this procedure, the rate can be written as follows: 91 AM At = S p. 'cal fl 1 'f - 0.1 f ' f 0 t t I J I J (gr/sec) (6.1) while, with respect to the gas flow i t is AM AG,-fl cal 0.1 (gr/cm3) (6.2) M is the weight of fines (g) transported in the time t (sec) from the distributor grid to the bed height, 1^, (cm), where the composition is equal to the in i t i a l weight f ract ion, X cal ' o f calcines due to the passage of G f cm3 of gas. S is the cross-sectional area of the two-dimensional bed (cm 2); "p is the apparent density of the cal ci nes-sand mixture, which has been defined in Eq. (5.3). From Table 6.2 i t is apparent that the rate of solids transported along the f luidized bed decreases with time as the lower section of the f luidized bed becomes depleted of f ines. The value calculated for (AM/AGf) are probably also a function of the weight fraction of calcines in the bed until a saturation value is reached (choking velocity of transport). At higher values of X°a-j , the rate of transport may become constant for a given size of fines in the bed. 92 6.1.2 Effect of Bed Composition on Strat i f icat ion A series of tests was carried out to determine the influence of bed composition on the s t ra t i f i c ia t ion of so l ids , using the sand of wide size range, -40/+140 mesh. Three different in i t i a l compositions of the sand-calcine mixtures were studied. Results of three minute s t rat i f icat ion tests for beds consisting i n i t i a l l y of 20, 40 and 60 wt-% calcines are shown in Figure 6.3. The strat i f icat ion profi les obtained are similar for the different compositions tested. These results indicate that for compositions between 20 and 60 wt-% calcines, the bed appears to maintain a constant com-position with time. This assumes that the three minutes of continuous f luidizat ion are suff icient to achieve a steady state of mixing, as indicated by the depleted zone in the three minute curve of Figure 6.2. The experimental results given in Figure 6.3 indicate that an axial gradient in the fines concentration of approxi-mately 0.7 g/cm could be expected in the f luidized bed reactor. 6.1.3 Effect of the Superficial Gas Velocity on  Strat i f icat ion Strat i f icat ion tests were also performed at three different superficial gas velocit ies using a bed containing 9 3 T 1 r 1 i 1 1 r——i r i 1 1 1 r E c D O T> C o CO cn <U c o o 80! 70 i u 0 = 3 0 cm/sec 3 min fluidization silica sand -40/+140 mesh B 64 • 6 0 wt. % calcines initially B 9 4 A 4 0 B 9 5 O 2 0 • ' U initial composition 1 0 2 0 3 0 Lm , Distance from the distributor (cm) F i g u r e 6.3. Particle size segregation in the f luidized bed as a function of the in i t i a l composition of the sand-calcines mixtures. 94 60 wt-% calcines. The results are shown in Figure 6.4. Here i t can be seen that s t rat i f icat ion of the particles decreases with increasing superficial gas velocity. In order to explain these results the individual mechanisms by which solids can be transported in the bed must be considered. From previous work in this area, i t is clear that there are three such mechanisms: 1) Upward transport as a suspension inside the rising gas bubbles, due to gas recirculat ion. For this mechanism to be important, however, the gas recirculation velocity inside the bubbles must signif icant ly exceed the terminal velocity of the calcine part ic les , which is about 0.25 cm/sec. This condition holds for the experiments reported here since the gas recirculation velocity, as calculated using the Davidson and Harrison [95] model, is approximately 8 cm/sec. The upward transport by this mechanism is a dynamic process in that a continuous stream of fine particles is dragged inside the bubbles from the wake (due to the pressure gradient that exists) and recirculated along the gas streamlines through the bubbles back to the emulsion. 2) Carry up in the wake of the bubbles due to the pressure gradient below the bubbles. 3) Transport through the emulsion phase as a result of gas percolation between the coarser sand part ic les. It is 95 100 Z3 X e C o Ifl CO c o o U5 c o o -0—O—o—o • A — A — A — A -T — i — i — r " i — r o. o-. A t nop of bed 60 wt. % calcines initially 3 min fluidization sand - 4 0 / + I40 mesh B 65 O u 0 = 10 cm/sec B 93 A u 0 = 20 B 64 • u 0 = 30 0 J L 10 J L 20 J L 30 ^ , Distance from the distributor (cm) Figure 6 . 4 . Particle size segregation in the f luidized bed as a function of the superficial gas velocity. 96 important to note that the direction of gas flow in this latter case may be either upward or downward, depending on the magnitude of the superficial gas velocity or al ternatively, Uo / u m f . Kunii and Levenspiel [96] have shown that the flow is downward when u 0 / u m f exceeds 6-11, and upward below this c r i t i ca l range. In the l ight of this knowledge, i t would appear certain that the observed strat i f icat ion is related at least par t ia l ly , to the third mechanism of transport, that i s , the direction of gas flow through the emulsion. The strongest evidence to this effect is the finding that marked s t r a t i f i c a -tion occurs experimentally for u 0 / u m f < 8-10 which coincides with the c r i t i ca l ratio determined by Kunii and Levenspiel. Thus at low flow rates when u 0 / u m f < 8-10, the flow of gas through the emulsion is upward, assisting the upward transport of fine particles (and strat i f icat ion) already provided by the gas bubbles and their wakes. If the gas flow is increased until the ratio u 0 / u m f , exceeds 8-10 the gas flow through the emulsion is reversed to the downward direct ion, thereby enhancing the return of fine particles to the lower regions of the bed where depletion occurs. It is conceivable that depletion of the lower regions of the bed accelerates as the fraction of fines in the emulsion diminishes. This could occur since a decrease in fines con-centration results in an increase in umf ; as the u m - c 97 increases more gas wil l percolate through the emulsion and so raise the capacity to carry the fine particles to the upper regions of the bed. Based on the results of these tests, i t was decided to use -40/+140 mesh sand in the f luidized bed reactor and work with a superficial gas velocity greater than 20 cm/sec to minimize s t ra t i f ica t ion . 6.2 Concentration Profi le of Reacting Solids in the  Fluidized Bed The special type of pneumatic injection system used in the pi lot f luidized bed reactor, was designed to avoid the sintering problem that can arise when a slow, steady flow of molybdenite is fed mechanically, such as by a screw feeder, into the bed. As mentioned ear l ier , the sintering can occur as a result of the highly exothermic roasting reaction near the feed entrance. In order to achieve an optimum design for the injec-tion system, knowledge of the efficiency of part icle d i s t r i -bution in the bed was considered to be essential . Since no information of this kind could be found in the l i terature a series of experiments was conducted in the two-dimensional f luidized bed. The experiments had a two-fold purpose: to determine how the particles in the feed were distributed by 98 the injection system and to measure the axial concentration profi le of the molybdenite in the f luidized bed. 6.2.1 Axial Concentration Profi le Measurements In these tests a delta input of MoS2 tracer of the same particle size as the calcines (-325 mesh) was in-jected into the f luidized bed through a dispersion nozzle (see Chapter 4, Figure 4.1). As in the previous tests, the f luidized bed consisted of a mixture of Mo03 calcines and sand. After injecting the tracer, the bed was immediately "frozen" by turning off a l l the gas flows. The mixture of sand-calcine was subsequently analyzed for sulphur, i . e . , MoS2, along the bed. A photograph of the f luidized bed after the input of tracer is shown in Figure 6.5. Two of the tests were conducted using a mixture of sand with a narrow size distr ibution (-40/+70 mesh) and calcines (-325 mesh). The results , which are given in Figure 6.6 show that the transference of MoS2 from the gas jet to the calcine-sand emulsion is very rapid, with about 50% of the tracer being transferred in the f i r s t 5 cm from the injection point. In Figure 6.6 i t can also be seen that the concentration of MoS2 tracer drops to nearly zero close to the surface of the bed. This indicates that al l the MoS.2 that was injected into the bed as a pulse has Figure 6.5. Typical tracer test in the two-dimensional f luidized bed after an input of MoS2 tracer (black) into the mixture of Mo0 3 and sand (whi te) . 100 remained in the bed and not been blown directly out inside the gas bubbles. The results also indicate that the transfer of solid becomes effective about 2 cm from the dispersion nozzle holes where the bubbles formed by the emerging jets have collapsed and coalesced with the gas from the distr ibutor. Two additional tests were conducted to determine the influence of the position of the injection point above the distributor on the concentration profi le of MoS2 using a wide size distribution sand (-40/+140 mesh). The injection point in these tests was located 12 cm above the gas distr ibutor , as compared to a position 1 cm above the distributor of the previously described tests. The results of the chemical analysis of the bed are shown in Figure 6.7. From a comparison of concentration profiles obtained using the dispersion nozzle at the two different levels i t can be seen that the general shape of the distribution of calcines with height are similar. However, in the latter case the calcines are distributed more over the whole bed and are present in signif icant concentrations up to near the bed surface. From Figure 6.7 no appreciable difference in the shape of the concentration prof i le of MoS2 can be seen for the different weight fraction of calcines in the bed and the different superficial gas velocit ies tested. The presence of tracer below the injection point is caused 2 0 , 101 CM CO c o D v_ C CD O C o o CL) O O 2 0 wt. % calcines : 3 0 cm / s e c 3 sec tracer input 10 2 0 l f , Fluidized bed height (cm) 3 0 Figure 6.6. Axial concentration profi le after a MoS2 tracer input. 1 0 2 r - 1 r 1 1 — r -B 9 4 © 4 0 w t . % calcines,u 0=30cm/sec B 6 5 O 6 0 " " u0= 10 " l f , Fluidized bed height (cm) Figure 6.7. Influence of the position of the dispersion nozzle on the concentration profi le of MoS2 in the f luidized bed. 103 by the downflow of emulsion due to displacement by r ising bubbles. As expected, the concentration profi le from the injection point to the top of the bed is different than that from the injection point down to the distr ibutor. This is apparently produced by the upward transport of solid (MoS2 in this case) along the f luidized bed as suspended solid inside the gas bubbles. The cumulative per cent of tracer transferred from the injection point up to the top of the bed for the four experiments reported above is plotted in Figure 6.8. It can be seen that in al l cases, most of the tracer trans-ference to the emulsion takes place near the injection point, but a more even distribution of the tracer is achieved when the injection point is further above the distributor leve l . For this reason, the injector in the f luidized bed reactor was positioned at the higher level above the distr ibutor . 6.3 Elutr iat ion of Calcines from the Fluidized Bed The basic principle of the recirculating f luidized bed designed for roasting molybdenite is that the calcines are continuously elutriated and transported from the bed in the gas stream, and recycled to feed. The driving force for this elutriation is the difference between the operating velocity and the terminal velocity of the calcine part ic les. T T i i p — i 1—~r—i j 1 1 1 1 1 r B 2 ^ "1 injection point at B 3 o J distributor level B 6 5 o 1 injection point at B 9 4 A J. 12 cm from distributor l m , Distance from the distributor (cm) Figure 6.8. Cummulative percent of MoS2 transferred along the f luidized bed. 705 The terminal velocity of the calcines can be c a l -. l a t e d from the relation given by Kunii and Levenspiel [ 9 7 ] 4 g d, 3p C . w9 d 0 . 5 (cm/sec) (6 .6) where C d is t h e drag coefficient which can be calculated from the v e l o c i t y independent group 3 o 4 q d p P d 3y P s ' p g , ^ ( 6 . 7 ) p a r t i c l e R e y n o l d s number is g i v e n by t h e relationship Re -ILfaJit p u (6.8) The a v e r a g e p a r t i c l e d i a m e t e r d~s , terminal velocity and R e y n o l d s number of t h e molybdenite concentrates and sand are tven i n T a b l e 6.3. The d e t a i l e d results of the size analysis a.-e p r e s e n t e d i n A p p e n d i x 7. The average part icle diameter i s tu'- :;n t o be t h e s i z e t h r o u g h which 50% of the solids pass. S i r e s .'ere d e t e r m i n e d by C o u l t e r Counter and Cyclosizer analysis 1 06 Table 6.3 Particle Size and Terminal Velocity of Molybdenite Concentrates Concentrate (cm) u t (cm/sec) Endako (-325 mesh) 0.8 x 10~3 4 x 10"2 0.34 Brenda " 1 .0 x 10"3 7 x 10" 2 0.45 B.C. Moly 1 .2 x IO"3 9 x IO"2 0.60 Kennecott " 3.8 x IO"3 28 x 10- 2 1 .90 Sand (-40/+140 mesh) 3.1 x IO"2 690 150 At the superficial gas velocit ies used in the pi lot f luidized bed reactor (u-b = 15 to 35 cm/sec), only the calcines will be elutriated while the sand wil l remain as a stable f luidized bed. For the sand the ratio u 0 / u t is between 0.021 and 0.05, whereas for the calcines u 0 / u t is between 125 and 370. 6.3.1 Elutriation Flux In order to achieve a steady state mass balance between feed and discharged product, an exact value of the elutr iat ion rate must be known. At operating conditions in the reactor, the ratio of elutriation rate to feed rate always was 1arger than 1 . 107 To determine the elutr iat ion f lux , F , defined as the flow of calcines discharged from the top of the bed per unit are per unit time, the rate of elutriat ion at steady state was measured for each of the molybdenite concentrates under investigation. Conditions of the tests are given in Table 6.4 and results (see Appendix 8) are plotted in Figure 6.9 St St St in terms of k = F S , where k is the elutr iat ion rate (g/sec) and S the cross-sectional area of the reactor. Table 6.4 Elutriation Tests Performed in the Fluidized Bed Reactor, 12.5 cm Diameter Temp. (°C) Uo (cm/sec) B.C. Moly calcines Brenda " Kennecott " Endako 330 - 580 330 - 550 520 - 570 235 - 530 13.8 - 33.2 15.6 - 25.3 23.6 - 29.6 14.4 - 23.6 From the tests, different curves for the e lutr ia -tion flux (or rate of elutriation) could be drawn for each of the four molybdenite concentrates used. It can also be seen that both particle size and superficial gas velocity pro-foundly influence the elutriation flux at a given value of u 0 . 108 2.0 o <D § l .5h C7> ' o X lux 1.0 c o o i_ U J - 0.5 L L . 0 ~TTM i r 5 reactor ^ Kennecott ( I ) o B.C. M o l y ( 2 ) • Brenda (3) A Endako (4 ) (4 ) (3 ) / o 3 reactor a B.C. Moly U T 160 140 120 ~ c £ 100 C7> O 8 0 a: 8 P r 6 0 4 0 . 2 0 c o J3 UJ 5 IQ 15 2 0 2 5 3 0 3 5 4 0 ( u o ~ u m f ) . E x c e s s Gas Velocity (cm/sec) 0 Figure 6.9. Elutriation rate and elutriat ion flux as a function of the superficial gas velocity. 109 For example, at u 0 = 25 cm/sec, for the f iner calcines of Endako the elutr iat ion flux is 2.4 times larger than for the coarser calcines of Kennecott. The value of k appears to depend exponentially on the superficial gas velocity. It should be noted also that the curves in al l cases begin at the terminal velocity of the par t ic les , between 0.04 and 0.28 cm/sec. To check on the probability that k varies expo-nentially with u 0-u f , a semi 1ogarithmic plot was drawn as shown in Figure 6.10. A linear relation is found for a l l four cases although a small deviation does exist at low values of u 0 . This may be due to the large scatter in the experi-mental data. Thus for each type of molybdenite concentrate tested, the general relationship k* = A e B ( u o ' u m f ) (g/min) (6.9) is found to hold for the rate of e lutr iat iqn. The experi-mental constants A and B for the concentrates tested are given in Table 6.5. The Calculated values of k using Eq. (6.9) agree well with the curves traced with the experi-tal points, with an error of less than 5%. men n o 10 9 8 7 « i 1 ! 1 ® Endako ^ Brenda o B.C. Moly Kennecott T 1 1—•—i r 2 0 2 5 Excess gas velocity Figure 6.10. Semi-logarithmic plot of the rate of elutr iat ion as a function of the superficial gas velocity. I l l Table 6.5 Elutriation Constants of Eq. (6.9) Calcines } - - -—__ A (g/min) B (sec/cm) Endako 7.80 0.107 Brenda 3.18 0.132 B.C. Moly 2.79 0.126 Kennecott 1.67 0.135 A general equation for the elutr iat ion flux of fines from a bimodal system of particles with a large d i f -ference in size may be expressed as p* _ A _B (u o-u f ) • -~ S . m f (g/cm 2sec) (6.10) Probably, the temperature influences noticeably the value of F* but the scatter of the experimental data does not permit this factor to be evaluated. The size of the reactor apparently has some influence on the value of k as can be seen in Figure 6.9 where data from the 7.5 cm reactor are plotted together with data from the 12.5 cm reactor. For the same superficial gas veloci ty , the elutr iat ion f lux , F , decreases from about 15% at u 0 -u m ^ = 20 112 cm/sec to about 50% at (uo-u -f) = 30 cm/sec. Lewis, Gi 11 i 1.and and Lang [ 9 8 ] found that the diameter of the reactor has no influence on the elutr iat ion rate for d^ > 10 cm (4 in) but for smaller diameters, the elutr iat ion rate decreases rapidly. Chapter 7 GAS AND SOLID DISTRIBUTION IN THE FLUIDIZED BED REACTOR 7.1 Gas Tracer Experiments Depending on the extent of mixing, different elements of gas may reside for different lengths of time inside the reactor. A measure of the distribution of residence times of gaseous elements in the reactor can be made by means of tracer studies. In this type of experiment, a tracer is injected into the reactor over a short time interval after which the concentration of tracer in the output stream is monitored as a function of time. Adjusted to unit cross-sectional area of the reactor, the concentration-time profi le is expressed as the internal age distribution function, C(t). From these values, the Reactor Dispersion Number (RDN) or inverse Peclet number for mass transfer can be calculated, which gives a measure of the extent of mixing of gas (or solid) inside the f luidized bed reactor. 113 114 7.1.1 Residence Time Distribution Function of Gas in  the Fluidized Bed Reactor Tracer tests were performed under normal operating conditions to determine the residence time distribution function of gas in the reactor and the extent of gas disper-sion in the axial and radial directions. S0 2 , acting as tracer, was injected in the form of a pulse into the bed through the dispersion nozzle (see Figure 4.4, Chapter 4), centred 15 cm above the distributor gr id. The output con-centration of S02 was then measured at the top of the reactor and continuously recorded with an infrared analyzer coupled to a chart recorder. Three different superficial gas ve loc i t ies , which covered the usual range of operating conditions, were tested. A typical recorded response is depicted in Figure 7.1 and data is given in Appendix 9. From the recorded concentration of the tracer in the outflow of gas, the exit age distribution function E(t) was calculated by means of a computer program (see Appendix 10). Values are plotted in Figure 7.2. To calculate the extent of gas dispersion in the bed, the "dispersion model" proposed by Levenspiel and Bishoff [99] was applied. This In fact , due to experimental d i f f i cu l t i es the S0 2 was injected into the gas stream on the entry side of the recycle gas preheater. 115 Figure 7.1. Response signal from S 0 2 pulse input recorded by the infrared analyzer. 