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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.Sc, U n i v e r s i t y of Concepcion, 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 THE  SUBMITTED  IN P A R T I A L  REQUIREMENTS  FOR THE  DOCTOR OF  in  FULFILMENT DEGREE  OF  PHILOSOPHY  the Department of METALLURGY  We  accept  req u i r e d  THE  this  thesis  as c o n f o r m i n g  to the  standard  UNIVERSITY  OF  B R I T I S H COLUMBIA  November,  1 9 7**  OF  In presenting this thesis in partial  fulfilment  of the  requirements for an advanced degree at The University of B r i t i s h 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 i n a n c i a l gain shall not be allowed without my written permission.  IGOR A . E . WILKOMIRSKY  Department of Metallurgy The University of B r i t i s h Columbia Vancouver, Canada Date  November 1974  ABSTRACT  The development of a new, r e c i r c u l a t i n g  fluidized  bed process for the roasting of molybdenite concentrates has been successfully completed in a bench scale p i l o t  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 l u i d i z e d bed consists of a mixture of  calcines and coarse sand with a wide size range. provides smooth f l u i d i z a t i o n effect  The l a t t e r  behaviour and an a t t r i t i o n  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 l u i d i z e d bed prevents build-up of material along the reactor walls. of f l u i d i z a t i o n  The reactor design was based on studies  c h a r a c t e r i s t i c s , gas and solid mixing, and  particle stratification  using a two-dimension f l u i d i z e d bed  operated at room temperature.  i i  The performance of the process on molybdenite concentrates 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 l u i d i z e d bed process is able to compete favourably with the existing multiple hearth process.  The f l u i d i z e d 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 l o o r area), and a capital cost that is lower by 50%. In addition,potentially levels of S0 bed process.  2  lower operating costs and higher  in the off gases may be realized by the f l u i d i z e d Tests showed that direct slurry feeding of MoS  concentrates is f e a s i b l e .  2  This would eliminate the need for  f i l t r a t i o n and drying steps prior to roasting. Batch kinetic studies in the f l u i d i z e d bed, and observations using hot stage and scanning electron microscopes indicate that the transformation of MoS to Mo0 2  3  is a complex  process which involves an i n 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  that s o l i d state diffusion is the rate l i m i t i n g  is possible  process.  The v o l a t i l i z a t i o n and subsequent condensation of Mo0 seems 3  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 i n a l  residual  sulphur in the calcines are the average residence time of molybdenite p a r t i c l e s in the f l u i d i z e d 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 i n a l composition is not strongly dependent on temperature.  iv  TABLE OF CONTENTS  Page ABSTRACT  ii  LIST OF TABLES . . . .  . . . . . . . . . . .  LIST OF FIGURES.  ix xii  ACKNOWLEDGEMENTS  .  .xviii  Chapter 1  2  INTRODUCTION  1  1.1  Minerals of Molybdenum. .  1  1.2  Ore Dressing of Molybdenum Minerals  1.3  Uses of Molybdenum  1.4  Extractive Metallurgy  . . . .  3 4  of Molybdenum . . . .  6  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  38  Two- and Three-Dimensional Models  v  Chapter  5  6  Page 4.2  P i l o t Plant Equipment  4.3  Experimental Techniques . . .  8  5 5  GAS BEHAVIOR IN TWO- AND THREE-DIMENSIONAL FLUIDIZED BEDS 5.1  Fluidization Properties of SandCalcines Mixtures  5.2  Gas Bubble Measurement in TwoDimensional Fluidized Beds  61 61 7 4  PARTICLE BEHAVIOR IN TWO- AND THREEDIMENSION FLUIDIZED BEDS  85  6.1  Particle S t r a t i f i c a t i o n  6.2  Concentration P r o f i l e of Reaction Solids in the Fluidized Bed  97  E l u t r i a t i o n of Calcines from the Fluidized Bed  103  6.3 7  43  .  85  GAS AND SOLID DISTRIBUTION IN THE FLUIDIZED BED REACTOR  113  7.1  Gas Tracer Experiments  113  7.2  Solid Tracer Experiments  129  KINETICS AND MECHANISM OF MOLYBDENITE OXIDATION 8.1 Batch Kinetic Oxidation of Molybdenite in the Fluidized Bed Reactor 8.2 8.3 8.4  139 139  Mechanisms and Morphology of Molybdenite Oxidation  152  Temperature of Particles During the I n i t i a l Stage of Transformation  163  Estimation of the Total Time of Transformation  167  vi  Chapter 9  Page CONTINUOUS ROASTING OF MOLYBDENITE CONCENTRATES IN THE FLUIDIZED BED REACTOR.  10  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  . 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 11  212  SUMMARY AND CONCLUSIONS.  SUGGESTIONS FOR FURTHER RESEARCH  221 .  224  NOMENCLATURE  225  REFERENCES  232  APPENDICES 1  .  241  2  .  243  3 4  246 . .  274  5  250  6 7  251 259  vii  Appendices  Page  264  8  9  .  10 11  270 . . .  • • •  12  . . . . . . . . .  277 278  13  280  14  . . . . . . . . . . . . .  15 16  267  285 ,  .  .  287 289  viii  ^  •  •  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 - MoS , Mo0 and Mo0 .  2.2  Thermodynamic Functions of Some Molybdenum  2  2  3  . .  14  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 i c a Sand And Ca1 ci nes Bubble Diameter Measurements in 12.5 cm Two-Dimensional Fluidized Bed  77  6.1  Effect of P a r t i c l e Size Stratification  86  6.2  P a r t i c l e Size S t r a t i f i c a t i o n of Time  5.3  6.3  Distribution as a Function  P a r t i c l e Size and Terminal Velocity of Molybdenite Concentrates  ix  63  90 106  Table 6.4  Page E l u t r i a t i o n Tests Performed in the Fluidized Bed Reactor, 12.5 cm Diameter  107  6.5  E l u t r i a t i o n Constants of Eq. (6.9)  Ill  7.1  Gas Tracer Experiments  7.2  Average Values of K / . v , and D, and D„ (be)b a,g r,g Chemical Composition of Molybdenite Concentrates Average Particle Size and Surface Area  119 n  n  125  3  8.1 8.2  141  of Molybdenite Concentrates  141  8.3  Batch Kinetic Experiments  142  8.4  Calculated Rate of Reactionand Reaction Rate Constant for Molybdenite Oxidation Kinetic Studies on Molybdenite Oxidation  150 153  Calculated Values of the Temperature at the Reacting Surface of MoS Particles During the I n i t i a l Oxidation  166  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 MoS  191  8.5 8.6  2  8.7  2  x  Ta ble 9.3  Page Material Balance for Fluidized Bed Roasting of MoS Concentrates in 12.5 cm Reactor. 2  10.1  192  Dimensions and Operating Performance of a 10 TPD Plant for the Roasting of MoS  214  10.2  Calculated Output of Reactors  214  10.3  Capital Cost for a Molybdenite Roasting Plant using a Multiple Hearth Furnace  2  10.4 10.5  . .  216  Capital Cost for a Molybdenite Roasting Plant using a Fluidized Bed Process  217  Summary of Operating Costs  218  xi  LIST OF FIGURES  Fi gure 2.1  Page Phase s t a b i l i t y S  2.2 4.1 4.2  4.3  4.4  4.5  (g)  "  0  "  M o  (s)  diagram for the system a  t  9  0  °  O  |  <  1  8  Gas bubble models in gas-solid f l u i d i z e d beds  28  Two-dimensional f l u i d i z e d bed model and tracer injection device  39  7.5 cm diameter, three-dimensional ( l e f t ) and two-dimensional (right) f l u i d i z e d bed pi exiglas models  41  7.5 cm diameter plexiglas f l u i d i z e d bed model and recirculating system operating in closed c i r c u i t  42  Schematic diagram of the 12.5 cm diameter f l u i d i z e d bed reactor and continuous recirculating system  44  Cyclone system used in the f l u i d i z e d bed reactor  46  4.6  Cyclone discharge rotary valve  48  4.7  Exploded view of r e c i r c u l a t i o n and discharge system Pneumatically operated venturi nozzle connected to the discharge valve (without insulation)  4.8  xii  49 50  Figure 4.9 4.10 4.11  Page Gas d i s t r i b u t i o n grid and dispersion nozzle 12.5 cm diameter f l u i d i z e d bed reactor and r e c i r c u l a t i n g system during operation  52 . . . .  54  General view of the f l u i d i z e d bed p i l o t plant for molybdenite roasting  56  Factor of volumetric increase of a s t a t i c bed of sand as a function of the wt-% of calcines in the mixture  62  Pressure drop through the bed as a function of the s u p e r f i c i a l gas velocity  65  Minimum f l u i d i z i n g velocity for mixtures of sand and calcines  67  5.4  Fluidized bed height as a function of the s u p e r f i c i a l gas velocity  70  5.5  Bed expansion factor for the f l u i d i z e d bed as a function of the s u p e r f i c i a l gas velocity  71  Pressure drop through the f l u i d i z e d bed in the p i l o t reactor as a function of temperature  73  5.7  Typical high speed pictures of twodimensional f l u i d i z e d bed model  75  5.8  Typical bubble measurements made on an "equivalent bubble diameter"  76  5.9  Bubble diameters measured in the twodimensional f l u i d i z e d bed model  78  5.10  Experimental and calculated values for the bubble diameter in the f l u i d i z e d bed  81  5.1  5.2 5.3  5.6  xii i  Figure 5.11 6.1  6.2 6.3  6.4  6.5  Page Bubble diameter calculated using empirical expressions  different 84  P a r t i c l e size segregation in the f l u i d i z e d bed as a function of the p a r t i c l e size d i s t r i b u t i o n of the sand.  87  Particle size segregation in the f l u i d i z e d bed as a function of the f l u i d i z a t i o n time.  . . .  89  P a r t i c l e size segregation in the f l u i d i z e d bed as a function of the i n i t i a l composition of the sand-calcines mixtures  93  Particle size segregation in the f l u i d i z e d bed as a function of the s u p e r f i c i a l gas velocity  95  Typical tracer test in the two-dimensional f l u i d i z e d bed after an input of MoS tracer (black) into the mixture of Mo0 and sand (white)  99  2  3  6.6 6.7  Axial concentration p r o f i l e after a MoS tracer input.  2  101  Influence of the position of the dispersion nozzle on the concentration p r o f i l e of MoS in the f l u i d i z e d bed  102  Cumulative percent of MoS along the f l u i d i z e d bed  104  2  6.8 6.9 6.10  7.1  2  transferred  E l u t r i a t i o n rate and e l u t r i a t i o n flux as a function of the s u p e r f i c i a l gas velocity  108  Semi-1ogarithmic plot of the rate of e l u t r i a t i o n as a function of the s u p e r f i c i a l gas velocity  110  Response signal from S0 pulse input recorded by the infrared analyzer . . . . . . . .  115  2  xi v  Figure 7.2  7.3  7.4 7.5 7.6  7.7  Page Exit age distribution function of tracer gas in the f l u i d i z e d bed as a function of dimensionless time  117  Reactor dispersion number of gas in the 7.5 and 12.5 cm diameter f l u i d i z e d bed reactors as a function of the s u p e r f i c i a l gas vel oci ty.  120  Calculated overall gas transfer c o e f f i c i e n t as a function of the f l u i d i z e d bed height . . . .  124  Calculated p r o f i l e of the axial dispersion coefficientofgas. . . . . . 126 Calculated p r o f i l e for the radial , dispersion c o e f f i c i e n t of gas along the f 1 uidi zed bed . . . . .  1 26  Overall axial dispersion coefficient of gas in the f l u i d i z e d bed as a function of the s u p e r f i c i a l gas velocity . . . . . . . . .  128  Response signal of the discharge point after a pulse input of MoS 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 l u i d i z e d bed height  136  7.8  2  7.11  7.12  Calculated radial dispersion c o e f f i c i e n t of solid as a function of the f l u i d i z e d bed height Average axial dispersion c o e f f i c i e n t of solids in the f l u i d i z e d bed as a function of the s u p e r f i c i a l gas velocity  xv  .  .136  138  Figure 8.1  Page Mole fraction of molybdenite oxidized as a function of time for batch tests  143  Computed values of the rate of reaction as a function of the fractional conversion of MoS  145  Rate of reaction as a function of temperature in the three stages of transformation  147  Influence of the oxygen partial pressure on the i n 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 l u i d i z e d bed reactor. .  156  Hypothetical view of the stages of oxidation of an MoS p a r t i c l e  162  Influence of temperature on the total time required for transformation of MoS to Mo0 in the f l u i d i z e d bed reactor  169  Influence of the s u p e r f i c i a l gas velocity on the sulphur content in the calcines  173  Sulphur content in the calcines as a function of the temperature of roasting  175  Influence of the average residence time of solids in the 7.5 and 12.5 cm diameter f l u i d i z e d bed reactors. . . . .  177  Sulphur content of the calcines as a function of the residence time of solids for different temperatures of roasting  179  Influence of the oxygen partial pressure on the sulphur content of c a l c i n e s . . . . . . . .  181  8.2  2  8.3 8.4  8.7  2  8.8  2  3  9.1 9.2 9.3  9.4  9.5  xv i  Figure 9.6  9.7  9.8  9.9  9.10  9.11  Page Influence of the temperature on the sulphur level of calcines fpr the Kennecott concentrates.  183  Sulphur content of calcines as a function of the average residence time of reaction for the Kennecott concentrates.  184  Sulphur level of calcines as a function of the residence time for low calcium concentrates. .  186  Sulphur content in calcines as a function of the roasting temperature for low calcium concentrates  188  Apparent rate of reaction of molybdenite in the f l u i d i z e d bed reactor as a function of the fraction of sulphur oxidized  190  S0 in the off gases as a function of the MoS feed rate. 2  2  10.1  10.2  10.3 10.4  10.5  .  195  Experimental and calculated residual sulphur in the calcines as a function of the average retention time and temperature . . . . . . . . . . . . .  205  Calculated reactor dimensions as a function of the sulphur level in c a l c i n e s , for d i f f e r e n t output capacities  209  Calculated level of S0 in the off gases as a function of the reactor capacity . . . . . .  211  Layout of a 10 TPD f l u i d i z e d bed plant for roasting molybdenite concentrates  213  Thermal balance for the f l u i d i z e d bed reactor using slurry feed . . .  220  2  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 a l 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 s k i l f u 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 e f f i c i e n t l y  traced the figures of this  t h e s i s ; and to other technicians who helped me in the several stages of this research. I am grateful and through i t  to The University of B r i t i s h Columbia,  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, C h i l e , 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 found at  the end of the nineteenth  application  as an a l l o y i n g  of molybdenum f i r s t of i t s  1 .1  extensive  Minerals  element  by C.W.  century  its  in s t e e l .  began i n d u s t r i a l l y  usage in m e t a l l u r g i c a l  Seheele,  first  practical  Production  in 1910,  as a r e s u l t  and e l e c t r i c a l  technology.  of Molybdenum  Molybdenum i s one of the l e s s common elements, with an average  concentration  Molybdenum-bearing five  main types . ( a )  in the  earth's  d e p o s i t s of i n d u s t r i a l  c r u s t of 0.001%  interest  [2]: Q u a r t z  veins  (b)  Pegmat i t e s  (c)  High  (d)  Metarporphic  (e)  Disseminated deposits of the p o r p h y r i c copper type  temperature  skarn  schist  1  bodies  and g n e i s s e s  are of  [1].  Most of the commercially producing deposits are located along the c o r d i l l e r a systems of North and South America, but as an element, molybdenum is widely distributed over most of the earth's crust. interest  The minerals of molybdenum of  are given in Table  industrial  1.1.  Table  1.1  principal Molybdenum Minerals • • j  . -  T—  '  Mineral  •  r  [3]  1  :  :  ••  :  Composi tion  Molybdenite  MoS  2  Wulfenite  PbMo0  3  Molybdite  Fe0  Poweli te  3  • 3Mo0  3  • nH 0 2  CaCMoWjO^  Ilsemanite  Mo0  2  • 4Mo0  Bilonesite  Mg Mo CU  Pateraite  Co Mo 0^  Of these, dply the f i r s t  3  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], Taken together,  followed by the USSR, Canada and C h i l e .  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 P h i l i p p i n e s .  1.2  Ore Dressing of Molybdenum Minerals Molybdenum-containing materials are concentrated  almost exclusively by f l o t a t i o n , which yields very high recoveries.  Gravitational  concentration and magnetic separa-  tion are used in very limited cases, mainly to separate i r o n . Molybdenite is easily f l o a t e d .  Ores containing a  few hundredths of 1% molybdenite y i e l d concentrates containing 85-90% MoS , with over 90% recovery [6]. 2  Porphyric copper  minerals, with 0.3 to 1.5% copper and 0.001  to 0.1% molybdenum  are an increasingly important source of molybdenite. minerals usually are treated by a c o l l e c t i v e f l o t a t i o n sulphides followed by a d i f f e r e n t i a l  flotation  These of  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% MoS . 2  An alternative practice is to depress the molybdenite  with starch.  The main impurities present in molybdenite  concentrates usually are S i 0 , CaC0 , A 1 0 , CaSO^, Cu S, 2  FeCuS , PbS, FeS 2  2  3  2  3  2  and F e 0 , depending upon the source of 2  3  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 l o t a t i o n  circuit  down to -200 mesh,  followed in some cases by a regrind of the primary molybdenite concentrates.  As a r e s u l t , the f i n a l molybdenite concentrates  are extremely f i n e , 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 d i r e c t l y by this  industry.  Table 1.2 gives a breakdown of usage of molybdenum in some of the western industrialized countries in 1971. With a s o l u b i l i t y of about 8% in iron [7], in steel exists mainly as a solid s o l u t i o n .  molybdenum  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 i t reaches 7.5 to 8.5%,  replacing tungsten.  steels  Molybdenum produces  5  Table  1.2  Molybdenum Usage in the Western Countries [8]  Structural  steels  45%  Stainless steels  20  Tool  11  apd high speed steels  Cast iron and steel mill  rolls  6  Super al1oys  5  Molybdenum metal  4  Chemical and Lubricants  8  Mi seel 1aneous  1 TOTAL  100%  several important modifications in s t e e l s :  it  imparts a fine  and uniform grain s i z e ; reduces the eutectoid decomposition temperature  range for hardening and tempering; and improves  the e l a s t i c l i m i t of steels and the wear and impact resistance. When alloyed with chromium and n i c k e l , molybdenum also eliminates  the tempering brittleness  of s t e e l s .  In cast i r o n , molybdenum reduces the grain s i z e , thus improving the wear resistance and high temperature  properties.  Cast iron containing s i l i c o n and molybdenum is employed as an acid-resistant material.  Molybdenum is widely employed in  6  high temperature nickel  and  and a c i d - p r o o f s t e e l s a l l o y e d with chromium,  cpbalt. Molybdenum m e t a l ,  point of 2620°C [ 9 ] ,  due to i t s  very high  second only to t u n g s t e n , i s  used in e l e c t r o n i c and e l e c t r i c a l  melting extensively  components, as f i l a m e n t  supports and g r i d s in e l e c t r o n i c tubes and as wire and ribbons in heaters thermal  f o r high temperature  furances.  