116 model assumes that the contribution to back-mixing by gas flowing unidirectionalTy (along the axis of the reactor in the present case) can be represented by a Fick-type of diffusion equation, as follows: 9 £ a t a ,g 5 2 C (7.1) e f is the bed voidage, and Dg the axial dispersion coefficient or eddy d i f fus iv i ty which characterizes the degree of mixing axial ly along the f luidized bed. Using the dimensi oni ess terms, 9 = t / T = u 0t/L^. and x = 1 ^ / L ^ , equation (7.1) can be transformed [100] to the following: <LP_ 80 Ui f J .9 3 X2 (7.2) The dimensionless group (D e f / u 0 L f ) is the Reactor Dispersion Number or inverse Peclet number for mass transfer The mean value or centroid of d istr ibut ion, ip, is given by tC(t)dt C(t)dt (7.3) tracer injection" T — ' - i r '-detection point gas plug flow — l i 1— 60 wt. % calcines T = 5 1 0 - 5 2 0 °C 9 u = 24.4 cm/sec o u =26.1 A u = 28.8 I 2 3 4 6 , Dimensionless time E x i t age d i s t r i b u t i o n f u n c t i o n of t r a c e r gas in the f l u i d i z e d bed as a f u n c t i o n of dimensionless time. 118 By calculating the variance of the distribution of the tracer, o 2 , using the relationship ,00 t 2 C(t)dt C(t) L dt (7 the RDN for the reactor can be determined. The for a closed vessel: o o = 2(RDN) - 2(RDN) 1 - e' ( 1/RDN) (7, while for an open vessel cr = 2 (RDN ) + 8(RDN ) where a2 = oz/t2 . For a perfectly mixed (backmixed) reactor, RDN = °° , whereas for no axial mixing (plug flow) RDN =0. In practice [101], a value of RDN > 0.2 indicates a large degree of mixing and RDN < 0.002 a small amount of mixing. It should be noted further that for a pulse input the internal age distribution function C(t) of Equation 7.1 and 7.2 is equivalent to the exit age distribution function E(t): 119 E(t) = C(t) E(e) = c(e) (7.6) Since the injection point of the tracer was located 15 cm from the distributor gr id , the reactor should be c lassi f ied as being between a closed vessel (with no mixing below the injection point) and an open vessel where the same amount of mixing exists at any point inside the reactor. The computed values of RDN for both cases are given in Table 7.1. Table 7.1 Gas Tracer Experiments 63 wt-% calcines -325 mesh (Brenda concentrate) 37 wt-% sand -40/+140 mesh Tracer: 16.5 1 S0 2 Input time: 5 sec Test No . Temp. (°c) G f ,T (1/min) (cm/sec) RDN Closed Vessel Open Vessel B.l25-1 520 198 24.4 0.626 0.537 B.125-2 51 0 219 26.1 0.626 0.536 B.125-3 515 232 28.8 0.598 0.531 The values of RDN indicate that under both condi-tions considered, extensive gas mixing exists, approaching CO cn _i 0.50 0.45 o A 12.5 reactor T= 510 - 5 2 0 °C 63 wt. % calcines closed vessel open vessel 7.5 reactor • T = 2 5 °C 15 wt. % calcines closed vessel 0.40. 15 Figure 7.3 20 2 5 3 0 u 0 , Superficial gas velocity (cm/sec) Reactor dispersion number of gas in the 7.5 cm and 12.5 cm diameter f luidized bed reactors as a function of the superficial gas velocity. 121 the backmixed condi t ion (RDN = °°) rather than plug flow (RDN = 0 ) . In addit ion the s u p e r f i c i a l gas ve loc i ty seems to have only a minor inf luence on the value of RDN over the range of u 0 tested (Figure 7.3) . In considering these r e l a t i o n -ships i t should be recognized that the response signal may have been af fected by three experimental l i m i t a t i o n s : I . Some m i x i n g o f t h e t r a c e r and t h e a i r c o u l d o c c u r i n t h e r e c i r c u l a t i n g gas p r e h e a t e r p r i o r t o e n t r y i n t o t h e r e a c t o r . 2. Gas m i x i n g c o u l d a l s o o c c u r i n t h e f r e e b o a r d o f t h e r e a c t o r b e t ween t h e t o p o f t h e bed and t h e d e t e c t i o n p o i n t a t t h e t o p o f t h e r e a c t o r . (Due t o t e c h n i c a l p r o b l e m s i t was n o t p o s s i b l e t o l o c a t e t h e d e t e c t o r a t t h e s u r f a c e o f t h e bed.) A l t h o u g h t h e R e y n o l d s Number f o r t h e gas i n t h e f r e e b o a r d i n d i c a t e d a l a m i n a r f l o w c o n d i t i o n (Re. = .1000-1300 a t u 0 = 15-35 cm/sec) t h e r e w o u l d be some a x i a l m i x i n g o f t h e gas i n t h e f r e e b o a r d . 3. The t i m e r e q u i r e d e x p e r i m e n t a l l y f o r t h e p u l s e i n p u t , 5 s e c , i s r e l a t i v e l y l o n g , a p p r o a c h i n g a s t e p i n p u t . T h i s w o u l d have t h e e f f e c t o f i n c r e a s i n g t h e v a l u e o f OQ and c o n s e q u e n t l y t h e RDN . In Figure 7.3 the ca lculated values of RDN are also compared with the RDN obtained in the 7.5 cm reactor using S0 2 as t race r . This comparison shows that in both cases extensive gas mixing occurs. The more rapid decrease in the value of RDN with increasing s u p e r f i c i a l gas ve loc i ty for the 7.5 cm reactor may possib ly be caused by the formation of gas s lugs . 122 7.1.2 Gas Transfer Between Phases After the gas enters the f luidized bed as bubbles, i t transfers to the particulate phase by a two-step process: bubble to cloud plus wake, and cloud plus wake to emulsion. These transfer processes can be characterized by the transfer coeff ic ients, K^ b ( . j b and K ^ c g j b , respectively. Then the overall gas transfer coeff icient between the gas bubbles and the emulsion is given by the relationship [102]: 1 1 1 (be)b N(bc)b K(ce)b (1/sec) (7.7) The extent of the gas distribution axially along and radial ly across the f luidized bed can be estimated from the axial and radial dispersion coefficients D and D which a,g r»g are given by the following relationships, derived by Kunii and Levenspiel [103]: a ,g r b i u° ui l-b (be)b (cm 2/sec) (7.8) D =0.2 r,g b (be)b (cm 2/sec) (7.9) where b is a coefficient given by 123 b = a 'mf 'mf U 0 - U mf-' (1 - 6 - a<5) + m a ( l - e m f ) (7 The factor m in Eq. (7.10) represents the extent of gas adsorption by the s o l i d . For a non-adsorbing s o l i d m = 0, whereas fo r a high surface area ca ta lys t - type s o l i d , m ~ 10. The use of expression (7.7) to estimate the value of the overa l l t ransfer c o e f f i c i e n t appears to be adequate for the purposes of ca lcu la t ion in th is work. No appreciable improvement was found by using other expression proposed by Chiba and Kobayashi [104] and Drinkenburg and Rietema [105]. The ca lculated values (see Appendix 10) of D , a , g D r and K ( t , e )b a r e P i t t e d in Figures 7.4, 7.5 and 7.6 and the average values for the f l u i d i z e d bed are given in Table 7.2. In these c a l c u l a t i o n s , the expression developed e a r l i e r in this work (Chapter 5, Eq. 5.10) for the axial bubble diameter p r o f i l e has been app l ied . It can be seen that d^ exerts a strong inf luence on the value of ^(be)b' Near the d i s t r i b u t o r where large numbers of small bubbles e x i s t , the overa l l gas t ransfer c o e f f i c i e n t is about 10 times larger than near the surface of the f l u i d i z e d bed (Figure 7.4) These resul ts are in agreement with the f indings of Kobayashi and Arai [106] who also reported a sharp decrease in the 1 24 Figure 7.4. Calculated overall gas transfer coefficient as a function of the f luidized bed height. 125 calculated axial prof i le of ^(be)b' ^ e individual values of K (b e )5 "f r° r a given value of d b agree well with the values of Kobayashi, Arai and Sunakawa [107]. Table 7.2 Average Values of K ( b e ) b and D and D r ^ g in the Fluidized Bed 63 wt-% calcines -325 mesh (Brenda concentrate) 27 wt-% sand -40/+140 mesh S0 2 tracer Run u o d b K(be)b D a ,g D No. (cm/sec) (cm) (sec- 1 ) (cm 2/sec) (cm 2/sec) B.125-1 24.4 3.2 46 180 1 28 B.125-2 26.1 3.5 42 220 132 B.125-3 28.8 3.9 37 213 141 The large value calculated for the overall gas transfer coefficient near the distributor grid would indicate that for a fast chemical reaction, much of the reaction would take place near the distr ibutor. This effect could be less s igni f icant , however, for the present case since some solid strat i f icat ion was always observed near the distributor grid (see Chapter 6, Figure 6.3) under the Operating conditions of the f luidized bed reactor. 1000 •Z 500 100 1 26 Figure 7,5 Calculated profi le of the axial dispersion coefficient of gas. Figure 7.6. Calculated profi le for the radial dispersion coefficient of gas along the f luidized bed. 127 As can be seen, in Figure 7.5, the axial dispersion coefficient of gas in the f luidized bed is also strongly affected by the bubble s ize , increasing by about 25 times from the distributor grid to the top of the bed. On the other hand, for the radial dispersion coeff icient the i n -fluence of d^ is somewhat smaller, with an increase in value proportional to the increase in bubble diameter along the bed (Figure 7.6). The superficial gas velocity has a minor influence on D but a larger effect on D at the r »9 a »9 top of the bed especially where the value of K ( b E ) D 1 S s m a l l (see Eq. 7.8 and 7.9). A decrease in the value of K(be)b corresponds to a decrease in and an increase of D along the f luidized bed. At the top of the bed where a, g the dispersion is more vigorous due to large bubbles formed, D increases by about 50% for a corresponding increase of a ,g 15% in the superficial gas velocity. The average value of the axial dispersion coefficient of gas, calculated for the f luidized bed, is plotted in Figure 7.7 as a function of superficial gas velocity. For purposes of comparison, data from other studies are also presented. In the calculat ion, i t was assumed that the solids were non-adsorbing (m = 0), since no information is available on the value of the adsorption coefficient m of the gas by the solid (Mo03 + sand), Eq. 7.10. The values of 1000 8 0 0 o CD CO CJ E o CO O cn c O o o c o y> \_ CD Q. CO D x < cf Q 6 0 0 4 0 0 2 0 0 100, 1—— Yoshi_da el al_ MS catalysis 150/xm He tracer , 25 °C Shurgerl ( 1 3 1 ) glass beads 250/i.m He tracer , 25 °C o \ > (130) Baern-: et al silica sand I75^.m H 2 tracer , 25 °C Present work O m = 5 © m = 0 T = 5I0-520 °C S 0 2 tracer 63 wt. %, -325 calcines 15 2 5 u 0 , Superficial gas velocity 3 5 F i g u r e 7.7. O v e r a l l a x i a l d i s p e r s i o n c o e f f i c i e n t o f g a s i n t h e f l u i d i z e d b e d a s a f u n c t i o n o f t h e s u p e r -f i c i a l g a s v e l o c i t y . 129 D n are in a l l cases lower than others that have been a »9 reported (Figure 7.7) . This d i f ference may be due to the choice of m in ca lcu la t ing D since i f moderate •.»9 adsorption is assumed (m = 5) the ca lcu la ted values of D a »9 are about 45% larger (Figure 7.7) . A value of m ~ 5 may be more r e a l i s t i c than m = 0 since the surface of the Mo03 p a r t i c l e s (Chapter 8) is i r r e g u l a r and may adsorb gas. Due to the large values of the overa l l gas t ransfer c o e f f i c i e n t K (be)b near the d i s t r i b u t o r g r i d , an in jec t ion level of the MoS2 feed well above the d i s t r i b u t o r appears to be most convenient to avoid s in te r ing problems. S inter ing occurs as a resu l t of the heat generated by the highly exothermic oxidat ion reaction when a high concentration of MoS2 p a r t i c l e s are contacted with a i r . Such concentra-tions are found experimentally near the in jec t ion point (Chapter 6) . 7.2 So l id Tracer Experiments The so l ids entering the reactor via the d ispers ion nozzle are rapid ly t ransferred to the f l u i d i z e d bed, pre-dominantly to the emulsion phase as was shown in Chapter 6. In order to make a comparative estimate of the extent of mixing of the MoS2 - (Mo0 3 + sand) ins ide the f l u i d i z e d bed reactor , two tests were performed at d i f f e r e n t s u p e r f i c i a l gas v e l o c i t i e s . 130 7.2.1 Residence Time Distribution Function of Solid in the Fluidizied Bed Reactor A pulse input of MoS2 tracer was injected through the dispersion nozzle into the f luidized bed at a point 15 cm above the distributor gr id . The bed consisted of a typical charge of M 0 O 3 and sand in the proportion required for high levels of conversion. Samples were taken from the discharge, as a function of time to calculate the Reactor Dispersion Number for the sol ids. The response to the input signal plotted as %MoS2 in the discharged calcines versus time, is shown in Figure 7.8. The computed residence time distribution function E(e) is given in Figure 7.9 as a function of dimensionless time e. From the calculated values of ^he variance o Q for the so l id , the value of the RDN was computed. For the superficial gas velocit ies of 17.4 and 22.5 cm/sec tested, the RDN are 0.74 and 0.75 respectively, which indicate a reactor which is vir tual ly back-mixed. The two-dimensional studies using MoS2 as tracer (Chapter 6) showed earl ier that the sol ids entering the reactor are completely transferred to the emulsion phase without any bypass as elutriated sol ids . The tracer experi-ments on the reactor now show that this feed material becomes completely mixed w i t h the rest of the solids comprising the bed. This finding was expected for this reactor design in 131 t , Time (min) Figure 7 .8 . Response signal at the discharge point after a pulse input of MoS2 tracer. 132 which the solids have a large average residence time. These results also confirm the previous findings of other studies [108]. The near backmixing condition is a fundamental cri terion for complete ut i l izat ion of the reactor and provides the added advantage that scale-up is considerably simpl i f ied. 7.8) can be explained by assuming that the tracer that enters the reactor is transferred into the bed at a singe! leve l . This assumption is only approximate for the present case however since only about 70% of the tracer is transferred into the bed close to the injection point, the remainder being transferred up into the bed by bubbles (see Chapter 6). That only a portion of the tracer is transferred near the injector strongly influences the exit age distribution function determined. Van Deemter [109] suggests that for the case where a tracer is injected at one point the subsequent mixing will be pro-duced by two mechanisms of solid flow: upward flow of solids in the wake of the bubbles (in this case, also inside the bubbles as suspended particles) (see Chapter 6), and down-ward flow by gravity. The equations governing the solids mixing are presented below: The shape of the response curve obtained (Figure f + Ps ( c s - c s ) = 0 (7.10) F + Ps <Cs " cs> = 0 (7.11) 1 33 \ backmixed V RDN = CO o u 0 = 22.5 cm/sec • u 0 = 17.4 detection point tracer — injection gas plug flow \S RDN =0 RDN =0.75 RDN = 0.74 Ji! 2 3 9 , D mensionless time 9. Dependence of solid residence time distribution function on dimensionless time. 1 34 where c g and C g are the fractions of tracer in the down-flow and upflow respectively; f g and F g the fraction of solid moving downwards and upwards; v $ and Vg their velocit ies and p s the transfer coeff icient of sol id between both phases. Van Deemter solved both equations simultaneously by numerical methods and plotted the concentration of tracer as a function of the parameter (F V s / l f ) t , where t is the time of measuring the pulse at a given leve l . His pre-dicted concentration profi le of tracer with time coincides closely with the response obtained in the present work, where the concentration of MoS2 in the elutriated solids approaches the completely backmixed condition monotonically. This seems to confirm also the existence of the two mixing mechanisms in the reactor and the influence of the bubble transport. 7.2.2 Solid Dispersion in the Fluidized Bed. The axial dispersion coefficient of solids in the f luidized bed was calculated using the model proposed by Van Deemter [109] and modified by Kunii and Levenspiel [110] 2 a>s 3 6 u m f a b U o - umf (7.12) whereas for the radial dispersion coefficient of s o l i d , the expression of Kunii and Levenspiel [110] was employed: 135 D = 0.1875 (7.13) The calculated values of D and D (Appendix 10) are a ,s r ,s plotted in Figure 7.10 and 7.11 respectively. The axial bubble diameter profi le developed in the present study (Eq. 5.10) has been ut i l ized in the calculat ion. The c a l -culated values show the influence of the bubble diameter on the mixing of so l ids . For a large value of d b (nearer the top of the bed), i t is expected that the amount of solid carried upwards in the wake and in the bubble - and returned by downward flow - would be larger than for small, slower bubbles. The influence of the superficial gas velocity on D„ _ is small in the range tested but large for D, . r, s a , s mainly in the upper levels of the bed. The large values of D3 compared with D _ indicate that mixing of solids is a , s r, s almost exclusively a result of the axial transport of solids in bubbles and their wakes. A limited amount of radial mixing appears to be produced by a diffusion type of mech-anism since the gas and solid flow directions are almost exclusively axial . to compare with the present results. In Figure 7.12, a plot of the average value of D calculated along the f luidized bed is compared with the work of Bart [111]. A large difference exists between the two sets of results. Few data exist on experimental values of D a ,s Figure 7.10 Calcu1ated solid as a height. axial dispersion coeff icient of function of the f luidized bed E o - 6 o c u 4 o u o TD O CC 6 0 wt. % calcines T = 5 5 0 °C u Q , c m / s e c 10 2 0 3 0 Fluidized bed height (cm) 4 0 5 0 Figure 7.11. Calculated radial dispersion coefficient of solid as a function of the f luidized bed hei gh t. 137 Attempts have been made to predict by means of a mathematical model the concentration prof i le of the MoS2 inside the reactor at steady state conditions; the pre-liminary results appear to confirm that the transfer of solid is very rapid close to the dispersion nozzle, decreasing rapidly towards the surface of the bed. The concentration profiles determined coincide approximately with the concen-tration profi les found in the pulse test of Chapter 6. 138 Figure 7.12. Average axial dispersion coefficient of solids in the f luidized bed as a function of the superficial gas velocity. Chapter 8 KINETICS AND MECHANISM OF MOLYBDENITE OXIDATION The successful scale-up from the pi lot plant unit to fu l l size plant operation depends to a large extent on rel iable kinetic data. It is of particular importance to know whether the overall reaction rate is controlled by mass transport and/or chemical reaction steps. In order to obtain the necessary kinetic information, batch kinetic experiments were performed in the f luidized bed reactor. Also, to elucidate details of the reaction of MoS2 to Mo03, samples oxidized under various conditions were observed under the scanning electron microscope. 8.1 Batch Kinetic Oxidation of Molybdenite in the Fluidized  Bed Reactor In these experiments a sample of 20 to 40 g of molybdenite concentrates, which had been dried and screened to -325 mesh, was fed into the reactor under normal f lu idiz ing 139 140 conditions. The reactor had been previously charged with a known amount of calcines, usually 2 Kg. Samples were then taken at given intervals of time from the discharge bin, while the rest of the calcines were kept recirculating con-tinuously through the reactor. Concentrates from four different sources -B.C. Moly, Endako, Brenda and Kennecott Mines - were used in this study. The chemical, - particles size - , and surface area - analysis of the feed samples at -325 mesh are given in Tables 8.1 and 8.2 and are detailed in Appendix 7. The particle size distribution was analyzed with a Coulter Counter and Cyclosizer, while the surface area was measured by the BET method using Krypton gas. The results of batch experiments using molybdenite from B.C. Moly and Brenda Mines are given in Table 8.3 and plotted in Figure 8.1 At 450"C, the extent of sulphur oxidation can be seen to be vir tual ly linear - the rate is constant - with time over the range covered. However at 480°C the i n i t i a l rate of oxidation is faster and no longer constant. Increasing the reactor temperature further, results in a sharp increase in both the extent and in i t i a l rate of oxidation. Between 526 and 595°C a high in i t i a l oxidation rate can be seen. The fraction of sulphur oxidized in this stage varies from 0.5 to 0.75, depending upon the temperature. Table 8.1 Chemical Composition of Molybdenite Concentrates Source %MoS2 %S %Cu %Pb %Ca0 %Fe %Si02 %A1203 Kennecott 88.72 35.55 0.17 0.025 0.60 0.47 6.50 _ _ B.C. Moly 88.42 36.43 0.09 0.30 0.40 1 .63 3.34 Endako 88.95 34.62 0.05 0.035 0.03 0.30 3.25 0.38 Brenda 93.58 37.50 0.03 0.01 5 0.05 0.47 2.68 1 .02 Table 8.2 Average Particle Size and Surface Area of Molybdenite Concentrates cTs, average diameter (cm) S s , surface area (cm2/g) Kennecott 3.5 x I O - 3 2.5 x 10* B.C. Moly 1 .2 x 10 r 3 3.5 x 10* Endako 0.8 x I O " 3 4.1 x 10* Brenda 1 .0 x I O " 3 3 . 8 x 1 0 * 142 Table 8.3 Batch Kinetic Experiments Exp. No Concentrate Sampl e (g) Temp. (°c) Po 2 (atm) B.66 B.C. Moly 42.05 450 0.21 B.38 t i 49.20 480 II B.19 n 23.93 526 II B.58 n 43.20 530 n B.59 II 40.23 557 II B.24 n 32.05 560 II B.34 II 38.75 595 II B.108 Brenda 19.75 554 0.36 B.97 •I 21 .60 557 0.54 B.109 II 17.05 557 0.61 Once the rapid in i t i a l oxidation stage has passed, a slower transformation regime begins to take hold. At temperatures between 560 and 595°C a transitional stage can be seen, lasting about four minutes and representing between 0.075 and 0.10 of the total fraction of sulphur oxidized. At lower temperatures, 560 to 526°C, the transitional stage becomes longer in time and accounts for about 0.15 to 0.20 of the total fraction of sulphur oxidized. At 480°C and 143 T J N X o CM CO 3 10 c o o o o B.C. Moly Concentrate B66 H 450 °C , air B38 O 480 °C " B 19 V 526 °C " B58 © 530 °C " B 59 • 557 °C " B'24 • 560 °C " B-I08A 554 ° C - 3 6 % 0 2 B 34 A 595 °C B97 £f 557 ° C - 5 4 % •" BI09 © 557 °C- 6 I % " I 15 20 25 30 t , Time (min) 35 40 Figure 8.1. Mole fraction of molybdenite oxidized as a function of time for batch tests. 