Due to i t s  low  neutron capture cross s e c t i o n , molybdenum metal  also used as s t r u c t u r a l Molybdenite  material  is u t i l i z e d  lubricant  due to the l a m e l l a r  structure  of MoS . 2  in atomic  is  reactors.  in a very pure form as a  nature  of the hexagonal  crystal  Sodium molybdate i s used as pigment and  dye, and molybdenum oxides Mo0 employed as c a t a l y s t s  3  and Mo0  2  in d e s u l p h u r i z a t i o n  are  increasingly  u n i t s at  petroleum  refineries.  1 .4  Extractive  Metallurgy  Molybdenite for  of Molybdenum  concentrates are the s t a r t i n g  the production of ferromolybdenum and several  material  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,  and calcium molybdate.  In  almost a l l  cases,  sodium molybdate  the f i r s t  in processes which convert molybdenite to a usable form  step  7  involves the oxidation of MoS to Mo0 . 2  cases alternative  1.4.1  3  hydrometallurgical  However, in some  routes may be used.  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 scale.  no longer iq use or are employed on a very  limited  Fluidized bed furnaces have not been successfully  u t i l i z e d although several attempts have been made to roast MoS in them on a p i l o t 2  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/m  2  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  s o l i d s , 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 compounds and ferro a l l o y s .  molybdenum  Technical grade molybdic oxide  is marketed in powder or briquette  form.  of standard products are given in Table  Table  Chemical analyses 1.3.  1.3  Chemical Analysis of Technical Grade Molybdenum Trioxide and Oxide Briquettes Tech.  M0O3  M0O3  Briquettes  Moo  79-90%  70-75%  Mo equivalent  53-60%  47-50%  Cu (maximum)  0.50%  0. 50%  S  (maximum) .  0.25%  0.25%  C  (maximum)  3  Si 0 , Al 0 , others 2  2  3  12% balance  balance  The sulphur content of both technical grade molybdic oxide and oxide briquettes  has to be s t r i c t l y  controlled for  steel  9 alloying purposes, with a maximum allowable l i m i t of 0.20 0.25% S.  to  Copper and lead contents also cannt exceed 0.5%.  Technical grade molybdic oxide and oxide briquettes  are  used d i r e c t l y for ferroalloys as alloying addition in structural s t e e l s .  In the l a t t e r case the oxide is added d i r e c t l y  to the ladle where i t  is reduced rapidly by the carbon in  the s t e e l . Pure molybdic oxide is obtained from technical grade material by v o l a t i l i z a t i o n  of the oxide above 600°C  to y i e l d a product with about 99.975% Mo0 . 3  The fine powder  and acicular crystals formed during condensation of the Mo0  3  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 d i r e c t l y as an addition to the steel furnace.  Molybdenum  metal is produced by hydrogen reduction of pure Mo0 or 3  ammonium paramolybdate at 950 to 1100°C. is a fine powder of about 0.1  The product obtained  to 6 microns in s i z e .  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 Tri oxi de At the present time, no hydrometallurgical process  is in industrial  operation to extract molybdenum d i r e c t l y from  molybdenite concentrates or ores. metallurgical  A l l the existing hydro-  processes are u t i l i z e d to obtain molybdenum  compounds by treating technical grade molybdic oxide. hydrometallurgical ores of molybdenum:  However,  processes are used to treat the oxidized 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 p u r i f i c a t i o n , 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  1.4.3  1.1  Alternative Processes for Treatment of Molybdenite Several alternative methods'of treating molybdenite  concentrates have been proposed.  Pyrometal1urgical processes  include decomposition of MoS in a plasma at 5000/8000°C 2  [74,75], aluminothermic reduction of salt electrorefining  [76]  M0S2  followed by fused  and electrooxidation of MoS -C  anodes in a fused s a l t bath [76].  2  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 s c a l e .  PYROMETALLURGICAL  Molybdenite  HYDROMETALLURGICAL  TREATMENT  Oxidized  Concentrates  >8535 M o S  Ca  2  MoO*  -  PbMoC, -  TREATMENT  Ores  F e ( M o O * ) • , , • >0. U 2  Mo  r o a s t i ng  T e c h n i c a l Molybdic Oxide >55% Mo, <0.25% S , <0.5% Ca  Briquetting pitch  with  Distillation  Thermi te reduction  Limestone roasting  Ferromolybdenum 58-645! Mo  Calcium molybdate >40% Mo  Mo  Oxide briquettes >52% Mo <12S C <0.5% Cu <0.25% S  Metallurgical  Figure  1.1  uses,  metallic powder >95% Mo <2.S% Fe <1.5% S1 < 0 . 5 S Cu  Lub r i cants 2.1*  85-0$  Principal  extractive  metallurgical  (_% r e p r e s e n t  process  the average  metal sheet, rod, wire It. 7*  pigments, fertilizers, ca ta1ys i s t s , 5.8*  to  treat  ores  and  u s e I n 1966 i n USA.)  concentrates  of  molybdenum.  Chapter 2  LITERATURE  2.1  REVIEW  The Roasting of Molybdenite 2.1.1  Physical and Thermodynamic Properties of Molybdenite and Molybdenum oxides 2.1.1.1 Natural  Physical Properties  molybdenite ( a - MoS ) has an hexagonal 2  crystal l a t t i c e with a lamellar structure, in which the molybdenum atoms l i e between two layers of sulphur atoms in the form of a trigonal Artificially structure  prismatic co-ordination polyhedron  [12].  prepared molybdenite ( 8 - MoS ) has a rhombodedral 2  [12]. Molybdenum dioxide has a distorted r u t i l e  structure  [13], while molybdenum trioxide has an orthorhombic c r y s t a l l o graphic structure.  The change from the hexagonal  structure  of a - MoS to the orthorhombic structure of Mo0 during the 2  3  roasting of the molybdenite plays an important role in the transformation process, as w i l l 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 - MoS  2s  Mo0 and Mo0 2  a - MoS  Mo0  Mo0  160.08  127.95  143.95  4.80/4.88  6.342  4.694  2  Molecular weight Density, gr/cc 25°C  [16]  3  2  3  Melting point °C  [17]  1650/1700  decomposes  795  Boiling point °C  [14]  decomposes  decomposes  1100  hexagonal  di st. ruti 1 eorthorhombi c  Crystal 1ographi c structure  Molybdenum trioxide pressures of the metal oxides. vapour pressure is 1 0  -2  has one of the highest vapour At 600°C, for example,  mmHg, whereas at 850°C i t  Kubaschewski, Evans and Alcock [14]  give the  expressions for the partial pressure of Mo0  3  the  is 22 mmHg.  following as a function of  temperature: (a)  from 298 to 1068°K:  log p = -15,230T  (b)  _1  - 4.02  log T + 27.16  mmHg  (2.1)  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 Mo0  3  molecules to a trimer form [15]  (Mo0 ) . 3  Calculated values  3  of the vapour pressure of Mo0 as a function of 3  temperature  are l i s t e d in Appendix 1.  2.1.1.2  Thermodynamic Functions  The s p e c i f i c heat (Cp)» enthalpy (AH° ) T  and free energy of formation  (AG° ) T  of formation are w,ell established  for molybdenum disulphide and molybdenum t r i o x i d e .  Few data  exist however for the less stable molybdenum dioxide. Table 2.2 equations -fo'rAC , A H ° are given.  and A G °  T  T  In  for these compounds  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 MoS to Mo0 following the reaction: 2  MoS,  (s)  + 7/2 0  2  3  Z  at 550°C, a usual temperature lated values of enthalpy,  Mo0  3  (s)  2  (.2.3)  for roasting operations, calcu-  free energy of reaction and  equilibrium constant are, respectively  i  + 2S0  Table  2.2  Thermodynamic F u n c t i o n s of Same Molybdenum Compounds  jTemperature Range, °K  S p e c i f i c Heat, c a l / m o l AC, = -19.7 + 3,15 • TO-  a - Mo S . ;  T  3  273 - 729  .[19]  P  MoCJj AC  McO = Ri  M  °ts)  +  ,  action  Ug) t  i S  U)  Mo  Cs)  +  *S  +  0  2(s)  «os  I  l ( s )  MoS  AH!  2  T - 3.08  =  -57,640 + 6.85T - 5.60 • IO"  T  3  -89,000 - 14.42TlogT - 0.2  - 0.503  2  • l'0~ T 3  -57,640 - 15.78TlogT + 5.60 • IO"  AH  s  800°X  =  -  1 3  °.°  0 0  (s)  + fo  2  t  Mo0  T  s  298 - 1068  [20]  2  Temperature Range, °K  3 ( s )  "  • 10  T"  1  5  2  273 - 729  [21]  [21]  - 0.252/ 10  T- • + 49 1  5  24T  [21]  C22]-  AG° = -140,100 - 4.6TlogT • 55.8 T  M0  . 10  + 87.86 T  2  T  3  2 ( s )  Mo0 ^. j  2  3  Thermodynamic F u n c t i o n , c a l / m o l  &&j =  M o  IO"  -- 3.85 + 3.10 •  AH° == -182,600 + 3.85T + 1.55 • 1 0 "  273 - 729  298 -  1300  298 -  1068  [235  3  T  AGy == -182,600 - 3.85TlnT - 155 • 1 0 "  3  +3.08 • 10  2  T  2  s  + 1.54 • 1 0  T"  s  [20]  1  T"  1  + 89 .77  [20]  17  AH AG  823°K  -295,920  cal/mol  823°K  -207,210  cal/mol  K823°K  4.98 x 10  These values indicate that the reaction occurs i r r e v e r s i b l y at this temperature,  with a large amount of  heat generated. Coudurier, Wilkomirsky and Morizot [18]  have c a l -  culated the phase s t a b i l i t y  diagram (Kellogg diagram)  the system  at 900°K, as shown in Figure 2.1.  - 0 - Mo ^  for  At the operating conditions of normal roasters with 1 to 3% S0  2  in the reactor and in the presence  stable compound is Mo0 . 3  of air the only  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 t e r a t u r e [24,25].  Butters [26]  has given a detailed  description of the problems of gas disposal and s o l i d  18  Figure 2 . 1 .  Phase s t a b i l i t y diagram f o r the system (g) " " ( s ) ° -  S  0  M o  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. [27]  Wright  has described the operations at Endako Mines where about  five million pounds of molybdenite concentrates are roasted yearly in a 16 f t . et al. [28]  diameter,  10 hearth furnace.  Lastovitskaya  have studied the different molybdenum compounds  formed during roasting in an industrial  multiple-hearth  furnace, and Coudurier, Wilkomirsky and Morizot [18]  investi-  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 p o s s i b i l i t y that a f l u i d i z e d bed with i t s inherent high efficiency and s i m p l i c i t y , 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. [29]  The f i r s t known attempt was by Deev and Smirnov  who used a small batch f l u i d i z e d bed to roast molybdenite,  but no information was given on the f e a s i b i l i t y  of the process.  In the U.S.S.R. there have been several plgnt studies of f l u i d i z e d bed roasting. [30]  pilot  Zelikman et al.  attempted to roast granulated molybdenite concentrates  20  using bentonite as a binder in a 1 m f l u i d i z e d bed at 2  580°C.  570-  Small pellets of 0.25 to 2 mm diameter were employed  to minimize the dust entrainment to the furnace.  prior to charging molybdenite  The feed rate was 1160 Kg/m  . day.  2  Dust  formed by a t t r i t i o n of the small granules and entrained by the gases amounted to 39% of the feed. partially it  This dust was only  oxidized and was granulated again before  to the reactor.  Thus the problem of dust  returning  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 e l e c t r i c a l particles inside the f l u i d i z e d bed.  f i e l d on the  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]  t r i e d to  roast coarse molybdenite concentrates of 35 to 60 mesh in a batch f l u i d i z e d 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 Mo0 . 3  Dust entrained in the off gases amounted to as much as 56% of the charge and contained 2 to 4% S.  The remaining c a l c i n e s ,  21  comprising 44% of the original of 0.3 to 0.5%.  charge, had a sulphur content  Their research apparently did not progress  beyond these batch experiments. Golanit, Korneeva and Stepanov [33], et al. [34]  and Gorin  have published work on the temperature  control  of f l u i d i z e d bed furnaces for molybdenite roasting. industrial  However,  application of the process in the U.S.S.R. does  not appear to have been achieved. et al. [35]  Indeed, Lastovitskaya  report in 1970 that "the fluidized-bed furnace  cannot be used since the calcines obtained would contain to 2% S."  In a recent attempt, Grigoriu and Balasanian  1.5 [36]  roasted coarse molybdenite concentrates of 0.1-0.2 mm diameter in a batch laboratory f l u i d i z e d 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 l u i d i z e d 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 l u i d i z e d 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 solved.  of p a r t i a l l y  converted dust been s a t i s f a c t o r i l y  22  2.1.3  Kinetics and Mechanism of Molybdenite Oxidation Heterogeneous gas-solid reactions such as the  roasting of molybdenite have usually been analyzed mathematically by the unreacted shrinking core model.  It  is assumed  that reaction occurs topochemically, beginning on the surface and moving toward the centre of the p a r t i c l e .  A partially  reacted p a r t i c l e 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 p a r t i c l e can be controlled by the following steps: (1)  Mass t r a n s f e r o f oxygen 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 the part icIe,  (2)  D i f f u s i o n o f oxygen through t h e formed oxide layer, i f the oxide i s not voI 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 the o x i d e / s u l p h i d e i n t e r f a c e ,  (4)  Diffusion the oxide  (5)  Mass t r a n s f e r o f S 0 t h r o u g h t h e g a s film surrounding the particle.  of the generated l a y e r , and  S0  2  at  through  2  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 l a t e r , this c l a s s i c a l picture of g a s - s o l i d reaction kinetics is inadequate for the molybdenite-oxygen reaction.  Table Kinetic  Workers  Studies  Buik 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 Temperature 380-550°C  Calistru  and  et  Nelson  & Belaevskaya  Ammann a n d  Loose  Lastovitskaya  Oxidation  et  al.  Important  oxygen  Oxidation  start  Oxidation  is  37  at  400°C 38  -0.2  mm  Chemical  linear  reaction  controls  at  initial  diffusion  controls  late  transformation  Pressed concentrate discs Temperature 3 6 0 - 6 4 0 ° C Oxidation in a i r  R a t e p a r a b o l i c f r o m 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 t o 5 4 0 ° C A c t i v a t i o n energy 16.8 Kcal/mol  Pressed concentrate discs Oxidation in a i r Temperature 4 0 0 - 6 0 0 ° C  D i f f u s i o n in s o l i d s t a t e at 400°C Chemical r e a c t i o n + D i f f u s i o n a l in State c o n t r o l at 500°C Chemical r e a c t i o n at 5 0 0 ° C  Thin concentrate layer O x i d a t i o n i n a i r and o x y g e n Temperature 5 2 5 - 6 3 5 ° C  Chemical r e a c t i o n to 70-80% o f A c t i v a t i o n energy  Concentrates Industrial multiple hearth furnace Oxidation with a i r  Detected Mo0 , different 2  Deev  Smirnov  Concentrates B a t c h f l u i d i z e d bed Oxidation with a i r  Below 4 0 0 ° C controlled;  & Balasanian  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 Temperature 4 5 0 - 6 0 0 ° C Oxidation with a i r  Chemical followed  hearth air  stages 39  Solid  Concentrates Pilot multiple Oxidation with  Grigoriu  Reference  Rate p a r a b o l i c from 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 Mo0 only product found  Coudurier, Wilkomirsky and M o r i z o t  and  Results  3  Buik c o n c e n t r a t e O x i d a t i o n in a i r Temperature 4 8 0 - 6 0 0 ° C Bulk p a r t i c l e s of O x i d a t i o n in a i r Temperature 5 0 0 - 6 5 0 ° C  al.  Ong  Zelikman  on M o l y b d e n i t e  Conditions  Ghen  Godfrey  2.3  furnace  controls overall convertion 35.3 Kcal/mole  M c O n , M o 0 6, M0O3 l e v e l s of furnace 9  2  40  solid  rate  42,43  up 44  at  Reaction sequence M o S - M o 0 - M o 0 . Oxidation c o n t r o l l e d by i n t r a - p a r t i c 1 e diffusion and c h e m i c a l c o n t r o l a t t h e end 2  2  35  3  reaction is chemically above t r a n s p o r t control  r e a c t i o n c o n t r o l s r a t e at b e g i n n i n g , by d i f f u s i o n i n l a t e r stages  CONTINUED  .18  29  36  T a b l e 2.3 Continued) Workers Zelikman Zelikman  & Voldman  Conditions  Important  Reference  Concentrates Continuous p i l o t f l u i d i z e d bed  Between 500 t o 650°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  Buik concentrate Solid state reaction MoS and M0O3  between  Rate f a s t 700°C, d e c r e a s i n g s h a r p l y decrease i n temperature  47  Bulk c o n c e n t r a t e Solid state reaction MoS and M0O3  between  Over 700°C, with MoS  2  Galateanu  Results  gaseous  M0O3  probably  with  react  2  48  2  Cardoen  Pressed cylinders 492 t o 6 7 0 ° C 0.003 t o 0.83 a t m o x y g e n  Chemical r e a c t i o n c o n t r o l t h e transformation A c t i v a t i o n energy 42.4 ± 1 Kcal/mole  41  25  Despite several investigations having been conducted, the kinetics and mechanism of molybdenite oxidation have only p a r t i a l l y been c l a r i f i e d .  Contradictory results have  been reported by workers using different techniques to study the oxidation k i n e t i c s . summarized in Table It  Findings of these researchers are  2.3.  is evident from these results that a c l e a r l y  defined mechanism for the reaction has not yet been established.  Nevertheless, the studies of Cardoen, Ammann and  Loose, C a l i s t r u , Zelikman and Belaevskaya and Grigoriu and Balasanian a l 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 Mo0  3  2.2  has not yet been proven.  Fluidized Beds Fluidization is an operation by which fine solids  are suspended in a f l u i d - l i k e through a bed or p a r t i c l e s .  state by passing a f l u i d upwards  26  In order to scale-up results from laboratory f l u i d i z e d bed roasting studies to f u l l  size industrial reactors,  in addition to obtaining information about the kinetics of the reaction, a basic knowledge of gas and s o l i d mixing and mass and heat transfer in f l u i d i z e d beds is required. When a gas flows upwards through  a  bed at low  flow rates, the bed remains s t a t i c 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 l u i d i z a t i o n condition."  Any further gas flow rate increase over this minimum condition produces gas bubbles that ascend through the bed in a f l u i d - l i k e manner. The gas bubbles ascending along the f l u i d i z e d bed are the major mechanism of s o l i d mixing in the bed and represent a separate "dilute phase" which is d i s t i n c t from the f l u i d i z e d solid i t s e l f ,  often referred to as the "emulsion  Penetration of gas from the bubble into the  phase."  particulate  phase forms the s o - c a l l e d "cloud" of the bubble, which plays an important role in the gas interchange between both gas and emulsion phase.  Mathematical modeling of f l u i d i z e d 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 l u i d i z e d beds up tp 1970 are well summarized in the books of Zenz and Othmer [49], [52]  Leva [50],  Zabrodsky [51],  and Harrison and Davidson  2.2.1  Kunii and Levenspiel  [53].  Gas Bubbles The importance of gas bubbles in f l u i d i z e d beds was  first  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 bubbles are not spherical as  findings: the  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 e f f e c t s . In p a r t i c u l a r ,  the Murray model (Figure 2.1-b) appears to  represent the experimental  findings more c l o s e l y .  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.  penetration of gas cloud gas bubble solid - g a s emulsion flow lines for downward motion of emulsion ggs recirculation flow lines inside bubble and cloud (flow is symmetrical on both sides) (a)  Davidson - Harrison  model  emulsion flow lines penetration of gas cloud gas bubble gas recirculation flow lines inside bubble and cloud (flow is symmetrical on both sides  solid-gas emulsion  wake  emulsion recirculation flow lines in wake  (b)  Figure 2.2.  Murray  model  Gas bubble models in g a s - s o l i d f l u i d i z e d 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 l u i d i z e d bed.  The fact  that bubbles coalesce, s p l i t and change in size and shape as they rise through a f l u i d i z e d bed has not yet been described  2.2.2  quantitatively.  Fluidized Bed Models Mathematical  models of the f l u i d i z e d bed as a whole,  attempt to predict the size and d i s t r i b u t i o n of bubbles across and along the bed.  The e a r l i e r  assumption that the  be,d was an homogeneous mixture of gas and s o l i d s , 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 This model considers a two-step  (emulsion)  phase.  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 i z e , " that i s assumed to be constant along the f l u i d i z e d bed.  However, this s i m p l i f i c a t i o n 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 l u 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 c a t a l y t i c f l u i d i z e d bed reactors [60].  Yoshida and Wen [61]  extended the application of the  model tp the simulation of roasting of s p h a l e r i t e . values agree .cfosely with some experimental et al. [62]  Calculated  data of Yagi  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 v a l i d i t y 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. apparently simple factor is most d i f f i c u l t  However, to  this  evaluate.  That these models appear to be unable to account  quantitatively  in a r e l i a b l e manner for the many phenomena occurring inside the f l u i d i z e d bed has precluded their application in the present work.  31  2.2-3  E l u t r i a t i o n of Solids E l u t r i a t i o n refers to the selective removal of  fine particles from a f l u i d i z e d bed reactor.  This phenomena  is p a r t i c u l a r l y important for f l u i d i z e d bed reactors that treat fine metallurgical f l o t a t i o n concentrates, since much of the feed material  may be carried over with the f l u i d i z i n g  gases rather than being recovered as a bed overflow. Non-steady state studies on the phenomenon have been made by Leva [64], [65],  Hanesian and Rankell [66]  elutriation  Osberg and Charlesworth  and others [67 ,68,69 ,70,71 ].  General relationships have been proposed by Wen and Hashinger [72]  •Vf  in the form  7  =  U o  Re.  f  V  which correlates the e l u t r i a t i o n  "mfJ  (2.4)  rate constant X* per unit  area of the bed per unit time with the variables of the f l u i d i z e d bed.  Yagi and Aochi[73] proposed a similar  expression in the form  K*d u  = f  .  1 I  (Re  P  )(Fr)  (2.5)  32  A general expression that accurately describes the  elutria-  tibn rate for a wjde range of conditions remains to be developed. actual  For design purposes, i t  seems safer to use the  data from the reactor to determine the value of the  rate of e l u t r i a t i o n ,  at a given condition.  Chapter 3  SCOPE OF THIS RESEARCH PROGRAM  This research was undertaken with the primary goal of developing a f l u i d i z e d bed process for the roasting of molybdenite concentrates. on  Normal constraints were placed  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 fluidized  bed  to the  process was chosen as a probable alternative  multiple hearth furnace since, as has been shown for the roasting of other metal sulphides, i t  has several  advantages: (1)  higher output hearth  (2)  s i m p l e r process equipment with possible reduction in capital costs  (3)  clqser control of operation  (4)  possibly less maintenance  (5)  higher thermal  per u n i t area o f  efficiency  33  inherent  34  . (6) (7)  greater  control  of gas  higher concentration • t h e r o a s t e r gases  emissions  o f SO2 i n  As discussed i'n Chapter 2, f l u i d i z e d beds have not yet been used on an industrial  scale to roast molybdenite  concentrates, unlike other metal sulphides (copper, z i n c , nickel and iron)-  The reasons for this lack of application  have also been outlined in Chapter 2.  B a s i c a l l y , attempts  at f l u i d i z e d bed roasting molybdenite have foundered on one or more Qf the d i f f i c u l t i e s centrates.  peculiar to molybdenite con-  F i r s t l y , the concentrates are extremely  fine,  haying an average p a r t i c l e size of about 10 microns, and are difficult  to handle.  Previous workers [82]  have attempted  to sol ve the problem of feeding such fine materials f l u i d i z e d bed by granulating the MoS . 2  into a  However these granules  break down again due to a t t r i t i o n inside the reactor and the fines, which r e s u l t , are blown out in the f l u i d i z i n g gas.  In order to prevent excessive losses of  the molybdenum, expensive dust c o l l e c t i o n systems are needed, with the dust being returned in granule form back to the f l u i d i z e d bed.  Secondly,  $he  vapour pressure of  M0O3,  the  product; of the molybdenite roasting reaction, is very high evefi at r e l a t i v e l y  modest temperatures  such as 600°C.  imposes an upper l i m i t on the temperature can  be conducted, since as the temperature  This  at which roasting 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 t i o n , further  results in low rates of oxida-  necessitating long retention  low levels of sulphur in the c a l c i n e s .  times to achieve  T h i r d l y , the sulphur  content of the calcines must be less than 0.25% and indeed closer to 0.1%, the level  routinely achieved in the  hearth furnace, i f the calcines are to be sold for  multiple metallurgical  uses. The f l u i d i z e d bed process developed in this work has been designed s p e c i f i c a l l y to overcome these major d i f f i c u l t i e s that have plagued the e a r l i e r workers.  Problems  associated with feeding the fine MoS p a r t i c l e s and recovering 2  the equally fine calcines from the bed have been solved by continuously recirculating the s o l i d s , and feeding them back into the reactor along with fresh MoS employing a s p e c i a l l y 2  designed pneumatic i n j e c t o r .  In this system the fines  elutri-  ated from the f l u i d i z e d bed are collected in high e f f i c i e n c y 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 l u i d i z e d beds consisting of coarse sand and Calcines.  The sand, which is not  elutriated,  prpyides the necessary a t t r i t i o n to keep calcine p a r t i c l e s separated.  In addition, build-up of sinter along reactor  walls has been avoided by employing a rotary scraper.  Finally,  36  low 'sulphur levels in the product calcines have been achieved by optimizing the roasting temperature and average retention time.  particle  Long retention times are, of course, possible  with [the solids r e c i r c u l a t i n g system incorporated in this process. The design of the process and attainment of reasonable operating conditions have not been reached  arbitrarily.  Considerable research e f f o r t was also expended to gain knowledge of the micro- and macro-processes occurring inside the f l u i d i z e d bed which affect  i t s operation.  Separate  experiments were conducted in two- and three-dimensional f l u i d i z e d beds, contained in p l e x i g l a s , to study the f l u i d dynamics of the gas (both in the bubbles and in the emulsion) and the mixing behavior of the s o l i d s .  The results of these  experiments are presented and discussed in Chapters 5, 6 and 7. To study the oxidation kinetics under f l u i d i z i n g 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 during roasting.  particl  This work involved observations with a hot-  stage- and a scanning electron-microscope, the results of wliich-are described in Chapter 8. F i n a l l y , the f l u i d i z e d bed process was tested and proved using a 12.5 cm diameter reactor in a f u l l y continuous, bench scale p i l o t plant.  The results of these t e s t s , in  which four different commercial concentrates were roasted, are given in Chapter 9.  The a b i l i t y of this f l u i d i z e d bed  37  process cussed  to in  roasting  compete Chapter  operation  with 10. have  the  multiple  Results been  hearth  obtained  scaled  up  in to  furnace the  make  is  dis-  continuous this  comparison.  Chapter 4  EXPERIMENTAL EQUIPMENT AND OPERATIONAL PROCEDURES  In order to obtain preliminary information on the gross behaviour of the f l u i d i z a t i o n of molybdenite concentrates, two and three-dimensional plexiglas models were f i r s t structed.  con-  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 p i l o t  4.1  plant.  Two and Three-Dimensional Models To study the p a r t i c l e transfer and bubble d i s t r i -  bution, as well, as p a r t i c l e segregation in the f l u i d i z e d bed, two-dimensional reactors were i n i t i a l l y one 7.5 cm* wide by  1 m  second .1 2.5 cm* wi de by  high by 1.25 cm 1 m  fabricated,  thick, and a  high and 1.25 cm  thick.  These  were constructed to represent a central section of the three-dimensional c y l i n d r i c a l reactors.  The 12.5cm wide,  Reactors were made of 3-in and 5-in diameter pipe.  38  tluid bed light source  / / /  recycling nozzle  • flowmeters  distributor ceramic balls  11 high speed movie camera  tracer injection • device  V  fluidizing air  recycling air  3-way valve CO  Figure 4.1.  Two-dimensional fluidi;:ed bed model and tracer injection device.  40  A general  two-dimensional reactor is shown in Figure 4.1. view of the equipment used in the f l u i d i z a t i o n  studies is  shown in Figure 4.2. To determine the main f l u i d i z i n g of t h e material  characteristics  used, a. 7.5 cm diameter by  reactor, shown in Figure 4.2, was employed. distributor pf j mm.  1 m  model  The gas  in the model has 42 holes, each with a diameter  To separate the entrained material  from the  f l u i d i z i n g bed gases, a cyclone system was also designed and constructed [83,84].  The basic principle of the f l u i d i z e d  bed roaster used in this work was the elimination of losses of material  from the entrained s o l i d s , which is a serious  problem i'n multiple hearth roasting, by continuously rec i r c u l a t i n g these fine solids to the reactor. purpose a venturi-type  For this  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 c o l l e c t i o n of solids was found always to be less than 80 to 90% of the material.  entrained  Therefore, to increase the c o l l e c t i o n e f f i c i e n c y  a second cyclone was designed to operate at about 1.5 th,e tangential  velocity of the main cyclone.  times  Figure 4 . 2 .  7.5 cm diameter, three-dimensional ( l e f t ) and two-dimension (right) f l u i d i z e d bed plexiglas models .  Figure  4.3.  7 . 5 cm d i a m e t e r 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 i n c l o s e d ci rcui t.  43  4.2  P i l o t Plant Equipment Based on the information obtained from the plexiglas  models, a 7.5 cm diameter by constructed and tested.  1 mm  high reactor was f i r s t  Subsequently, due to the problems  associated with the size of this reactor, a l a r g e r , diameter reactor was constructed.  12.5 cm  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 c o n t r o l l e r s .  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 g r i d , measured the average bed temperature.  Another thermocouple at the top of the reactor  monitored the temperature of the gases at the To avoid build-up of material  outlet.  along the walls of  the reactor, a 6 rpm rotating arm scraper was i n s t a l l e d . It  vya^ driven from the top of the reactor by a 1/24  HP high  torque motor. Air for f l u i d i z i n g the bed was preheated in an electrical  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 b e d 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 .  reactor  45  1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21 j 22. 23. 24. 25. 26. 27 28. 29. 30. 31. 32. 33. 34. 3,5. 36.  Recycling Air Flowmeter Distributor Air Flowmeter Air Blowers Compressed Oxygen, Sulphur Dioxide and Nitrogen 12.5 cm Diameter x 1 m length Reactor 6 rpm Scraper Device Distributor P l a t e , 54 holes, 1 mm Diameter Recycling Nozzle, 3 holes, 2.5 mm Diameter Gas Calming Section of 1.25 cm ceramic beads Fluidizng Air Preheater, 7.5 cm Diameter x 45 cm length Recycling Air Preheater, 3.8 cm Diameter x 30 cm length Modified Venturi Nozzle for recycling and feeding the solids Discharge Flap Valve for Calcines Solenoid for operating the Flap Valve Automatic Adjustable Time for Solenoid Bin and 2.50 Diameter Screw Feeder for Molybdenite Adjustable Speed Motoreducer, 0.25 to 25 rpm Product Discharge Bin Compensating Pressure Chamber, 12.5 cm Diameter x 14 cm height Rotary Star Valves, 44 mm Diameter and 38 mm Diameter Main Cyclone, 4 cm Diameter Secondary Cyclone, 3.2 cm Diameter Coil Heated Pipes and Cyclones Gas Sampling Cooler for S0 Analyzer Sulphur Dioxide Infrared Analyzer Water Ejector Scrubber, 2.5 cm Diameter x 20 cm height Water Tank, 25-1 High Pressure Centrifugal Pump Exhaust Gases Chromel-Alumel Thermocouples Water Manometer Electromagnetic Vibrators S0 Analysis Recorder Slurry Tank, 1-1 A g i t a t o r , 200 rpm High Pressure Venturi for Slurry Feeding 2  2  Figure 4.4.  Legend.  46  47  f i l l e d with ceramic beads.  The gas distributor was made of  316 stainless s t e e 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 d i s -  charged through an expanded cone at the apex of each cyclone, to decrease re-entrainment  of s o l i d s .  To avoid back flow  of gases and s o l i d from the r e c i r c u l a t i n g system, a rotary valve system was designed, consisting of a stainless steel rotor rotating  against a Teflon lining  cooling proved to be s u f f i c i e n t  (Figure 4.6).  Air  to avoid softening of the  Teflon l i n i n g even at very high discharges rates, resulting reasonably long l i f e .  in  The valves were rotated at 6 rpm  (main cyclone) and 3 rpm (secondary cyclone), by means of a high-torque, 1/24  HP motors.  S o l i d s , discharged from the cyclones, were collected together  through a 1 inch pipe and passed continuously to  a distribution  valve.  This l a t t e r valve had an internal  flapper connected to a solenoid so that the s o l i d material could be directed either recycling system.  toward the product bin or the  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 pressureabsorbing chamber was connected to the upper section of the  48  Figure 4.6.  discharge valve. assembly.  Cyclone discharge rotary  valve.  The system is shown in Figure 4.7  A l l components were made of  and insulated externally  with  0.6 cm  before  316 stainless s t e e l , ceramic wool to avoid  heat losses. The calcines to be recycled to the reactor were discharged by the flap valve to a s p e c i a l l y 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  The venturi  insulated with ceramic wool.  nozzle  system operated with a i r which was preheated in a separate e l e c t r i c preheater  similar to that used for the  fluidizing  Figure 4 . 7 .  Exploded view of r e c i r c u l a t i o n and discharge system. 1. 2. 3. 4. 5. 6. 7.  D i scha rge b e l l P r e s s u r e chamber Distribution valve D i v i d e d d i s c h a r g e cone Venturi nozzle Screw f e e d e r inlet F l a p p e r v a l v e arm  50  Figure 4.8.  Pneumatically operated v e n t u r i nozzle connected to the discharge valve (without i n s u l a t i o n ) . 1. 2. 3. k. 5. 6.  Venturi nozzle Secondary a i r loop A i r preheate r Recycled calcines pipe Discharged calcines pipe Divided discharge cone  51  air.  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 d i s t r i b u t o r and bends upwards at the center of the reactor, ending 15 cm above the d i s t r i b u t o r .  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 finally  discharged into the f l u i d i z e d 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 v e r t i c a l , 4 rpm rotary arch-breaker was i n s t a l l e d . 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 recirculated to the bed.  nozzle to be  A 300 watt electromagnetic  vibrator was clamped to the discharge system, providing a smooth, continuous flow of solids through the r e c i r c u l a t i o n system.  ure  4 . 9 .  Gas d i s t r i b u t i o n 1. 2. 3.  grid  Gas distributor Dispersion nozzle Thermocouples  and d i s p e r s i o n  nozzl  53  Two rotary  vane blowers provided air to the  f l u i d i z e d bed and the recycling system. measured with two flowmeters.  Air flows were  Compressed oxygen, sulphur  dioxide and nitrogen, when required, were connected to the r e c i r c u l a t i o n l i n e , and metered through gas flowmeters. To avoid condensation of gaseous lower oxides of molybdenum, a l 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.  All  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 ejector  water tank.  The water  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 content in the gases. Were f i n a l l y  From the water scrubber the gases  scrubbed with a solution of sodium carbonate  to absorb the S 0 . 2  The aqueous solution was continuously  2  Figure  4.10.  1 2 . 5 cm d i a m e t e r f l u i d i z e d b e d r e a c t o r r e c i r c u l a t i n g system d u r i n g operation.  and  55  recirculated by means of a small centrifugal  pump connected  to a reservoir tank containing 150 1 of s o l u t i o n . A view -of the 12.5cm reactor and r e c i r c u l a t i n g system is shown in Figure 4.10 and an overall view of the p i l o t plant in Figure  4.3  4.11.  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 Mo0 was prepared at the desired composition prior to 3  being charged into the bed.  A high speed camera and a 35 mm  camera with high speed f i l m , using angular illumination, was employed to record the instantaneous size and distribution ofbubblesinthebed. For p a r t i c l e distribution measurements, a tracer of 2 to 5 gr of MoS was added to a tracer injection device 2  (Figure 4.1).  After a free bubbling period of the f l u i d i z e d  bed the recycling gas was diverted to the tracer for 3 seconds by means of a 3-way valve. then shut off simultaneously.  injector  A l l gas flows were  After the t e s t , samples of  the sand-calcines mixture were taken every 2 cm from the distributor to the top of the bed, screened to separate the  Figure  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 plant f o r molybdenite roasting.  pilot  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 s o l i d d i s t r i b u t i o n along the bed, as well as the p a r t i c l e size d i s t r i b u t i o n . 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 l u i d i z e d bed.  4.3.2  P i l o t Plant Tests Gas tracer studies were performed at 510-520°C  to reproduce the actual conditions of operation. input of S0  2  A delta  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 conditions.  The downstream detecting point was located at the  top of the reactor.  