144 below, the fractional conversion follows smoother curves with no abrupt change in rate. Above 526°C, in al l cases, a slower reaction stage was observed at high conversions. In this latter stage an apparent constant or very slowly decreasing rate of oxidation exists. This probably persists until the complete conversion of MoS2 to Mo03 is attained. It is possible that i f further data were available at 480°C for longer times, similar behaviour would be found. An increase in the oxygen partial pressure was found to accelerate the in i t i a l rate of reaction although a slower rate was s t i l l found at high conversions, as can be seen in Figure 8.1. In the batch experiments, the Brenda concentrates were tested only with enriched a i r , so that i t is not clear whether the higher conversions obtained were due to the increase in oxygen concentration alone or due to a difference in react iv i ty . B.C. Moly and Brenda concentrates are very similar in terms of size of particles and specif ic surface area, as shown in Table 8.2. The rates of reaction reported by Ammann and Loose [112] are 10 to 25% higher than the values found in the present work. They used a thin layer technique in which a sample, with an area of 2.5 cm2 and a depth of about 45 par t ic les , was reacted. Although their bed of particles was shallow, considerable heat effects may s t i l l have been 145 Mole fraction of Mo S 2 oxidized Figure 8.2. Computed values of the rate of reaction as a function of the fractional conversion of MoS2. 146 produced in between particles with the result that the actual temperature of the reacting particles was higher than the reported values. From the experimental curves of Figure 8.1, the rate of reaction was computed as a function of the fraction of sulphur oxidized. For curved portions of the conversion-time plots, slopes were obtained by graphical differentiat ion using a mirror. The results are plotted in Figure 8.2. It can be observed that above 526°C the i n i t i a l , constant rate of oxidation for B.C. Moly concentrates prevailed up to 0.50-0.55 of the total fraction of sulphur oxidized. Results plotted in Figure 8.3 and given in Table 8.4 show that the in i t i a l rate of reaction increases rapidly with temperature. Increasing the oxygen partial pressure in the reacting gas increases the in i t i a l reaction rate in roughly a linear fashion, as shown in Figure 8.4. This finding may mean that the in i t i a l rate is controlled by the chemical reaction between the oxygen and MoS2 as follows: M o S 2 ( s ) + 7 / 2 ° 2 * M o ° 3 ( s ) + 2 S ° 2 (8.1) although transport control involving oxygen is also possible. The rate of this reaction based on the surface area of solid may be expressed as 147 Figure 8.3. Rate of reaction as a function of temperature in the three stages of transformation. 148 dt J 7 K s u 0 2 ( 8 - 2 ' where S s is the specif ic surface area of MoS2 (cm 2/mole), -dMoS 2/dt the rate of reaction (mole f ract ion/sec) , k g the specif ic rate constant based on the surface area of MoS2 (cm/sec) and the concentration of oxygen (mole/cm 3). The specif ic rate constant was calculated for each experiment as a function of the temperature for the i n i t i a l stage. Values are given in Table 8.4. The Arrhenius plot ..'of In kg vs 1/T (Figure 8.5) gives a reasonably straight l ine . The activation energy for this in i t i a l period of transforma-tion is then calculated to be 25.3 Kcal/mole. This value is somewhat lower than the value of 35.3 Kcal/mole calculated by Ammann and Loose [113], and also lower than the value given by Cardoen [114], of 42.4 Kcal/mole. It, however, is higher than the value of 16.8 Kcal/mole reported by Ong [115] < or that given by Zelikman and Belaevskaya [116] of 13.7 Kcal/ mole. The best agreement was found with the value of Cardoen and Sepulveda [117] of 25 ± 4.2 Kcal/mole. The magnitude of the activation energy lends support to the theory that the rate of chemical reaction is controllinci for the f i r s t stage of oxidation. This has been postulated by other workers [118,119]. Figure 8.4. Influence of the oxygen p a r t i a l pressure on the i n i t i a l rate of reaction. T a b l e 8.4 C a l c u l a t e d Ra te o f R e a c t i o n and R e a c t i o n Ra te C o n s t a n t f o r M o l y b d e n i t e O x i d a t i o n Exp.. No M o S 2 C o n c e n t r a t e Temp. ( ° C ) n I n i t i a l Regime T r a n s i t i o n Reqime F i n a l Regime p 0 2 (atm) - d M o S 2 / d t (mole f r a c t / s e c ) k s ( c m / s e c ) - d M o S 2 / d t (mole f r a c t / s e c ) - d M o S 2 / d t (mole f r a c t / s e c ) B.66 B . C . M o l y 4 5 0 0.21 2.00 X io-1*' 4 .1 X i o - 5 2.00 X io-* B.38 •I 4 8 0 II 4.67 X io-1* - 4 .75 X I O " 5 2.67 X I O " * B.19 II 526 II 9.63 X I O " " 2 .00 X io-* 2.94 X io-* 5.0 x I O - 5 B.58 u 530 II 1.11 X i o - 3 2 .31 X 1 0 - * 3.02 X io-* 5.9 x 1 0 " s B.59 II 557 II 1 .58 X To- 3 3 .29 X io-* 3.34 X io-* 6.2 x 1 0 " 5 B.24 n . 5 6 0 II 2.03 X i o - 3 4 .23 X io-* 5.01 X io-* 6.3 x 1 0 " s B.34 II 5 9 5 II 2.17 X i o - 3 4 .54 X io-* 4.17 X io-* 6.5 x I O " 5 B . 1 0 8 B r e n d a 554 0.36 4.46 X i o - 3 5 .27 X 1 0 " * 5.00 X io-* 1.2 x 1 0 " 5 B . 1 0 9 n 557 0.54 5.00 X 1 0 " 3 3 .94 X 1 0 " * 3.44 X 1 0 " * 6.0 x 1 0 " * B.97 H 557 0.61 5.53 X i o - 3 3 .85 X 1 0 " * 2.67 X I O " * 5.4 x 1 0 - * o 10 o CD e o c o o IO 4 o Q> D rr 6 0 0 T , Temperature (°C) 550 5 0 0 4 5 0 T Brenda Concentrate j- B 108 A 36 % 0 2 B 97 'ff 54 % " B 109 © 6 1 % " B.C. Moly Concentrate 10 o v a A B 66 B 38 B I 9 B 58 B 59 B 24 B 34 air E= 25.3 kcal / mole .10 1.20 l / T x 10 - 3 1.30 ( ° K ~ ' ) Figure 8 . 5 . Arrhenius p l o t of the rate constant. 1 52 In the transition stage, the effect of temperature on the oxidation rate decreases (Figure 8.3) until only a weak dependence exists in the f inal reaction period. The magnitude of the rates in this stage, and their weak tempera-ture dependence may indicate that a diffusional process, perhaps through solid Mo03 is rate l imit ing. This final stage for the B.C. Moly concentrates occurred between 0.70 and 0.85 fraction of MoS2 oxidized, as can be seen in Figure 8.3. In the case of the Brenda concentrates, the range was 0.925 to 0.975 approximately, and the rate of reaction was slower. The sharper transition of an apparent chemical reaction control to an apparent diffusion control regime observed for the Brenda concentrates may have been the result of the thermal effects discussed in the next section of this chapter. 8.2 Mechanisms and Morphology of Molybdenite Oxidation From the data obtained in the above kinetic experiments, evidence was obtained that two different steps seem to control the oxidation of MoS2. To study the nature of the reactions in these two regimes, samples of natural molybdenite with dimension of 1 to 2 cm by 0.5 - 0.6 mm thick were oxidized in a hot stage microscope. A large air T a b l e 8 . 5 S c a n n i n g E l e c t r o n M i c r o g r a p h s o f Samples o f M o S 2 O x i d i z e d i n the Hot S t a g e M i c r o s c o p e S a m p l e No. Temp. °C Gas T ime S e c . O b s e r v a t i o n s (a) 25 a i r — Sample o f M o S 2 f rom a l a r g e p i e c e (2 cm s i z e ) b e f o r e o x i d a t i o n . The l a y e r e d s t r u c t u r e o f t h e m o l y b d e n i t e can be o ) s e r v e d . C.b) 550 a i r 30 G e n e r a l v iew o f a sample o x i d i z e d i n a i r f o r 0 . 5 min a t 5 5 0 ° C . The o r i g i n a l smooth s u r f a c e i s c o v e r e d w i t h s m a l l n e e d l e l i k e c r y s t a l s o f o x i d e . (c ) 550 0 2 30 G e n e r a l v i e w o f a s a m p l e o x i d i z e d a t 5 5 0 ° C i n p u r e o x y g e n . C r y s t a l s a r e s t a r t i n g to n u c l e a t e p r e f e r e n t i a l l y a t the f a u l t s and c r a c k s o f t h e M o S 2 s u r f a c e . (d) 600 o2 180 I d e n t i f i c a t i o n t e s t done under the e l e c t r o n s c a n n i n g m i c r o p r o b e on s l a b - l i k e c r y s t a l s f o u n d on t h e s u r f a c e o f a l l s a m p l e s . A n a l y s i s shows t h a t t h e y a r e p u r e M o 0 3 . The back o f sample i s u n r e a c t e d M o S 2 . Ce) 500 a i r 600 Sample o x i d i z e d a t low t e m p e r a t u r e . No a p p a r e n t s u r f a c e a l t e r a t i o n i s o b s e r v e d . Some c r y s t a l g r o w t h on a r e s t r i c t e d s c a l e seems to have t a k e n p l a c e a t the c r a c k s . S m a l l p a r t i c l e s a p p e a r t r a n s f o r m e d to o x i d e . Cf) 550 a i r 180 A f t e r a s h o r t p e r i o d o f t i m a , c r y s t a l s o f Mo0 3 a r e f o r m e d a t t h e s u r f a c e o f M o S 2 i n an i r r e g u l a r p a t t e r n i n d i c a t i n g some p r e f e r e n t i a l s i t s o f n u c l e a t i o n . S l a b - l i k e c r y s t a l s a r e a p p a r e n t l y f o r m e d by o t h e r m e c h a n i s m s , as v a p o r c o n d e n s a t i o n o f M o 0 3 , and a r e g r o w i n g up f rom the s u r f a c e o r n o t a t t a c h e d ( a r r o w s ) . L a r g e r c r y s t a l s a r e g r o w i n g f r o m the s u r f a c e c r a c k s . (g) 550 a i r 300 The s u r f a c e o f M o S 2 i s n e a r l y c o v e r e d w i t h c r y s t a l s o f Mo0 3 w i t h a d i s t i n c t r h o m b o h e d r a l s t r u c t u r e . L a r g e c r y s t a l s a r e g r o w i n g normal to t h e s a m p l e f rom a c r a c k . T h i s i n d i c a t e s c l e a r l y t h a t g a s e o u s Mo0 3 was e v o l v i n g f rom t h e c r a c k , w h i c h was p r o b a b l y a t a h i g h e r t e m p e r a t u r e t h a n the r e s t o f t h e s a m p l e . The M o 0 3 c r y s t a l s a p p a r e n t l y f o r m s p o n t a n e o u s l y by c o n d e n s a t i o n ( a r r o w s ) . (h) 550 a i r 300 G e n e r a l v iew o f t h e o x i d i z e d sample s h o w i n g c r y s t a l s o f M o 0 3 t h a t a p p a r e n t l y a r e a t t a c h e d l o o s e l y , i f a t a l l , to the s u r f a c e l a y e r o f o x i d e . CO 550 a i r 180 C l o s e up o f c r y s t a l s o f M o C 3 a t the s u r f a c e o f M o S 2 . C l e a r l y , the c r y s t a l s a r e g r o w i n g a t random and a r e a p p a r e n t l y not a t t a c h e d t o t h e s u r f a c e . A g a i n a v a p o r m e c h a n i s m seems to p r o d u c e t h e s e c r y s t a l s . 550 a i r 180 S m a l l p a r t i c l e o f M o S 2 w h i c h shows t h a t c r y s t a l s of M o 0 3 grow by c o n d e n s a t i o n of Mo0 3 away f rom t h e s u r f a c e . P a r t i c l e s of M o 0 3 a r e f o r m i n g i n any d i r e c t i o n by v a p o r e j e c t i o n -c o n d e n s a t i o n c l o s e to t h e s u r f a c e ( a r r o w s ) . CONTINUED T a b l e 8 .5 ( C o n t i n u e d ) S a m p l e No. Temp. ° C Gas T ime S e c . O b s e r v a t i o n s U) 550 0 2 30 In t h e p r e s e n c e o f p u r e o x y g e n , i t seems t h a t the c r y s t a l o f Mo0 3 f o r m e d a r e v a p o r i z e d a g a i n , p r o b a b l y due to a ho ; t e r s u r f a c e o f t h e s a m p l e due to t h e f a s t c h e m i c a l r e a c t i o n i n p u r e o x y g e n t h a n i n a i r . (1) 600 a i r 600 A t h i g h t e m p e r a t u r e , and l o n g t i m e (10 m i n ) , t r a n s f o r m a t i o n i s c o m p l e t e . C r y s t a l s o f Mo0 3 grow away f rom o r i g i n a l c r a c k . A t t h e s u r f a c e , the p a c k e d c r y s t a l s t r u c t u r e seems w e l d e d t o g e t h e r , p r o b a b l y by v a p o r d e p o s i t i o n o f M 0 O 3 a n d / o r i o n i c m o b i l i t y a t t h i s t e m p e r a t u r e . tm) 25 a i r — Sample o f m o l y b d e n i t e c o n c e n t r a t e ( B . C . M o l y ) b e f o r e r o a s t i n g . P a r t i c l e s a r e f l a t and t h i n , r e s e m b l i n g t h i n i r r e g u l a r s l a b s . (n) 560 a i r 5400 Sample r o a s t e d i n t h e f l u i d i z e d bed r e a c t o r shows a s u r f a c e c o v e r e d w i t h r h o m b o h e d r a l c r y s t a l s o f M 0 O 3 . Some sma'i l c r y s t a l s seem to be b r o k e n f r o m t h e o x i d i z e d p a r t i c l e s . Co) 580 a i r 1800 Sample o x i d i z e d i n the f l u i d i z e d bed r e a c t o r . A g a i n p a c k e d c r u s t o f M 0 O 3 c r y s t a l s c o v e r s the p a r t i c l e . CP) 605 a i r 2700 C l o s e up o f a s a m p l e r o a s t e d i n t h e f l u i d i z e d bed r e a c t o r . The o r i g i n a l c r y s t a l s o f M 0 O 3 f o r m e d a r e m e l t e d f o r m i n g an homogeneous c r u s t o f M 0 O 3 . 155 flow rate was used to avoid mass transfer control in the gas phase. The samples were heated rapidly (1 to 1.5 min) and the temperature was measured with a bui l t - in thermocouple located at the contact between the hot stage and the sample. During the reaction period, samples were continuously observed directly through the microscope at 300 x magnificati After the desired time, the sample was cooled rapidly (~30 sec) and further observed using the scanning electron mi croscope. The oxidation was studied at temperatures of 500 to 600°C with both air and pure oxygen as oxidizing gases. Samples were also taken from the f luidized bed roaster for observation under the scanning electron microscope. A summary of the conditions and observations are given in Table 8.5. Scanning electron micrographs are given in Figure 8.6 (a) to (p). From the observations of oxidized samples, the following mechanism for the oxidation reaction seems probable 1) At low temperature (500°C) oxide grows evenly over the surface of the sample and the Mo03 formed does not distort the original MoS2 particles (Figure 8.6 (e)). 2) At higher temperatures (550 to 560°C), rapid formation of Mo03 crystals occurs at the surface at some preferential nucleation points (Figure 8.6 ( f ) ) . Unattached 1 56 (d)2IOO x| 600 °C ,0 2 180 sec 5/i.m Figure 8.6. Scanning electron micrographs of molybdenite samples oxidized in the hot stage microscope [ and in the f luidized bed reactor. 1 57 F i g u r e 8.6 ( C o n t i n u e d ) 1 58 F i g u r e 8.6 ( C o n t i n u e d ) 1 59 Figure 8.6 (Continued) 160 or loosely held oxide crystals capable of motion were always found on the surface (Figure 8.6 (f) and (!•)). This fact and the presence of large crystals growing from cracks and small particles (Figure 8.6 (g) and (j)) suggest that an important mechanism of Mo03 crystal formation during the i n i t i a l , rapid oxidation, is the vaporization of Mo03 at the reacting interface of MoS2, followed by condensation. Due to the large amount of heat evolved (-295.9 Kcal/mole at 550°C), local areas of high temperature may exist , which are capable of generating Mo03 vapour. This may condense on cooler crystals of Mo03 or form new Mo03 crystals spontane-ously. In support of this mechanism are observations of part ial ly oxidized samples which have shown that some of the particles appear to be melted (Figure 8.6 (p)). 3) In particles roasted in the f luidized bed reactor, the layer of Mo03 crystals covers the original MoS2 particles completely (Figure 8.6 (n) and (o)), leaving apparently few pores for oxygen transport toward the reacting MoS2/Mo03 interface. In al l cases, small, separate crystals of Mo03 were found to be present with the larger particles of transformed MoS2. The former may have been broken from the parent Mo03 material as a result of the at tr i t ion action of the sand during the f luidizat ion (Figure 8.6 (n)). 161 4 ) At high temperature (600°C) the Mo03 crystals formed, seem to melt, probably on the overheated surface of the reacting MoS2 particles (Figure 8.6 (p)). In this case, a more dense and uniform surface of Mo03 appears to be the result . In al l cases, between 500 and 600°C there is evidence that after the fast i n i t i a l step of Mo03 crystal formation on the surface of the par t ic le , the layer of oxide crystals does not grow outward further from the or iginal ly formed surface (Figure 8.6 (h) and (1)). Rather i t appears to grow inward toward the inter ior , slowly closing the space between the Mo03 crystals , perhaps by a condensation process until a relat ively impervious layer builds up. This hypothetical transformation mechanism at the surface of a MoS2 particle is schematically depicted in Figure 8.7. It may be possible that at the MoS2/Mo03 inter-face some Mo02 could be formed by the solid state reaction 6 M o 0 3 ( s ) + M o S 2 ( s y ^ 7 M o 0 2 ( s ) + 2 S0 2 (8. Nevertheless, X-ray analysis performed on samples from the f luidized bed reactor, at high levels of conversion (>99%) did not indicate the presence of Mo02. This suggests that Mo02, i f i t does exist during the transformation, lasts a 162 (I ) initial "chemical" regime (2) initial "chemical" regime ffThTTnThTTTTJm (3) transition regime Figure 8.7. Hypothetical view of the stages of oxidation of an MoS2 p a r t i c l e . 1 63 very short period of time and exists probably in a thin layer at the very surface of the MoS2 core. 8.3 Temperature of Particles During the Init ial Stage of  Transformati on The fact that a fast reaction occurs with a large evolution of heat at the beginning of the transformation means that an increase in the temperature of the reacting surface can be expected. This of course has the effect of increasing the vapour pressure of the Mo03 being formed An estimation of the particle temperature can be made by performing a heat balance on the part ic le . For s impl i f icat ion, the following assumptions have been made: 1) The particles are spherical with a diameter, v 2) The particles are suspended in air transferring the heat generated by conduction and radiation. 3) The particles are small enough that internal temperature gradients are small, and individual particles are at a uniform temperature. 4) During the in i t i a l stage, the oxidation and heat generation is a steady state process, that i s , 8q /9t = 0. 164 The heat generated by the reaction , q, i s : q =AHj ks C (cal/cm 2sec) (8.4) The heat dissipated by convection, q c > is given by: q c = h $ a s (Ts - T ) (cal/cm 2sec) (8.5) where h g is the heat transfer coeff icient between the MoS2 particle and the surrounding air (cal/cm 2sec ° C ) , a $ is the surface area of the par t ic le , and T and T are s g the temperature of the particle and the bulk gas, respectively ( °C) . The heat dissipated by radiation is given by the expression: q r = £ o a s k 100 - 9 100 I J I J (cal/cm 3sec) (8.6) where e is the particle emissivity, a is the Stefan-Botzmann constant. At steady state, the heat balance for a single particle becomes: 165 (A HT) ks C g . h s a s <Ts " V + e ° *s 100 T M (8 .7) The heat transfer coefficient h $ can be estimated assumi conduction into a stagnant medium where ng Ii d N = - J §. u k where kg is the thermal conductivity of air at temperature T . g The temperature of the reacting particle can then be calculated by successive iterations of the equation h s Ts + e a 100 aj ks C g + h s T g + e o ' T ^ Too (8 .8) The computer calculated values of T $ for different tempera-tures of the bulk gas are given in Table 8.6 and Appendix 11. The following values were used to compute the particle temperature. 166 Table 8.6 Calculated Values of the Temperature at the Reacting Surface of MoS2 Particles During the Init ial Oxidation T CC) T cc) P M 0 O 3 a t T s g (mm H ) 9 500 520 0.056 520 547 0.068 550 581 0.091 580 630 0.221 600 650 0.435 d s = 10"3 cm (Brenda MoS2 concentrates) kg = 4.32 x 10"6 (cal/cm sec °C) e = 0.8 c = 1.4 x IO" 1 3 (cal/cm 2sec °K*) Values of kg were taken from the experimental data (Figure 8.5). From the calculated values, i t appears that the particle temperature is substantially higher than its surroundings. This increase is magnified at higher temperature: for a gas temperature of 500°C, the surface would be about 520°C (Table 8.6), whereas at 600°C the surface will reach 6.50°C. 167 It should be noted further that these calculated values are bound to be too low since they do not take into account the continuous increase in rate of oxidation of the particle surface, and hence heat evolution, as the surface temperature r ises . 8.4 Estimation of the Total Time of Transformation The total time of transformation x (min) and the time to reach a given transformation level in the f i r s t "chemical" stage of oxidation, t 1 (min), were determined from the curves, suitably extrapolated, in Figure 8.1, Values measured are given in Table 8.7.. Values of x as a function of the temperature are presented in Figure 8.7. The best f i t line for this plot is given below: x = T° * 7 6 5 - T + 520 (min) (8.9) where T is in °C. This relationship is valid for molybdenite concentrates with an average diameter of ~10 microns. Between 520 and 560°C, Eq. (8.9) agrees to within ±5% of the plot in Figure 8.7. 168 Table 8.7 Measured Values of Time of Reaction in "Chemical and "Diffusional" Regimes "Chemical" Regime "Diffusional" Regime Temp. °C po 2 atm. XMoS2 mol . f ract . ox. mi n XMoS2 mol . f rac. ox. T mi n 458 0.21 0.400 30.0 1.0 „ * * * (180) 480 0.21 0.370 15.0 1 .0 * * 150 526 0.21 0.475 7.5 1.0 * 120 530 0.21 0.475 6.5 1.0 * 100 554 0.36 0.750 2.8 1.0 * 95 557 0.21 0.525 5.0 1.0 * 85 557 0.54 0.775 2.5 1.0 * 90 557 0.61 0.800 2.35 1.0 * 85 560 0.21 0.600 5.0 1.0 * 80 595 0.21 0.650 5.0 1.0 * 70 * Extrapolated ** Approximate *** Estimated The measured values of t' and x can then be used for scale-up purposes, as i t will be shown in Chapter 10. 169 2 0 0 i . | I50| c o e 00 c •£ 1001 0) e o o 5 0 B.C. Molv Concentrate 12 4 5 0 °C ., air O . 