Glass wool f i l t e r s  were used to prevent  the entrance of solids to the analyzer during the t e s t s . Solid residence time d i s t r i b u t i o n tests were performed with, the aid of a special injector which could  58  feed 20 gr of tracer MoS d i r e c t l y to the discharge valve 2  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 s o l i d product were then taken p e r i o d i c a l l y  for analysis while the rest of the calcines were kept continuously r e c i r c u l a t i n g . Continuous experiments were also performed to test the operation and f e a s i b i l i t y of the p i l o t plant process at steady state.  Fpr these experiments, a charge consisting  of coarse s i l i c a sand and previously roasted c a l c i n e s , 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 a t t r i t i o n effect overcoming possible sintering problems, and r e l a t i v e l y smooth f l u i d i z a t i o n 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 v e l o c i t i e s , the sand was not elutriated with the c a l c i n e s .  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 a n a l y s i s . During the t e s t s , steady state conditions were checked continuously by the temperature the S0  2  recording and by monitoring  analysis on a recorder coupled to the S0 At the completion of the t e s t ,  2  analyzer.  the total calcines  from the bed were discharged and weighed, as were the total calcines roasted during the t e s t .  The scrubber tank was  also discharged and the solids weighed after f i l t e r i n g drying,  and  F i l t e r e d l i q u i d was p e r i o d i c a l l y analyzed for  molybdenum. E l u t r i a t i o n tests were conducted during or at the end of roasting experiments by taking samples from the d i s charge bin during 1 to 5 minutes. Slurry feeding to the reactor was investigated a single test.  A small , high pressure venturi  in  nozzle and a  pinch valve were used to control the drop-by-drop flow of slurry from an agitated operated the venturi  tank to the bed.  Compressed air  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, d i r e c t l y 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 l u i d i z i n g conditions of  sand-Mo0  3  calcine mixtures, a study was conducted in the  three-dimensional, plexiglas f l u i d i z e d bed.  The mixtures used  and conditions of the experiments are l i s t e d in Tables  5.1  and 5.2 as well as in Appendix 5. The volume of the s t a t i c beds containing mixtures of Mo0 calcines with 1 Kg of sand was measured. 3  were f i r s t  The beds  packed with an electromagnetic vibrator until  volume reading became constant. are plotted in Figure 5.1.  The experimental  the  results  The s t a t i c bed height of the  sand-calcine mixture is given by  Lm = i f.  "m,sil  61  (cm)  (5.1)  62  Figure 5.1  Factor of v o l u m e t r i c i n c r e a s e of a s t a t i c bed of sand as a f u n c t i o n of the wt-% of c a l c i n e s in the m i x t u r e .  63  Table  5.1  Minimum F I u i d i z a t i o n  3-inch  diameter,  Silica  sand -40/+140 mesh;  Mo0  3  wt-%  three-dimensional f 1 u i d i z e d bed d  c a l c i n e s -325 mesh;  gas  si 1  *cal  c a l c i n e s in sand:  Un, superficial  Experiments  = 0.028  cm  = 0.001  cm  0 to 1 00  velocity:  0 to 30 cm/sec  FI ui di zi ng gas :  Air,  Table  25°C,  1 atm  5.2  Density of Pure and Bulk S i l i c a Sand and C a l c i n e s  p , Density (gr/cm )  e » Voidage ( f r a c t i on)  2.65  -  1 .58  0.404  4.70  -  1 .24  0.724  3  Silica  sand  (quartz)  Silica  sand, b u l k ,  -40/+140 mesh  M0O3  C a l c i n e s of Mo0 , 3  bulk,  -325 mesh  0  64  where  L  .,  is the height of the s t a t i c sand bed, and f.  is the volumetric expansion factor. The following expressions were applied to calculate the average p a r t i c l e s i z e , density and voidage:  A  Ps  £  =  *s11  X  sil  +  *cal  X  cal  =  Psil  X  sil  +  Peal  X  cal  o = e  An alternative  X  0 > s 1 l  S 1 1  +  £  0  j  C  a  l  ^  ^  X  r  '  (5.2)  c  m  ^  <-> 5  3  (5.4)  c a l  relationship often recomended for the calcu-  lation of the average p a r t i c l e size based on the volume/ surface ratio  [85],  's  =  X si...1 \ x cal, d . • d , sil cal  ( - ) 5  +  could not be used since i t gave erroneous values of the minimum f l u i d i z a t i o n  velocity.  For example, even at 20  wt-% of c a l c i n e s , the calculated value for  u ^ m  was 0.035  cm/sec, compared with 5 cm/sec measured experimentally.  5  Figure 5.2.  Pressure drop through the bed as a function of the s u p e r f i c i a l gas v e l o c i t y .  66  The c h a r a c t e r i s t i c  curves of pressure drop  the f l u i d i z e d  bed as a f u n c t i o n  velocity,  the d i f f e r e n t  for  are given in Figure 5.2.  of the s u p e r f i c i a l  gas  s a n d - c a l c i n e mixtures  In  all  through  tested,  c a s e s , with the  exception  of the pure c a l c i n e s bed, a constant weight of sand ( 1 . 5 Kg) was used, and the weight of c a l c i n e s was v a r i e d . these c o n d i t i o n s , a p r o g r e s s i v e i n c r e a s e of i n c r e a s i n g weight f r a c t i o n  Under  AP  with  of c a l c i n e can be seen.  The  unusual shape of the curve f o r pure c a l c i n e s i s probably due to severe channeling which was observed even at very superficial  gas v e l o c i t i e s .  a f t e r a few minutes  Particles  of o p e r a t i o n ,  apparently  the humidity  of the a i r ,  The humidity  problem does not e x i s t ,  fluidized all  bed r e a c t o r .  began to  low  agglomerate  as a r e s u l t  and p o s s i b l y e l e c t r o s t a t i c of c o u r s e , in  effects. the  The unusual curves were observed in  cases where the c a l c i n e content exceeded 50 wt-%.  be seen in Figure 5.2,  at  50 and 60 wt-%  pressure drop decreases at due to the onset of  plotted  u^,  high s u p e r f i c i a l  values of the minimum  were determined  in Figure 5.3.  calculated  calcines gas  As can  the  velocities  agglomeration.  Experimental velocity,  of  values of u ^  by Kunii and Levenspiel  fluidization  from Figure 5.2,  and are  Also presented in Figure 5.3  are  obtained from the expression given [85]:  CD  Wt. % calcines Figure 5.3.  Minimum f l u i d i z i n g velocity for mixtures of sand and c a l c i n e s .  68 _ 2  mf  U  =  d Pc P, "s 1650 u ' '  The agreement between experimental u  m f  is only f a i r  "Re  q  (5,  < 20  and calculated values for  between 0 and roughly 60 wt-% c a l c i n e s , and  poor above this range.  This discrepancy may result from  the fact that the empirical equation does not take into consideration e l e c t r o s t a t i c attraction which as shown by Baerns [87], It  between small  can be s i g n i f i c a n t l y  is also possible that the interaction  between  particles large.  particles  in this system involving a bimodal p a r t i c l e size d i s t r i b u tion was responsible for the observed difference measured and calculated u ^.••The experimental  between  results  suggest that the smaller particles  (Mo0 ) play a more  important  characteristics of the  role in the f l u i d i z a t i o n  bed than the coarse particles  3  (sand).  From Figure 5.3,  can be seen that the measured values of u up to about 50 wt-%;  m f  above this value the u  it  drop rapidly 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 l a t t e r values for calculation purposes on the bed reactor.  fluidized  69  5.1.2  F l u i d i z e d Bed Expansion Direct  the f l u i d i z e d velocity. 5.4,  measurements were made of the height  bed as a f u n c t i o n  The r e s u l t s  where i t  of the s u p e r f i c i a l  of these t e s t s  can be seen that above u  are given i n 0  be the consequence of s l u g  with u . 0  bed height  from Figure 5.4  by the f o l l o w i n g  L  m f  -  From Figure 5.4, f u  = Lf/Ljpf m f  for  The bed height computed by the  the  T h i s may  can be r e l a t e d  at  to the  minimum static  relationship:  1.048  L  m  (cm)  the bed expansion  superficial  was c a l c u l a t e d .  Figure  formation.  The measured values of the bed height fluidization  gas  ~ 25 c m / s e c ,  expansion of the bed i s no longer l i n e a r  of  gas v e l o c i t i e s  Values are p l o t t e d  under f l u i d i z e d  (5  factor, greater  in Figure  than 5.5.  c o n d i t i o n s then can be  relationship  The voidage of the bed at minimum c o n d i t i o n s was c a l c u l a t e d as  follows:  fluidization  25  20 Lm  E  f  o  AA-A-A_A  CD  15 h Lm  •o M  S  10 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, 0  5  10 u  0  15  , Superficial  ga:; velocity  20  25  30  (cm/.sec) o  Figure  5.4.  Fluidized  bed  height  as  a  function  of  the  superficial  gas  velocity  —•  k  1—:  -i  —  r  r  silica s a n d - 4 0 / 4 - 1 4 0 mesh calcines - 3 2 5 mesh air 2 5 °C , I atm. wt. % calcines o 10  u  0  , Superficial  gas  velocity  Figure 5.5. Bed expansion factor for the f l u i d i z e d bed as a function of the superficial gas v e l o c i t y .  72  e  m f  = 1.048  where, as shown in Eq. 5.7,  e  (5.9)  0  the factor 1.048 was the volumetric  ratio between the bed at minimum f l u i d i z i n g conditions and the s t a t i c bed.  5.1.3  Pressure Drop and Temperature At 25°C, the pressure drop through the f l u i d i z e d  bed (u  0  > u ) mf  Figure 5.2.  remained approximately constant as shown in  However, a sharp decrease in the measured value  of AP was always observed in the f l u i d i z e d 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 t e r a t u r e , is possibly caused by an increase in the "bed f l u i d i t y . " by M i i , Yoshida and Kuni [88] value of u  m f  This has been suggested  who found a decrease in the  with increases in the bed temperature.  Channeling  in the bed also may play an important r o l e , but no direct evidence was obtained which could substantiate this p o s s i b i l i t y . From the experimental  data obtained, between 150  and 550°C the pressure drop through the f l u i d i z e d bed decreases, on the average, by 10 cm of water for every 100°C increase in  temperature.  73 60i  O CM  X  •  e  Of  A  50 CD X)  V  o> Nl  40  O  A  A A ® A T  ® 2  •  30 j  ®  •  CL  o  ZS-zv Test no. o B 64 A B 68 B 69 ^ B 70 V B 71  CD to  20  CD  O O t-  10  0  0  •  B B ® B ^Q, B & B  73 105 114 116 119  cm/sec  •  JL  200  _L  400 Temperature  Figure 5 . 6 .  600  of bed (°C)  Pressure drop through the f l u i d i z e d bed in the p i l o t 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 l u i d i z e d bed, experiments were conducted in the 5-inch, two-dimensional model.  Flash pictures were  taken in each test (see Chapter 4, Figure 4.1) bubble diameter measured from the pictures.  and the  The conditions  used in the experiments are given in Table  5.3.  Typical pictures for different compositions of the bed and s u p e r f i c i a l gas v e l o c i t i e s are shown in Figure 5.7.  The adherence of the fine Mo0  3  cover of the f l u i d i z e d bed made i t clearly the bubble diameter.  calcines to the p l a s t i c  difficult  to  identify  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 l u i d i z e d bed.  measurements are shown in Figure 5.8.  Typical bubble Average measured values  of the bubble diameter are plotted in Figure 5.9  (a)  to  (d)  and given in Appendix 6. From the experimental Figure 5.9,  it  values of d  b  given in  can be concluded that, despite the  fairly  large s c a t t e r , the bubble diameter is not a function of the average size of the particles in the bed. with the results of Geldart [89]. a few smaller values of d  h  This finding agrees  When pure sand was used,  were measured probably due to  Figure 5.7.  Typical high speed pictures of two-dimensional f l u i d i z e d bed model .  76  Figure 5.8.  Typical bubble measurements made on an "equivalent bubble diameter."  77  Table  5.3  Bubble Diameter Measurements in  12.5 cm  Two-Dimensional Fluidized Bed Gas: L m:  A i r , 25°C, 1 atm 28 ± 2 cm  High speed film ASA 400, f . 1 . 8 , 1/1000 sec  Run  Calci nes  No.  (wt-%)  B.106-A.1  f  u  (1/min)  (cm/sec)  G  0  0  20  20  A.2  n  29  30  A.3  II  38  40  20  14  15  II II  20  20  •I  29  30  II  38  40  B.106-B.1 B.2 B.3 B.4 B.106-C.1 C.2 C.3  40  14  15  II  20  20  M  29  30  38  C.4 B.106-D.1 D. 2 D.3  60  . 14  40 15  n  29 29  20 30  38  40  II  D.4 D.5  "  6  10  B.106-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  5  <L>  E  o  u = 10 cm/sec Run no. wt. % calcines A B-I06-D-5 60 • B-I06-E-1 80  (a)  0  — X2  o> O i_  CU  >  < l-o  0  (b)  u = 20 cm/sec Run no. wt. % calcines E o o B-I06-AI 0 Cl> A BI06-B-2 20 OJ |._ ^ g. ir>^_r.o 40 t o A B- I06-D2 "O • B- I06-E-3 Q> 4 0  X) X)  x>  3  OJ  o>  O i_ CD >  21  0 0  10 l  Figure 5.9.  20 f  , Fluidized  30  JL  40  bed height (cm)  Bubble diameters measured in the f1ui di zed bed model.  two-dimensional  u  0  = 30  Run 7 E 6  E o  5  T  cm/sec  np_.  wt. %  O BI0S-A-3  0  A  BI06-B-3  20  ©  B-I06-C-3  A  B-I06-D-3  •  B106-E-4  T  " i — — r (c)  colcines  (d)  u  y A  0  = 40  cm/sec  Run no. O B-I06-A-4  0  10 f  Figure  ,  0  BI06-B-4  20  ©  BI06-C-4  40  A  B-I06-D-4  60  •  B-I06-E-4  80  30  Fluidized 5.9  calcines  A  20 l  wt. %  bed  40 height  (Continued)  (cm)  50  80  more gas (from 50 to 25%)being required for f l u i d i z a t i o n and proportionately  alone  less for bubble formation over the range  of v e l o c i t i e s studied (20 to 40 cm/sec). From the measured values of d , a b e s t - f i t  line  b  was traced for each case.  No attempt was made however to  treat the data obtained at u scatter.  0  = 15 cm/sec due to the large  The data could be f i t t e d by the following  relation-  ship.  d  where d l.f  b  b  = 0.00265  l  f  'mf  (cm)  (5.10)  is the diameter of the bubble in cm at the distance  (cm) from the d i s t r i b u t o r g r i d ; N^ is the number of holes  per unit area in the d i s t r i b u t o r and u cm/sec.  0  and u  m f  are in  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 experimental 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 for lower values of u . 0  data, but is poor  This discrepancy does not  the usefulness of Eq. 5.10  in calculation for the  affect fluidized  bed reactor however since the range of s u p e r f i c i a l gas v e l o c i t i e s 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 l u i d i z e d 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 l u i d i z e d bed model, to a threedimensional reactor, i t must be assumed that wall in the model are not excessive.  effects  This assumption cannot be  v e r i f i e d in this work but is necessary due to the  difficulty  of measuring bubble diameters in a three-dimensional  fluidized  bed. Equation 5.10 can be compared with the empirical expressions found in the  (a)  literature:  Kobayashi, Arai and Chiba  d. = 1.4-1- d p b f s ys  (b)  1  Geldart  ( d  b  I  h  mf J  Uo  (5.11)  ^ m f  [91]  >  = 1.43  [90]:  0.4  1 19  + 0.027 l J  f  (u  0  - u ) mf  (5.12)  83  (c)  Chiba, Terashima and Kobayashi  i 0.  d  b  "  d  b  1 .25 0  The calculated values of d  b  [92]:  286  + 1  (5.  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 l u i d i z e d beds larger than about 1 m in diameter. The lack of agreement with the relation proposed by Kobayashi et al. suggest that the l a t t e r equation is in gross error.  At only 9 . 6 cm from the d i s t r i b u t o r  their  equation predicts a bubble diameter that is equal to the width of the f l u i d i z e d bed. observed.  Such conditions were never  84  Figure  5.11.  Bubble diameter c a l c u l a t e d empirical expressions.  using  different  Chapter 6  PARTICLE BEHAVIOR IN TWO- AND THREE-DIMENSIONAL FLUIDIZED BEDS  6 .1 P a r t i c l e  Stratification  From early tests performed in the three-dimensional f l u i d i z e d bed consisting of mixtures of calcines and a narrow size range of sand, i t became clear that s i g n i f i c a n t p a r t i c l e size  stratification  was further  could occur during f l u i d i z a t i o n .  It  recognized that this phenomenon could present  problems in the f l u i d i z e d bed reactor since the hoped for a t t r i t i o n effect of the sand, that i s , to prevent sticking of the c a l c i n e s , and the smooth f l u i d i z a t i o n  properties  (expected with a homogeneous mixture of calcines and sand) might not be r e a l i z e d . In order to determine the operating conditions under which  stratification  experiments were undertaken.  could be avoided, a number of An experimental  approach was  necessary due to the lack of information in the l i t e r a t u r e 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 p a r t i c l e  size d i s t r i b u t i o n , bed composition and s u p e r f i c i a l gas velocity on the extent of s t r a t i f i c a t i o n  has been assessed.  experiments were carried out in the two-dimensional  The fluidized  bed.  6.1.1  Effect of Particle Size Distribution S t r a t i f i cation Two ranges of size d i s t r i b u t i o n 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  I n i t i a l composition of mixture = 20 wt-% c a l c i n e s , -325 mesh Superficial gas v e l o c i t y , Fluidization time,  u  0  =30 cm/sec  t = 3 min  Test B.l Sand (-40/+70) mesh, 520 g Test B.95 Sand (-40/+140) mesh, 520 g  Prior to each t e s t ,  the bed was charged with a homogeneous  mixture of sand and calcines and f l u i d i z e d for three minutes.  87  lOOp  1  Bl  B-95 80  o o  1  1  r  1  = 30 cm/sec initial : 2 0 wt. % c a l c i n e s 3 min f l u i d i z a t i o n  90  tn tu c  1  © sand - 3 0 / + 7 0 mesh O - 4 0 / + 140 "  I top of b e d  70  60  I  •o c  °  501  a> 4 0 o o  30 I  5  initial composition  20  K  ho  j_o-o-o-o-"i o-=o—-—O—O „  O  ^_  p t o p  o f b e d  |  o-  I0h  10 \  Figure 6.1.  w  , Distance  20 from  the distributor  J  I  L  30  (cm)  Particle size segregation in the f l u i d i z e d bed as a function of the p a r t i c l e size d i s t r i b u t i o n of the sand.  88  The axial  size d i s t r i b u t i o n of s o l i d s , expressed  in terms of weight per cent of calcines versus f l u i d i z e d bed height, after  three minutes of f l u i d i z a t i o n , is shown  in Figure 6.1 for the two size ranges of sand.  It  can be  seen that for the narrow size d i s t r i b u t i o n 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 l a t t e r case, no segregation  exists in the upper two-thirds of the f l u i d i z e d bed, whereas in the lower one-third of the bed, a gradient in the calcine concentration towards the d i s t r i b u t o r is evident.  A small  increase in the calcines fraction can also be seen adjacent to the d i s t r i b u t o r .  This is due to the fact that some material  remains between the holes of the d i s t r i b u t o r where is absent.  The mechanism of s t r a t i f i c a t i o n  fluidization  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 c a l c i n e s , -325 mesh, in which samples were f l u i d i z e d 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  Figure 6.2.  