4 8 0 " V 5 26 " © 5 30 " 55 7 " • 5 6 0 " A 5 9 5 " 0' Brenda Concentrate A 554 °C 0 557 " ® 557 " _ J 36 % 0 2 54 % " 61 % " 4 5 0 5 0 0 550 T , Temperature (°C) 6 0 0 Figure 8.8. Influence of temperature on the total time required for transformation of MoS2 to Mo03 in the f luidized bed reactor. Chapter 9 CONTINUOUS ROASTING OF MOLYBDENITE CONCENTRATES IN THE FLUIDIZED BED REACTOR The main objective of the present research program was to study the technological f e a s i b i l i t y of operating a f l u i d i z e d bed reactor for roasting molybdenite concentrates. In this work, four d i f f e r e n t concentrates have been tested, the composition and ch a r a c t e r i s t i c s of which are given in Chapter 8, Table 8. The results of these tests, obtained mainly in the 12.5 cm diameter reactor, are presented and discussed in this chapter. Prior to each experiment the reactor was charged with a fixed amount of sand (3000 g) of -40/+140 mesh. As was shown in Chapter 6, the use of this size d i s t r i b u t i o n minimized the extent of p a r t i c l e s t r a t i f i c a t i o n during f l u i d -zation. Next a given quantity of calcines which had been roasted previously was charged. In addition some MoS2 was added to boost the temperature during the heating period. The precise amount of calcines was altered as needs required to vary the residence time of solids inside the reactor. 170 171 The experiments which were then conducted, were run continuously for a period of 8 to 10 hours. 9 .1 Operating Conditions Studied The variables of the process that were studied in the 12.5 cm reactor and the ranges investigated are given in Table 9.1. In most of the experiments, concentrates from B.C Table 9.1 Variables Investigated for the Roasting of Molybdenite in the 12.5 cm Reactor Variable Range Temperature, T Overage residence time of so l ids , t Superficial gas veloci ty , u 0 Oxygen partial pressure, P n (J 2 Feed rate of MoS2, F 0 Particle size of MoS2, d s Calcium in Feed 500 to 595°C 3 to 38 hrs 15 to 34 cm/sec 0.04 to 0.8 atm 133 to 2650 Kg/m2 day -100 to -10 microns 0.4 to 0.04% Moly were used, but the optimum conditions were determined using the Brenda concentrates. The quantitative influence 172 of each of the operating variables was measured in terms of its effect on the roasting ef f ic iency, i . e . , the residual * sulphur in the calcines (see Appendix 15). The complete results of the 89 experiments performed employing the 12.5 cm diameter f luidized bed reactor are given i n Appendi x 13. 9 . T. 1 Superficial Gas Velocity The superficial gas velocity appears to have a minor influence on the residual sulphur content of the calcines, as can be seen in Figure 9.1. For values of u 0 ranging from 20 to 35 cm/sec, no appreciable difference in the residual sulphur level was found: for example, for t= 27 hrs, at u 0 - 20 cm/sec the calcines contained about 0.43% sulphur, whereas at u 0 = 30 cm/sec the sulphur was about 0.40%. This finding is in agreement with the tracer studies (Chapter 7) where i t was shown that neither the gas transfer c o e f f i c i -ent K (h e )b n o r t n e R D N varies appreciably over this range qf superficial gas velocity. The range of superficial gas velocity investigated in this work was dictated by operational l imitat ions. Below -u 0 " 18 cm/sec the bed f luidizat ion was not suff icient to permit smooth operation of the rotary scraper inside the bed, while at u 0 > 35 cm/sec the elutr iat ion rate of calcines The reported values of the sulphur in calcines represents the average of 3 to 6 samples taken during the roasting process. 173 1.0 0.8 B.C. Moly Concentrates T = 5 5 0 ± 2 ° C a i r , J atm. to CD o 0.6 o t = 2 0 hr ± 1 0 % I 0.4 t = 27 hr ±10 % c CD O v_ CD Q_ 0.2 0 2 0 2 5 3 0 3 5 u 0 , Superficial Gas Velocity (cm/sec) Figure 9.1. Influence of the superficial gas velocity on the velocity on the sulphur content in the calcines. 174 was higher than the discharge capacity of the rotary valves of the cyclones. This resulted in an accumulation of sol id inside the cyclone system. The lower l imit of u 0 > 18 cm/sec cannot be readily overcome, but the upper l imit can be in -creased by employing faster or larger rotary valves at the cyclone discharge. 9.1.2 Roasting Temperature Temperature has a reasonably strong influence on the residual sulphur in the calcines for a retention time of calcines in the bed of 11 hr, as can be seen in Figure 9.2. However, the effect becomes less pronounced at longer times; for retention times greater than 15 hr, for example, the residual sulphur reaches a steady value of 0.4% for the B.C. Moly concentrates. This value could not be lowered even by increasing the roasting temperature up to 580°C, the highest value tested. This apparent lower l imit is due, as will be shown later , to the high content of calcium in the concentrates. For an average residence time below ~23 hr, i t is not possible to reach the lower l imit of sulphur in the calcines even by increasing the temperature to 580°C. For longer retention times, t > 27 hr, the lower l imit of sulphur can be achieved even at 520-525°C. This suggests that between 520 to 580PC, the average residence time of reaction CO cu c _o o o C O c CU o cu Q_ .0 0.5 B.C. Moly Concentrate u0=28.5 cm /sec T ±. 10 % (hr) A 3.5 hr 9 8 'hr O I I ® 22 A 27 " t = 27 hr 0 500 520 540 560 Temperature (°C) 580 600 Figure 9 . 2 . Sulphur content in the calcines as a function of the temperature of roasting. 176 has a larger effect on the f inal sulphur content of calcines than does the temperature of roasting. Severe problems arose from the sintering of calcines inside the reactor, particularly at high temperatures. Sinter-ing was f i r s t noticed at about 560°C, when the bed thermo-couple readings began to fluctuate from their steady state values. The process was found to be completely inoperative at temperature in excess of 580°C, where large pieces of sintered material formed and blocked the rotating scraper. For this reason, only three tests have been conducted at temperatures in excess of 580°C (Figure 9.2). These experiments clearly show that for practical operation of the f luidized bed reactor, the roasting tempera-ture must not exceed 560°C. The optimum temperature appears to be 550°C for the reactor to operate continuously with no severe problems. 9.1.3 Average Residence Time of Reaction The average residence (or retention) time of solids inside the reactor was calculated as the quantity of calcines inside the bed at any time, W (constant at steady state), divided by the feed rate of fresh MoS2, F 0 : 177 10, . CO c o o _3 to c o v . CL) CL \520 °C 5 4 0 °C B.C. Moly Concentrate reactor © 520 °C , air A 540 °C , air reactor 0.4 o 524 °C , air A 550 °C , air 0.2 0.1 0 10 15 2 0 25 T , Average time of residence (hr) 30 Figure 9.3. Influence of the average residence time of solids in the 7.5 and. 12.5 cm diameter f luidized bed reactors. 178 (hr) (9.1) The average residence time has a drastic effect on the sulphur content in the product calcines, as can be seen in Figure 9.3. Here, results of studies performed in the 7.5 cm and 12.5 cm diameter reactors are compared: at 540r550 6C p for t = 5 hr, the calcines contain about 2.5%S, while for t = 25 hr, the sulphur level drops to 0.4%. The decrease in sulphur content is very rapid from the i n i t i a l value of 36.43% down to about 1.5%, below which a much slower decrease takes place until the sulphur content reaches 0.4%. These results are in good agreement,with the batch experiments (Chapter 8) and seem to confirm the existence of two different kinetic regimes of transformation: a fast i n i t i a l regime followed by a slow f inal one. In Figure 9.4, the results of experiments conducted at three different temperatures are plotted. It can be seen again that for T = 55Q-575°C, for t > 20 hr, no difference in the residual s,ulphur content in the calcines was found; and in the three cases, i t appears that for t" > 20 hr, a lower l imit value of ~0.4%S is approached asymtotically. 9.1.4 Oxygen Partial Pressure 1 ' " " " " ' 1 ! 1 ~ The influence of the oxygen partial pressure on the residual sulphur in the calcines at high conversion levels 179 B.C. Moly Concentrates t , Average Time of Residence (hr) Figure 9.4. Sulphur content of the calcines as a function of the residence time of solids for different temperatures of roasting. 180 (over 97% of the sulphur of the MoS2 oxidized) was found to be negligible as can be seen in Figure 9.5. The sulphur con-tent remains constant for oxygen partial pressures above about 0.1 atm. For lower values of P n , the percent sulphur U 2 in the calcines increases sharply. For example, at P n = u 2 0.05 atm, the sulphur level in the calcines increases to 0.8%, whereas for P Q z = 0.04 atm the sulphur in the calcines has reached 1.6%. In this case, the oxygen efficiency in the reactor approaches 100% since the calculated values of the gas transfer coefficient K (b e )b ^ s e e C n a P t e r is large enough to assure a high degree of oxygen transfer between the gas bubbles and the emulsion phase. It may happen, there-fore, that the bed becomes depleted, i . e . , starved, of oxygen in the upper region at P n < 0.1 atm. u 2 The negligible effect of the oxygen partial pressure above ~0.1 atm for high levels of conversion is in agreement with the previous findings from the batch tests (Chapter 8) where the time required to achieve a fu l l con-version was found to be v i r tual ly independent of the oxygen partial pressure. 9.1.5 Calcium Content of the Molybdenite Concentrates As shown in Figure 9.4, i t appears that a lower limit of ~0.4% residual sulphur was reached for the calcines roasted from B.C. Moly concentrates. Further sulphur 2.5, 181 CO CD c \> D O 2.0 .5 B.C. Moly Concentrate T =25 hr ±10% T =550 ±2 °C u 0 = 28.5 c m / s e c sz CL ZD CO c CD a CD CL .0 0.5 0 0 20 40 60 Percent Oxygen in Reacting Gas 80 Figure 9.5 Influence of the oxygen partial pressure on the sulphur content of calcines. 182 elimination could not be effected either by increasing the residence time (Figures 9.3 and 9.4) or by increasing the temperature of roasting (Figure 9.2) or by enriching the roasting air with oxygen (Figure 9.5). A similar effect was found for the roasting of Kennecott concentrates, where a l imit of about 0.45% residual sulphur was reached. Again, i t appeared that no substantial decrease in the residual sulphur in the calcines could be achieved, as shown in Figures 9.5 and 9.6. The l imit of the sulphur elimination appears to be linked to the high calcium content in both the molybdenite concentrates from B.C. Moly (0.3%) and Kennecott (0.4%). The calcium is generally found in molybdenite concentrates as CaC03 [120], which at the roasting temperature decomposes to CaO. The CaO formed can subsequently react with the S02 generated (or M o S 2 ) during the M o S 2 roasting to form Ca SOn, following the reaction: CaO + S0 2 + i'0 2 * CaSO* (9.2) This reaction can occur at 550°C due to i ts favourable free energy of reaction ( A G ^ Q O Q =-67.25 Kcal). The CaSOit formed is stable at this temperature and decomposes appreciably only above 1200°C [121]. Other reactions might also occur during roasting, such as: 183 Figure 9 . 6 . Influence of the temperature on the sulphur level of ca lc ines for the Kennecott concentrates. 184 Figure 9 . 7 . Sulphur content of ca lc ines as a funct ion of the average residence time of react ion for the Kennecott concentrates. 185 CaC03 + S0 2+i0 2 * CaSCU + C0 2 (9. which similarly result in the formation of the stable CaS0\ compound. In order to elucidate the influence of the calcium content on the residual sulphur levels of the discharged calcines at high levels of transformation, two molybdenite concentrates low in calcium were roasted. 9.1.6 Roasting of Low Calcium Molybdenite Concentrates Two molybdenite concentrates with low average calcium contents were tested: Brenda, with 0.04% Ca and Endako, with 0.03% Ca. The Brenda concentrates were standard concentrates leached industr ial ly with a chloride solution of C u C l 2 , FeCl 3 and CaCl 2 at 120°C to remove most of the copper, lead and calcium. The Endako concentrates were industr ial ly leached with a dilute solution of HC1 (pH 3-4) for the sole purpose of removing calcium. Results of the tests performed using both concentrates are given in Appendix 13. The results of roasting these concentrates showed a dramatic decrease in the level of sulphur in the calcines. The residual sulphur in the calcines is plotted in Figure 9.7 as a function of the average residence time. It can be seen 186 0.301 1 r 1 1 1 1 r -t , Average Residence Time (hr) Figure 9.8. Sulphur level of ca lc ines as a funct ion of the residence time for low calcium concentrates. 187 that at 550°C for t > 20 hr, the sulphur levels have de-creased below 0.2% for both Endako and Brenda, and for t > 30 hr a sulphur content of less than 0.1% was achieved in roasting the Brenda concentrates. These values are signif icant ly lower than the maximum industrial sulphur specification of 0.25% S and compare favourably with the average sulphur content of calcines roasted in multiple hearth furnaces. In the latter process the average sulphur content in the calcines is about 0.1 to 0.15% [122]. An additional leaching of the Brenda concentrates with 0.5 N HC1 at 100°C was carried out. Then, an additional sl ight decrease in the sulphur content of the calcines after roasting was obtained, as shown in Figure 9.7. This decrease may represent removal of residual calcium which was not completely leached in the standard chloride process. As was found for the B.C. Moly concentrates, an increase of the oxygen partial pressure seems to have no noticeable influence on the sulphur content of the calcines at such high conversion levels (99.7% of the sulphur oxidized). Similar ly, at these conversion levels , temperature over the range of 525 to 550°C seems to exert a minor effect on the residual sulphur content as can be observed in Figures 9.7 and 9.8. Only a sl ight decrease in the sulphur content of the calcines was found by increasing the temperature from 524 to 550°C. This finding has important practical consequences; 188 0.25 CO CD I 0.201 o O Brenda Concentrate— — A O A ® > air 6 0 % 0 2 J chloride leached 0 air | double 6 0 % 0 2 J leached Endako Concentrate • air , H Cl leached i r 7 = 16 hr xz w 0.151 c CD O k_ <D Q_ 0.10 -= 2 0 hr 0.05 1 5 2 0 5 3 0 5 4 0 5 5 0 Temperature (°C) 5 6 0 F i g u r e 9 . 9 . S u l p h u r c o n t e n t i n c a l c i n e s a s a f u n c t i o n o f t h e r o a s t i n g t e m p e r a t u r e f o r l o w c a l c i u m c o n c e n t r a t e s . 189 a working temperature of ~530-540°C is more desirable since for the same percent sulphur elimination, i t is far from 560°C, the temperature at which sintering commences. These experiments show the feas ib i l i t y of the f luidized bed process to produce calcines of low sulphur contents by using standard low-calcium molybdenite concen-trates. It represents a major improvement over other f luidized bed processes attempted previously. For comparison purposes, an a r t i f i c i a l average rate of transformation -(dMoS 2/dt) was calculated as rMoS 2 dMoS dt = mole fraction MoS2 oxidized average residence time Calculated values of r"MoS2 are plotted in Figure 9.10. As expected, a sharp decrease in the apparent rate of reaction occurs with increasing fraction of sulphur oxidized for each concentrate. For example for the roasting of B.C. Moly concentrates, to lower the %S in the calcines from 0.8 to 0.4%,the ayerage residence time of reaction must be increased from 10 to about 25 hr. This in turn decreases the apparent rate of reaction from 8.1 x 10 - I f to 6.8 x 10"^ (mole fraction of MoS2 oxidized/min). When the minimum level of sulphur is reached, any further increase in t" wil l only decrease the value of r"|yj0<5 » that i s , decrease the output of the reactor without any further decrease in the sulphur content of the, calcines. T — i — i — i — i — i — i — i — i — i — i — i — i — i — i — T — r 0 4 '— 1 — ' — i — i — i — i — i — i i i i • » « • • Ii I 0.970 ^ 0.980 0.990 1.000 X s , Fraction of sulphur oxidized Figure 9 .10. Apparent rate of reaction of molybdenite in the f luidized bed reactor as a function of the fraction of sulphur oxidized. 191 For low calcium concentrates, at high levels of conversion (< 0.1% S in calcines) the output of the f luidized bed reactor is given by: c = 9.6 rM c MoS2 (kg/hr) (9.4) where r [ l ] o S z is the experimentally determined value for the concentrates tested, at the desired level of conversion. 9.2 Optimum Operating Conditions Based on the previous f indings, the optimum operat-ing conditions of the f luidized bed process for molybdenite roasting have been determined. The range of operations is given in Table 9.2. Table 9.2 Optimum Range of Operating Conditions for Fluidizing Bed Roasting of MoS2 Variables Minimum Value Maximum Value Temperature Average residence time Superficial gas velocity MoS2 feed rate Oxygen partial pressure Particle size MoS2 Calcium in MoS2 Sulphur in calcines 520°C 20 hr 18 cm/sec 200 Kg/day/m2 air - 3y 550°C 30 hr 35 cm/sec 300 Kg/day/m2 air 50y 0.05% 0.2% 192 9.3 Material Balance on the Process At proper working conditions, the eff iciency of the cyclone system was between 98.5 and 99% (see Appendix 13), while that of the scrubber system was about 98%. Thus an overall collection efficiency in excess of 99.9% was achieved. An average material balance on the process is given in Table 9.3. Table 9.3 Material Balance for Fluidized Bed Roasting of MoS2 Concentrates in 12.5 cm Reactor Base 1 day Feed rate: 3600 gr/day MoS2 concentrates 90% MoS2 Calcines with 0.12% S T = 550°C u 0 = 24.4 cm/sec t = 26.7 hrs In (gr) Out (gr) Feed Air Calcines Solid in Scrubber Solid in Solution Gases Mo S 0 2 N2 1942.6 1314.0 26205.6 66530.4 1727.8 3.5 863.9 195.1 0.4 97.4 19.2 1300.3 23944.0 66530.4 193 The solids collected in the scrubber were calcu-lated for a 99% efficiency of collection in the cyclones, at an elutriation rate of 85 gr/min (120 Kg/day). This represents a ratio of elutriat ion rate to feed rate of: * : F 0 = 34:1 A small fraction of the molybdenum goes into solution. This is made up of molybdenum dissolved from the collected calcines in the S0 2-saturated acid solution of the water scrubber, as well as the condensed lower oxides of molybdenum formed during roasting, known as "molybdenum blue." This compound apparently contains molybdenum in two valence states [123], IV and V, and has a probable formula of Mo 80 23 • n H20 [124]. It is very soluble in water and has a dew point of about 120-150°C. To prevent the con-densation of this product, the gas lines and cyclones were kept well above 250°C. The total molybdenum dissolved represents about 1% of the charge, and can be recovered from the solution as ammonium molybdate by precipitation with NrUOH as follows MoO,, + 2NHt -> (NrU)2 MoO^  + 2 H 2 0 (9.5) The solid not collected in the cyclone system amount to 1-2% of the total calcines recirculated, but depend on the particle size of the molybdenite concentrates and the 194 efficiency of the cyclone system. Occasionally, the secondary cyclone became fouled, probably by a closed discharge apex, and the collection efficiency dropped to 80-85%. Normally, however, over 98.5% efficiency was achieved. The calcines collected in the water scrubber contained usually a s l ight ly higher content of sulphur than the discharged calcines. This suggests that some MoS2 was backfed due to gas leakage through the rotary valves (see Appendix 13). The overall recovery of molybdenum in the process is over 99.9%, including the molybdenum in solution. By comparison, the overall recovery of molybdenum in the multiple hearth process is about 98.5% [125]. The S02 in the off gases ranges from 0.6 to 3.4%, depending upon the feed rate and the total gas flow rate.. For high levels of conversion at 550°C and u 0 = 25 cm/sec, the off gases contain 0.8 to 0.9% S0 2 . This value is val id only for this 12.5 cm reactor; substantial increases can be achieved in a larger reactor with large output, as wil l be shown in Chapter 11. In Figure 9.11 some calculated values of the S02 content are given, from several experiments using different feed rates. The results are given also in Appendix 13. The recorded signal of S0 2 in the infrared analyser at steady state is given in Appendix 16. 195 F0 , Mo S 2 feed rate (kg/m x day) Figure 9.11. S0 2 in the off gases as a function of the MoS2 feed rate. 196 9 . 4 Slurry Feed Injection The possib i l i ty of feeding slurry instead of solid molybdenite concentrate is attractive, since this would permit the direct feeding of the discharge from the f i l t e r s in the concentration plant, thereby eliminating a dryer. Furthermore, slurry feeding would eliminate sqlid handling which can present problems due to the tendency of molybdenite to compact. To test the appl icabi l i ty of slurry feeding to the f luidized bed process, one experiment was performed in which a 50 wt-% solid suspension of molybdenite concentrate in scrubber l iquid was injected directly into the f luidized bed reactor. The arrangement used is shown in Chapter 4, Figure 4 . 4 . Analysis showed (Exp. No. 120, Appendix 13) thait t,he sulphur content of the calcines from this test was similar to that obtained with solid feeding. Apparently no problems arise from slurry feeding although certainly several more experiments are requirecj to test i t f u l l y . Careful control must be exerted over the feed rate with this method to avoid a violent vaporization of large amounts of water in a short time in the reactor. Belton and Jordan [126] have reported that the presence of water vapour enhances the vola t i l i za t ion of Mo03 as M 0 O 3 • H20 when molybdenum metal is oxidized 197 between 1 200 and 1 500°C. However, such effects have not been shown to exist at roasting temperatures. An alternative form of feeding slurry to the reactor could be used by preheating the operating air for the slurry injection venturi. This would produce a flash vaporization of the slurry drops inside the expansion section of the venturi, and avoid feeding the l iquid directly into the reactor. Chapter 10 INDUSTRIAL APPLICATION OF THE FLUIDIZED BED PROCESS FOR MOLYBDENITE ROASTING From the successful results obtained in this study, i t is conceivable that the f luidized bed process could be applied on a larger industrial scale to replace the multiple hearth furnace for molybdenite roasting. This work has shown, for example, that the f luidized bed process is capable of producing calcines which have suff ic ient ly low sulphur contents 0.085 - 0.15% - to be acceptable for use either directly as an alloying agent or in ferroalloy production. These sulphur levels are comparable to those presently realized employing multiple hearth furnaces. In this chapter i t further will be shown that there are positive advantages in using the f luidized bed roaster compared to the multiple hearth furnace. These advantages include substantially lower capital costs, possible reduction in operating costs and simpler solids and l iquid handling equipment. In order to make this comparison, the 198 199 size of f luidized bed reactor, the collecting equipment and circulation system required for industrial purposes, have been estimated based on experience with the small pi lot plant reactor. Details of these calculations are the main subject of this chapter. 10.1 Scale-up of Fluidized Bed Reactor The ab i l i ty of the industrial scale process to achieve low sulphur calcines will depend upon the operating variables and the internal dimensions of the reactor. For design purposes, i t wil l be assumed that the concentrates to be roasted in the f luidized bed will meet the following conditions: 1. P a r t i c l e s i z e o f c o n c e n t r a t e s : -325 mesh, d < 15 m i c r o n s . ' s 2. C a l c i u m c o n t e n t o f c o n c e n t r a t e s : 0.04$ maximum. Both conditions are representative of standard molybdenite concentrates produced by f lotation and subsequent leaching operations, e.g. those tested from Brenda Mines and Endako Mines. The following additional conditions will be assumed in the reactor design: 1. S u p e r f i c i a l gas v e l o c i t y : 25 cm/sec a t t h e r o a s t i n g t e m p e r a t u r e . 2. B a c k m i x e d r e a c t o r ( a s d e m o n s t r a t e d i n C h a p t e r 7 u s i n g t r a c e r s ) . 200 10.1.1 Sulphur Content of Discharged Calcines For a backmixed reactor 9 the exit age d i s t r i b u t i o n function E(t) of the p a r t i c l e s at the discharge, which i s equal to the internal age d i s t r i b u t i o n function C(t) in this case, is given by the expression: E(t) = C(t) = I e~t/T (10.1) t The f r a c t i o n of MoS2 transformed (as Mo03) at the discharge of the reactor i s given by the relationship XMo03 t = T F'(t) E(t) dt (10.2) t = 0 where F(t) is the transformation function of an individual p a r t i c l e of MoS2 which depends on the nature of the rate l i m i t i n g st;ep (s). In this c^se, the rate c o n t r o l l i n g steps are not f u l l y understood but are thought to involve, f i r s t , chemical control for a brief period, during which the rate of oxidation proceeds rapidly. This i s followed by a longer period in which the oxidation rate i s considerably slower, possibly as a resu l t of s o l i d state d i f f u s i o n . Discussion of the kinetics has been presented in Chapter 8. The lack of a precise understanding of the kinetics in the later stage of 201 transformation makes i t v i r tual ly impossible to calculate F(t) in Eq. (10.2) from theoretical considerations alone. Therefore in order to obtain an expression for scale-up purposes, a f i r s t attempt was made to empirically f i t the experimental data of %S vs retention time, obtained in this work, to a polynomial expression of the form %S = a - bt + c t 2 - dt 3 + • • • "' (10 By regression analysis, the following relationship was found: %S = 0.31435 - p.01052t + 0.00011237t2 (10 The calculated values of %S in the calcines using Eq. (10.4) are plotted in Figure 10.1, together with the experi-mental data from the f luidized bed reactor for Brenda con-centrates. As can be seen, the agreement is reasonably good (to within ±10%) pver the range 16 < t < 36. It is important too to note that Eq. (10.4) is valid only for a temperature of 550°C, at which the data was or iginal ly obtained. For other temperatures, different relationships have to be developed, using the appropriate experimental data.. A second attempt to f i t the experimental data to a mathematical expression was made by assuming that the slow 202 step in the latter stages of oxidation was diffusiopal in nature. It was further assumed that the mathematics which ho]d for diffusion of gas through porous solid also applied to the oxidation of MoS2 at these high conversion levels. Then for slab-shaped particles of MoS2, x M 0 S 2 v a p l ' e s with time in the following way X M o S 2 = 1 XMo03 ~ 1 (10.5) The fraction of solids not transformed.,^ <. , discharged r' • 2 in the calcines from the reactor is given, therefore by substituting Eq. (10.5) which is the expression for F(t ) , into Eq. (10.2) to y ie ld* t = x MoS • - 1 - V 1 1 s i I e-tn dt t (10.6) t=0 Integrating this expression and neglecting the terms smaller than t " 3 / 2 , Eq. (10.6) becomes: 7t The lower l imit of the integral should, s t r ic t ly speaking, be a non-zero value of time corresponding to the onset of the final stage of oxidation. However, since the f inal stage occupies over 90% of T , the integral as written is a reasonable approximation. 203 MoS = 1 + e T / t I -.1 (10.7) This relationship was f i t ted to the experimental data of Figure 9.7. for the Brenda concentrates by successive i terat ions. The resulting equation, in terms of the residual sulphur in the calcines, is as follows: = 1.35 1 + - T / t 3 - 1 (10.8) The total time of reaction, T ,can be expressed by the empirical correlation developed in Chapter 8 as a function of the temperature (Eq. 8.9). In this form, the sulphur content in the calcines is given by the relationship: %S = 1.35 1 + 1 JL(yO . 7 6 5 T + 520 ) (T°- 7 6 5 - T + 520 + 1 (10.9) 204 Computed values of %S in the calcines using Eq. (10.9) are "plotted in Figure 10.1 for temperatures of 524 and 550°C. The agreement between calculated and experimental values for the Brenda concentrates can be seen to be reasonably good (±8%) for retention times in the range, 16 < t" < 36 hr. The discrepancy observed at lower retention times is not serious since a successful industrial scale reactor would have to operate with t > 20 hr to produce low sulphur calcines. Based pn the comparison outlined above, i t is clear that either Eqs. (10.4) or (10.9) could be used to scale-up the f luidized bed reactor. However, since Eq. (10.9) gives a marginally better f i t to the experimental data, i t was given preference. The analysis can be carried one step further i f the latter stage of oxidation is diffusion controlled as suspected, and i f Eq. (10.5) is a reasonable mathematical description of this process. Under these condition the dependence of %S on average MoS2 particle s ize , d~s, can be included in Eq. (10.9) as follows 205 t , Average time of reaction (hr). Figure 10.1. Experimental and calculated residual sulphur in the calcines as a function of the average retention time and temperature. 206 ; 2 I J O . 7 6 5 . j + . 5 2 g J ( 1 0 . 1 0 ) However, insuff ic ient data was obtained in the present investigation to check the re l i ab i l i t y of this equation. 10.1.2 Reactor Design and Simulation of Operating  Performance Based upon the sulphur levels desired in the calcines, the dimensions of a reactor can be calculated. Since the optimum ranges of roasting temperature and superficial gas velocity are very narrow, the following values have been used in the calculations: T = 550°C u o = 25 cm/sec The average particle size of concentrates is taken to be -325 mesh, d~s = 10 microns. The concentrates are assumed to be 100% MoS2, with 0.04% Ca. The most important scale factor used in the calcula-tions of the reactor size is L / R . , which has been taken to m d be 1 i n i t i a l l y ; this seemingly arbitrary choice has been = 1.35 1 + exp -j 10"?d2 0.765 T + 520 207 dictated by the need to maintain the stat ic bed height, L m , such that the pressure drop through the bed will not be excessive (see Chapter 5). Consequently the reactor diameter required tp achieve a given value of if for a given feed rate, F 0 , is given by the expression where F 0 t is the total amount of calcines in the f luidized bed and p s is the apparent density of the solids (Eq. (5.4)). The height of the reactor, R-j , which includes freeboard above the f luidized bed, should then be roughly 2.5 times the reactor diameter based on experience with the small pi lot reactor. The % SQ2 in t h e off gases can be calculated in terms of the feed rate of MoS2 and the total rate of gas flow through the- bed,. G f , as follows 0.333 (m) (10.11) R1 ;=• 2.5 Rd (m) (10.12) %S02 = 1 2.775 F.0 (10.13) G f + 0.12775 F 0 208 Calculated curves of per cent residual sulphur in the calcines vs reactor diameter (height) are presented in Figure 10.3 and Appendix 12 for three different feed rates, and values of L m /R d = 1 and 0.5. It can be seen that the reactor diameter increases rapidly as sulphur levels decrease. For example, to obtain 0.1% sulphur in the calcines at a feed rate of 10 TPD, a reactor with a diameter of 2.3 m and a height of 5.8 m is required; whereas for 0.08% in the calcines a 2.5 m diameter reactor with a height of 6.3 m is needed. For a shallower bed, that is L m / = 0-5, the reactor diameter required for 0.1% sulphur in the calcines is 2.8 which increases to 3.1 m for 0.08% sulphur. In any case, the size of the reactor is very modest compared with i ts industrial equivalent, the multiple hearth furnace. A 10 TPD f luidized bed reactor with 1.4 m deep bed that produces calcines with 0.1% S, wi l l be 2.8 m diameter by 6.8 m high while for the same throughput a 5.2 in diameter by 10.6 m high ten hearth roaster is required. The calculated values of %S02 in the roasting gases are plotted in Figure 10.3 as a function of the reactor capacity. It can be observed that for a capacity over ~1 0 TPD of molybdenite concentrates, the %S02 in the off gases in -creases almost l inearly with increasing reactor capacity; at lower capacities the S0 2 content in the gases decreases sharply. For L m /R d = 1, the calculated values are about 40% higher than for L m /R r i = 0.5, as would be expected. The RI , R e a c t o r length (m) 0.051 1 —i L__ i I 10 1.5 2.0 2.5 3.0 3.5 Rd , Reactor diameter (m) o to Figure 10.2. Calculated reactor dimensions as a function of the sulphur level in calcines, for different output capacities. 210 calculated va]ues of S0 2 in the off gases of an industrial f luidized bed of over ~5 TPD is greater than 2.5% under any conditions, and i t is about 4% S0 2 for a 20 TPD reactor with a shallow bed. These compare favourably with the level of S02 found during standard operations of multiple hearth furnaces of "1.5% S0 2 , or 3-3.5% S0 2 for spray water operation [127]. In the 12.5 cm fluidized bed reactor the L.. /R. ratio was higher, approximately 3, with the result that the off gases were proportionally higher in S0 2 than for L /R . = 0.5 to 1 . m d The predicted concentration of MoS2 at the injec-tion point for a 10 TPD is about 70%, which is rather high. To avoid resultant sintering problems i t is suggested that four separate dispersion nozzles, positioned around the axis of the reactor, be used. 10.1.3 Two-Stage Operation Since 65 to 75% of the sulphur in the molybdenite concentrates is removed in the f i r s t 8 to 10% of the total time of transformation, a two-stage operation could be used. The f i r s t stage would be a small reactor where the rapid in i t i a l stage of oxidation would be completed. The off gases from this reactor would contain up to 10% S0 2 , depending 8 T = 550 °C iO JUL m 0.1 % S in calcines </> 6 CD CO O CP 10 15 C , Reactor output (ton Mo S 2 / day ) 20 25 Figure 1 0 . 3 . Calculated level of S 0 2 in the off gases as a function of the reactor capacity. 212 upon the capacity of the reactor. The calcines would t h e n be discharged to a second, larger reactor where the transfor-mation would be completed. It is important to note, however, that on ly a small (~8-T0%) decrease in volume of the second reactor would be achieved by adopting such a two-stage operation. Thus the apparent advantage of obtaining a roaster gas with a high S0 2 concentration by this method is question-able from the capital investment point of view. 10.2 F l u i d i z e d Bed Plant for Molybdenite Roasting A proposed layout of a plant for roasting molybdenite concentrates a t a rate of 10 TPD is depicted schematically in Figure 10.4. A slurry feed consisting of 50% MoS2 solids has been assumed; the same basic layout would apply however to the case of solids feeding. The calculated dimensions of the reactor and predicted process performance are given in Table 10.1. The throughput of the 10 TPD f luidized bed reactor per unit area of reactor cross-section is compared to the equivalent multiple hearth in Table 10.2. It can be seen that the net output of the f luidized bed is from 3.3 to 5 times larger than the output of the multiple hearth when the latter is calculated on the basis of f loor area, or 33 to 50 times larger based on total hearth area. Fluidized Bed Reactor -Fuel-Off Gases Gases Elutriated Calcines Fluidized Bed Combustion: -Chamber 6 Air Blcwer •ligh Efficiency Cyclone System V distribution Valve Molybdenite Slurry Feed do Calcines Recycling System /Pump Feeding System Preheated Air *< Make-up Water Water Scrubber System Pump Figure 10.4. Layout of a 10 TPD f luidized bed plant for roasting molybdenite concentrates. 214 Table 10.1 Dimensions and Operating Performance of a 10 TPD Plant for the Roasti ng of MoS2 Capacity " ~ 1 1 ' — • : 10 TPD MoS2 concentrate Concentrates particle size : -325 mesh, d~s = 10 microns Calcium in concentrates : < 0.04% Calcines : 0.10% S Reactor dimensions : 2.8 m diam. x 6.3 m height Recirculation ratio : 3.41:1 Gas flow rate 19 m3 air/min Superficial gas velocity : 25 cm/sec Temperature of the bed : 550°C Off gases : 4.3% S0 2 Cyclones efficiency : 95% Scrubber efficiency : 95% Overall collecting efficiency : 99.75% Solids from scrubber to : 0.3 5 Kg/min calcines (0.70 Kg slurry 50% sol ids / min) reactor Table 10.2 Calculated Output of Reactors r~" 1 . Kg/m2 day per individual hearth Kg/m2 day total hearth area Multiple hearth 480 48 Fluidized Bed 2400 -7 1260 10.2.1 Capital Costs The capital cost of instal l ing a 10 TPD f luidized bed roaster and a multiple hearth furnace with a similar capacity are detailed in Tables 10.3 and 10.4 respectively. The equipment l isted for the multiple hearth process are standard for this type of plant [128], whereas the equipment required for the f luidized bed process has been estimated based on information obtained in the pi lot f luidized bed reactor and t h e previous calculations. The equipment costs f o r 1974 have been taken to be 1.3 times the 1971 prices N 2 9 J . Type 316 stainless steel was assumed to cost three times the p r i c e o f plain carbon steel . A comparison of the capital costs in Tables 10.3 and 1C.4 indicates clearly that the f luidized bed process can be expected to enjoy a considerable advantage over the multiple hearth furnace. The capital cost of the f luidized bed p r o c e s s is 60% less than that of the multiple hearth p r o c e s s . This reduction in cost is due to two factors: less diversi f ied equipment, and a much smaller reactor. 1C . 2 .2 Operating Costs A comparison of operating costs for the f luidized bed and multiple hearth processes is given in Table 10.5. In arriving a t many of these f igures, values from a rel iable 216 Table 10.3 Capital Cost for a Molybedenite Roasting Plant using a Multiple Hearth Furnace Number of Units: Equipment Description Total Cost 1 F i l t e r , 6' x 4' discs $ 6,500 1 Multiple hearth dryer, 10' <p x 4 hearth 200,000 1 Multiple hearth furnace, 16' , 30' height by 10 hearths, complete, in -stalled and with controls [128] 1 ,050,000 1 Rotoclone collector (SS 316), 45,000 CFM 30,000 2 Bins, 90 %, MoS2 46,000 1 Bin, 10 %, MoS2 6,500 8 Screw conveyors, 3 to 5 inches. <f> 60,000 5 Bucket elevators, 6 to 35 feet 40,000 3 Bins, 25 T, Mo03 35,000 2 Multiclone, bank of 4 each (SS 316) 40,000 1 Electrostatic precipitator 16,000 CFM (SS 316) 200,000 1 Stack, 4'6" x 148' (SS 316) 35,000 1 Impact Mi 1.1 5,000 1 Packer Unit 10,000 CAPITAL COST -1 ,785,000 Install ation 370,000 TOTAL CAPITAL COST $2,155,000 217 Table 10.4 Capital Cost for a Molybdenite Roasting Plant using a Fluidized Bed Process Number of Units: Equipment Description Total Cost 1 F i l t e r , 6' x 4' discs $ 6,500 1 Blower, lobular, 500 CFM 5,000 1 Cylindrical reactor tank (SS 316, 1/2" thick) , 2.8 mt x 6,3 mt 45,000 1 Combustion system and chamber 20,000 2 Multiclones, bank of 4 (SS 316) 40,000 8 Discharge valves for cyclones (SS 316) 8,000 1 Discharge system (SS 316) 10,000 1 Slurry feeding system 5,000 1 Agitation tank for MoS2 slurry 25,000 1 Bin, IQ %, MoS2 6,500 3 Bins, 25 %, Mo03 35,000 1 Scrubber system 55,000 1 Stack (SS 316, ?> x 150') 25,000 1 Scraper system for reactor (SS 316) 15,000 - Piping, insulation for reactor 20,000 1 Screw conveyors [3"$) 5,000 - Instrumentation 25,000 1 Packer Unit 10,000 CAPITAL COST ' 1 •~ 382,000 Installation 440,000 TOTAL CAPITAL COSTS ~ $830,000 218 Table 10.5 Summary of Operating Costs 10 TPD Molybdenite roasting plant Multiple Hearth Process Fluidized Bed Process Labour 1 tf/lb 1 tf/lb Mai ntenance 3 tf/lb 2 tf/lb Fuel 1 <t/lb 0.04 <t/1 b Power 0.1 tf/lb 0.1 tf/lb Gas cleaning 1.5 tt / l b 1 tf/lb Total Operat-ing Costs 6.7 tf/lb -4.2 tf/lb industrial source [128] have been used. Labour has been taken to be the same for both processes, even though lower values of maintenance and gas cleaning costs for the f luidized bed are probable. This is because the reactor requires less attention, e.g. continuous cleaning of the discharge ports, than the multiple hearth furnace. In addition there is no expense for the operation of electrostatic precipitators. The current prices for the molybdenite concentrates are ~$2.0/lb, while the technical molybdic oxide is ~$2.30/lb, leaving a margin of about $30 tf/lb. Since the operating 219 cost in the f luidized bed process is only about 60% of the multiple hearth process, amortization of the plant can be undertaken over about 8-1/2 years against 10 years for the multiple hearth process. The thermal balance for the f luidized bed reactor using slurry feed is given in Figure 10 .5 . It can be seen that a negligible amount of fuel is required in this process: 0.04 <£/lb against 1 <£/lb in the multiple hearth furnace. For the feeding of solid molybdenite concentrates into the fluidized bed reactor, no fuel is required, and indeed a large surplus of heat is generated (4.31 x 105 Kcal /hr) . From this analysis, i t seems clear that substantially lower capital investment and operating costs can be achieved using the f luidized bed process rather than a multiple hearth furnace for molybdenite roasting. 220 vaporized water from scrubber slurry 2.0 X \tf kcal/ hr discharged calcines 2.5 X IO2 kcal/hr [preheat air 1.8 x io kcal/hr calcines 2.5 XIO3 kcal/hr vaporized water from slurry feed /4.26 X io kcal/hr total heat to the reactor 6.60 X |0 5 kcal/hr heat of reaction 6.58 XIO 5 kcal/hr heat losses in reactor and lines 3.28 X |0 4 kcal/hr fuel 2.5 X io 3 kcal/ hr Figure 10.5. Thermal balance for the f luidized bed reactor using slurry feed. Chapter 11 SUMMARY AND CONCLUSIONS The present research work on t h e development and operation of a f]uidized bed process for the roasting of molybdenite concentrates can be summarized as follows: 1) Without pretreatment, molybdenite concentrates have been roasted successfully in a novel type of recirculat i f luidized bed reactor which uses a mixture of calcines and coarse sand. Molybdenite concentrates and recirculated calcines are continuously fed %o the reactor using a pneumatic injection system, while the buildup of material inside the reactor is avoided using a rotary arm scraper. 2) To avoid strat i f icat ion of the calcines and inert material (sand) inside the reactor, the latter should have a wide size d istr ibut ion, such as -40/+140 mesh; a lso , the f luidized bed reactor should be operated with a super-f i c i a l gas velocity in excess of 18 cm/sec at the roasting temperature. 