m  ,  Distance from  the distributor (cm)  Particle size segregation in the f l u i d i z e d bed as a function of the f l u i d i z a t i o n time.  90  Table  6.2  Particle Size S t r a t i f i c a t i o n  as a Function of Time  I n i t i a l composition of mixture = 20 wt-% calcines -325 mesh Superficial gas v e l o c i t y , u = 30 cm/sec Sand A -40/+70 mesh, 520 g 0  Test No  Fluidization time t (sec)  Total gas flow G (1) f  Rate of Fines Transport AM/At  AM/AG  (g/sec)  f  (g/cc gas)  B.4  15  6.21  2.989  7.22 x IO"  B.2  30  1 2.42  1 .416  3.42 x IO"  3  B.l  180  74.52  0.472  1.14  3  x 10~  3  The results in Table 6.2 and Figure 6.2 show c l e a r l y that the s t r a t i f i c a t i o n  of fines in this bimodal system of  particles with a large difference in p a r t i c l e size is a very rapid phenomenon, being v i r t u a l l y  completed in three minutes.  The time-averaged rate of fines transport  through  the f l u i d i z e d 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.  can be written as follows:  Adopting this procedure, the rate  91  AM At = S  fl 'f 'cal t I J 1  p.  f  0.1  'f 0  t  I  fl cal  (6.1)  (gr/cm )  (6.2)  J  while, with respect to the gas flow i t  AM AG,-  (gr/sec)  is  0.1  3  M is the weight of fines (g) transported in the time from the d i s t r i b u t o r grid to the bed height, 1^, the composition is equal to the i n i t i a l X  cal '  o  f  calcines due to the passage of  weight  (sec)  (cm), where  fraction,  G cm of gas. 3  f  is the cross-sectional area of the two-dimensional "p  t  S  bed (cm ); 2  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 l u i d i z e d bed decreases with time as the lower section of the f l u i d i z e d bed becomes depleted of f i n e s .  The value calculated for  (AM/AG ) are probably f  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° -j , the rate of a  transport  may become constant for a given size of fines in the bed.  92  6.1.2  Effect of Bed Composition on S t r a t i f i c a t i o n  A series of tests was carried out to determine influence of bed composition on the s t r a t i f i c i a t i o n using the sand of wide size range, -40/+140 mesh. different  initial  were studied.  of s o l i d s , Three  compositions of the sand-calcine  mixtures  Results of three minute s t r a t i f i c a t i o n  for beds consisting i n i t i a l l y are shown in Figure 6.3.  tests  of 20, 40 and 60 wt-% calcines  The s t r a t i f i c a t i o n  are similar for the different  the  p r o f i l e s obtained  compositions tested.  These  results indicate that for compositions between 20 and 60 wt-% c a l c i n e s , the bed appears to maintain a constant composition with time.  This assumes that the three minutes of  continuous f l u i d i z a t i o n  are s u f f i c i e n t to achieve a steady  state of mixing, as indicated by the depleted zone in the three minute  curve of Figure  The experimental  6.2.  results given in Figure 6.3  indicate  that an axial gradient in the fines concentration of approximately 0.7  g/cm could be expected in the f l u i d i z e d bed  reactor.  6.1.3  Effect of the Superficial Gas Velocity on Stratification  Stratification different  tests were also performed at  three  s u p e r f i c i a l gas v e l o c i t i e s using a bed containing  93  r  1  T  1  i  1  1  r——i  r  i  1  1  1  r  u = 3 0 cm/sec 3 min fluidization silica sand - 4 0 / + 1 4 0 mesh B 6 4 • 6 0 wt. % calcines initially B94 A 40 0  80!  70  i  B 95  O  20  •  E  'U  initial composition  c D O T>  C  o  CO  cn <U c  o o  1  L m  Figure  6.3.  20  0  , Distance  from  30  the distributor (cm)  P a r t i c l e size segregation in the f l u i d i z e d bed as a function of the i n i t i a l composition of the sand-calcines mixtures.  94  60 wt-% c a l c i n e s . it  The results are shown in Figure 6.4.  can be seen that s t r a t i f i c a t i o n  Here  of the particles decreases  with increasing s u p e r f i c i a l gas v e l o c i t y .  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 r i s i n g  gas bubbles, due to gas r e c i r c u l a t i o n . to be important,  For this mechanism  however, the gas r e c i r c u l a t i o n  inside the bubbles must s i g n i f i c a n t l y exceed the  velocity terminal  velocity of the calcine p a r t i c l e s , which is about 0.25  cm/sec.  This condition holds for the experiments reported here since the gas r e c i r c u l a t i o n v e l o c i t y , 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 p a r t i c l e s .  It  is  95  100  T — i — i — r  " i — r  o. o-  -0—O—o—o  Z3 X  . A  e • A — A — A — A C  o  t  Ifl  nop of bed  CO  c  o o  U5  c  o o  6 0 wt. % calcines initially 3 min fluidization sand - 4 0 / + I40 mesh B 65 O u = 10 cm/sec B 93 A u = 20 B 64 • u = 30 0  0  0  0  10 ^  Figure 6 . 4 .  J 20  L  J  L  30  J  L  , Distance from the distributor (cm)  P a r t i c l e size segregation in the f l u i d i z e d bed as a function of the s u p e r f i c i a l gas v e l o c i t y .  96  important to note that the direction of gas flow in this l a t t e r case may be either upward or downward, depending on the magnitude of the s u p e r f i c i a l gas velocity or Uo/u .  Kunii and Levenspiel [96]  m f  is downward when critical  u /u 0  m f  alternatively,  have shown that the flow  exceeds 6-11,  and upward below this  range. In the l i g h t of this knowledge, i t would appear  certain that the observed s t r a t i f i c a t i o n partially,  is related at least  to the third mechanism of transport, that i s , the  direction of gas flow through the emulsion. evidence to this effect  is the finding that marked s t r a t i f i c a -  tion occurs experimentally for with the c r i t i c a l  The strongest  u /u 0  < 8-10  m f  which coincides  ratio determined by Kunii and Levenspiel.  Thus at low flow rates when  u /u 0  m f  < 8-10, the flow of gas  through the emulsion is upward, assisting the upward transport of fine p a r t i c l e s (and s t r a t i f i c a t i o n ) the gas bubbles and their wakes. until  the ratio  u /u 0  m f  If  already provided by  the gas flow is increased  , exceeds 8-10 the gas flow through  the emulsion is reversed to the downward d i r e c t i o n , thereby enhancing the return of fine p a r t i c l e s 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  u  mf  ; as the  u  m-c  97  increases more gas w i l 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 t e s t s , i t was decided to use -40/+140 mesh sand in the f l u i d i z e d bed reactor and work with a s u p e r f i c i a l gas velocity greater than 20 cm/sec to minimize  6.2  stratification.  Concentration P r o f i l e of Reacting Solids in the Fluidized Bed The special type of pneumatic injection system used  in the p i l o t f l u i d i z e d 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 e a r l i e r ,  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 i n j e c tion system, knowledge of the efficiency of p a r t i c l e d i s t r i bution in the bed was considered to be e s s e n t i a l .  Since no  information of this kind could be found in the l i t e r a t u r e a series of experiments was conducted in the two-dimensional f l u i d i z e d bed.  The experiments had a two-fold purpose:  to  determine how the p a r t i c l e s in the feed were distributed by  98  the injection system and to measure the axial  concentration  p r o f i l e of the molybdenite in the f l u i d i z e d bed.  6.2.1  Axial Concentration P r o f i l e Measurements In these tests a delta input of MoS tracer of 2  the same particle size as the calcines (-325  mesh) was i n -  jected into the f l u i d i z e d bed through a dispersion nozzle (see Chapter 4, Figure 4.1).  As in the previous t e s t s , the  f l u i d i z e d bed consisted of a mixture of Mo0 calcines and 3  sand.  After injecting the tracer,  "frozen" by turning off a l l sand-calcine  the bed was immediately  the gas flows.  The mixture of  was subsequently analyzed for sulphur,  MoS , along the bed. 2  i.e.,  A photograph of the f l u i d i z e d 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 d i s t r i b u t i o n calcines (-325  mesh).  (-40/+70 mesh) and  The r e s u l t s , which are given in  Figure 6.6 show that the transference of MoS from the gas 2  jet  to the calcine-sand emulsion is very rapid, with about  50% of the tracer being transferred in the f i r s t the injection point.  In Figure 6.6 i t  5 cm from  can also be seen  that the concentration of MoS tracer drops to nearly zero 2  close to the surface of the bed.  This indicates that a l l  the MoS. that was injected into the bed as a pulse has 2  Figure 6.5.  Typical tracer test in the two-dimensional f l u i d i z e d bed after an input of MoS tracer (black) into the mixture of M o 0 and sand (whi te) . 2  3  100  remained in the bed and not been blown d i r e c t l y out inside the gas bubbles. The results also indicate that the transfer s o l i d becomes effective  of  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 Two additional  distributor.  tests were conducted to  determine  the influence of the position of the injection point above the distributor  on the concentration p r o f i l e of MoS using 2  a wide size d i s t r i b u t i o n  sand (-40/+140 mesh).  The  injection  point in these tests was located 12 cm above the gas d i s t r i b u t o r , as compared to a position 1 cm above the d i s t r i b u t o r  of the  previously described t e s t s . 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 s i m i l a r .  However,  in the l a t t e r case the calcines are distributed more over the whole bed and are present in s i g n i f i c a n t  concentrations  up to near the bed surface.  no appreciable  From Figure 6.7  difference in the shape of the concentration p r o f i l e of MoS can be seen for the different weight fraction of calcines 2  in the bed and the different s u p e r f i c i a l gas v e l o c i t i e s  tested.  The presence of tracer below the injection point is caused  101  20,  2 0 wt. % calcines 3 0 cm / s e c 3 sec tracer input :  CM  CO  c o D  v_ C CD O C  o o  CL) O O  10 l , Fluidized f  Figure 6.6.  20 bed  height  30 (cm)  Axial concentration p r o f i l e after a MoS tracer input.  2  1 0 2  r-  1  r  1  B 9 4 © 40wt.% B65 O 60 "  l Figure 6.7.  f  , Fluidized  bed  1  — r -  calcines,u =30cm/sec " u = 10 "  height  0  0  (cm)  Influence of the position of the dispersion nozzle on the concentration p r o f i l e of MoS in the f l u i d i z e d bed. 2  103  by the downflow of emulsion due to displacement by r i s i n g bubbles.  As expected, the concentration p r o f i l e from the  injection point to the top of the bed is different than that from the injection point down to the d i s t r i b u t o r .  This  is apparently produced by the upward transport of s o l i d (MoS in this case) along the f l u i d i z e d bed as suspended 2  s o l i d 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 It  can be seen that in a l l  plotted in Figure 6.8.  cases, most of the tracer trans-  ference to the emulsion takes place near the injection point, but a more even d i s t r i b u t i o n of the tracer is achieved when the injection point is further above the distributor l e v e l . For this reason, the injector in the f l u i d i z e d bed reactor was positioned at the higher level above the d i s t r i b u t o r .  6.3  E l u t r i a t i o n of Calcines from the Fluidized Bed The basic p r i n c i p l e of the r e c i r c u l a t i n g f l u i d i z e d  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. this e l u t r i a t i o n  The driving force for  is the difference between the operating  velocity and the terminal  velocity of the calcine p a r t i c l e s .  T  B B B B  T  i  2 3 65 94  ^ o o A  i  p — i 1—~r—i  Figure 6.8.  1  1  1  1  1  "1 injection point at J distributor level 1 injection point at J. 12 cm from distributor  l , Distance m  j  from  the distributor  Cummulative percent of MoS along the f l u i d i z e d bed.  2  (cm)  transferred  r  705  The terminal  velocity of the calcines can be c a l -  from the relation given by Kunii and Levenspiel [ 9 7 ]  .lated  0.5  4  g d, 3p 9 w  where from  C the  (cm/sec)  C . d  (6.6)  is t h e drag coefficient which can be calculated  d  velocity  independent group  3 P  o P  particle  4 q d  p  d  Reynolds  s  '  p  g, ^  (6.7)  3y  number  is g i v e n b y t h e relationship  Re -ILfaJit p  The a v e r a g e p a r t i c l e Reynolds tven  number  s  , terminal v e l o c i t y  and  of t h e molybdenite concentrates and sand are  i n T a b l e 6.3.  The d e t a i l e d  tu'- :;n t o be t h e s i z e  Sires  d~  diameter  a.-e p r e s e n t e d i n A p p e n d i x is  (6.8)  u  .'ere d e t e r m i n e d  7.  results of the size analysis  The average p a r t i c l e  through  by C o u l t e r  diameter  which 5 0 % of the solids pass. Counter and Cyclosizer analysis  1 06  Table 6.3 Particle Size and Terminal Velocity of Molybdenite Concentrates  Concentrate Endako (-325 Brenda  mesh) "  B.C. Moly Kennecott  "  Sand (-40/+140 mesh)  t (cm/sec) u  (cm) 0.8 x 10~  3  4 x 10"  2  0.34  1 .0 x 10"  3  7 x 10"  2  0.45  1 .2 x IO"  3  9 x IO"  2  0.60  2  1 .90  3.8 x IO"  3  3.1  2  x IO"  28 x 10690  At the s u p e r f i c i a l gas v e l o c i t i e s used in the p i l o t bed reactor  150  fluidized  (u-b = 15 to 35 cm/sec), only the calcines will  be elutriated  while the sand w i l l remain as a stable  bed.  For the sand the ratio  0.05,  whereas for the calcines  u /u 0  is between 0.021  t  u /u 0  t  fluidized and  is between 125 and  370.  6.3.1  E l u t r i a t i o n Flux In order to achieve a steady state mass balance  between feed and discharged product, an exact value of the elutriation  rate must be known.  At operating conditions in  the reactor, the ratio of e l u t r i a t i o n was 1arger than 1 .  rate to feed rate always  107  To determine the e l u t r i a t i o n f l u x , F , defined as the flow of calcines discharged from the top of the bed per unit are per unit time, the rate of e l u t r i a t i o n  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 e l u t r i a t i o n  rate  (g/sec) and  S  the cross-sectional area of the reactor. Table 6.4  E l u t r i a t i o n Tests Performed in the Fluidized Bed Reactor, 12.5 cm Diameter  Temp. (°C)  Uo  (cm/sec)  B.C. Moly calcines  330 - 580  13.8 -  33.2  Brenda  "  330 - 550  15.6 -  25.3  Kennecott  "  520 - 570  23.6 -  29.6  235 - 530  14.4 -  23.6  Endako  From the t e s t s , different tion flux  (or rate of e l u t r i a t i o n )  curves for the  elutria-  could be drawn for each of  the four molybdenite concentrates used.  It  can also be seen  that both p a r t i c l e size and s u p e r f i c i a l gas velocity profoundly influence the e l u t r i a t i o n  flux at a given value of u . 0  108 2.0  ~TTM  r  5 ^ reactor Kennecott ( I ) o B.C. M o l y ( 2 ) • Brenda (3) A Endako (4)  o  <D  §  i  3  l.5h  C7>  160  (4)  (3)  140  /o  reactor a B.C. Moly  120 ~ c  £ 'o  100  lux  X  1.0  C7>  O  8 0 a: c o  c o  o  60  8  i_  J3 UJ  P  UJ  - 0.5  40.  LL.  r 0  U  20  T  5  IQ ( u  o~ mf) u  Figure 6.9.  15 .  20  Excess  25  30  Gas Velocity  35 (cm/sec)  E l u t r i a t i o n rate and e l u t r i a t i o n flux as a function of the superficial gas v e l o c i t y .  40 0  109  For example, at u  0  = 25 cm/sec, for the f i n e r calcines of  Endako the e l u t r i a t i o n  flux i s 2.4 times larger than for the  coarser calcines of Kennecott.  The value of  k  appears to  depend exponentially on the s u p e r f i c i a l gas v e l o c i t y . be noted also that the curves in a l l  It should  cases begin at the  terminal  velocity of the p a r t i c l e s , between 0.04 and 0.28 cm/sec. To check on the probability that nentially with u -u 0  f  ,  shown in Figure 6.10.  k  varies expo-  a semi 1ogarithmic plot was drawn as A linear relation  is found for  all  four cases although a small deviation does exist at low values of u .  This may be due to the large scatter in the experi-  0  mental data.  Thus for each type of molybdenite concentrate  tested, the general  relationship  k* = A e  B ( u o  ' mf u  )  (g/min)  is found to hold for the rate of e l u t r i a t i q n . mental constants  A  and  are given in Table 6.5. Eq. (6.9) mental  B  (6.9)  The experi-  for the concentrates tested  The Calculated values of  k  using  agree well with the curves traced with the experi-  points, with an error of less than 5%.  no  10 9 8  7  «  i  1  !  1  1—•—i  r  ® Endako ^ Brenda o B.C. Moly Kennecott  20 Excess Figure 6.10.  1  T  25 gas  velocity  Semi-logarithmic plot of the rate of e l u t r i a t i o n as a function of the s u p e r f i c i a l gas v e l o c i t y .  Ill Table 6.5 E l u t r i a t i o n Constants of Eq. }  Calcines  (6.9)  -—__  - -  B (sec/cm)  A  (g/min) 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 e l u t r i a t i o n  flux of  fines from a bimodal system of p a r t i c l e s with a large  dif-  ference in size may be expressed as  p*  Probably, the temperature  _ A _B (u o-u  ~ S  .  f  )  m f  •-  (g/cm sec)  (6.10)  2  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 s u p e r f i c i a l gas v e l o c i t y ,  the e l u t r i a t i o n  F , decreases from about 15% at u - u ^ = 20  flux,  0  m  112  cm/sec to about 50% at (uo-u - ) = 30 cm/sec. f  Lewis, Gi 11 i 1.and  and Lang [ 9 8 ] found that the diameter of the reactor has no influence on the e l u t r i a t i o n  rate for d^ > 10 cm (4  but for smaller diameters, the e l u t r i a t i o n rapidly.  in)  rate decreases  Chapter 7  GAS AND SOLID DISTRIBUTION IN THE FLUIDIZED BED REACTOR  7.1  Gas Tracer Experiments Depending on the extent of mixing,  elements of gas may reside for different inside the reactor.  different  lengths of time  A measure of the d i s t r i b u t i o n of  residence times of gaseous elements in the reactor can be made by  means  experiment,  of tracer studies.  In this type of  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, concentration-time p r o f i l e is expressed as the age d i s t r i b u t i o n function, C(t).  the  internal  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 l u i d i z e d 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 d i s t r i b u t i o n function of gas in the reactor and the extent of gas dispersion in the axial  and radial  directions.  S 0 , acting as 2  tracer, was injected in the form of a pulse into the bed through the dispersion nozzle (see Figure 4.4, centred 15 cm above the d i s t r i b u t o r g r i d . centration of S0  2  Chapter  4),  The output con-  was then measured at the top of the reactor  and continuously recorded with an infrared analyzer coupled to a chart recorder.  Three different  s u p e r f i c i a l gas  v e l o c i t i e s , which covered the usual range of operating conditions, were tested. depicted in Figure 7.1  A typical recorded response is  and data is given in Appendix 9.  From the recorded concentration of the tracer the outflow of gas, the exit age d i s t r i b u t i o n function  in 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 f a c t , due to experimental d i f f i c u l t i e s the S0 was injected into the gas stream on the entry side of the recycle gas preheater. 2  115  Figure 7.1.  Response signal from S 0 pulse input recorded by the infrared analyzer. 2  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:  5 C 2  9 £ a t  e  f  a ,g  is the bed voidage, and D  (7.1)  the axial dispersion  g  c o e f f i c i e n t or eddy d i f f u s i v i t y which characterizes the degree of mixing a x i a l l y along the f l u i d i z e d bed. the dimensi oni ess terms, 9 = t / T = u t/L^. 