3) For the bimodal system of particles of coarse sand and fine calcines, the concentration of the calcines 221 222 in the bed has a signif icant ly larger effect on the f l u i d i z a -tion properties than the sand for mixtures containing over 20 wt-% calcines. For the range 20-60 wtr% calcines, an empirical relationship was obtained to estimate the size of the gas bubbles along the f luidized bed. 4) Using the pneumatic injection system for the feed and a long average residence time for the so l ids , e .g. over 15 hours, the reactor is v ir tual ly backmixed for the solids with a large dispersion of gas along the f luidized bed. However, a small concentration profi le seems to exist for the M Q S 2 from the feed point to the top of the bed at steady state conditiops. For the rate of elutriat ion of fines from ^he bed, an exponential expression of the super-f i c i a l gas velocity was found to hold. 5) The transformation of MoS2 to M 0 O 3 is a complex process which involves an in i t i a l fast stage which seems to be chemical reaction controlled. This is followed by a f inal slow, and apparently solid diffusion controlled regime. Vaporization of M 0 O 3 and subsequent condensation seems to play an important role in the kinetics of oxidation. 6) The optimum operating conditions of the f luidized bed reactor l i e within the following ranges: 223 S u p e r f i c i a l gas v e l o c i t y : 18 t o 35 cm/sec T e m p e r a t u r e o f t h e bed : 520 t o 550°C A v e r a g e r e s i d e n c e t i m e o f s o l i d s : 20 t o AO h r s 7) To obtain calcines with a low sulphur concentra-t ion, the molybdenite concentrates must not contain more than about 0.04% calcium, with an average part icle size of MoS2 less than 15 microns and 100%-325 mesh. 8) For scale-up purposes, a semi-empirical re lat ion-ship based on the average size of molybdenite part icles and the temperature of roasting has been used. The scale-up indicates that industrial size f luidized bed reactors should be capable of throughputs that are 30 to 50 times greater than those obtained in the mu1tipie hearth furnace per unit area of hearth. The predicted values of S0 2 in the off gases from the large f luidized beds (5 to 20 TPD) should range from 3 to 4%. 9) The capital cost and operating costs of a f luidized bed process appear to be about 50% and 60%, respec-tively of the capital and operating costs of a conventional multiple hearth process of the same capacity. 224 SUGGESTIONS FOR FURTHER RESEARCH 1) A fu l l size pi lot plant of 0.3-0.5 ton/day is required to assess on a larger scale the feas ib i l i t y of this process for possible application in industry. 2) More information is required concerning the strat i f icat ion process for different sizes and densities of part ic les, as well as the influence of external variables on phenomenon such as the f lu id character ist ics, superficial gas velocity and bubble size and distr ibut ion. 3) Additional research on the solid and momentum transfer of a pneumatic injection system discharged into a f luidized bed reactor is needed. 4) Further studies on the kinetics of the MoS2 roasting in a f luidized bed seems desirable, as well as a study of the f inal stage of sulphur elimination. 5) More research on the elutr iat ion phenomenon is needed to predict from given characteristics of the f lu id and solids in a f luidized bed reactor the rate of elutriat ion and the size distribution of the elutriated material. N O M E N C L A T U R E b adsorption factor of gas ( - ) C capacity of reactor (ton/day) C concentration of tracer gas (cm 3 / l t ) CQ concentration of oxygen in the 2 reaction gas (gr-mole/cm3) C p , s ' C p s s h e a t c a P a c i t y ° f solid gas (cal/gr-mole) C s concentration of tracer solid (%) C° concentration of MoS2 at feeding point {%) C* concentration of MoS2 in elutriated calcines (%) C(t ) , C(e) internal age distribution function of tracer as a function of time t and dimensionless time 6 ( - ' ) dg bubble diameter (cm) d £ cloud diameter (cm) 225 226 axial dispersion coefficient of solid Da q axial dispersion coefficient of , y gas ° r - radial dispersion coefficient of solid Dr q radial dispersion coefficient of gas dc a 1 average diameter of calcine parti cles dS 7 - ; average diameter of sand particles a v e r a g e diameter of MoS2 particles ds diameter of MoS2 particles D diffusion coefficient E(t) , E ( 6 ) exit age distribution function of tracer in the bed as a function of time t and dimensionl ess time 9 f expansion factor of f luidized bed f. volumetric factor of f luidized bed f voidage factor of f lu idizat ion rs volumetric fraction of downflow phase F(t) transformation function of an individual MoS2 part icle F* elutriation flux (cm2/sec) (cm2/sec) (cm2/sec) (cm2/sec) (cm) (cm) (cm) (cm) (cm2/sec) ( - ) ( - ) ( - ) ( - ) • : (cm3/cm3) ( - ) (gr/cm2sec) 227 volumetric fraction of upflow phase (cm3/cm3) feed rate of molybdenite concern trates (gr/min), (Kg/hr) (Ton/day) gas flow rate at 25°C (1/sec) (m3/min) gas flow rate at roasting tempera- (1/sec) ture (m3/min) free energy change number of holes in the gas d i s t r i -butor gri d (Kcal/mol ) ( - ) heat transfer coeff icient between gas and particles (cal/cm 2sec °C) heat of reaction (Kcal/mole) elutriation rate (gr/min) (Kg/hr) thermal conductivity of gas (cal/cm sec°C) rate of constant for the chemical reaction based on unit area of sol ids (cm/sec) jas transfer coeff icient between >ubble and cloud (sec - 1 ) gas transfer coefficient between cloud and emulsion (sec - 1 ) 228 overall transfer coefficient of gas between the gas phase and the emulsion phase ' ( sec - 1 ) f luidized bed height (cm), (m) stat ic bed height (cm), (m) minimum f lu idiz ing bed height (cm), (m) static bed height of sand (cm), (m) adsorption coefficient ( - ) number of holes per unit area of the gas distributor ( - ) Nusselt number ( - ) partial pressure of oxygen (atm) total pressure (atm) Peclet number for mass transfer ( - ) heat flow (cal/sec cm2) heat losses by convection (cal/sec cm2) heat losses by radiation (ca./sec cm2) recirculation ratio (.-• ) reactor diameter reactor height Reynolds number with respect to the gas Reynolds number with respect to the sol id reactor cross sectional area specif ic surface area of mplybde ni te concentrates surface area of bubble temperature of the bed, average temperature of the gas temperature of reacting MoS2 parti cles time average residence time of solids in the bed superficial gas velocity of gas^  at minimum f lu idiz ing conditions superficial gas velocity 229 (cm), (mt) (cm), (mt) ( - ) ( - ) (cm 2), (m2) (cm2/mole-gr) (cm2) (°C) (°C) (°C) (sec), (min), (hr) (min), (hr) (cm/sec) (cm/sec) 230 bubble r ising velocity emulsion downflow velocity bubble volume (cm/sec) (qm/sec) (cm 3) superficial velocity of downflow of solids (cm/sec) superficial velocity of upward flow of solid (cm/sec) total weight of solids in the bed (gr), (Kg), (Ton) weight fraction of calcines in t h e bed (gr), (Kg), (Ton) weight fraction of sand in the bed ' (gr), (Kg), (Ton) Fraction of MoS2 in the discharged ca1ci nes ( - ) fraction of Mo0 3 in the discharged calcines volumetric fraction of wake in the bubbles ( - ) fraction of the bed consisting of gas bubbles ( - ) voidage of static bed ( - ) voidage of bed at minimum f lu idiz ing conditions emmissivity coefficient ( - ) ( ^ ) 231 e f voidage of f luidized bed ( - ) n col lection eff iciency {%) 6 dimensionless time ( T ) y g viscosity of gas (poise) P s density of solid (gr/cm 3) P g density of gas (gr/cm 3) P s average bulk density of solids (gr/cm 3) P c a l average bulk density of calcines (gr/cm 3) P s i 1 average bulk density of sand (gr/cm 3) a Stefan-Boltzmann constant (cal/cm 2sec °K*) a 2 variance of distribution (sec 2) i total time of transformation (min), (hr) ip centroid of distribution ( - ) Y stoichiometric factor of ^ M0S2/M0O3 ( - ) R E F E R E N C E S Zelikman, A . N . , Krein, O.E. and Samsonov, G.V. , "Metallurgyof.Rare Metals," 2nd E d . , Izdatel'stro Metal 1urgiya, Moscow (1964), 69. Butterf ie ld, J .A . and Morgan, D.L. , CIM Bul le t in , No, 8 (1972), 45-48. Northcott, L . , "Molybdenum," Butterworths, London (1 956) , 5 Scholz, E .A. , CIM Bul let in , No. 1 (1969), 23. Scholz, Ibid., p. 22. Zelikman, A . N . , Krein, O.E. and Samsonov, G.V., op. cit., p. 70. Ibid. 3 p. 66. Butterf ie ld, J .A . and Morgan, D.L. , op. cit., p. 70. Kubaschewski, 0. , Evans, E.LI, and Alcock, C . B . , "Metallurgical Thermochemistry," 3rd ed . , Pergamon, London (1967), 334-335. Northcott, L . , op. cit., p. 18-19, Sutulov, A . , "Molybdenum Extractive Metallurgy," University of Concepcion, Chile (1965), 176. Gerasimov, Ya. I., Krestovnikov, A.N. and Shakhov, A . S . , "Chemical Thermodynamics in Nonferrous Metallurgy," Vol. I l l (NASA t rans l . ) , Washington, D.C. (1965), 66. 232 233 13. Gerasimov, Ya. I., Krestovnikov, A.N. and Shakhov, A . S . , Ibid., p. 58. 14. Kubaschewski, 0 . , Evans, E. LI. and Alcock, C . B . , op. c i t . , p. 414. 15. Gerasimov,. Ya . ,1, , Krestovnikov, A.N. and Shakhov, A . S . , op. o i t . , p. 61 . .16. Gerasimov, Ya. I., Krestovnikov, A.N. and Shakhov, A . S . , Ibid, j p. 67. 17. Gerasimov, Ya. I., Krestovnikov, A.N. and Shakhov, A . S . , Ibid., p. 68. 18. Condurier, L . , Wilkomirsky, I., and Morizot, G . , Trans. IMM, Sec. C, 79, (1970), C-34-40. 19. Kelley, K.K. , "Contribution to the Data on Theoretical Metallurgy," U.S. Bureau of Mines, Bu l l . 406 (1937), 371 . 20. Wicks, C.E. and Block, F . E . , "Thermodynamic properties of 65 elements," U.S. Bureau of Mines, Bu l l . 605 (1 963), 79. 21. Kelley, K.K., op. cit., p. 386. 22. Wicks, C .E . and Block, F . E . , op. oit., p. 79. 23. Gerasimov Ya. I,, Krestovnikov, A . N . , and Shakhov, A . S . , op. c i t . , p. 59. 24. Wheeler, E . S . , AIME Techn. Publ. (1944), 1718-1723. 25. 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Tech., 38 (1966), 1 1 69. 241 APPENDIX 1 VAPOUR PRESSURE OF Mo03 AS A FUNCTION OF TEMPERATURE Temperature (°C) 243 APPENDIX 2 ENTHALPY AND FREE ENERGY OF REACTION *sS(g) + ° 2 2 S ° 2 ( g ) 298-2000°K AH° = 0172,630 - 1.49% + 0.712 x 1O - 3 T 2 + 0.336 x 10 T - 1 AG-J = -173,240 + 34.62T M o S 2 ( s ) + 7/2 0 2^ M o 0 3 ( s ) + 2S02 A H " =, -297,590 - 4.49T + 7.862 x 10~3 T 2 + 3.191 x 105 T - 1 A G j = -266,840 + 12.75T x logT - 1.35 x IO - 3 T 2 + 36.44T + 1.54 x 10~5 T" 1 M o S 2 ( s ) + 3 ° 2 2 M o ° 2 ( s ) + 2 S0 2 298-1100°K A G ? = -224,340 + 9.82T log T + 0.2 x IO"3 T 2 t + 2.56T c a l cal/mol cal/mol 298-1068°K cal cal 2 (s) + 0 2 + M ° ° 3 ( s ) 298-1068°K 244 AG T = -45,500 + 2.93TlogT _ 1.55 x 10~3 T 2 + 1.54 x 105 T - 1 cal 6 M o 0 3 ( s ) + M o S 2 ( s ) I 7 M o 0 2 ( s ) + 2 S0: 298-1068°K A G T = 30,660 - 7.76TlogT + 9.5 x 10"3 T 2 -9.24 x 10"5 T~ 1 - 200.84T cal APPENDIX 3 246 MODIFIED VENTURI NOZZLE FOR SOLID INJECTION INTO THE FLUIDIZED BED REACTOR oil dimensions in millimeters 150 247 APPENDIX 4 FLUIDIZING CHARACTERISTIC OF SILICA SAND - M0O3 CALCINES MIXTURES 7.5 cm diameter plastic reactor Ottawa sand (-40/+140 mesh) M 0 O 3 (-325 mesh) G f U o AP L f (It/min) (air 25°C) (cm/sec) (cm H20) 1 (cm) 1) 10 wt-% M 0 O 3 Lm = 8.5 cm m Lmf = 8 ' 8 .5 2.0 2.2 8.5 .8.6 3.4 4.9 8.7 12.0 4.6 7.0 8.8 16.1 6.2 7.7 8.9 24.0 9.0 8.4 8.9 32.2 12.0 9.0 9 40.8 15.0 9.1 9.3 2) 20 wt-% Mo03 Lm = 8 - 7 c m Lmf = 9 ' ° c m 5.0 2.0 5.1 8.7 7.0 2.7 3.7 8.7 8.6 3.4 8.9 8.8 12.0 4.6 8.4 9.0 16.1 ' 6 . 2 8.4 9.3 24.0 9.0 9.0 9.5 32.2 12.0 9.5 9.9 40.8 15.1 10.3 10.3 248 ( I t / m i n ) ( a i r 25°C) u 0 (cm/sec) AP (cm H 20) (cm) 3) 40 w t - % Mo03 lm = 11.7 cm L m f = 12 .2 cm 5.0 7.0 8.6 12.0 16.1 24.0 32.2 40.8 2.0 2.7 3.4 4.6 6.2 9.0 12.0 15.1 7.3 11 .0 12.8 12.5 12.6 12.3 12.8 12.9 12.0 1 2.0 12.2 12.4 12.7 13.5 13.7 14.0 4) 50 w t - % MoO ' L m = 1 5 - 2 c ™ Lmf = 1 15 .3 cm 5.0 7.0 8.6 12.0 16.1 24.0 32.2 40.8 2.0 2.7 3.4 4.6 6.2 9.0 12.0 15.1 16.7 16.3 17.1 17.2 17.2 17.5 17.6 17.0 15.4 15.4 15.3 15.2 15.2 15.5 16.2 16.6 5) 60 w t - % MoO3 L m = 19.2 cm Lmf - 20 .4 5.0 7.0 8.6 12.0 16.1 24.0 32.2 40.8 2.0 2.7 3.4 4.6 6.2 9.0 12.0 15.1 20.3 22.9 22.7 23.2 22.3 22.1 22.0 20.3 19.5 20.4 20.4 20.5 20.5 21 .0 21 .3 21.7 249 ( I t / m i n) ( a i r 2 5 ° C ) U o (cm/sec ) AP (cm H 2 0 ) L f (Cm) 6) 0% M o 0 3 ( 100 wt-% O t t a w a S a n d - 4 0 / + 1 4 0 ) L m = 8 c m L m f = 8 .1 cm 5 . 0 7 . 0 8 . 6 1 2 . 0 1 6 . 1 2 1 . 0 2 4 . 0 3 0 . 0 3 2 . 2 4 0 . 8 4 9 . 6 7 5 . 0 2 . 0 2 . 7 3 . 4 4 . 6 6 . 2 8 . 0 9 . 0 1 1 . 0 1 2 . 0 1 5 . 1 2 2 . 0 2 7 . 5 1 .6 2 . 8 3 . 5 5 . 6 7 . 3 9 . 0 9 . 4 9 . 8 1 0 . 0 1 0 . 0 10 .1 1 0 . 1 8 . 0 8 . 0 8 . 0 8 . 0 8 .1 8 .1 8 .1 8 . 2 8 . 4 8 . 8 9 . 3 1 0 . 8 7) 100 wt-% M 0 O 3 L m = 2 7 . 5 cm ( - 3 2 5 ) Lmf = 2 7 - 8 cm 2 3 5 6 9 11 1 6 . 1 21 24 1 . 0 1 .4 2 . 0 2 . 7 3 . 5 4 . 2 6 . 2 8 . 0 9 . 0 4 . 5 9 . 0 5 . 2 5 . 6 7 . 6 6 . 0 6 . 6 1 0 . 2 8 . 6 2 7 . 5 2 7 . 5 2 7 . 6 2 7 . 7 28 2 8 . 3 2 9 . 5 31 38 250 APPENDIX 5 AVERAGE MEASURED DIAMETER OF BUBBLES FROM HIGH SPEED PICTURES 1 , Test wt-% u 0 ^b' c m No Calci nes cm/sec Distance from the gas distributor, cm 0-5 5-10 10-15 15-20 20-25 25-30 3Q-35 B.106-A.1 A.2 A.3 0 II 20 30 40 0.5 0.9 1.3 1 .2 1 .8 2.3 1.0 2.6 2.8 1 .4 1 .6 6.1 6.3 B.106-B.1 B.2 B.3 B.4 20 II II M 15 20 30 40 0.6 1.4* 1 ,4 1 - e ! 2.3* 2 " ° * 2.5 4.4* 1 .6 3.3 5.6 2.4 3.2 4.6* - -B.106-C.1 C.2 C.3 C.4 40 II n n 15 20 30 40 1.5 1.2 1 .0* 1.6 1.9 2.1 1 .8* 1.9 1.9 1 .7 2.5* 3.7 3.4 2 - 2 * 3.7 3.6 3.0 2.6 4.4 4.0 2.2 -B.T06-D.1 D.2 D.3. D.4 D.5 60 II II II II 15 20 30 40 10 1 .9 0.9 1.2 0.8 2.6* 1 .2 1.6 1 .4 4.8 1.8 2.6 2.8 0.5 2.4 3.6 3.0 1.1 4.4 2.1 3.9 4.5 2.4 4.5 5.2 2.0 3.1 5.8 B.106-E.1 E. 2 E.3 E.4 E.5 80 II n II 10 15 20 39 40 1.2 0.9 1 .8 1.9 2.0 2.6 1 .2 2.4 2.2 2.4 2.7 1.6 3.7 1 .8 2.8 1.9 2.9 3.7 4.2 2.9 4.2 2.8 3.2 4.3 T Average from two different picture measurements. APPENDIX 6 STRATIFICATION OF SOLIDS TEST RESULTS RUN B.I 20 wt-% Mo03 - 325 mesh 80 wt-% sand -40/+70 mesh L = 25 cm m G f = 28.2 lt/min U0 = 30 cm/sec time of test = 180 sec time of delta input = 3 sec Sample No Di stance cm S i l i ca sand gr Calcines gr wt-% B.l-1 0-2 35.5 4.45 11.1 2 2-4 28.7 0.55 1.9 3 4-6 26.3 0.04 0.1 4 6-8 52.0 0.06 0.0 5 8-10 51 .0 0.07 0.0 6 10-12 48.5 0.17 0.2 7 12-14 73.1 0.15 0.1 8 14-16 43.8 0.12 0.1 9 16-18 53.5 0.15 0.1 10 18-20 49.2 0.08 0.0 11 20-22 34.7 8.90 20.4 12 22-24 7.6 30.70 80.0 13 24-26 5.9 32.70 84.7 14 26-28 6.5 44.80 87.4 CONTINUED 252 APPENDIX 6 (Continued) RUN B.2 20 wt-% Mo03 -325 mesh 80 wt-% Ottawa s i l i c a sand -40/+70 mesh 1 = 25 cm m G f = 28.2 It/min U0 = 30 cm/sec time of test: 30 sec Sample . No Di stance,, cm S i l i ca sand gr Calcines MoS2 % . gr wt-%. B .2-1 0-2 43.2 5.00 10.3 1.00 2 2-4 53.6 0.85 1.6 6.72 3 4-6 47.2 0.45 0.9 11 .40 4 6-8 57.8 0.83 1 .4 6.75 5 ' 8-10 50.0 . 6.16 10.9 1 .05 6 10-12 33.8 9.45 21 .9 0.26 7 12-14 36.4 13.45 31 .1 0.12 8 14-16 34.2 13.42 31 .0 0.01 9 16-18 40.0 16.58 29.3 0.0 10 18-20 33.8 15.22 31 .3 — 11 20-22 38.8 17.37 30.8 0.0 12 22-24 26.3 12.96 33.0 0.0 13 24-26 23.1 1 5.30 39.8 0.0 CONTINUED 253 APPENDIX 6 (Continued) RUN B.4 20 wt-% Mo03 -325 mesh 80 wt-% Ottawa s i l i c a sand -40/+70 mesh 1 = 25 cm m G f = 28.2 It/min U0 = 30 cm/sec time of test - 15 sec time of delta input = 3 sec Sample No Distance cm S i l i ca sand gr Calcines MoS2 % gr wt-% B.4-1 0-2 57.61 2.05 3.4 0.03 2 2-4 45.33 0.11 0.2 16.75 3 4-6 50.72 0.13 0.3 10.40 4 • 6-8 54.00 0.15 0.3 7.70 5 : 8-10 54.71 1 .58 2.8 3.70 6 10-12 43.18 6.39 12.9 1 .50 7 12-14 42.91 12.82 23.0 1 .20 8 14-16 39.58 13.79 25.8 1 .00 9 16-18 40.08 16.86 29.6 0.93 10 18-20 38.63 16.05 29.4 0.85 11 20-22 34.34 1 3.75 28.6 0.95 12 22-24 22.85 9.88 30.1 0.88 13 24-26 10.50 5.61 34.8 0.80 CONTINUED APPENDIX 6 (Continued) RUN B.64 60 wt-% Mo03 -325 mesh 4.0 wt-% Ottawa sand -40/+140 mesh G f = 28.2 It/min U0 = 30 cm/sec time of test = 180 sec Sample No Distance cm S i l i ca sand gr Calcines gr , wt-% B.64-1 0-2 27.76 19.05 40.7 2 2-4 21 .82 16.60 43.2 3 4-6 16.89 18.26 51 .9 4 6-8 15.73 18.69 54.3 5 8-10 15.33 20.15 56.8 6 10-12 14.11 20.63 59.4 7 12-14 11.26 17.00 60.2 8 14-16 12.93 17.47 57.5 9 16-18 15.32 24. 52 61 .6 10 18-20 14.72 24.33 62.3 11 20-22 13.13 23.19 63.8 12 22-24 12.12 21.51 63.9 13 24-26 12.82 23.43 64.6 14 26-28 13.82 25.52 64.8 15 28-31 14.18 28.82 67.0 CONTINUED 255 APPENDIX 6 (Continued) RUN B.65 60 wt-% Mo03 -325 mesh 40 wt-% Ottawa sand -40/+140 mesh 1 = 31 cm m G f = 10 It/mi n U0 = cm/sec time of test = 180 sec time of delta input = 3 sec Sample No Distance cm S i l i ca sand gr Calcines MoS2 gr Wt-% % B.65-1 0-2 31 .56 17.57 35.8 0.08 2 2-4 40.05 6.63 13.5 0.08 3 4-6 42.72 5.91 12.1 0.05 4 6-8 34.60 . 8.25 19.3 0.08 5 8-10 16.39 21 .29 56. 5 0.53 6 10-12 10.50 25.23 70.6 1 .83 7 12-14 8.93 22.95 72.0 1 .60 8 14-16 7.55 21 .56 74.1 1.21 9 16-18 9.32 26.82 74.2 0.80 10 18-20 9.74 28.43 74.5 0.73 11 20-22 8.17 24.91 75.3 0.60 12 22-24 8.52 26.93 76.0 0.53 13 24-26 8.65 28.72 76.9 0.43 14 26-28 12.62 54.84 81.3 0.25 15 28-31 9.95 51 .63 83.8 0.05 CONTINUED APPENDIX 6 (Continued) RUN B.93 60 wt-% .M0O.3 -325 mesh 40 wt-% Ottawa s i l i c a sand -40/+140 mesh 1 • = 31 cm m G f = 19 It/min U 0 = 20 cm/sec time of test = 180 sec Sample No Distance cm S i l i ca sand gr Calcines gr wt-% B.93-1 0-2 26.84 17.69 39.7 2 2-4 28.00 12.56 31 .0 3 4-6 20.84 17.44 45.6 4 6-8 16.73 22.84 57.7 5 8-10 14.11 23.82 62.8 6 10-12 11 .44 21 .86 65.6 7 12.-14 10.26 22.22 68.4 8 14-16 10.51 20.00 65.6 9 16-18 13.33 26.35 66.4 10 18-20 12.67 24.71 66.1 11 20-22 13.25 25.71 66.0 12 22-24 1 2.37 26.70 68.3 13 24-26 9.46 22.53 67.7 14 26-28 5.76 13.25 69.7 15 28-31 3.07 8.01 72.3 CONTINUED 257 APPENDIX 6 (Continued) RUN B.94 40 wt-% Mo03 -325 mesh 60 wt-% Ottawa S i l i c a sand -40/+140 mesh 1 = 31 cm m G f =28.9 It/min U 0 = 30 cm/sec time of test = 180 sec time of delta input = 3 sec Sample No Distance cm S i l i c a sand gr Calcines MoS2 gr wt-% %. B.94-1 0-2 39.48 12.74 24.4 0.08 2 2-4 35.39 9.57 21 .9 0.35 .3 4-6 31.32 11 .72 27.2 0.15 4 6-8 28.22 12.45 30.6 0.53 5 8-10 26.72 14.55 35.3 1,15 6 10-12 29.12 17.78 37.9 2.20 7 12-14 22.32 15.62 41 .7 1.85 8 14-16 21 .95 1 5.77 41 .8 1 .88 9 16-18 25.59 18.93 42.6 1.13 10 18-20 22.85 21.93 49.0 0.80 11 20-22 19.15 19.00 49.8 0.85 12 22-24 19.54 20.83 49.1 1 .00 13 24-26 12.75 11 .66 47 T8 0.55 14 26-28 4.07 3.83 48.5 0.40 15 28-31 1.92 2.05 51.6 0.35 CONTINUED APPENDIX 6 (Continued) RUN .3.9.5.. 20 wt-% M0O0 -325 mesh 80 wt-%.-Ottawa', s i l i c a sand -40/+140 mesh G f = 28.9 It/min U0 = 30 cm/sec time of test = 180 sec Sample No Dis tance cm S i l i ca sand gr Calcines gr wt-% B.95-1 0-2 42.46 6.80 13.8 2 2-4 37.72 4. 21 10.0 3 4-6 39.03 6.50 14.3 4 6-8 39.62 7.57 16.0 5 , 8-10 41.42 9.00 17.9 6 10-12 39.85 9.13 18.1 7 12-14 43.96 10.72 19.6 8 14-16 35.12 9.48 21 .3 9 16-18 43.86 11 .00 20.1 10 18-20 42.59 11.11 20.7 11 20-22 43.97 1 2.09 21 .6 12 22-24 42.17 11 .83 21 .9 13 24-26 43.41 12.00 21 .7 14 26-28 43.82 12.33 22.0 15 28-34 111.94 40.00 25.0 259 APPENDIX 7 SIZE DISTRIBUTION OF STANDARD OTTAWA SILICA SAND, COULTER COUNTER SIZE ANALYSIS OF MoS2 CONCENTRATES, AND CYCLOSIZER ANALYSIS OF MoS2 CONCENTRATES SIZE DISTRIBUTION OF STANDARD OTTAWA SILICA SAND mesh, US Std. wt-% ~ T - 1—: 1 d s , mm -40/+50 63.2 0.359 -5O/+70 26.1 0.254 -7U/+100 9.2 0.180 -100/ + 140 1.5 0.127 100.0 d s = 0.311 mm PARTICLE SI2 E DISTRIBUTION S A M P L E IDENTIFICATION ^y b.^.ilL5^.i niriIi ( H o S? ) B' e n d a • Density * * - 8 0 g / c c LIQUID: Density 0 , 8 2 1 9 . g / c c V i s c o s i t y . 7 , 6 8 0 c p s ; Preparation 0 . 35 g sampl e were, . d i s p e r s e d ..u. I t r a son ixaJ Iy ._w. i th js imul ta neous ._s t i r r ing f o r 5 m i n u t e s i n 30 ml S E O I S P E R S E - E C C - 1 ? » J < D A T E 9-26-1974* BY R. H o l l ANALYSIS NO. 1^86 RATE 112 a t 30 C LiH Tt.t.r +rtr i 1 j t~ L f lie tta s s . M i-rt-j- Tt"ii£- iiii Iii! :: j ai; L i i r • c b i —'—1—-t-1*0 in: i^:: LHi •;ii"f S I : i : ; Hi i Hi'; !"T:; 9b i i i i i i i i .! i •.:•.: i . : : 9o : ;. t.:. /tiiiiti" :'t:i: i i i :rii H i : " :;M-:p-r LlL-i H ! is:o •if: Liif it Li Li5i tit "till'--T-HH 80 •rr=t4r #LtfLS •Hii. Hi? ii±i r: 3rrr Hi i j : i t4:r.r. i i U l iH :. _i_:.. U : i l : ; iii: :; l : H i : V : 1 l-i Lii'L I Tin - iiir.|: t ' !"'"tM" * •-. '• l-f 60 iiii . . - t i i - i • i , 60 LLiiBt :::i LM-: il}: Liu ::+. :.'t: riLHlHi • H i - \: : 1 •V 50 m\ rt'.: i i i i i : .- 4:l4-t:. • : • so ; : : : Hr? ilHT Hift ll i iLi : : :V i i i i tftf p. ;-•!•::' : 40 :.r.tt-: rv£ i i i : i i i i mi:. ::f:t-:T-i-:.t. r. j . • t-' -i f 5 1 . : : t l , . t i f f i : i~ i i iL - i i i i - i i i - i:i i l i pi i i i i irxL: • ;-:::t.-:;.v-• » • M * i i i 1 ' •*-tlS5S H_i" r it t i 11-iH ' T- f : : : : j r _ g : i i : » :ii: •iv : H i : .fRljl IHitT •Li: i : Li i ± f S r : : i : i i i i i i i i ; ";t-li-Jt::. " : is 5§ • § 3 r LH; 'ti' iiii 113 m J t r r tri r >A___L____ ni'- • ir UJ 2 U C/5 tn < > < D 5 u 100 80 60 50 40 1 0.8 0.6 0.5 0.4 0.3 EQUIVALENT SPHERICAL DIAMETER, MICRONS. MICRON DATA LABORATORIES, INC. 22 Secord Dr., St. Cat larines, Ont., Canada L2N IK8 Phone (116) 934-0629 ro P A R T I C L E S I Z E D I S T R I B U T I O N S A M P L E IDENTIFICATION M o l y b d e n i t e c o n c e n t r a t e ( H o S 2 ) II.C. Moly  Density *».80 g / c c LIQUID: Density .Q.-.821? g / c c Viscosity .7..680 cps Preparation 0 -35 9 sample were d i s p e r s e d u l t r a s o n i c a l 1 y w i t h s i m u l t a n e o u s s t i r r i n g f o r 5 ._"»nutes i n 30 ml SEOISPERSE-S ;. GC-19924 D A T E 9-19-197^  B Y R. H o l t A N A L Y S I S N O . li»78 RATE 112 a t 30 C :::|:;:: r - r rrrr iv iff: I i S f f j i f •ifi 1 R 1 M ~t - i~r -f" : ; t : ; f n ~—"Pico : • • • I W i f f . ff Iril .-,-4-1 ::n i f f f f f f f i l i i i i n - ; i f i f iiii ;|t:::!:: iii: it-n ; f f f : ; i ; i i i f H i f ; ' s m :::: It4i-::il ::!: i f f ! iii: iff k|3 f f e i n io iff': i ...-4-...-. :::: • ; -; filii" i f f f • :: / : Hill';::; j f f i r \ r r t 70 mm .tiff fff-.fr • .i-i 70-:.;4_!. : : . : : . ; \ i i muii •Hi f f f f r : . i § r r h l : m- 1 _;V._4— i f\*o i::i : . j . : : : : U i f f f M i t fr jr .r . : : : 60 i-fii^ : i i i V : if:: m i . . : : : : i f f . 1 - - j -ffffffi •fffH-: : : '• so> Mm i f f ; fff.T.' : : ... I-.J 50 : ru f i f f - tffrj inf. i f i i i f-ffiT i : i f i'i'.ff; - Jr4- i -i -r j r r r r r *(j i f f : t f l f i : i i :• i E f f i f i 1. tiff i H : i f f f i f t f ii. '• '• '• 30 \ : : \ t i M f f f 30 : i : ; : : : : iiiibii i ' f ff . fffri :r£::::;._: rcr-l-:-. :::> f . f H -• .rr ; r : f : i : . " 1 " i o :::: Hf:i r f r F E i f : . : ; i 4 TO .iff: * 4 - L i -f t f f " f f - i i . i> f i f f i " i i i : '•|* ^ i i i . r C r r r t n . : - . i i f i —A f f § P + 3 lit- • r t : : T : i f f f . f H f r F :::: tO S O . 40 1 0.8 0.6 O.S 0.4 EQUIVALENT SPHERICAL DIAMETER. MICRONS. MICRON DATA LABORATORIES, INC. 22 Secord Dr., St. Catharines, Ont., Canada L2N IK8 Phone (416) 934-0629 PO 264 APPENDIX 8 ELUTRIATION TESTS RESULTS A) B.C. Moly Calcines Temp. G f , 25 °C G f ,T°C Uo , » i k* °C 1t/mi n lt/min cm/sec gr/mi n 540 81 226 29.7 98.2 540 90 253 33.2 105.2 540 76 214 28.0 74.6 540 70 198 26.0 61 .0 540 66 187 24.6 52.8 480 77 204 26.4 53.0 460 77 202 26.5 87.8 523 81 224 29.4 113.6 560 81 232 30.6 98.9 505 62 170 22.3 46.7 539 76 214 28.2 84.0 560 80 231 30.2 101 .0 525 81 223 29.4 96.7 550 77 218 28.5 51.0 550 77 218 28.5 104.2 550 77 218 28.5 106.1 550 77 218 28.5 106.4 550 69 198 26.1 95.2 550 62 177 23.3 68.6 550 88 247 32.3 136.9 574 68 199 26.0 56.4 574 70 206 27.0 68.4 574 73 213 28.1 73.2 550 68 195 25.4 82.8 550 72 205 26.8 90.0 550 82 233 30.7 111.8 560 41 114 15.0 11.1 550 50 138 18.0 14.1 550 77 218 28.5 73.2 540 56 153 20.0 51,0 550 52 143 18.6 36.0 CONTINUED 265 APPENDIX 8 (Continued) Temp. G f ,25°C G f ,T°C Uo, T k* °C 1t/mi n 1t/mi n cm/sec gr/min 560 55 154 20.3 38.3 534 51 137 17.9 23.0 450 77 198 25.9 92.4 486 61 165 22.4 60.5 342 60 141 18.7 29.9 540 77 215 27.8 56.6 550 62 177 23.3 60.9 550 68 195 25.4 76.7 540 72 203 26.5 84.4 510 76 207 27.4 85.5 560 63 184 24.0 42.2 555 58 168 22.0 56.9 580 72 214 27.9 57.4 567 69 198 25.9 79.0 556 66 191 25.1 54.6 556 64 184 24.2 49.2 500 . 72 195 25.4 55.4 546 70 199 26.1 86.8 546 74 212 27.6 91 .2 518 63 175 23.5 96.3 446 62 163 21 .4 48.0 430 43 103 13.8 17.4 504 65 178 23.7 82.8 540 65 186 24.3 81 .5 538 65 183 24.0 83.6 330 66 1 56 20.6 81 .2 540 66 187 24.6 107.5 550 66 191 25.0 88.9 550 77 218 28.0 115.2 518 50 132 17.5 40.4 500 70 190 25.0 72.5 480 86 228 30.0 129.5 550 67 193 25.3 90.5 B) Brenda Calcines 550 67 193 25.3 85.4 551 66 192 25.2 79.0 540 62 177 23.2 69.5 530 44 119 15.6 26.6 546 65 188 24.4 75.6 CONTINUED 266 APPENDIX 8 (Continued) Temp. G f , 25 °C G f ,T°C Uo, T k* °c 1t/mi n 1t/mi n cm/sec gr/min 520 64 177 23.2 54.7 480 64 171 22.4 53.8 470 72 189 24.8 48.8 490 48 125 16.7 47.1 490 60 154 20,2 56.0 490 52 134 17.6 50.8 490 72 193 25.3 84.1 360 62 154 20.2 19.5 360 57 142 18.7 19.3 340 72 171 22.7 33.4 330 83 192 25.3 42.5 550 63 179 23.7 56.7 555 65 187 24.7 65.0 555 66 193 25.3 59.8 C) Kennecott Calcines 570 63 184 24.2 39.8 550 72 2Q8 27.5 61 .0 525 77 208 27.5 74.0 520 82 225 29.6 73.2 544 63 180 23.6 49.6 550 65 190 24.9 56.1 550 72 205 26.8 68.1 550 83 235 31 .0 100.0 544 65 188 24.6 67.1 D) Endako Calcines 235 64 109 14.4 26.5 300 64 116 15.4 35.2 350 64 123 16.1 28.0 400 64 145 19.1 56.0 425 64 161 21 .2 77.4 450 67 173 22.7 50.3 495 64 173 22.7 89.8 503 64 174 22.9 44.7 410 67 154 20.1 44.8 430 62 147 19.4 63.3 470 64 160 21 .1 113.4 51 5 64 175 23.0 105.8 530 64 179 23.6 120.4 267 APPENDIX 9 • • RECORDED SIGNAL FROM THE INFRARED ANALYSER TO A S0 2 PULSE INPUT 268 9 9 1 2 M - S ' ° N IHVHJ KOrmoHiimii IT - r> i o T 7 =. 7 7 / / / . - . . . ^ P J T C t ^ r p ? / / A * - * j.r/c9ic:: U .L- : - 37 << : --0S--08:-- - 1 4 (Mi*-™ - r h mis B 0 -269 70/ £ 1 trzr - 0 C — H -0f -• :r -rp r -h -1 uniiti -06- - r / H -' : > ' :• ' • ' trzr::t.rr !' - • • • • • ^ 4 R = ; ^ - t-^i; •^Ift #r 7^7 — en '&?<>•:&#.: fit m m 7v; -—-Of-1332! j r r ' 4o— 70- —i r-to g-270 APPENDIX 10 COMPUTER PROGRAM FOR CALCULATING THE RDN OF FLUID BED REACTOR AND GAS TRANSFER COEFFICIENTS c i M I C H I G A N TF.F MI MAL S Y S T E M UF ' TRAN 0 ( 4 1 3 3 6 ) ' M A I N 1 0 r l 6 - 7 4 11 2 6 : 3 3 RAGE PU C C S Y M P C L S U S E D I N CO MPUT AT I ON 6 . 0 0 0 7 . 0 0 0 7 . 0 0 0 7 . 0 0 0 V c — ? c C c A A L P H AC V P I C A G E C O E F F I C I E N T CF C l . C U C - W A K E AND F M L L S I LIN A B S O R P T I O N C O E F F I C I E N T CF THE C A S I N THE S O L I D P . 0 0 0 9 . 0 0 0 1 0 . 0 0 0 C C c A DCA ADI S ADS AV f R A G E S l / F . D I A M E T E R C F C A L C I N E S P A R T I C L E S , CM A V E R A G E S U E D I A M E T E R CF S C L I D P A R T I C L E S IN R E A C T O R , CM A V E R A G E D E N S I T Y UF S O L I D I N R E A C T O R , G R / C U . C M 1 1. 0 0 0 1 2 . 0 0 0 1 3 . 0 0 0 C c c. AOS I ADT A F F AV FE AGE S U E D I A M E T E R OF S I L I C A SAND P A R T I C L E S , CM D I M t f T I O N L F S S T I M E E X P A N S I O N F A C T O R CF F L U C U E C BED R E S P E C T TO THE S T A T I C B E D 1 4 . 0 0 0 1 5 . 0 0 0 1 6 . 0 0 0 V c c A F I ALM V t . L U M E T R I C F A C T O R C F I N C R E A S E CF BED R E S P E C T OF S I L I C A SAND BED S T A T IC BED HC IGHT , CM 1 7 . 0 0 0 • 1 8 . 0 0 0 1 9 . 0 0 0 c c . . c A K C E A K P C MASS T R A N S F E R C O E F F I C I E N T OF GAS BETWEEN C L O U D AND E M U L S I O N , C M / S E C MASS T R A N S F E R C O E F F I C I E N T OF C A S BETWEEN B U B f i L F AND C L C U D , 2 0 . 0 0 0 2 1 . 0 0 0 2 2 . 0 0 0 c c c A K B E C M / S E C MASS T R A N S F E R C O E F F I C I E N T OF G A S B E T W E E N B U B B L E AND E M U L S I C N , C M / S EC 2 3 . 0 0 0 2 4 . 0 0 0 2 5 . 0 0 0 c. c c A K E E B A K C E B O V E R A L L C O E F F I C I E N T OF GAS I N T E R C H A N G E B E T W E E N GAS B U B B L E S AND E M U L S I O N , I / S E C ' C O E F F I C I E N T C F GAS I N T E R C H A N G E BETWEEN C L O U D AND E M U L S I O N , 2 6 . 0 0 0 2 7 . 0 0 0 2 8 . 0 0 0 c c r AKBC E I / S E C C O E F F I C I E N T OF GAS I N T E R C H A N G E B E T W E E N G A S B U B E L E S AND rt n u n . i / S F C 2 9 . 0 0 0 3 0 . 0 0 0 3 1 . 0 0 0 c c c A K h E S AL A L O SC.Lt ! ) I N T E R C H A N G E B E U E E N G A S B U B B L E S ANC W A K E . l / S E C F L U I C I Z E D F E D H E I G H T , CM F E E D I N G P O I N T , CM 3 2 . 0 0 0 3 3 . 0 0 0 2 4 . 0 0 0 c c c AL F A L P H A A L S TOTAL H E I G H T GF F L U I D I Z E D B E D , CM F R A C T I O N OF WAKE IN GAS B U B B L E S S T A T I C BED H E I G H T CF S I L I C A S A N D , CM 3 5 . 0 0 0 3 6 . 0 0 0 3 7 . 0 0 0 c c c ANH APDCA A P D S I N C " 3 F - K R O L E S PER U N I T A R E A IN C A S D I S T R I B U T O R , I / S O . C M A P P A P E N T D E N S I T Y OF C A L C I N E S , G R / C U . C M A P P A F E N T D E N S I T Y CF S I L I C A S A N D , G R / C U . C M 3 8 . 0 0 0 . 3 9 . 0 0 0 4 0 . 0 0 0 c c c AT". BN A V E R A G E R E S I D E N C E T I M E OF C A L C I N E S IN B E D , M I N T O T A L NUMBER CF B U B B L E S AT L E V E L L IN THE F L U I D I Z E B E D , CU.C .M 4 1 . 0 0 0 4 2 . 0 0 0 • 4 3 . 0 0 0 c c c BNT C G NUMBER OF B U B B L E S P A S S I N G A G I V E N H E I G H T O F ' R E A C T O R S , I . / S E C C O N C E N T R A T I O N . OF R E A C T I N G G A S , G R / C U . C M 4 4 . 0 0 0 4 5 . 0 0 0 4 6 . 0 0 0 c c c • CGB CGE C GM C C N C E N T RAT ION OF R E A C T IN G GAS I N S I D E B U B R L E S .MOL . G R / C U . C M C O N C E N T R A T I O N OF R FACT I N G C A S IN E M U L S 1 0 N , M O L . G R / C U . C M • C O N C E N T R A T I O N OF R EAC T I N G G A S , M C L . G R / C U . C M 4 7 . 0 0 0 4 8 . 0 0 0 4 9 . 0 0 0 c •• c c CK B C P " 0 S CPMOX C O E F F I C I E N T GF R I S I N G V E L O C I T Y OF B U B B L E S , ( 0 . 9 0 ) HE AT C A P A C I T Y OF M G S 2 . C A L / N C L - D F G . K HEAT C A P A C I T Y OF M 0 0 3 , C A L / M O L - D E G . K 5 0 . 0 0 0 5 1 . 0 0 0 5 2 . 0 0 0 c c c C S 9 CS EH C O N C E N T R A T I O N OF R E A C T I N G S O L I D IN GAS P H A S E AT ANY P O I N T IN F L U I D I Z E D BED C D N C F N T R A T I O N O F R E A C T I N G S C L I D IN GAS P H A S E AT F E E D I N G 5 3 . 0 0 0 5 4 . 0 0 0 5 5 . 0 0 0 • c c c C S B E P O I N T , AS F A C T I O N OF T O T A L C A L C I N E S IN GAS P H A S E AT T H I S P C I N T C O N C E N T R A T I O N OF R E A C T I N G SHL. ID IN GAS P H A S E AT TOP OF B E D 5 6 . 0 0 0 5 7 . 0 0 0 5 8 . 0 0 0 ^ 271 MICHIGAN TERMINAL SYSTEM ' -OF TP. AN G( 4 12 2C) CSE CSEE csro CCNCENTPAT ION OF REACTING SOLID PCINT IN FLU 1CIZED BEC CONCENTRATION OF REACTING SOLID CCNCENTPATION OF REACTING SOLID MAI N 10-16-74 IN EMULSION PHASE AT ANY 11:26:31- PAGE P002 E MIL SI ON EMULSIUN PHASF PHASE TOP BED FEEDING 59. 000 60. 000 61.000 62. OOJ "C PCn>rTT~AS"F R A C T 1 ON 0~F C A L C 1 N E S AT THIS PCINT C C T G CONL ENTRAT ION CF TRACE0 AT M EA SLR I NG POINT, CU.CM C TRACER/LT GAS _ ~t CT5 tm^TT ' R T r T l W T j F SOLID TpACEft -'AT SAHPTlNd POINT, K T . K E R CfcNT C DAG AXIAL DISPERSION COEFFICIENT OF GAS IN FLUIDIZED BED, C SCO/SEC bl,000 64.000 65. 000 ~C —DAS A X U U D I S P C P " S 1C ' N I COEFFICIENT OF SCLIDS I N iHt b. t U , SU . CM/ b te. C DB DIAMETER OF G A S BUBBLE,- EQUATORIAL, CM c DBO ! iini_ L__ B LL?iL L- R C-L^—-Tl9 Hf0*_ CJ* ~C" DC- ' OIAMETTk " O F BUBBLE"' S~" CL'Cub , ""EO U A T pp. I A L , CM C D C A DENSITY C F CALCINES, CR/CL.CM C D E F F X I T OR FES1LTNCE TIME. DISTRIBUTION OF TRACER 66.000 67. 000 68. OOP 69.000 70. 000 71. 000 72.000 73. 000 74. 000 LTA F ACT ION 0 F BEL) CONSISTING OF CAS BUBBLES DH NUMBER CF HOt.FS IN GAS DISTRIBUTOR DIFF MOLECULAR D I_F F US I UN C CE F F I C I ENT CF GAS, SO.CM/S EC ~CTF"FT"E F F £ CT IV ED I FPUS ION CO E F FI CIE N T ' OF GAS TN THE FLUIDIZED FEDi SC.CM/SFC DMMOS MOLAR DENSITY 75.000 76. OOJ 77.000 OF MOLYBDENUM CI SUL P H I DE . GR-MO L/CU .CM 78.000 7?. 000 8 0.000 ~D>MOX MOLAR DENSITY CF MOLYBDENUM TR 10 X ID E , GR - MOL/CU . C M M i n t n p « ! T v r F MriYRDFMIM r.I SULP FI DE . GR/CU .CM DMOX DENSITY OF MOLYBDENUM TR I r X I DE ,GR /C U. C M 81. 00) 82.000 C ORG RADIAL DfSPERSIONXOEFF1CIENT OF GAS C SO.CM/SEC C DCS RACIAL DISPERSION COFFFICIFNT CF SOLIDS IN FLUIDIZED BED, IN THE BED,SO.CH/SEC 84. 000 85.000 86.000 T~~ r~DS"' AVERAGE SIZE DIAMETER CF M:CS 2 PARTICLES IN FEED, CM • r C DSI DENSITY OF SILICA SAND, GP/CU.CM C EA VCIDAGE AT CLOUD AMD kAKE OF GAS BUBBLES, CONSTANT C EF VOIDAGE CF E MULS I ON AT F LUT6 IZ ING "COND IT JUNS C EFAM EXPANSION FACTOR OF BED MINIMUM FLU1DIZING CONDITIONS, C CONSTANT FOR 10 TO 60 WT .PERCENT OF CALCINES A _ _ _ 87. 000 88.000 8'9. 000 90. 000 91.000 92.000 C. E MF VCID/GE OF BEC AT MINIMUM FL UI C I Z \ NG CONC IT I QNS C EMI EMISSIVITY FACTOP. OF MCS2 PARTICLES, 10.E) EO VOJ_DAGE__OF BEC AT STATIC CONDITIONS 93. 000 94.000 95. 000 C EOCA STATIC VOIDAGE OF C EOSI STAT IC VO I DA GE OF. C ER ELUTRI AT ICN RATE, CALCINES BED SILICA SAND BED GR/SEC 96.000 97.000 98. 000 ~ C " F R K ELUTRIATION RATE COEFFICIENT, GP/SO.CM*SEC C FO F F EO RATE CF . MOLYBDENITE CONCENTRATE, GR./MIN C FOMD FEED RATE CF MOLY B DEN IT F CCN C ENT RATE PER UM IT AREA OF 99.000 100. 000 101. 000 C REACTOR., TfJN/SU.MT.»DAY C FV FLUIDIZATION FACTOR CF VOICAOE, CONSTANT FOR 10 TO 60 C WT.PERCENT OF CALCINES 102.000 103.000 104.000 ~J G ACCELERATION CF GRAVITY, CM/SO.SEC ~ C GF GAS FLOW RATE AT 25 CEG.C, LT/MIN C GF T GAS FLOW RATE AT WORKING TEMPERATURE, C U . CM/ S E C 105.000 106. 000 107.000 C GK THERMAL CONDUCTIVITY OF AIR C; GL TOTAL GAS AT LEv EL L IN THE C GM TOTAL GAS'FLPV, RATE AT 25 01 . ~< 4.32*10**-£l ,CAL7CM*2*SEC + CEGC 108.000 FLUIDIZED EEC, BU.CM 109. 000 ;G.C FOR ThE SIMULATED FLUIDIZED 110.000 BFCV F EACT03. CU.MT/M IN GMF MINIMUM FLUICIZING GAS FLCW RATE, CU.CM/SEC CP GAS FLOW RATE OF R EC IRC U L A T I CN INSIDE BUBBLES, CU.CM/SEC 111. 000 112.000 113.000 272 S Y S T E M F O R T R A N G I 4 1 3 2 6 ) R A I N 1 0 - 1 6 - 7 4 1 1 : 2 6 : 3 3 P A G E P 0 0 3 c GT T O T A L G A S I N R E D AT A N Y T I M E , CU.CM. 1 1 4 . 0 0 0 : c HR G H E I G H T O F INITIAL B U R R L E S F O R M A T I O N , CM 1 1 5 . 0 0 0 i c H E A T H E A T O F R F A C T I U N F OR CX1 C A T ICN•CF M C S 2 TC M 0 0 3 , C A L / M O L . G R 1 1 6 . 0 0 0 ! C H P HE A T T R A N S F E R C O E F F I C I E N T G A S - P A R T I C L E , C 4 L / S O . C M * S E C * D E C . C 1 1 7 . 0 0 0 • C P T L T A L P R E S S U R E , A T M . ' l i e . ooo 1 c P G P A R T I A L P R E S S L R E O F R E A C T I N G G A S , A T M 1 1 9 . 0 0 0 c P N P A R T I A L P R E S S U R E OF N I T R O G E N , A T M 1 2 0 . 0 0 0 c PO R AR I 1 AL PRtSSUPR Of- C X I GEN, AT M 1 2 1 . 0 0 0 c PR P R A N D L M U M P E R 1 2 2 . 0 0 0 r P S O P A R T I A L P R E S S U R E O F S U L P H U R C I O X I D E , A T M 12 3 . JOO c P T K P A R T I C L E T R A N S F E R C O E F F I C I E N T BEJWFFN S O L I D IN B U B B L E S A N D 1 2 4 . 0 0 0 c C L O U D A N D W A K E B A S E D I N U N I T A P E A O F B U B B L E S U R F A C E , 1 2 5 . 0 0 0 c C R / S C . C M * S E C 1 2 6 . 0 0 0 c O B V l . L U M E T R I C G A S F L U X P E R U N I T A P E G F B E D F E P . U N I T T I M E , 1 2 7 . 0 0 0 c C U . C M / S O . C M * S E C 1 2 8 . 0 0 0 r C S SOLID F L U X P E R U N I T A R E A C F B E C PER. U N I T T I M F , GR / SC . CM * S E C 1 2 9 . 0 0 0 C P. GAS C O N S T A N T 1 3 0 . 0 0 0 c RD R E A C T O R riAMETEP, CM ,MT 1 3 1 . 0 0 0 c R DN P E A C TOR D I S P E R S I O N N U M B E R 1 3 2 . 0 0 0 c R E G F E Y N U L O S N U M B E R F O R G A S 1 3 3 . 0 0 0 c R E P R F Y N O L C S N U M B E R F U R S C L I C S 1 3 4 . 0 0 0 c P.R R E C I R C U L A T I O N R A T I O 1 3 5 . 0 0 0 c R.S R E A C T O R C R O S S S E C T I O N , SO.CM 1 3 6 . JOO c R T D E X I T A G E D I S T R I B U T I O N FUf> C T I C N CF G A S , C I M E N T I O N L E S S 1 3 7 . 0 0 0 c Sb S U R F A C E O F B U R B L E , SC.CM 1 3 8 . 0 0 0 c S B K SI E F A N - B O L T Z M ANN C O N S T A N T . ( 1 . 3 5 6 * 1 0 * * - 1 2 . C A L/C M 2*SE C * K * * 4 ) 1 3 9 . 0 0 0 c S B N N U M B E R OF B U R B L E S P E R U N I T A R E A C F B E D . l / S C . C M 1 4 0 . 0 0 0 c S C A L S U L P H U R C O N T E N T O F D I S C H A R G E D C A L C I N E S F R CM T H E R E A C T O F 1 4 1.OOJ c SGK V O L U M E O F S C L I ' J P E R UNIT V O L UR L O F G A S INSIDE C A S B U R P L E S 1 4 2 . O O J c SH S H E R W O O D N U M B E R 1 4 3.JOO c SK R A T E C O N S T A N T F D R C H E M I C A L » E A C T I O N B A S E D ON U N I T AR F A O F 1 4 4 . 0 0 0 c S C L I D , C M / S E C 1 4 5 . 0 0 0 c S K S E L U T R I A T I O N R A T E B A S E D I N U N I T V C L U M E OF G A S AT 1 4 6 . 0 0 3 c R E C I R C U L A T I O N S U P E R F I C I A L G A S V E L O C I T Y , C O N S T A N T , G P / C L . C M 1 4 7 . 0 0 0 c S 0 2 REP C E N T S U L P H U R U I O X I L ' E I N C F F G A S E S 1 4 8 . 0 0 0 c S P S P E C I F I C S U R F A C E A R E A O F M O L Y B D E N I T E C O N C E N 1 R A T E , S 0 . C M,/ G R 1 4 9 . 0 0 0 c S P F S P H E R I C I T Y F A C T O R O F S O L I D 1 5 0 . 0 0 0 c SR S L U R R Y R E C Y C L E T C R E A C T O R AS 5 0 5 S C L I OS.KG/MIN 1 5 1 . 0 0 0 c SS S P E C I F I C S U R F A C E A R E A C F M O L Y R D - N I T E C O N C E N T R A T E , SO. C M / M C L E 1 5 2 . 0 0 0 c T E T E M P E R A T U R E O F E M U L S I O N P H A S E I N F L U I C I Z F G B E D , D E G . C 1 5 3 . 0 0 0 c T F 3 MFAN T E M P F R A T L R E O T THE F L U I C I Z E C B E D , D E G . C 1 5 4 . 0 0 0 c T G T E M P E R A T U R E OF T H E GAS IN F L U I D I Z E D B E D , O E G . C 1 5 5 . 0 0 0 c T H T I M E , S E C 1 5 6 . 0 0 0 c TN H T O T A L N U M B E R C F BUBrtLES IN F L U I D I Z E D B E D AT ANY T I M E 1 5 7 . 0 0 0 c 7 R S T E M P E R A T U R E AT THE S U R F A C E O F T H E R E A C T I N G M 0 S 2 , D E G . C 1 5 8 . 0 0 0 c T S MEAN T E M P E R A T U R E O F S O L I D IN F L U I D I Z E D B E D , D E G.C 1 5 9 . 0 0 0 c T K T E M P E R A T U R E OF- THE RE'.CTCR W A L L S , D E C . C 1 6 0 . J O O c UB B U B B L E R I S I N G V E L O C I T Y , C M / S E C 1 6 1 . 0 0 0 r U P R PURREE V E L O C I T Y R E S P E C T T C E M U L S I O N , C M / S E C 1 6 2 . 0 0 0 c U F E M U L S I O N ' V E L O C I T Y , CM/SEC 1 6 3 . 0 0 0 c U F S U P E R F I C I A L GAS V E L O C I T Y T H R O U G H E M U L S I O N R E L A T I V E T O T H E 1 6 4 . 0 0 0 c SOLID I N THE t - H J L S I C N , C M / S E C 1 6 5 . 0 0 0 c U"F S U P E R F I C I A L CAS V E L O C I T Y REFFREu T O T H E E M P T Y R E A C T O R A T 1 6 6 . 0 0 0 c M I N I M U M F L U I D I Z I N G C C N C I T I C N S , C M / S E C 1 6 7 . 0 0 0 c UN M U S S E L T N U M B E R 168 .OOO 273 ' IC HI GAN T E RMIrAL SYSTEM FOPTFJN GI41336) MAIN "10-16-74 UR u r VAC S U P E R F I C I A L G A S V E L O C I T Y F E F F P . P E D TO T H E EMPTY R F A C T O R AT WOPK I N G T E M P E R A T U R E C M / S F C A V E R A G E S U P E R F I C I A L GAS V E L O C I T Y O F R E C I R C U L A T I O N I N S I D E G A S F M J B B L E S , C M / S E C ""TTHT SCL l O T EMULSION RELATIVE TO THE VB VC VK L I NEAR V E L O C I T Y "OF" F M U L S I O N , CM/SEC V A P J A N C E O F D I S T R I B U T I O N , SE C . **2 , MI N. **2 'VCluFE 0-" B U ? OLE",* CU.CM-""' '"' " — 1  V O L U M E OF H U B B L E 1 S C L C U D , C U . C M P A T E C O N S T A N T FOR CHFM IC AL R E A C T I O N B A S E O IN U N I T VOLUME OF 11:26:2; 169, 170. 171. 172. PAGE P004 0 0 0 0 0 0 0 0 0 0 0 0 T7T 174 175 " 5 o a r 0 0 0 0 0 0 C SOLID, MCL/CU.CM.*SEC C VMAIP MOLAR VOLUME CF AIR, (APPARENT), CC/GR-MCL _C YJ?f'__ MQ LAP VOLUME CF NITRCGEN, CC /GP.-MOL C VKCX" "MOLAR VOLUME "CF 6YlCTt'NTTr/GP-"Mclt C V«SO Mf LAP VOLUME CF SULPHUR DIOXIDE, CC/GR-MCL C ' VP VOLUME OF A SINGLE PARTICL E, CU.CM "T76". 177, 178. 0 0 0 0 0 0 0 0 0 179. 180. 181. 0 0 0 0 0 0 0 0 0 VS VT VOLUME OF SCLID PER UNIT MASS O F PART I CL E S , VTB VW SPECIFIC CU .CM/GR TO TAL VOL UM E CF_ FLUlrlZFC PFC, CU. t C TA l" B "0" t'B L CSV OTU W I ( C T C U ID I Z EC" BED, CU.CM"" TOTAL CLCUD VOLUME IN FLUIDIZED 3E0, CU.CM VOLUME OF BUBBLE WAKE, CU.CM 182. 183. 184. 0 0 0 0 0 0 0 0 0 .CM 105. 186. 18 7. 0 0 0 0 0 0 0 0 0 C V T O T A L WEIGHT Cr SOLID IN "THE FLUIDIZED" B E D , GR, KG~" C WCA W E I G H T O F CALCINES IN FED, GR J J^'''L F_!:"-' !i: E' c ut Ai: h^i?tl T_- 0 fL_ /L Ii''_ ( A P P A R E N T ) , G R / ' H O L ' C W-'"OS MOLECULAR"WE I GUT OF M 6 L V R~D E N L M DI MJOMTT b~E , OPT""" 0L"~ C WMMO MOLECULAR WEIGHT O F MOLYBDENUM, GR/MOL _C WM*"TX MOLECULAR WEIGHT CF MCLYP.nRKLM. TRICXIDE, GR/MOL 188, 18 9. 190. 191, 192, 193. v. !•• r x WMSO _WF VP V WS I _XCA X LL XP XS I Mi. L E'CUL AF MOLECULAR WEIGHT OF F R A C 1 1UN OF WEIGHT FRACTICN OF WE IGHT FRACTION OF IGHT OF UX IGLN-" GR TFol '. ' WEIGHT OF SULPHUR CICX ICE, GR/MQL :'i_SLfiGJ-I__?^ T!.£-l!:_' GR SPFCfFIC WEIGHT OF SOLID PER*-UNI T VOITJM E~P A R TIC L E S, GR/CU.TM" WMCHT OF SILICA SA.NC IN BED, GR IGHT REACTION OF CALC I N F S IN BED CLOUDS AT LEVEL L FN THE FLUIDIZED SCLID TRANSFORMED SILICA SAND IN BED 194 195. 196, 197, 198, 199, 0 0 0 0 0 0 OOP 0 0 0 0 0 0 0 0 0 _ 0 *00 0 0 0 0 0 0 OOP 0 0 0 0 0 0 BED 200. 201. 202. "203. 204. 205. 206. 000 OOP OQi) POO 000 00 0_ OOP {UN «FTN .SPl,'NCh- = C C N P P f i F I L C "~."UT ICN""acCIMS i •c 274 C R f SPO NS f c TESTS 'Of- GAS TRACER INPUTS IN FIVE INCHES REACTOR \\7 AND G A S A N D ' S n t i n O I S P R K S I ON_ COEFFICIENT CALCULATIONS IN THE BCD 2 3 8 C 24 0 _Q I.Ls.L.P.FRFORMf 0 IN WURKING CONDITIONS^  TYPICAL, OF HIGH CON VER T I ON 24 1 C LEVELS OF MOLYBDENITE TI I 7-ICL I b OENUN T R I>JX I UV ~' ' " ~2A~ C MOLYBDENITE CONCENTRATE FROM "BRENDA MINES".-32b MF SH PAR TI C LE S. 2^ 3 24 4 C ' — , — _ ._ • ; J^-246 C TEMPERATURE CF FLUIDIZED BED,510 TO 520 CEG C 247 C TRACER SUIPHUR DIOXIDE 248 C INPUT TIME FIVE SECONDS . . . 249 P VOLUMF. TRACER 16.5 LITERS 250 C 252 INTEGER TI . • 253 ? • 01 MENS I CN TM (20) ,CTG( 20) ,ADT( 18) ,OEF( 20) ,RTD( 18) ,AT( ICOi 1 254" 3 REALH5.10) TIME,EFAM,ALPHA,ALS,RS,G,GF,UMF,WCA,WSI,ERK,AFI,AFE,DSI 255 1 .DCA.ADCA.AOSI .APOSI , APP CA , F V , DH , WM SC , WM OX , VM SO , VMCIX , VMA I P , RD , CK R , 2*6 2HSP0.P.AC ! — 25T" 4 R E AD ( 5 j 20) (CTG( Jl ,J = 1,18) 258 _5 PEAD(5,20) ( TM( J), J=l. 18 ) ? 5 g 6 10 f 0RMAT(8F":i0.4) ~ ~ ~ : |g-g-7 20 FORMAT!RF10.2) 2 6 j C CALCULATION CF GAS RESIDENCE TIME DISTRIBUTION FUNCTION IRTD) 262 v. - . . 2 6 3 8 SUMC = 0 2 b l > 9 • SUMTC-0 ' 265 10 DO 40 J=l,l( 266 11 SUMC=SUMC+CTG!J) 267 12 40 SUMTC = SUMTC-«C.TG( J.)*TMU ) 26 B 13 A*TIME*SUMC "" "" ""' " 14 TMEAN=SUMTC/SUMC 269 2 7 0 _J5 DO 5 0 J = l , 1 8 . 