0  equation (7.1)  can be transformed [100]  <LP_  80  The dimensionless group (D  and  Using  x = 1 ^/L^,  to the following:  .9  Ui  f e /u f  J  0  3  (7.2)  X2  L) f  is the Reactor  Dispersion Number or inverse Peclet number for mass transfer The mean value or centroid of d i s t r i b u t i o n , ip, is given by  tC(t)dt (7.3) C(t)dt  T — '  - i  '-detection point  r  — l  i  1—  6 0 wt. % calcines T = 5 1 0 - 5 2 0 °C  9 u = 24.4 cm/sec o u =26.1 A = 28.8  tracer injection"  u  gas  plug flow  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 o f t r a c e r gas i n t h e f l u i d i z e d bed as a f u n c t i o n o f d i m e n s i o n l e s s t i m e .  118  By calculating the variance of the d i s t r i b u t i o n of the tracer, o , using the relationship 2  ,00  t C(t)dt 2  (7 C(t) L  dt  the RDN for the reactor can be determined.  The  for a  closed vessel:  o  = 2(RDN) - 2(RDN)  o  1 - e' ( 1 / R D N )  (7,  while for an open vessel  cr  where  a  2  = o /t z  2  .  = 2 (RDN ) + 8(RDN )  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  the internal  should be noted further that for a pulse input age d i s t r i b u t i o n function  C(t)  of Equation 7.1  and 7.2 is equivalent to the exit age d i s t r i b u t i o n function E(t):  119  E(t)  = C(t) (7.6)  E(e) = c(e)  Since the injection point of the tracer was located 15 cm from the d i s t r i b u t o r g r i d , the reactor should be c l a s s i f i e d 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 37 wt-% sand -40/+140 mesh Tracer: 16.5 1 S0 Input time: 5 sec  concentrate)  2  Test No .  Temp.  B.l25-1  RDN  f ,T (1/min)  (cm/sec)  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  (°c)  G  Closed Vessel  Open Vessel  The values of RDN indicate that under both conditions considered, extensive gas mixing e x i s t s , approaching  CO  cn  0.50  12.5 reactor T= 510 - 5 2 0 °C 6 3 wt. % calcines o closed vessel A open vessel  _i  0.45 •  7.5 reactor T = 2 5 °C 15 wt. % calcines closed vessel  0.40. 15  20 u , Superficial 0  Figure  7.3  25 gas velocity  30 (cm/sec)  Reactor dispersion number of gas in the 7.5 cm and 12.5 cm diameter f l u i d i z e d bed reactors as a function of the s u p e r f i c i a l gas v e l o c i t y .  121  the backmixed c o n d i t i o n  (RDN = °°) r a t h e r  (RDN = 0 ) .  the s u p e r f i c i a l  In  addition  than plug  flow  gas v e l o c i t y  seems t o  have only a minor i n f l u e n c e on the value of RDN over the of u  0  tested  ships i t  (Figure  7.3).  In c o n s i d e r i n g these  range  relation-  should be recognized that the response s i g n a l may  have been a f f e c t e d  by three experimental  limitations:  I .  Some m i x i n g o f t h e t r a c e r a n d t h e a i r could occur i n the 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 into the 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 between t h e top o f t h e bed and t h e d e t e c t i o n p o i n t at the top of the reactor. (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 to locate the detector a t the surface of t h e bed.) Although t h e Reynolds Number f o r t h e g a s i n t h e f r e e b o a r d indicated a laminar flow c o n d i t i o n (Re. = .1000-1300 a t u = 1 5 - 3 5 c m / s e c ) 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 the gas i n t h e f r e e b o a r d . 0  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 the pulse input, 5 sec, i s r e l a t i v e l y long, approaching a step input. This would 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 a n d c o n s e q u e n t l y t h e RDN .  In  Figure 7.3  the c a l c u l a t e d values of RDN are  also compared with the RDN obtained using S 0  2  as t r a c e r .  in the  7.5  This comparison shows that i n  cases extensive gas mixing o c c u r s .  gas  cm r e a c t o r may p o s s i b l y be caused by the  of gas s l u g s .  both  The more r a p i d decrease  in the value of RDN with i n c r e a s i n g s u p e r f i c i a l f o r the 7.5  cm r e a c t o r  velocity formation  122  7.1.2  Gas Transfer Between Phases After the gas enters the f l u i d i z e d 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 coefficients,  K^ .j b(  and  b  K^ j c g  b  , respectively.  Then  the overall gas transfer c o e f f i c i e n t between the gas bubbles and the emulsion is given by the relationship  1  1  (be)b  N  (bc)b  K  1 (ce)b  [102]:  (1/sec)  (7.7)  The extent of the gas d i s t r i b u t i o n a x i a l l y along and r a d i a l l y across the f l u i d i z e d bed can be estimated from the axial and radial  dispersion coefficients  D a,g  and D r»g  which  are given by the following r e l a t i o n s h i p s , derived by Kunii and Levenspiel [103]:  rbi °i u  a ,g  Dr,g  l-b  =0.2  u  (be)b  b  (be)b  where b is a c o e f f i c i e n t given by  (cm /sec)  (7.8)  (cm /sec)  (7.9)  2  2  123  b = a  The f a c t o r  'mf  'mf  U  -  0  U  mf-'  (1 - 6 - a<5) + m a ( l - e  (7  )  For a non-adsorbing s o l i d  m = 0, whereas f o r a high s u r f a c e area c a t a l y s t - t y p e The use of expression (7.7)  of the o v e r a l l for  f  m i n E q . (7.10) represents the extent of gas  adsorption by the s o l i d .  m ~ 10.  m  transfer  coefficient  the purposes of c a l c u l a t i o n  to estimate  solid,  the value  appears to be adequate  in t h i s work.  No a p p r e c i a b l e  improvement was found by using other expression proposed by Chiba and Kobayashi [104] and Drinkenburg and Rietema  [105]. The c a l c u l a t e d values  D  r  and  K  (t, )b  a  r  e  e  Pitted  (see Appendix 10) of  i n Figures 7 . 4 , 7.5 and 7.6  and the average values f o r the f l u i d i z e d Table 7 . 2 . earlier  in t h i s  d^  bed are given in  In these c a l c u l a t i o n s , the expression developed  bubble diameter that  D , a, g  exerts  work (Chapter profile  5, Eq. 5.10) f o r the a x i a l  has been a p p l i e d .  It  can be seen  a strong i n f l u e n c e on the value of  Near the d i s t r i b u t o r  where l a r g e  gas t r a n s f e r  ^(be)b'  numbers of small bubbles  exist,  the o v e r a l l  coefficient  i s about 10 times  larger  than near the surface of the f l u i d i z e d  bed (Figure  7.4)  These r e s u l t s are in agreement with the f i n d i n g s of Kobayashi and Arai  [106] who also reported a sharp decrease i n the  1 24  Figure 7.4.  Calculated overall gas transfer c o e f f i c i e n t as a function of the f l u i d i z e d bed height.  125  calculated axial p r o f i l e of of  K  (b )5 e  "f ° r  r  a  ^(be)b'  given value of  d  b  ^  e  individual values  agree well with the  values of Kobayashi, Arai and Sunakawa [107].  Table 7.2 Average Values of K (  b e  )  and D  b  and D ^ r  g  in the Fluidized Bed 63 wt-% calcines -325 mesh (Brenda concentrate) 27 wt-% sand -40/+140 mesh S0 tracer 2  Run No.  u  o  (cm/sec)  b (cm) d  (sec- )  D a ,g (cm /sec)  K  (be)b 1  2  D (cm /sec) 2  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 c o e f f i c i e n t near the d i s t r i b u t o r grid would indicate that for a fast chemical reaction, much of the reaction would take place near the d i s t r i b u t o r .  This effect could be less  s i g n i f i c a n t , however, for the present case since some s o l i d stratification  was always observed near the d i s t r i b u t o r grid  (see Chapter 6, Figure 6.3)  under the Operating conditions  of the f l u i d i z e d bed reactor.  1000  1 26 •Z 500  100  Figure 7,5  Calculated p r o f i l e of the axial dispersion c o e f f i c i e n t of gas.  Figure 7.6.  Calculated p r o f i l e for the radial dispersion c o e f f i c i e n t of gas along the f l u i d i z e d bed.  127 As can be seen, in Figure 7.5,  the axial dispersion  c o e f f i c i e n t of gas in the f l u i d i z e d bed is also strongly affected by the bubble s i z e , increasing by about 25 times from the distributor grid to the top of the bed. other hand, for the radial fluence o f  d^  On the  dispersion c o e f f i c i e n t the i n -  is somewhat smaller, with an increase in  value proportional to the increase in bubble diameter along the bed (Figure 7.6).  The s u p e r f i c i a l gas velocity has a  minor influence on D r  but a larger effect on D at the »9 a »9  top of the bed especially where the value of small  (see E q . 7.8 and 7.9).  K(b )D  1  E  S  A decrease in the value of  K(be)b corresponds to a decrease in and an increase of Da, g along the f l u i d i z e d bed. At the top of the bed where the dispersion is more vigorous due to large bubbles formed, Da ,g increases by about 50% for a corresponding increase of 15% in the s u p e r f i c i a l gas v e l o c i t y . The average value of the axial dispersion c o e f f i c i e n t of gas, calculated for the f l u i d i z e d bed, is plotted  in  Figure 7.7 as a function of s u p e r f i c i a l gas v e l o c i t y .  For  purposes of comparison, data from other studies are also presented.  In the c a l c u l a t i o n , i t was assumed that the solids  were non-adsorbing (m = 0),  since no information is  on the value o f the adsorption c o e f f i c i e n t by the solid (Mo0 + sand), E q . 7.10. 3  available  m of the gas  The values of  1000  800  1—— Yoshi_da el al_ MS catalysis 150/xm He tracer , 25 °C Shurgerl ( 1 3 1 ) glass beads  o  He  CD CO  CJ  E o  600  tracer , 25 °C  o \ al> Baern-: et silica H  250/i.m  ( 30) 1  sand  I75^.m  tracer , 25 °C  2  CO  O cn  400 c  Present work O m =5 © m= 0 T = 5I0-520 °C S 0 tracer 63 wt. % , -325 calcines 2  O  o o c o y> \_  CD Q. CO  200 D x <  cf Q  100,  15 u  Figure  7.7.  0  25  35  , Superficial gas velocity  Overall axial d i s p e r s i o n c o e f f i c i e n t o f gas i n the f l u i d i z e d bed as 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 gas v e l o c i t y .  129  D a »9 n  are in a l l  reported  cases lower than others that have been  (Figure 7 . 7 ) .  choice of  m  T h i s d i f f e r e n c e may be due to  in c a l c u l a t i n g  a d s o r p t i o n i s assumed (m = 5) are about 45% l a r g e r be more r e a l i s t i c Mo0  3  particles  D s i n c e i f moderate •.»9 the c a l c u l a t e d values of  (Figure 7 . 7 ) .  than  m = 0  (Chapter 8)  A value of  m ~ 5  s i n c e the s u r f a c e of  is i r r e g u l a r  level  K  (be)b  of the MoS  above the d i s t r i b u t o r  to be most convenient to avoid s i n t e r i n g problems. occurs as a r e s u l t of the heat generated by the exothermic of MoS  2  7.2  may  transfer  injection appears Sintering  highly  o x i d a t i o n r e a c t i o n when a high c o n c e n t r a t i o n  particles  are contacted with a i r .  t i o n s are found e x p e r i m e n t a l l y (Chapter  a »9  the  gas  near the d i s t r i b u t o r g r i d , an  feed well  2  D  and may adsorb g a s .  Due to the l a r g e values of the o v e r a l l coefficient  the  Such c o n c e n t r a -  near the i n j e c t i o n  point  6).  S o l i d T r a c e r Experiments The s o l i d s e n t e r i n g  nozzle are r a p i d l y t r a n s f e r r e d  the r e a c t o r v i a the d i s p e r s i o n to the f l u i d i z e d  bed, p r e -  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 MoS reactor,  2  -  (Mo0  3  + sand) i n s i d e the f l u i d i z e d bed  two t e s t s were performed at d i f f e r e n t  gas v e l o c i t i e s .  superficial  130  7.2.1  Residence Time Distribution Function of Solid in the F l u i d i z i e d Bed Reactor A pulse input of MoS tracer was injected through 2  the dispersion nozzle into the f l u i d i z e d bed at a point 15 cm above the d i s t r i b u t o r g r i d . charge of  M0O3  The bed consisted of a typical  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 s o l i d s . The response to the input signal plotted as %MoS  2  in the discharged calcines versus time, is shown in Figure 7.8.  The computed residence time d i s t r i b u t i o n  function  E(e) is given in Figure 7.9 as a function of dimensionless time e.  From the calculated values of ^he variance  o  for the s o l i d , the value of the RDN was computed.  Q  For the  superficial gas v e l o c i t i e s of 17.4 and 22.5 cm/sec tested, the RDN are 0.74 and 0.75 reactor which is v i r t u a l l y  respectively, which indicate a back-mixed.  The two-dimensional studies using MoS as tracer 2  (Chapter 6) showed e a r l i e r that the sol ids entering  the  reactor are completely transferred to the emulsion phase without any bypass as elutriated s o l i d s .  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  Figure 7 . 8 .  (min)  Response signal at the discharge point after a pulse input of MoS tracer. 2  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  c r i t e r i o n for complete u t i l i z a t i o n  of the reactor and provides  the added advantage that scale-up i s considerably s i m p l i f i e d . The shape of the response curve obtained  (Figure  7.8) can be explained by assuming that the tracer that enters the reactor is transferred into the bed at a singe! l e v e 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 d i s t r i b u t i o n function determined.  Van Deemter [109] suggests that for the case where a tracer is injected at one point the subsequent mixing will be produced by two mechanisms of s o l i d 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: f  +  Ps  F  +  Ps < s " s>  ( c  s -  C  c  c  s  )  =  =  0  0  (7.10)  (7.11)  1 33  o u • u  \ backmixed V RDN = CO  0  0  = 22.5 = 17.4  cm/sec detection  point  tracer — injection  gas RDN =0.75  RDN = 0.74  plug flow  \S RDN  =0  Ji! 2 9 , D mensionless  9.  3 time  Dependence of s o l i d residence time function on dimensionless time.  distribution  1 34  where  c  g  and  C  are the fractions of tracer in the down-  g  flow and upflow respectively;  f  and  g  solid moving downwards and upwards; v e l o c i t i e s and both phases.  p  v  F  g  and  $  the fraction of V  g  their  the transfer c o e f f i c i e n t of s o l i d between  s  Van Deemter solved both equations simultaneously  by numerical methods and plotted the concentration of tracer as a function of the parameter  (F  V /l )t s  f  , where  the time of measuring the pulse at a given l e v e l .  t  is  His pre-  dicted concentration p r o f i l e of tracer with time coincides closely with the response obtained in the present work, where the concentration of MoS in the elutriated solids 2  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 c o e f f i c i e n t of solids in the  f l u i d i z e d bed was calculated using the model proposed by Van Deemter [109]  and modified by Kunii and Levenspiel  [110]  2  a>s  3  6  whereas for the radial  u  a m f  b  U o  (7.12)  - mf u  dispersion c o e f f i c i e n t of s o l i d ,  expression of Kunii and Levenspiel [110]  was employed:  the  135  D  = 0.1875  The calculated values of  D  a ,s plotted in Figure 7.10 and 7.11  (7.13)  and D  (Appendix 10) are r ,s respectively. The axial  bubble diameter p r o f i l e developed in the present study (Eq. 5.10)  has been u t i l i z e d  in the c a l c u l a t i o n .  The c a l -  culated values show the influence of the bubble diameter on the mixing of s o l i d s . top of the bed), i t  For a large value of  d  b  (nearer  the  is expected that the amount of s o l i d  carried upwards in the wake and in the bubble - and returned by downward flow - would be larger than for small, slower bubbles. D„ _ r, s  The influence of the s u p e r f i c i a l gas velocity on  is small in the range tested but large for  mainly in the upper levels of the bed. Da , s 3  compared with  Dr, s_  D,a , s.  The large values of  indicate that mixing of solids is  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 mechanism since the gas and solid flow directions are almost exclusively axial . Few data exist on experimental to compare with the present r e s u l t s . of the average value of  D  values of  D  a ,s In Figure 7.12, a plot  calculated along the f l u i d i z e d  bed is compared with the work of Bart [111]. exists between the two sets of r e s u l t s .  A large  difference  Figure  E o  -  7.10  6  Calcu1ated axial dispersion c o e f f i c i e n t of solid as a function of the f l u i d i z e d bed height.  6 0 wt. % calcines T = 5 5 0 °C u , cm/sec Q  o  c  u  4  o u  o TD O  CC  10  20 Fluidized  Figure 7.11.  40  30 bed  height  (cm)  Calculated radial dispersion c o e f f i c i e n t of s o l i d as a function of the f l u i d i z e d bed hei gh t.  50  137  Attempts have been made to predict by means of a mathematical model the concentration p r o f i l e of the MoS  2  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. profiles determined coincide  The concentration  approximately with the concen-  tration p r o f i l e s found in the pulse test of Chapter 6.  138  Figure  7.12.  Average axial dispersion c o e f f i c i e n t of solids in the f l u i d i z e d bed as a function of the superficial gas v e l o c i t y .  Chapter 8  KINETICS AND MECHANISM OF MOLYBDENITE OXIDATION  The successful scale-up from the p i l o t plant unit to f u l l  size plant operation depends to a large extent on  r e l i a b l e 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 l u i d i z e d bed reactor. A l s o , to elucidate details of the reaction of MoS to Mo0 , 2  3  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 l u i d i z i n g  139  140  conditions.  The reactor had been previously charged with a  known amount of c a l c i n e s , usually 2 Kg.  Samples were then  taken at given intervals of time from the discharge b i n , while the rest of the calcines were kept r e c i r c u l a t i n g continuously through the reactor. Concentrates from four different  sources -  B.C. Moly, Endako, Brenda and Kennecott Mines - were used in this study.  The chemical, - p a r t i c l e s 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 d i s t r i b u t i o n was analyzed with a Coulter Counter and C y c l o s i z e r , 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 v i r t u a l l y  linear - the rate is constant - with  time over the range covered.  However at 480°C the i n i t i a l  of oxidation is faster and no longer constant. the reactor temperature  further,  in both the extent and i n i t i a l 526 and 595°C a high i n i t i a l  Increasing  results in a sharp increase  rate of oxidation.  Between  oxidation rate can be seen.  The fraction of sulphur oxidized in this stage varies from 0.5 to 0.75,  depending upon the  rate  temperature.  Table 8.1 Chemical Composition of Molybdenite Concentrates  Source  %MoS  %S  2  %Cu  %Pb  %Ca0 %Fe  %Si0  2  %A1 0 2  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  cT , average diameter (cm) s  Kennecott  3.5 x  B.C. Moly  1 .2 x 1 0  Endako  0.8 x  IO"  3  Brenda  1 .0 x  IO"  3  IO  - 3  r 3  S , surface area s  (cm /g) 2  2.5 x 10* 3.5 x 10* 4.1 x 10* 3.8x10*  3  142  Table 8.3 Batch Kinetic Experiments  Exp. No  Concentrate  B.66  B.C. Moly  42.05  450  B.38  ti  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  19.75  554  0.36  B.108  Brenda  Sampl e  (g)  Temp.  Po (atm) 2  (°c)  0.21  B.97  •I  21 .60  557  0.54  B.109  II  17.05  557  0.61  Once the rapid i n i t i a l  oxidation stage has passed,  a slower transformation regime begins to take hold. temperatures between 560 and 595°C a transitional  At  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  TJ N  X  o  CM CO  3 10  c o o o  B.C. B66 B38 B 19 B58 B 59  o  B-I08A 554 ° C - 3 6 % 0 B 9 7 £f 557 ° C - 5 4 % •" BI09 © 557 °C- 6 I % " 2  Moly Concentrate H 450 °C , air O 480 °C " V 526 °C " © 530 °C " • 557 °C " B'24 • 560 °C " B 34 A 595 °C  I  15  20  25  30  35  40  t , Time (min)  Figure 8.1. Mole fraction of molybdenite oxidized as a function of time for batch t e s t s .  144  below, the fractional  conversion follows smoother  with no abrupt change in  curves  rate.  Above 526°C, in a l l  cases, a slower reaction stage  was observed at high conversions.  In this l a t t e r stage an  apparent constant or very slowly decreasing rate of oxidation exists.  This probably persists until  of MoS to Mo0 2  3  is attained.  It  the complete conversion  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 found to accelerate the i n i t i a l  pressure was  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 , it  so that  is not clear whether the higher conversions obtained  were due to the increase in oxygen concentration alone or due to a difference in r e a c t i v i t y .  B.C. Moly and Brenda  concentrates are very similar in terms of size of particles and s p e c i f i c 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 cm and a depth of about 45 2  p a r t i c l e s , was reacted.  Although their  bed of p a r t i c l e s was  shallow, considerable heat effects may s t i l l  have been  145  Mole  Figure 8.2.  fraction  of  Mo S  2  oxidized  Computed values of the rate of reaction as a function of the fractional conversion of MoS . 2  146  produced in between particles with the result that the actual temperature of the reacting p a r t i c l e s 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 p l o t s , slopes were obtained by graphical using a mirror.  differentiation  The results are plotted in Figure 8.2.  can be observed that above 526°C the i n i t i a l ,  It  constant rate  of oxidation for B.C. Moly concentrates prevailed up to 0.55 of the total fraction of sulphur oxidized.  0.50-  Results  plotted in Figure 8.3 and given in Table 8.4 show that the initial  rate of reaction increases rapidly with Increasing the oxygen partial  reacting gas increases the i n i t i a l  pressure in the  reaction rate in roughly  a linear fashion, as shown in Figure 8.4. mean that the i n i t i a l  temperature.  This finding may  rate is controlled by the chemical  reaction between the oxygen and MoS as follows: 2  M o S 2  (s)  +  7  /  2  °  2  *  M o  ° (s) 3  +  2 S  °  2  (8.1)  although transport control involving oxygen is also possible. The rate of this reaction based on the surface area of s o l i d may be expressed as  147  Figure 8.3.  Rate of reaction as a function of temperature the three stages of transformation.  in  148  dt  where  S  s  -dMoS /dt 2  J  7 s K  u  0  ( 8 2  is the s p e c i f i c surface area of MoS  - ' 2  (cm /mole), 2  2  the rate of reaction (mole f r a c t i o n / s e c ) ,  k  the  g  s p e c i f i c rate constant based on the surface area of MoS  2  (cm/sec)  and  the concentration of oxygen  (mole/cm ). 3  The s p e c i f i c rate constant was calculated for each experiment as a function of the temperature stage. In k  g  Values are given in Table 8.4. vs  1/T  The activation  (Figure 8.5)  for the  initial  The Arrhenius plot ..'of  gives a reasonably straight  energy for this i n i t i a l  tion is then calculated to be 25.3  period of  transforma-  Kcal/mole.  is somewhat lower than the value of 35.3  line.  This value  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] mole.  of 13.7  Kcal/  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.  by other workers [118,119].  This has been postulated  Figure 8.4.  Influence of the oxygen p a r t i a l the i n i t i a l r a t e of r e a c t i o n .  pressure  on  Table Calculated  Rate  of  Reaction  and  Reaction  Initial Exp..  MoS  2  Temp.  n p  0  2  8.4 Rate  B.66  Concentrate  B.C.  Moly  (°C)  450  (atm)  -dMoS /dt  0.21  (mole  for  2.00  X io- *'  Oxidation  -dMoS /dt  s (cm/sec)  2  k  fract/sec)  Molybdenite  Transition Reqime  Regime  2  No  Constant  (mole  fract/sec)  Final  Regime  -dMoS /dt 2  (mole  fract/sec)  4 .1  X io-  5  2.00  X io-*  4 .75  X  5  2.67  X  2 .00  X io-*  2.94  X io-*  5.0  x I O  2 .31  X  10-*  3.02  X io-*  5.9  x 10"  3  3 .29  X io-*  3.34  X io-*  6.2  x  1  B.38  •I  480  II  4.67  X io- *  B.19  II  526  II  9.63  X  530  II  1.11  X io-  557  II  1 .58  X To-  . 560  II  2.03  X io-  3  4 .23  X io-*  5.01  X io-*  6.3 x 1 0 "  595  II  2.17  X io-  3  4 .54  X io-*  4.17  X io-*  6.5  x IO"  3  5 .27  X  10"*  5.00  X io-*  1.2  x 10"  3 .94  X  10"*  3.44  X  10"*  6.0  x 10"*  3 .85  X  10"*  2.67  X  IO"*  5.4  x 10-*  B.58 B.59 B.24 B.34 B.108 B.109 B.97  u II  n  II  Brenda n  H  1  IO""  554  0.36  4.46  X io-  557  0.54  5.00  X  557  0.61  5.53  X io-  3  1 0" 3  3  -  IO"  IO"* -  1 0"  5  s  5  s  5  5  o  T ,  550  600  10  Temperature (°C) 500  T  o CD  e o c o o o  IO  4  Q>  D  rr  j-  B 108 B 97 B 109  Brenda Concentrate A 36 % 0 'ff 54 % " © 6 1% " 2  B.C. Moly B 66 o B 38 v B I9 B 58 B 59 a B 24 A B 34 10  Concentrate air E= 2 5 . 3  1.20  .10  l/T Figure 8 . 5 .  Arrhenius  plot  of  x  the  -3  10  rate  kcal / mole  1.30 (°K~')  constant.  450  1 52  In the transition stage, the effect of on the oxidation rate decreases (Figure 8.3) weak dependence exists in the f i n a l  temperature  until  only a  reaction period.  The  magnitude of the rates in this stage, and their weak temperature dependence may indicate that a diffusional perhaps through s o l i d Mo0  3  is rate l i m i t i n g .  process,  This final  stage for the B.C. Moly concentrates occurred between  0.70  and 0.85 fraction of MoS oxidized, as can be seen in 2  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 d i f f u s i o n control regime observed for the Brenda concentrates may have been the result of the thermal of this  8.2  effects discussed in the next section  chapter.  Mechanisms and Morphology of Molybdenite Oxidation From the data obtained in the above kinetic  experiments, evidence was obtained that two different seem to control the oxidation of MoS . 2  steps  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  Table Scanning  Electron  Micrographs  of  Samples  of  8.5 MoS  2  Oxidized  in  the  Hot  Stage  Microscope  Temp. °C  Gas  Time Sec.  25  air  —  Sample of the  C.b)  550  air  30  General surface  (c)  550  0  30  General view of a sample o x i d i z e d at 550°C in pure o x y g e n . C r y s t a l s are nucleate preferentially a t t h e f a u l t s and c r a c k s o f t h e M o S surface.  (d)  600  Sample No. (a)  2  o  Observations o f M o S f r o m a l a r g e p i e c e (2 m o l y b d e n i t e c a n be o ) s e r v e d . 2  cm s i z e )  view of a sample o x i d i z e d i n a i r is covered with small n e e d l e l i k e  before  f o r 0.5 min c r y s t a l s of  oxidation.  at 5 5 0 ° C . oxide.  The  The  layered  structure  original  smooth  starting  to  2  2  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 f o u n d on t h e s u r f a c e o f a l l samples. A n a l y s i s shows of sample i s u n r e a c t e d M o S .  m i c r o p r o b e on that they are  slab-like crystals pure M o 0 . The back 3  2  Ce)  500  air  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 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 t o h a v e t a k e n 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 .  alteration is observed. Some p l a c e at the c r a c k s . Small  180  A f t e r a s h o r t p e r i o d of t i m a , c r y s t a l s of Mo0 are formed at the s u r f a c e of MoS i n an irregular 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 sits of n u c l e a t i o n . Slab-like crystals 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 , a s v a p o r c o n d e n s a t i o n o f M o 0 , and a r e growing up f r o m t h e s u r f a c e o r n o t a t t a c h e d (arrows). L a r g e r c r y s t a l s are growing from the surface cracks.  300  The s u r f a c e o f M o S 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 of Mo0 with a d i s t i n c t rhombohedral s t r u c t u r e . Large c r y s t a l s are growing normal to the sample from 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 gaseous Mo0 was e v o l v i n g f r o m t h e c r a c k , w h i c h was p r o b a b l y at a higher temperature than the r e s t of the s a m p l e . The Mo0 crystals apparently form 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 (arrows).  3  Cf)  550  air  2  3  2  3  (g)  550  (h)  550  air  300  General view of the o x i d i z e d attached l o o s e l y , i f at a l l ,  sample to t h e  CO  550  air  180  C l o s e up o f c r y s t a l s o f M o C a t random a n d a r e a p p a r e n t l y to p r o d u c e t h e s e c r y s t a l s .  at the s u r f a c e of M o S . C l e a r l y , the c r y s t a l s are growing not a t t a c h e d to the 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  550  air  180  S m a l l p a r t i c l e o f M o S w h i c h shows t h a t c r y s t a l s of M o 0 away f r o m t h e s u r f a c e . P a r t i c l e s of M o 0 are forming in 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 (arrows).  air  3  3  3  showing surface  crystals l a y e r of  of Mo0 oxide.  3  that  apparently  are  2  2  3  3  g r o w by c o n d e n s a t i o n of M o 0 a n y d i r e c t i o n by v a p o r ejection3  CONTINUED  Table  Sample No.  U)  Temp. °C  550  (Continued)  Time Sec.  Gas  0  8.5  2  30  Observations In t h e p r e s e n c e o f p u r e o x y g e n , a g a i n , p r o b a b l y due t o a ho ; t e r in pure oxygen than in air.  i t seems t h a t t h e c r y s t a l s u r f a c e o f t h e s a m p l e due  of Mo0 formed are vaporized the fast chemical reaction 3  to  At high temperature, 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 is complete. C r y s t a l s of Mo0 g r o w away f r o m o r i g i n a l crack. A t t h e s u r f a c e , t h e p a c k e d c r y s t a l s t r u c t u r e seems welded 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 this temperature.  (1)  600  air  600  tm)  25  air  —  (n)  560  air  5400  Sample r o a s t e d i n c r y s t a l s of M 0 O 3 .  Co)  580  air  1800  Sample covers  CP)  605  air  2700  C l o s e up o f M0O3 formed  3  Sample of m o l y b d e n i t e t h i n , resembling thin  concentrate (B.C. irregular slabs.  Moly)  before  roasting.  Particles  are  flat  and  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 Some s m a ' i l c r y s t a l s seem t o be b r o k e n f r o m t h e o x i d i z e d particles.  o x i d i z e d in the the particle.  fluidized  bed  reactor.  a sample r o a s t e d i n the f l u i d i z e d a r e m e l t e d f o r m i n g an h o m o g e n e o u s  Again  packed  crust  bed r e a c t o r . The c r u s t of M 0 O 3 .  of  M0O3  original  crystals  crystals  of  155  flow rate was used to avoid mass transfer control in the gas phase.  The samples were heated rapidly  and the temperature  (1 to 1.5  was measured with a b u i l t - i n  min)  thermocouple  located at the contact between the hot stage and the sample. During the reaction period, samples were continuously observed d i r e c t l y  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 a i r and pure oxygen as oxidizing gases. Samples were also taken from the f l u i d i z e d bed roaster  for  observation under the scanning electron microscope. A summary  of  the conditions  given in Table 8.5. in Figure 8.6  (a)  and  observations  are  Scanning electron micrographs are given  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 Mo0 formed does not 3  distort  the original MoS p a r t i c l e s 2  2)  At higher temperatures  formation of Mo0  3  preferential  (Figure 8.6  (e)).  (550 to 560°C), rapid  crystals occurs at the surface at some  nucleation points (Figure 8.6  (f)).  Unattached  1 56  (d)2IOO x|  6 0 0 °C ,0 180 sec 5/i.m 2  Figure 8.6. [  Scanning electron micrographs of molybdenite samples oxidized in the hot stage microscope and in the f l u i d i z e d bed reactor.  1 57  Figure  8.6  (Continued)  1  Figure  8.6  (Continued)  58  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 Mo0 crystal formation during the 3  initial,  rapid oxidation, is the  vaporization of Mo0 at the 3  reacting interface of MoS , followed by condensation. 2  Due  to the large amount of heat evolved (-295.9 Kcal/mole at 550°C), local areas of high temperature may e x i s t , which are capable of generating Mo0 vapour.  This may condense on  3  cooler crystals of Mo0 or form new Mo0 crystals spontane3  ously.  3  In support of this mechanism are observations of  p a r t i a l l y oxidized samples which have shown that some of the particles appear 3)  to be melted (Figure 8.6  (p)).  In particles roasted in the f l u i d i z e d bed reactor,  the layer of Mo0 crystals covers the original MoS p a r t i c l e s 3  2  completely (Figure 8.6  (n) and (o)),  leaving apparently few  pores for oxygen transport toward the reacting MoS /Mo0 2  interface.  3  In a l l cases, small, separate crystals of Mo0  3  were found to be present with the larger particles of transformed MoS . 2  The former may have been broken from the  parent Mo0 material as a result of the a t t r i t i o n action 3  of the sand during the f l u i d i z a t i o n  (Figure 8.6  (n)).  161 4)  At high temperature  formed, seem  (600°C) the Mo0  to melt, probably on the overheated surface  of the reacting MoS p a r t i c l e s  (Figure 8.6  2  (p)).  case, a more dense and uniform surface of Mo0  3  the  crystals  3  In  this  appears to be  result. In a l l  cases, between 500 and 600°C there is  evidence that after the fast i n i t i a l  step of Mo0  formation on the surface of the p a r t i c l e , crystals does not grow outward further formed surface (Figure 8.6  (h)  between the Mo0  3  until  Rather i t  appears  slowly closing the space  impervious layer builds up.  This  transformation mechanism at the surface of a  MoS p a r t i c l e  is schematically depicted in Figure  2  It  originally  c r y s t a l s , perhaps by a condensation process  a relatively  hypothetical  the layer of oxide  from the  and (1)).  to grow inward toward the i n t e r i o r ,  crystal  3  may be possible that at the MoS /Mo0 2  face some Mo0  2  8.7. 3  could be formed by the s o l i d state  6 Mo0  3 ( s )  + MoS  2 ( s  y^ 7 Mo0  2 ( s )  + 2 S0  interreaction  (8.  2  Nevertheless, X-ray analysis performed on samples from the f l u i d i z e d bed reactor, at high levels of conversion (>99%) did not indicate the presence of Mo0 . 2  Mo0 , i f 2  This suggests that  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 o x i d a t i o n of an MoS p a r t i c l e . 2  1 63  very short period of time and exists probably in a thin layer at the very surface of the MoS core. 2  8.3  Temperature of Particles During the I n i t i a l 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 Mo0 being formed 3  An estimation of the p a r t i c l e temperature can be made by performing a heat balance on the p a r t i c l e .  For  s i m p l i f i c a t i o n , the following assumptions have been made: 1)  The particles are spherical with a diameter,  2)  The p a r t i c l e s are suspended in air  v transferring  the heat generated by conduction and r a d i a t i o n . 3)  The particles are small enough that  internal  temperature gradients are small, and individual p a r t i c l e s are at a uniform temperature. 4)  During the  initial  stage,  the oxidation and  heat generation is a steady state process, that i s , 8 q / 9 t = 0.  164  The heat generated by the reaction  q =AHj k  C  s  (cal/cm sec)  where  h  = h  c  a  $  (T  s  - T )  s  is the heat transfer  g  MoS p a r t i c l e  c >  is given by:  (cal/cm sec)  (8.5)  2  coefficient  and the surrounding a i r  2  (8.4)  2  The heat dissipated by convection, q  q  , q, i s :  is the surface area of the p a r t i c l e ,  between the 2  and  T  and s  the temperature  of the p a r t i c l e  a  (cal/cm sec ° C ) , T g  $  are  and the bulk gas, respectively  (°C). The heat dissipated by radiation  is given by  the expression:  k  q  r  =  £ o  a  s  I  where  e  100  is the p a r t i c l e  Botzmann constant.  J  -  9 I100J  emissivity,  At steady state,  a single p a r t i c l e becomes:  (cal/cm sec) 3  a  is the Stefan-  the heat balance for  (8.6)  165  (A H T )  k  s  C  g  . h  s s <s a  T  "  V  +  °*  e  TM  100  s  (8.7)  The heat transfer c o e f f i c i e n t  h  can be estimated assumi ng  $  conduction into a stagnant medium where  Ii d N = - J §. u k where  k  g  is the thermal conductivity of air at  temperature  T .  g The temperature of the reacting p a r t i c l e can then be calculated by successive iterations of the equation  h  s  T  s  + e a  100  aj  k  s  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 MoS Particles During the I n i t i a l Oxidation 2  T CC) g  T  cc)  PM0O3  a  t  520  0.056  520  547  0.068  550  581  0.091  580  630  0.221  600  650  0.435  s  = 10"  k  g  = 4.32 x 10"  3  s  (mm H ) 9  500  d  T  cm (Brenda MoS concentrates) 2  6  (cal/cm sec °C)  0.8  e  =  c  = 1.4 x IO"  Values of  k  g  13  (cal/cm sec °K*) 2  were taken from the experimental  data (Figure  From the calculated values, i t appears that the p a r t i c l e temperature  is substantially higher than i t s 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.  8.5).  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 p a r t i c l e surface, and hence heat evolution, as the surface temperature  8.4  rises.  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 first  "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  are presented in Figure 8.7.  The best f i t  temperature  line for this  plot is given below:  x  where  T  is in °C.  = T° *  7 6 5  - T + 520  (min)  (8.9)  This relationship is valid for molybdenite  concentrates with an average diameter of ~10 microns. Between 520 and 560°C, Eq. (8.9) the plot in Figure 8.7.  agrees to within ±5% of  168  Table 8.7 Measured Values of Time of Reaction in "Chemical and "Diffusional" Regimes  "Chemical" Regime Temp.  po  2  MoS mol . f r a c t . ox. X  2  "Diffusional" Regime  MoS mi n mol . f r a c . ox. X  2  °C  atm.  458  0.21  0.400  30.0  1.0  480  0.21  0.370  15.0  1 .0  526  0.21  0.475  7.5  1.0  530  0.21  0.475  6.5  1.0  554  0.36  0.750  2.8  1.0  557  0.21  0.525  5.0  1.0  557  0.54  0.775  2.5  1.0  557  0.61  0.800  2.35  1.0  560  0.21  0.600  5.0  1.0  595  0.21  0.650  5.0  1.0  *  T  mi n „*** (180) ** 150 * 120 * 100 * 95 * 85 * 90 * 85 * 80 * 70  Extrapolated  ** Approximate  ***  Estimated  The measured values of t'  and x can then be used for scale-up  purposes, as i t w i l l be shown in Chapter 10.  169  200i  .|  I50|  c o  e 00  c  •£  1001  0)  B.C. Molv Concentrate 12 4 5 0 °C ., air  e  O  o o 5 0  .480  "  V  5 26  "  ©  5 30 55 7  " "  •A  5 60 59 5  " "  Brenda A 0 ®  0'  554 557 557  Concentrate °C 36 % " "  54 % 61 %  0 " "  2  _J  450  500  550  600  T , Temperature (°C)  Figure 8.8.  