2 7 1 16 A O T ! J ) = T M ( J ) / T M F A N 2 7 2 17 D E F l J i = C T G ( J ) / A 2 7 3 18 5 0 R T O I J ) = T M E A N * D E F t J ) . ' 274 1 9 . WP IT El 6, 6 5 ) ^ ; 2 7 5 " 20 65 '. F O R M A T ( I X , 1 D I M E N T I O N L E S S T I M E 1 / ) 2 7 6 21 W R I T E 1 6 . 6 0 ) ( A P T ! J ) , J - 1 , 1 8 ) 2 7 7 2 2 6 0 FORMAT ( I X , 1 8 F 6 . 3 ) 2.78 23 > I P I T E ( 6 , 6 6 ) „ 2 7 9 2 4 6 6 F O R M A T ! I X , ' F X I T A G E DI S TP.. F UNC T. ' / I -r 2 8 0 25 W P I T F ( 6 , 7 ' O ) r k T [ ) i j i , J = I , la) ; ~ : ; " ; IST~ 26 70 FORMAT!IX,18F6.3) 282 Q CALCULATION OF REACTCR DISPERSION NUMBER OF GAS I RDN ) 283 C T " " — " " R- 284" 27 R-0 . 2 8 5 2 8 0 = 0 • 2 8 6 2 9 ' DO 800 J = l , 1 8 ' ' ~i : ! ! 287" 30 0=0+ACT(J)**2.*RTD(J) 2 8 3 31 • 8 0 0 F = R + R T 0 ! J ) 32 ' VAC=0/R-1. 289 290 C. ITERATION LOOP FOR CALCULATION THE DISPERSION NUMBER 291 33 R DN= VAD/ 2. ' 2 9 2 34 1 DO 170 M = l,100 ~ ' 1 1 i ~ — 29T" 35 VA0D = 2.*RDN-2.*RDN«*2.*(1.-EXP(-1./RCN)) 294 36 IF ( (_V AO- VA D D K L F.J.I) GO TO 8 0 29 5 37 170 CONTINUE " ~ 1 ~ ' ~~296~ 8 V.PITEI6.75) 3? 75 FORMAT(IX, ' Q F A C T O R . D I S P • N O . ' / ) 40 80 WRITE(6,90) RDN '41 90 F0RMATIF6.4) 42 WP IT EJ 6 , 200 I 43 wVi'fEffiVfi'sr 2 9 7 2 9 8 29 9 300 301 302 C CALCULATION OF THE FLUIDIZING PROPERTIES OF THE BED 303 C 30-H 275 C C C A l C U L A H O N S OF T H E F L U I D I Z I N G C C N D I T I C N S OF T i iT HT fj ARE 1 ' ' o^TT C MADE A S S U M I N G AND I D E A L R E H A V O I R FOR G A S E S A T THF W O R K I N G C O N D I 306 C T I C N S OF P R E S S U R E AND T E M P E R A T U R E O F T H E RED. 307 44 G F T = G F * ( 550.+?73. )/27?.*( 1000./60.) 308 45 ALMf=1.01"ALS*AFI 3 0 9 46 A L F = A L M F * A F E • 3 1 0 4 7 XCA= 74 UF=UM'F/ETMF ' C 84 DIFFE-DIFF»EMF P / UH-0 .00 2b?«- AL" ( (UU-UNF)/ANH) 3 3 7 68 UB = GfJ-UMF«-0. 711*IG*OR)**Of5 '. > 1 : syg-69 DELTA=(U0-UMr)/U6 3 3 g 70 UE = UMF/EMF-1 ALPHA».UO/ I 1. -CELT A-ALPHA*CELTA >r-ALPHA»UMF ) • 5 A 0 71 EF = UE /R S 72 VP=((4./3. )*3.14*(UB/2.1**3,)*(1.-ALPHA) 2u2 J_3 _UBR = 0. 7U*(G*PB)**'0.5 ... . 3 ^ 3 344 13 bL = bH/ISUU,»Ulll*RS , n 76 BN=Gl/VP 346 C VOLUMETR'IC GAS FLOW ' R'ATE 'OF RECIRCULATION INSIDE BUBBLES WAS CAICU " 1 ' 3vT C LATED EY MEANS CF THE EXPRESSION CF CAVICSON AND HARRI SON -348 34 9 77 GR=3.*UF*EMF*3.14*(DB/2. ) * * 2 T '• 1 3 5 0 78 SB = 4. *3. 14*<DP/2. ) * * 2 . 3 5 1  79 VC_M0.6/D8**0.5 )*V9 • 3 5 2 80 US = ALPHA*DE.LTA*UB/ I 1 . - D E LT A- AL PH A* C EL T A ) 1 : : : ~ : J 5 3 -81 DC-I 6.*VC/3. 141**0. 333 3 5 4 _82 X C L = 2 . / 3 . * T B / U P * ( 8 N « U B / A L »_*_( DC/T) B ) *_*3_. /J_0 0_. 3 5 5 C DIFFUSION C C E F F n c I E N t OF' GAS" C AL CULAT E D EY MEANS CF THE RFLATIONSH : ^ K T C IP OF ANCRUSSOW \ 3 5 7 C r- - r : : ., ' 3 5 8 83 OIFEM 0. 06G6*( 2 73,+ 550. I *»1. 78* (T.+ I WMS 0 + W MCX 1**0 . 5 ) ) / ( R* (VMSO **0 . : '• 35~9 1333+VMAIR**0.333)**2.*(WMOX*WMSO1**0.5) 3 6 0 36 1 3 f e 2 CALCULATION OF GAS I NT PR C HAN GE COEFFICIENTS BETWEEN GAS BUBBLE, 363 CLOUDS AND EMULSION 3 6 4 48 XSI = WSI/ IW'CA + WSI ) i l I 312 313 49 ECSI = l.-APD SI/OSI 50 51 EpCA-J.-APDCA/DCA -E0 = EDS I*XSI + FCCA*XCA 314 31 5 316 52 EMF=FV*'-n 53 EA-1.20*EMF . . . . . . 31 I 31 3 319 54 C AALPH = 11 . -E A) / (1 .-t.MF ) * ( EME/ EA )**2. THE AVERAGE RECIRCULATION VELOCITY PF GAS INSIDE BUtiKLES WAS CAL CU C C LATED USING THE EXPRESSION OF LEUNG AMD SANDRORD 320 55 UR= IUMr*EA)/( AAIPH*FMF ) 321 32 2 56 MO-ALF/0.5 1 1 - 323 324 325 57 TI=TIO. 58 UO=GFT/R$ 59 ANHJ-CH/RS 326 32 7 32 8 60 00 100 K=1,TI 61 Y = FLOAT I K ) 62 SUUM--0 ' — . '32 9 330 . 33 1 63 . AL = SUUM + 0. 5*Y 64 AT (K>AL/0.5 65 AOSj-DCA^ XCA + IVS I*XS I 332 333 334 66 ADIS-APCA*XCA + ADS! *XSI C BUBBLE DIAMETER CALCULATOR USING THE EXPRFSSION DEVEIOP-'D I N THI S C c PRESENT WORK FROM TWO-DIMENSIONAL EPOS EXPERIMENTS ' ' 335 276 f " 85 86 87 C AKBCR = 4. 5*UMr/DB +5.855-1 0 IFFF**0.5 )*< G«*0.25 I /{ Dfi»«• 1. 251 AKCF E = 6 . 78*! FM.F*D1 FFE*Ufi/CB**3 .0 1**0 .5 AKPFfi=l./tl./AKBC8*l./AKCEfl) CA LC ULATI ON OF MASS TRANSFER COEFFICIENTS FOR GAS IN FLUIOIZEC BED 365 36 6 36 7 368 369 370 V 88 AKBC=0.975*DI FF**0 .5 » < C,/DP ) **0 . 25 r 89 . 90 . AKCE^l. 13*( (DIFFE'EMF* (UBR-UF) I / CB )*<!0 .5 * < ll.-UF/UB)/ (1. +2 . *U F/U b 1 1 >**0.1666 AKPE=1,/(1./AKBC+1./AKCE) 371 372 373 1 91 c c CALCULATION OF OISPEPSIUN COEFFICIENTS OF GAS ' DRG=0.2*DP**2.*AKBEB/DELTA 374 375 376 92 93 94 200 e = AL,PHA*EMF-UHF*( 1.-D EL T A-AI.PHA*0 EL TA ) /( LO-UMF) +AC»ALPHA*(1.-EMF ) OAG=3.21*(UO*Ufi/AKBEB) FORMAT! 2X, « DI ST. • ,2X,' BHL.DI A. ' ,2X CCD. CI A. • ,2X, • NO.BPL .' , IX, • B-C 377 373 379 95 115 1TF..CP.', IX, 'C-E TR.C P.', IX, 'B-C TR . C F ' , 1 X , • B-C MTC •', 1 X * ' C-F MTC',1 2X , ' B-E MTC , IX,' AX.DISP.CCEF ' , IX, • P A C . D I SP . CO EF . • / / ) FORM AT(2 X,1 SOL ID INTEPCM.U-WAKE' ,5X,'SOL ID AXIAL DISP.COEFF.•,5X, 380 381 362 96 97 300 2'SOLIO RADIAL 0 ISP.COEFF.•/) WPITC(6 ,300) FORMAT! IX, 12IP 7. 4, 2X>/I 383 384 385 ' .99 99 500 C WF !T E 16,500) AL,DB,DC,3N,AKBCB,AKCEB,AKBEB,AKBC,A KGE,AKBE,DAG,DRG FORMAT! IX, 12F10.2) CALCULATION OF SOLID INTERCHANGE BFTWFCN THE BUBBLES AND THF WAKE 386 387 388 100 C c AKEBS= (3* ( 1 .-EMF ) *UKF )/ 1 ! 1.-DELTA )*EMF*DB) CALCULATION OF AXIAL AND RADIAL DISPFRSICN COEFFICIENTS FOR SO L IDS IN THF PCD US ING THE. KUNI I AND LEVENSPIEL MODEL 389 390 39 1 101 102 c DAS= ( ALRHA**2.*EMF*DB* ( UC-UMF 1**2.)/ 1 3 . * CE L T A*L/MF ) DP S =0 .1875* ( DELTA/ (T .-DELTA ) )*(UMF*OB/EMF). 392 393 394 103 104 116 ' KPIIFtt.Ufcl AKEBS.DAS.DRS FORMAT! ix, 3F10.4) 395 396 5 9 7 106 107 STOP ' ENO 39 8 399 APPENDIX 11 COMPUTER PROGRAM TO ESTIMATE THE TEMPERATURE REACTION PARTICLES OF MoS2 ' ~' J.CCMPILE C . CALCULATION CP THE T CMPE RA TUP.E AT THE REACTING SURFACE CP MGLYBP6-C NI TE PARTICLES Al THE INITIAL CHEMICAL REGIME OF TRANSFORMATION. G ^ ' '' ~ ^C SIMULATION CH ROASTING CCMilTIUNS IN AIR FOP PART ICLFS IN EMULSION C MO L Y CPE NI TE CONCENTRATE FROM "BROKCA MINES".-325 MESH PARTICLES. ; ' ' ' . ' C : — . . ' ' i . • C . ' • . .1 DIMENSION I 1 20) , SM 2C) ; 2 P F f D 1 5 , ; j ) AC, FGiUN . V 5. PO. EM, I, SBO.CF , A , S . ' '3 READ (5,20)NA 4 READ ( 5, 10) ( TI J) , J = l",NA ) 5 PE AD ( 5,10) (SK ( J I, J = l ,Na ) • 6 10 FORMAT ( 5F10. 5) 7 • . 20 FORMAT I1I1CI B 30 F C RM AT (B Fl 0. £ )  • 9 00 2C0 J = I ,NA ' ' n U R I T FI ft. 210 ) II 210 ECRM AT 15X,'TEMP.PEP.' ,5X,'TEMP.PART.SURF•• ,5X, ' H f c A I F t A l l . -1.HEATH'/) 12 C P M C S = ( 1 9 . 7 + 3 . 1 5 / l O . * * : . * ( ( T ( J ) » 2 7 3 . ) - 2 9 E . ) ) / 1 6 J . 0 9 13 HE ATP=2 S 75 90*4. 4 9* T < •)) - 7 . 86 2 10 •> <• ( - 3 . ) ) » I T I J ) ) *g2.-3 .91 9* 1110.)** 15. >MT < J I I T * ( - l . I 14 . ki=L*S • 15 GK=AC*FG  16 SBK=S60*CF 17 HP=GK*LN/DS 18 CGM=PO*ll./224GQ.)*(273./(T(J)+273.))  19 TPS = T UI+50.0 20 DO 300 M=l,50 21 TP A= [-EM I*ShK»[ 1 TR S + 273 . ) / 1 00. ) ** 4.+ ( HE AT R* SK ( J ) * CC N * C S ) / A S + 1E,P*T ( J) • EMI"SEK* ( 1 T ( J ) +2 73 . )/ 100. >««-4. )/ HP 22 IF ( (7RS-TPA). LE. 1.0) GG 70 310 23 T RS = T R A  24 300 CONTINUE '. ' 25 310 WPITE(6,320) T ( J ) ,TRS,HtAIR,CPMOS 26 320 FORMAT (4 F1 5 . 2 )  27 2C0 CUNHNLE 28 STOP 29 ENC APPENDIX 12 COMPUTER PROGRAM TO CALCULATE THE REACTOR SIZE AND PERFORMANCE 279 f tCCMP!LF "" . - - •• — - - . -C 6 C . 7 C FLUIDIZED BED REACTOR FOR V:,H VBDEN ITE ROASTINC 9 ^ c t ALCOLATTnN"'.TF~"R^ n C PRL-DICTICN OF SLLRHUR LEVELS IN DISCHARGED CALCINES AND SIMULATION 12 C CF_REACTCR PERFORMANCE • 13 C MOLYBDENITE CONCENTRATE: 15 C PARTICLE S I Z E - -325 MESH. 16 c - _ _ ^ _ _ _ _ _ _ • - i c ^ r j ? l r - H T O - • • — -j-jr-C 18 _C CYCLONE SYSTEM EFFICIENCY = 9 r . ? ; IP C " "SCRUBBERETFTCI E N'C'Y'"""=""" 9 5"'}I : ' ' ' 2"0"~ C SLIP ERF.GAS VELOC. AT ROASTING TEMPERATURE =25 CM/SEC 21 C FOASTING WITH AIR ? 2 1 DIMENSION T(20) ,FOl20),DS<20),ATR( 20) 24 _2____ Rt AC (5 ,_1 1 J )_XC A ,_X_S I_, DCA,DSI,AFE,EFAM,ERK 25 3 " R.-.ADI 5, 33.!) NN V NC i'MOTivM ' "" 2o 4 RtAD(3,221) < FC'( I ) , 1= 1,NN ) • ' o. • 27 '5 RE_A_0_(_5 ._220_)_l T ( J ) ,J=1,NC) 28 ~6 "Rt AD( 5", 220 ) (OST"K) ,"'K = 1 ,NDI \ — 29~" 7 RF AD(5 ,550 ) ( ATR (L l,L= 1,NM ) 30 J I 0 FO R MA JJ_7 F1 0. 2J_ ' ; 31 3 30 FORMAT! 4110V 1 ~ 32" 10 220 FORMAT (3F10.2 ) 33 11 221 FORMATI5FI0.2) '_ 34 12 550 FORMAT ( 6F10.2 ) ' ' ' 3"5"" 13 DO 200 1 = 1 , M.N 36 _ _ 14 WRITEI6.20) F0< n ' ; 37 I - - WRITE ( 6 , 40) ~ "' " . ' 33"" It DO 300 J=l , NC 39 I 17 WRITE I 6 , 25) T<J) 40 ! 19 WRITE(6,30) OS(K) • «2 20 DO 500 L = I , N M A3 2 1 w R . IT! (4, 3") I ir-(L)' ~ 44 22 20 FORMAT(5 X,'iFED RATE =',F8.2,IX,•GR.MIN•) 45 _ 23_ 25 FORMAT! 5XL • TE MP . = • ,F R. 2!_1XJL'DEG.C • ) . • 46 I ' " 2 4 " "30 " F_*MAT ("5"X,'"'MOi'2'^S f/E = " 'TFff . 2 7 1 X 7 *MIC P T H ' ' ZT" j . 2 5 35 • FORMAT <5X, • AV.RFS.T IME=' , FS.2 , IX, • MIN' 1 48 (~. : 26 40 FORMAT! 3X, "PER CENT S ' , 3 X , ' CH ARGE ' , 3 X , • R E AC T . DI AM . < ,3X , • FLU I D BED 49 " j " " " 1HF IGHT ' , 3X , • GAS FLOW RAT E 1, 3X, 'P ER "CENT 502* 73X7' RECTRCTRATT 0"* ,3X~J 50"" ' 2'SLURRY RECIRC.'/) 51 27 AD_S-DCA*XCA+DSI *XSI 52 "28 SC AL"= 1.3 5i (1 " +( EXP i-0 .1 *D"_"<K7*<""( TTj ) ) **0 . 765-TTj") +52C.) /ATRTTTJTI ' ~ "5T~ 1*(2./((T(J))**0.765-T(Jl+520.1-1,)) 54 29 W=ATF.(L)*FO( I )/ 1000. 55 C RD:AH = l : l " " " ' Sb~ 30 RD= ( <4.*F0f I ) *ATR <L )/( ADS*3. 14) 1**0.333)/100. 57 31 _ P.S = 3.J4* (R_D/2 . ) **2. • 58 32 " GM = '(64.*RS"/T"4";) " "~ , " 59' 33 GF = 64.*RS/0.0 14 60 34 ALF=AFE*EF A M * K l i 61 35 S02 = (FU( I)*0.3o5*22.4*100./64.TA<"GF^ ; ' 6_" 36 ER = ERK*RS*fcO.* 10000. 63 37 RR = EF/FO(I ) . 64 y 38 39 40 41 42 i 3 45 44 "*5 45 50 0 46 400 47"" 30 0 48 20 0 49 50 SR= 2 . * (0 .0 50*RR I 65~" FCM=FO(I )/60. 66 CSE0=(F0M*100.)/(FUMf ER/60.) • 67 WRITEI6.45) SCAL,w,RD,ALF,GM,S02,RR,SR 68 WRITE (6, 55) CSEC 69 FCRMAT13F15.3) 70 7 D » ATTR G"7"5 T 1 71~ C0NT1NUE 72 CONTINUE 73 "CUNT INUE " ; TCT CONTINUE 75 _STOP . 76 " ;•;.} ' 1 : '• 77" APPENDIX 13 EXPERIMENTAL DATA OF OPERATION FLUIDIZED BED REACTOR, 12.5 CM DIAMETER Molybdenite coriceniTQhe. source Run N -T °C t hr <V Uo /sec I- -\0 /onset u* 97. /TIKI Fo 3 7 . /min IA Fo mice P o 2 atm. calcines %502 gale; dt |0-3 %S solid SCfUbb Cyclone. 1 S a u t t e r oxidized B . C . Moly B.7. 570 2.0 70 27 A 0.99 82 2.5 33 1 12 0.21 2.22 0.98 7.6149 0.9507 it B.8 580±5 1 .4 "7 O i C O O A • 1 .07 p. 7 2.5 35 1 i ? 0.21 3.56 0.91 2.6028 0.9110 11 B, 1-2 575 3.2 64 24'. 8 0.74 62 2.5 25 1 12 0.21 3.22 1 .21 4.8394 0.9195 .11 B.13 541 4.3 72 26.6 . 0.90 .75 2.5 30 1 12 0.21 1 .74 0.91 3.7353 0.9565 II B.14 532 11.7 72 26.5 0.90 75 2.5 30 1 1 2 0.21 1 .20 0.91 1.3857 0.9700 II B.15 560 9.0 81 30.6 1 .45 121 2.5 48 1 12 0.21 1 .42 0.78 1 .7861 - 0.9545 n B.I 7 582 3.2 81 31 .4 1 .60 132 2.5 53 1 12 0.21 1 .57 0.78 4.6914 99.4 0.9607 ii B.18 540 9.7 81 30.0 1 .36 113 2.5 45 1 12 0.21 1 .33 0 .-78 1 .6668- 99.2 0.9667 u B.20 511 13.2 81 25.4 0.7-9 67 2.5 27 1 12 0.2:1 0.93 0.78 1 .2363 : 0.9767 it B.21 524 13.2 81 29.4 1 .25 104 2.5 42 1 12 0.21 0.87 0.78 1 .2088 0.9782 . . II B.22 524 10.4 81 29.4 1 .25 104 2.5 42 1 12 0.21 0.90 0.78 1.5640 0.5 .0.9775 it B.23 550 10.4 81 30.2 1 .38 115 2.5 46 1 12 0.21 0.78 0.78 1.5686 0.9805 II B.-26 588 3.7 77 30.0 1 .36 113 2.5 45 1 12 . 0.21 1 .74 0.86 4.9050 0.9665 it B.27 600 4.3 77 30.5 1 .44 120 2.5 48 1 12 0.21 1 .48 0.86 3.7038 0.9630 II B..30 524 10.7 77 27.8 1 .05 88 2.5 35 1 12 0.21 0.95 0.86 1 .5747 0.9762 . I I B.32 570 10.7 77 29.4 1 .25 104 . 2.5 1 42 1 21 0.21 0.62 0.86 1 .5382 0.9845 II B.33 595 10.7 77 30.4 1 .43 121 2.5 48 1 21 0.21 0.61 0.86 1 .5386 0.9847 II B.36 551 10.7 77 28.7 1 .16 97 2.5 39 1 21 0.21 0.67 0.86 1.5364 0.9832 II B.39 550 15.0 77 28.5 1.14 95 2.5 38 1 21 0.21 0.-50 0.86 1.0972 0.48 99.6 0.9815 II B.40 550 21 .4 77 28.5 1 .14 95 2.5 38 1 21 0.21 0.43 0.86- 0.7728 0.28 97.3 0.9892 II B.41 550 22.6 77 28.5 1.14 95 2.5 38 1 21 0.21 0.48 0.86 0.7278 0.33 97.3 0.9880 it B.42 . 550 15.3 77 28.5 1 .14 95 2.5 38 1 21 0.21 0.53 0.86 1.0725 0.28 0.9867 ii B.43 550 20.8 77 28.5 1.14 95 2.5 38 1 21 0.21 0.56 0.86 0.7888 0.8760 ti B.44 550 20.8 65 24.5 0.71 59 2.5 24 1 21 0.21 0.60 1 .04 0.7880 0.9850 it II B.45 550 20.8 83 31 .0 1 .53 128 2.5 51 1 21 0.21 0.56 0.76 0.7888 0,9860 . n B.46 550 26.0 77 28.5 1.14 95 2.5 38 1 21 0.21 0.42 0.86 0.6342 97.8 0.9895 II B.47 550 26.5 77 28.5 1.14 95 2.5 38 1 21 0.21 0.39 0.86 0.6189 0.9902 it B.48 550 26.5 77 28.5 1 .14 95 2.5 38 1 21 0.36 0.39 0.86 0.6189 98.5 0.9902 II B.49 550 26.7 77 28.5 1 .14 95 2.5 38 1 21 0.59 0.40 0.86 0.6125 98.1 0.9900 n B.50 550 28.0 77 28.5 1.14 95 2.5 38 .1 21 0.21 0.40 0.86 0.5892 95.6 0.9900 O o ccocentKihe, Source Run N-T •c t hr = '/mii 1 Uo cry /sec I- -10 V/t /onset K 9<y. Fo /•run •Tiler: % atm. <7oS calctrcc. gates cit nnole/hr x |0-3 solid scrubb. er % 1 Sautter oxidized B.C. Moly "B.51 550 28 62 23.6 0.64 5-4 2.5 22 1 21 0.21 0.44 1 .08 0.5886 0.9890 B..52 550 28.0 58 21.0 0.46 39 2.5 16 1 .21 0r21 0.48 1.17 0.5880 0^.9880 II B.53 550 28.0 87 32.5 1 .88 158 2.5 63 1 21 0.21 0.44 0.71 0.5886 99.4 0.9890 11 B.54 573 -1 5.3 77 •2-9.4 : 1 .24 104 2.5 42 1 21 0.21 0.49 0.86 1.0736 0.9877 . II B.55 524 8.0 77 27.5 1.01 84 2.5 34 1 21 0.21 1.04 0.86 2.0291 0.-9740 n B .56 573 22.7 77 29.4 1.24 104 2.5 42 1 21 0.21 0.38 0.86 0.7283 0.9905 II B.57 ' 425 23.3 77 ~27.5 1.01 "84 2.5 34 1 21 0.21 0.41 0.86 0.7069 0.33 99.8 0.9897 t i B.-60 574 22.0 77 29.6 1.28 105 2.5 42 1 21 0.21 0.39 0.86 0.7428 0.46 99.9 - 0.9902 II B.61 550 23.3 77 28.5 1.14 95 2.5 38 1 21 0.52 0.41 0.86 0.6996 0.37 97.9 0.9897 II B.62 550 23.3 77 28.5 : 1.14 95 2.5 38 1 21 0.11 0.57 0.80 0.7041 0.30 98.5 0.9857 H 3.63 549 23.7 77 28.5' 1.14 95 2.5 38. 1 21 0.044 1.55 0.61 0.6866 0.52 98.7 0.9612 u B.68 550 25.0 77 28.5 1.14 95 2.5 38. 1 21 0.46 0.47 0.86 0.6588 0.9882 Cchloride leached) -B.C. Moly B.:69 530 13.4 77 28.1 1.08 90 2.5 36 1 21 0.21 0.52 0.86 1.2337 99.3 0.9870 u t B . 7 0 550 8 77 28.5 1.14 95 2.5 38 1 21 0.21 0.64 0.86 2.0500 99.6 0.9840 u B.71 550 19.8 77 28.5 1:14 95 2.5 38 1 21 0.21 0.42 0.86 0.8315 0.9895 11 B.72 550 19.8 77 28.5 1.14 95 2.5 38 1 21 0.52 0.41 0.86 0.8317 0.9897 B.C. Moly B.73 550 23.7 77 28.5 1.14 95 2.5 38 1 21 0.83 0.38 0.86 0.6903 97.1 0.9905 11 B . 7 4 550 23.7 77 28.5 1.14 95 2.5 38 1 21 0.055 0.67 0.86 0.6924 0.9832 11 B.75 550 23.7 50 17.8 0.27 22 2.5 11 1 21 0.21 0.57 1 .36 0.6940 94.2 0.9857 11 B.76 550 23.7 22.0 0.53 1 45 2.5 18 1 21 0.21 0.66 1.18 0.6926 95.8 0.9835 II - B.77 524 22.4 77 27.8 1.05 88 2.5 35 1 12 0.21 0.62 0.86 0.7341 0.9845 11 B.78 550 22.4 77 28.5 1.14 95 2.5 38 1 12 0.21 0.62 0.86 0.7343 0.9845 . II B.79 550 15.5 77 28.5 1.14 95 2.5 38 1 12 0.21 0.82 0.83 1.0532 0.9795 11 B.81 500 9.9 77 27.4 0.99 83 2.5 33 .1 12 0.21 1.13 0.80 1.6122 1.48 0.9717 II B.82 550 9.9 77 28.5 1.14 95 2.5 38 1 12 0.21 0.85 0.84 1.6588 2.18 0.9787 II B.83 524 8 7 7 27.8 1.05 88 2.5 35 .1 12 0.21 1.28 0.7.6 2.0166 99.5 0.9680 Molybdenite coooinrrar-e source Run T "C t hr —-—i Uo /sec / m i r Fo mice _tm. <?oS calo'nej gates c i t %$ solid saubb er Cyclone. 1 ^ M „ S 2 oxidized B.C. Moly B.84 5 0 2 23 77 2 7 . 5 1 .00 83 2.5 33-. 1 12 0.21 0.42 0.86 0 . 7 1 7 0 9 8 . 3 0 . 9 8 9 5 II B.85 511 23 77 2 6 . 8 0.92 78 2.5 31 :1 12 0.21 5 . 3 7 0.86 0 . 7 1 7 9 9 7 . 8 0 . 9 9 0 7 ll B.86 560±3 2 2 . 7 77 2 9 . 2 1.21 101 5.0 20 :1 12 t).21 4), 37 1.71 0 . 7 2 8 4 9 7 . 7 0 . 9 9 0 7 li B.87 550±3 1 1 . 3 77 2 9 . 2 1.21 101 1 0 . 0 10 :1 12 0.21 0.56 3.45 1 . 4 2 9 4 9 8 . 6 0 . 9 8 6 0 tl B.88 . 560±3 1 5 . 2 77 2 9 . 2 1.21 101 7.5 14 : . l 12 0.21 0.44 2.55 1 . 0 8 6 8 9 7 . 9 0 . 9 8 9 0 it B.-89 ~ 5 7 5 24 77 2 9 . 5 1.27 1 0 6 2.5 4 3 . 1 1 2 0.21 0-45 0.86 0 . 6 8 6 6 9 8 . 8 0 . 9 8 8 7 BRENDA-B.90 5 5 0 1 5 . 4 64 2 4 . 2 0.89 - 75- 2.5 30 :1 10 0.21 0.218 1.04 1 . 0 8 1 0 0.13 9 8 . 4 0 . 9 9 4 5 II B.91 5 7 5 1 5 . 4 64 2 5 . 9 1.08 .•91 2.5 36. :1 10 0.21 0 . 2 8 2 1 . 0 4 . 1 . 0 7 9 2 9 4 . 5 0 . 9 9 2 9 11 B.92 5 5 0 21 .0 64 2 4 . 2 0.39 7 5 2.5 30 :1 10 0.21 0.14.1 1.04 0.7971 9 8 . 4 0 . 9 9 6 4 II B.96 _ 5 5 0 1 7 . 2 64 2 4 . 2 0.89 75 2.5 30 :1 10 0.21 0 . 1 8 2 1 .04 0 . 9 6 6 4 0.3.6 9 6 . 7 0 . 9 9 5 4 ii B.98 5 5 0 • 19.4 64 2 4 . 2 0.89 75 2.5 30 •1 10 0.21 0.131 1 .04 0 . 8 5 9 2 0.25 9 6 . 5 0 . 9 9 6 7 H B.99 550 1 9 . 4 64 _ 4 . 2 0.89 75^ 2.5 30 1 10 0.21 0.131 1 .04 0 . 8 5 9 2 0.27 95.1 0 . 9 9 6 5 II B . l 00. .524 3 4 . 5 6 4 2 3 . 2 0.8O 6 8 2-5 27 1 IQ 0.21 0.115 1 .04 0 . 4 7 9 3 97.0 0.9 9 7 1 II • B.-101 -526 27 64 2 3 . 4 0.81 69 2.5 27 1 10 0.21 0.119 1.04 0 . 6 1 5 4 0.14 97.2 0 . 9 9 7 0 tl B. 1 02- 536 27 64 2 3 . 7 0.85 71 2.5 28 1 10 0.21 0 . 1 0 5 1 .04 0 . 6 1 5 6 0.14 97.5 0 . 9 9 7 3 it B.-103 549 27 64 24.1 0.88 74 2.5 30.. 1 10 0.21 0 . 1 0 8 1 .04 0 . 6 1 6 2 0.21 99.2 9 7 . 6 0 . 9 9 7 3 II B . l 05 5 5 0 2 0 . 6 64 2 4 . 2 0.89 75 2.5 30 1 10 0.21 0 . 1 2 3 1 .04 0 . 8 0 3 9 0.38 98.8 98.2 0 . 9 9 6 9 if B..110 5.50 2 0 . 6 64 2 4 . 2 0.89 75 2.5 30 1 10 0.60 0.130: 1 .04 0 . 8 0 3 8 0 . 9 9 6 7 BRENDA < Cdou.led 1eached) B . n i 5 6 0 1 7 . 6 64 2 4 . 2 0.89 74 2.5 30 1 10 0.21 0 . 1 3 0 1 .04 0. 9 4 9 4 0 . 9 9 6 7 II 8 . 1 1 2 550 1 7 . 6 64 2 4 . 2 0.89 75 2.5 30 1 10 0.60 0.118 1 .04 0 . 9 4 9 5 0 . 9 9 7 0 BRENDA II B . 1 1 3 5 5 0 3 8 . 0 64 2 4 . 2 0.89 75 1 .25 6 0 : 1 10 0.21 0.085 0.52 0 . 4 3 7 6 9 8 . 2 97.9 0 . 9 9 7 8 « B . 1 1 4 550 2 0 . 7 64 2 4 . 2 0.89 75 2.50 3 0 : 1 10 0.21 0.132 1.04 0 . 7 9 7 3 9 8 . 0 99.0 0.9967 KENNECOTT B . 1 1 5 5 5 0 17.1 64 2 4 . 2 0.51 4 2 2.50 1 7 : 1 28 0.21 0.65 1.04 0 . 9 6 4 5 9 5 . 3 0 . 9 8 3 7 ll B . 1 1 6 5 7 0 17.1 64 2 4 . 5 0.53 45 2.50 1 8 : 1 28 0.21 0.56 T.04 0 . 9 6 5 0 97.0 0.9860 Molybdenite s o u r c e N°-T t . • • h r Uo Cry /sec 1- -10 V I . /mil Fo /min mice ?0Z atm. calcines %so gases _ d r i o S z at mcle/hr _ | 0 " 3 %S solid SCrufcb e r 01 \ Sautter ^ M o S 2 oxidized KENNECOTT 8 .117 524 17.1 64 23 .2 0 .44 37 2 . 5 l i . r l 28 0.21 0 . 7 3 1 .04 0 . 9 6 2 4 9 2 . 9 0 . 9 8 1 7 II B . l 18' 550 : 2 0 . 0 64 " 2 4 . 2 -0.51 42 2.-50- 17:1 28 0.21 0 .54 1 .04 0 . 8 2 2 0 9 4 . 6 0 79865 11 B . l 1 9 . 550 • .2-5,8 64 2 4 . 2 : 0.51 4 2 2 .50 1"/ :1 28 - 0.21 0 .48 1 .04 0 . 6 4 1 5 9 6 . 3 0 . 9 8 8 0 II B.120 54 -9 25..8 . 64 ' 2 4 . 2 0.51 42 2 .50 17:1 28 0.21 0 .42 1 .04 0 . 6 4 2 5 0 . 9 8 9 5 : ENDAKO B .126 520 118 64 2 3 . 2 1.12 -96 2 . 5 3-7:1 8 0.21 0 . 3 2 1 . 0 4 1 .389 9 9 . 5 0 .9920 -11 B.127 520 1 5 . 4 64 2 3 . 2 1 .12 96 2 . 5 37:1 8 0.21 0 . 3 0 1 .04 1 . 0679 9 9 . 5 0 . 9 9 2 5 11 8 .128 550 1 5 . 4 64 2 4 . 2 - 1.1-5 96 2 . 5 38:1 8 0.21 0 .28 1 .04 1 .0793 99 9 9 . 5 0 .9930 11 B . 1 2-9 550 2 4 . 0 64 2 4 . 2 1 .15 96 2 . 5 3*.: 1 8 0.21 0 . 1 4 8 1 .04 2 . 6 9 1 8 9 9 . 5 0 . 9 9 6 2 ro oo -P=. APPENDIX 14 CALIBRATION CURVE OF THE SCREW FEEDER FOR MoS2 APPENDIX 15 SULPHUR ANALYSIS OF M0O3 CALCINES Samples of 0.1 gr of calcines for %S < 0.5 or 0.01 gr for %S > 0.5 were used to determine the total sulphur content of the calcines. The sulphur was analyzed by iodimetry of the gaseous products of samples fused in a Leco induction furnace. The analysis proceeds according to the reactions: KI03 + 5KI + 6H6 * 6KC1 + 3I 2 S0 2 + I2 + 2H20 * H 2S0 4 + 2HI where the free iodine is used as an indicator by the coloring of a starch solution. For each experiment, three to six samples of calcines were taken and analyzed separately or as a common sample. The minimum level of sulphur that could be analyzed by this method was checked periodically using the standard iron and tin accelerators for the "Leco" analyzer. The range -of confidence of the standard was ±0.02%S. Samples of calcines were also analyzed independently at Can Test laboratories and 287 288 two at Endako mines. The results of those analyses are given be!ow. Percent Sulphur in Calcines Can Test This Thesis A%S 0.80 0.86 + 0.06 0.82 0.88 + 0.06 0.03 0.05 + 0.02 0.40 0.55 + 0.15 1.10 0.96 -0.14 0.54 0.54 0.00 0.78 0.82 + 0.04 0.78 0.82 + Q.04 0.82 0.90 + 0.08 1.41 1.78 +0.37 0.56 0.58 + 0.02 0.46 0.39 -0.07 0.42 0.39 -0?07 0.59 0.49 -0.10 0.54 0.42 -0.12 Endako This Thesis L\%S 0.28 0.280 0.00 0.19 0.148 -0.04 The range of variation of the above sulphur analyses were: Calcines 0.5 - 1%S ~ ±0.09%S Calcines <0.5%S ~ ±0.045%S APPENDIX 16 INFRARED ANALYSIS OF S02 AT STEADY STATE 289 — ^ PUBLICATIONS A. Sutulov and I. Wilkomirsky, "Purification of • molybdenite concentrates by c h l o r i -nation in a f l u i d bed system", J . of • Engineering, 26_, No..l, Chile (1963). F. Concha and I. Wilkomirsky, "Recovery of Sulphur from low grade ores", J. of Engineering,' • Santiago of Chile, _2_7,- No. 3 (1964). J.P. Hager and I. Wilkomirsky, "Galvanic c e l l studies using a molten oxide electrolyte, Part I, Thermodynamic properties of the Pb-Ag system", Transactions of the AIME Soc, 242, No. 2 (1968). I. Wilkomirsky and G. Morizot, "Theoretical and practical action of the wetting agent Aocem-Cu", J. of Minerals, Santiago •'< of Chile, 23, No. 100 (1968). L. Coudurier, I. Wilkomirsky and F. Morizot, "Molybdenite roasting and rhenium vo l a t i l i z a t i o n in a multiple hearth furnace", Institution of Mining and Metallurgy Soc, Sec C, Vol. 80, March (1970). L. Coudurier and I. Wilkomirsky, "Principles of the Extractive Metallurgical Processes", University of Concepcion Press, Chile, 565 pages, (1971). 

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