Influence of temperature on the total time required for transformation of MoS to Mo0 in the f l u i d i z e d bed reactor. 2  3  Chapter 9  CONTINUOUS ROASTING OF MOLYBDENITE CONCENTRATES IN THE FLUIDIZED BED REACTOR  The main o b j e c t i v e of the present research program was to study the t e c h n o l o g i c a l f e a s i b i l i t y fluidized  of o p e r a t i n g a  bed r e a c t o r f o r r o a s t i n g molybdenite  concentrates.  In t h i s work, four d i f f e r e n t concentrates have been t e s t e d , the composition Chapter mainly  and c h a r a c t e r i s t i c s of which are given i n  8, Table 8.  The r e s u l t s of these t e s t s , obtained  i n the 12.5 cm diameter  discussed  in this  r e a c t o r , are presented and  chapter.  P r i o r to each experiment with a f i x e d amount of sand was shown i n Chapter minimized zation. roasted  charged  (3000 g) of -40/+140 mesh.  6, the use of t h i s s i z e  As  distribution  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  fluid-  Next a given q u a n t i t y of c a l c i n e s which had been p r e v i o u s l y was charged.  added to boost the temperature The  the r e a c t o r was  In a d d i t i o n  some MoS  2  was  during the heating p e r i o d .  p r e c i s e amount of c a l c i n e s was a l t e r e d  as needs r e q u i r e d  to vary the residence time of s o l i d s i n s i d e the r e a c t o r .  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  500 to 595°C  Overage residence time of s o l i d s , Superficial gas v e l o c i t y , u  0  Oxygen partial pressure, P (J 2 n  Feed rate of MoS , F 2  t  3 to 38 hrs 15 to 34 cm/sec 0.04  to 0.8 atm  133 to 2650 Kg/m  2  day  0  -100 to -10 microns  P a r t i c l e size of MoS , d 2  s  0.4 to  0.04%  Calcium in Feed  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 e f f i c i e n c y , i . e . ,  *  sulphur in the calcines  the residual  (see Appendix 15).  The complete results of the 89 experiments performed employing the 12.5 cm diameter f l u i d i z e d bed reactor are given i n Appendi x 13.  9 . T. 1 Superficial Gas Velocity The s u p e r f i c i a l gas velocity appears to have a minor influence on the residual sulphur content of the c a l c i n e s , 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: u  0  for example, for  t=  27 hrs, at  - 20 cm/sec the calcines contained about 0.43% sulphur,  whereas at  u  = 30 cm/sec  0  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 )b  n  o  r  t  e  n  e  R D N  varies appreciably over this range  qf superficial gas v e l o c i t y . The range of s u p e r f i c i a l gas velocity in this work was dictated by operational u  0  investigated  limitations.  " 18 cm/sec the bed f l u i d i z a t i o n was not s u f f i c i e n t  Below to  permit smooth operation of the rotary scraper inside the bed, while at u  0  > 35 cm/sec the e l u t r i a t i o n  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 B.C.  Moly  Concentrates  T = 550 ± 2 ° C a i r , J atm. 0.8  to CD  o  o  t = 20  0.6  t = 2 7 hr  I  hr  ±10%  ±10 %  0.4  c CD O  v_  CD  Q_  0.2  0 20 u , Superficial 0  Figure 9.1.  25 Gas  30 Velocity  35  (cm/sec)  Influence of the s u p e r f i c i a l gas velocity on the velocity on the sulphur content in the c a l c i n e s .  174  was higher than the discharge capacity of the rotary of the cyclones.  valves  This resulted in an accumulation of s o l i d  inside the cyclone system.  The lower l i m i t of  u  0  > 18 cm/sec  cannot be readily overcome, but the upper l i m i t can be i n 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 However, the effect for retention  9.2.  becomes less pronounced at longer times;  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 value tested.  This apparent  up to 580°C, the highest  lower l i m i t is due, as will be  shown l a t e r , 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 i m i t of sulphur in the calcines even by increasing the temperature longer retention  to 580°C.  For  times, t > 27 hr, the lower l i m i t of sulphur  can be achieved even at 520-525°C.  This suggests that  between 520 to 580 C, the average residence time of P  reaction  B.C. Moly  Concentrate  u =28.5 c m / s e c T ±. 10 % (hr) A 3.5 hr 9 8 'hr O I I 0  CO  cu  c  ®  _o  o  o  .0  A  22 27 "  CO  c 0.5  CU o  cu  t = 27 hr  Q_  0  500  520  540 Temperature  Figure 9 . 2 .  560  580  600  (°C)  Sulphur content in the calcines as a function of the temperature of roasting.  176  has a larger effect on the f i n a l  sulphur content of calcines  than does the temperature of roasting. Severe problems arose from the sintering of calcines inside the reactor, p a r t i c u l a r l y at high temperatures. ing was f i r s t  Sinter-  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 c l e a r l y show that for practical operation of the f l u i d i z e d bed reactor, the roasting temperature 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 divided by the feed rate of fresh MoS , F : 2  0  state),  177  10, B.C. Moly Concentrate \520 °C  reactor © 5 2 0 °C , air A 5 4 0 °C , air reactor o 524 °C , air A 5 5 0 °C , air  5 4 0 °C  . CO  c o o  _3 to c o v. CL)  0.4  CL  0.2  0.1  0  10  15  20  25  T , Average time of residence (hr) Figure 9.3.  Influence of the average residence time of solids in the 7.5 and. 12.5 cm diameter f l u i d i z e d bed reactors.  30  178  (9.1)  (hr)  The average residence time has a drastic effect on the sulphur content in the product c a l c i n e s , 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 C 6  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%.  results are in good agreement,with  These  the batch experiments  (Chapter 8) and seem to confirm the existence of two kinetic regimes of transformation: followed by a slow f i n a l  a fast  initial  different  regime  one.  In Figure 9.4, the results of experiments conducted at three different again that for  temperatures  T = 55Q-575°C,  are plotted. for  It  can be seen  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 i m i t value of ~0.4%S is approached asymtotically.  9.1.4  Oxygen Partial 1  '  "  "  "" '  Pressure 1  !  1  ~  The influence of the oxygen partial pressure on the residual sulphur in the calcines at high conversion levels  179  B.C. Moly  t,  Figure 9.4.  Average  Time  of  Concentrates  Residence  (hr)  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 MoS oxidized) was found to 2  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  in the calcines increases sharply.  , the percent sulphur  n  2  U  For example, at P = u n  2  0.05 atm, the sulphur level whereas for P reached 1.6%.  Qz  in the calcines increases to  0.8%,  = 0.04 atm the sulphur in the calcines has In this case, the oxygen efficiency in the  reactor approaches 100% since the calculated values of the gas transfer c o e f f i c i e n t  K  (b )b  ^  e  s e e  P  C n a  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-  f o r e , that the bed becomes depleted, i . e . , oxygen in the upper region at  P u  < 0.1  n  starved, of  atm.  2  The negligible effect of the oxygen partial pressure above ~0.1  atm for high levels of conversion i s  in agreement with the previous findings from the batch tests (Chapter 8) where the time required to achieve a f u l l version was found to be v i r t u a l l y  con-  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  181  2.5,  B.C.  2.0  Moly  Concentrate  T =25 hr ±10% T =550 ±2 °C u  0  = 28.5  cm/sec  CO CD  c  \> D  O  .5  sz CL ZD  CO  c a  CD  .0  CD  CL  0.5  0  0  Figure 9.5  20 40 60 Percent Oxygen in Reacting Gas  80  Influence of the oxygen p a r t i a l pressure on the sulphur content of c a l c i n e s .  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 i m i t 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 i m i t 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 CaC0 to CaO. S0  2  3  [120], which at the roasting temperature decomposes The CaO formed can subsequently react with the  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 * CaSO*  (9.2)  2  This reaction can occur at 550°C due to i t s favourable free energy of reaction  ( A G ^ Q O Q  =-67.25 Kcal).  is stable at this temperature above 1200°C [121]. roasting, such as:  The CaSOit formed  and decomposes appreciably only  Other reactions might also occur during  183  Figure 9 . 6 .  Influence of the temperature on the l e v e l of c a l c i n e s f o r the Kennecott  sulphur concentrates.  184  Figure 9 . 7 .  Sulphur content of c a l c i n e s as a f u n c t i o n of the average residence time of r e a c t i o n f o r the Kennecott c o n c e n t r a t e s .  185  CaC0  3  + S0 +i0 2  2  * CaSCU + C0  (9.  2  which s i m i l a r l y 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: Endako, with 0.03% Ca.  Brenda, with 0.04% Ca and  The Brenda concentrates were standard  concentrates leached i n d u s t r i a l l y with a chloride solution of C u C l , FeCl 2  3  and CaCl  2  copper, lead and calcium. industrially  at 120°C to remove most of the The Endako concentrates were  leached with a dilute solution of HC1 (pH  for the sole purpose of removing calcium.  3-4)  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 c a l c i n e s . The residual sulphur in the calcines is plotted in Figure as a function of the average residence time.  It  9.7  can be seen  186  0.301  1  r  t, Figure  9.8.  1  Average  1  Residence  1  1  Time (hr)  Sulphur l e v e l of c a l c i n e s as a f u n c t i o n of the r e s i d e n c e time f o r low calcium c o n c e n t r a t e s .  r-  187  that at 550°C for t > 20 hr, the sulphur levels have decreased 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  s i g n i f i c a n t l y lower than the maximum industrial  sulphur  s p e c i f i c a t i o n of 0.25% S and compare favourably with the average sulphur content of calcines roasted in multiple furnaces.  hearth  In the l a t t e r process the average sulphur content  in the calcines is  about 0.1  An additional  to 0.15%  [122].  leaching of the Brenda concentrates  with 0.5 N HC1 at 100°C was carried out.  Then, an additional  s l i g h t 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). S i m i l a r l y , at these conversion l e v e l s ,  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 s l i g h t decrease in the sulphur content  of the calcines was found by increasing the temperature 524 to 550°C.  This finding has important  from  practical consequences;  188  r  i  Brenda  Concentrate— —  A  O A ®  0.25  chloride leached  > air 6 0 % 02 J  0 CO CD  Io  air0 % 0 6  2  J| double leached  Endako Concentrate  •  0.201  air ,  H Cl leached  7 = 16 hr  O  xz  w  0.151  c  -= 2 0 hr  CD O  k_  <D  Q_  0.10  0.05  1  520  530  540 Temperature  Figure  9.9.  550  560  (°C)  S u l p h u r c o n t e n t i n c a l c i n e s as a f u n c t i o n o f the r o a s t i n g temperature f o r low calcium concentrates.  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 f e a s i b i l i t y of the f l u i d i z e d bed process to produce calcines of low sulphur contents by using standard low-calcium molybdenite concentrates.  It represents a major improvement over other f l u i d i z e d  bed processes attempted previously. For comparison purposes, an a r t i f i c i a l rate of transformation -(dMoS /dt) 2  r  MoS  dMoS dt  2  Calculated values of  average  was calculated as  = mole fraction MoS oxidized average residence time 2  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 f o r 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 1 0  - I f  to 6.8 x 10"^ (mole fraction  of MoS oxidized/min).  When the minimum level of sulphur  is reached, any further  increase in t" w i l l only decrease the  2  value of  r"|yj<5 0  » 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  04'— —'—i—i—i—i—i—i i i i • » « 0.970 ^ 0.980 0.990 X , Fraction of sulphur oxidized 1  •  • Ii I 1.000  s  Figure 9 . 1 0 .  Apparent rate of reaction of molybdenite in the f l u i d i z e d 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 l u i d i z e d bed reactor  is given by:  = 9.6 r  c  M  (kg/hr)  c  MoS  (9.4)  2  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 i n d i n g s , the optimum  ing conditions of the f l u i d i z e d bed process for roasting have been determined. given in Table  operat-  molybdenite  The range of operations is  9.2.  Table  9.2  Optimum Range of Operating Conditions for  Fluidizing  Bed Roasting of MoS  2  Variables  Minimum Value  Temperature Average residence time Superficial gas velocity MoS feed rate Oxygen partial pressure P a r t i c l e size MoS Calcium in MoS Sulphur in calcines  520°C 20 hr 18 cm/sec 200 Kg/day/m air - 3y  2  2  2  2  Maximum Value 550°C 30 hr 35 cm/sec 300 Kg/day/m air 50y 0.05% 0.2%  2  192  9.3  Material Balance on the Process At proper working conditions, the e f f i c i e n c y of the  cyclone system was between 98.5 and 99% (see Appendix 13), while that of the scrubber system was about 98%. overall  Thus an  c o l l e c t i o n 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 MoS Concentrates in 12.5 cm Reactor 2  Base 1 day Feed rate:  3600 gr/day  MoS concentrates 90% MoS 2  2  Calcines with 0.12% S T = 550°C u  = 24.4 cm/sec  0  t = 26.7 hrs In Feed  (gr) Air  Out Calcines  Solid in Scrubber  (gr) Solid in Solution  Gases  Mo  1942.6  1727.8  195.1  S  1314.0  3.5  0.4  1300.3  863.9  97.4  23944.0  0  2  N  2  26205.6 66530.4  19.2  66530.4  193 The solids collected in the scrubber were calculated for a 99% e f f i c i e n c y of c o l l e c t i o n in the cyclones, at an e l u t r i a t i o n  rate of 85 gr/min (120 Kg/day).  represents a ratio of e l u t r i a t i o n * : F  0  This  rate to feed rate of:  = 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 -saturated acid solution of 2  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 0 3 • n H 0 [124]. 8  2  2  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,, + 2 N H t -> (NrU)  2  MoO^ + 2 H 0  (9.5)  2  The s o l i d not collected in the cyclone system amount to 1-2% of the total calcines r e c i r c u l a t e d , but depend  on  the p a r t i c l e 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 c o l l e c t i o n efficiency dropped to 80-85%. however, over 98.5% e f f i c i e n c y was achieved.  Normally,  The calcines  collected in the water scrubber contained usually a s l i g h t l y higher content of sulphur than the discharged c a l c i n e s .  This  suggests that some MoS was backfed due to gas leakage through 2  the rotary valves (see Appendix 13). The overall recovery of molybdenum in the process is over 99.9%, including the molybdenum in s o l u t i o n .  By  comparison, the overall recovery of molybdenum in the  multiple  hearth process is about 98.5% [125]. The S0  2  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 the off gases contain 0.8 to 0.9% S 0 .  0  = 25 cm/sec,  This value is v a l i d  2  only for this 12.5 cm reactor; substantial  increases can  be achieved in a larger reactor with large output, as w i l l be shown in Chapter 11. values of the S0  2  In Figure 9.11  content are given, from several  using different feed rates. Appendix 13.  some calculated experiments  The results are given also in  The recorded signal of S0  2  in the  infrared  analyser at steady state is given in Appendix 16.  195  F Figure 9.11.  , Mo S  0  2  feed  rate  (kg/m x day)  S0 in the off gases as a function of the MoS feed rate. 2  2  196  9.4  Slurry Feed Injection The p o s s i b i l i t y of feeding slurry instead of  solid molybdenite concentrate is attractive,  since this would  permit the direct feeding of the discharge from the  filters  in the concentration plant, thereby eliminating a dryer. Furthermore, slurry feeding would eliminate s q l i d handling which can present problems due to the tendency of molybdenite to compact. To test the a p p l i c a b i l i t y of slurry feeding to the f l u i d i z e d bed process, one experiment was performed in which a 50 wt-% solid suspension of molybdenite concentrate in scrubber l i q u i d was injected d i r e c t l y into the f l u i d i z e d 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 s i m i l a r to that obtained with s o l i d 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 Belton and Jordan [126]  reactor.  have reported that the  presence of water vapour enhances the v o l a t i l i z a t i o n Mo0 as 3  M0O3  • H0 2  when molybdenum metal is oxidized  of  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 a i r 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 i q u i d d i r e c t l y into the reactor.  Chapter 10  INDUSTRIAL APPLICATION OF THE FLUIDIZED BED PROCESS FOR MOLYBDENITE ROASTING  From the successful results obtained in this study, it  i s conceivable that the f l u i d i z e d bed process could be  applied on a larger industrial  scale to replace the  hearth furnace for molybdenite roasting.  multiple  This work has  shown, for example, that the f l u i d i z e d bed process is capable of producing calcines which have s u f f i c i e n t l y low sulphur contents  0.085 - 0.15% - to be acceptable for use either  d i r e c t l y 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 w i l l be shown that  there are positive advantages in using the f l u i d i z e d 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 i q u i d handling equipment.  In order to make this comparison, the  198  199  size of f l u i d i z e d bed reactor, the c o l l e c t i n g equipment and c i r c u l a t i o n system required for industrial  purposes, have  been estimated based on experience with the small p i l o t reactor. of this  10.1  plant  Details of these calculations are the main subject chapter.  Scale-up of Fluidized Bed Reactor The a b i l i t y  of the industrial  scale process to  achieve low sulphur calcines w i l l depend upon the variables and the internal  operating  dimensions of the reactor.  For  design purposes, i t w i l l be assumed that the concentrates to be roasted in the f l u i d i z e d bed will meet the  following  conditions: 1.  P a r t i c l e size of concentrates: - 3 2 5 m e s h , d < 15 m i c r o n s . ' s  2.  Calcium content 0 . 0 4 $ maximum.  of concentrates:  Both conditions are representative  of standard molybdenite  concentrates produced by f l o t a t i o n  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 c m / s e c at t h e r o a s t i n g temperature.  2.  Backmixed r e a c t o r (as demonstrated i n Chapter 7 using t r a c e r s ) .  200  10.1.1  Sulphur Content of Discharged  For a backmixed r e a c t o r function  Calcines  the e x i t age d i s t r i b u t i o n  9  E ( t ) of the p a r t i c l e s at the d i s c h a r g e ,  equal to the i n t e r n a l age d i s t r i b u t i o n f u n c t i o n t h i s c a s e , i s given  which i s  C(t)  in  by the e x p r e s s i o n :  (10.1)  E ( t ) = C ( t ) = I e~  t/T  t  The  f r a c t i o n of MoS  2  transformed  of the r e a c t o r i s given  (as Mo0 ) at the discharge 3  by the r e l a t i o n s h i p t=T  X  Mo0  (10.2)  F'(t) E(t) dt  3  t=0  where  F ( t ) i s the transformation  p a r t i c l e of MoS limiting  st;ep  are not f u l l y  2  f u n c t i o n of an i n d i v i d u a l  which depends on the nature of the r a t e  (s).  In t h i s c^se,  the rate c o n t r o l l i n g steps  understood but are thought to i n v o l v e ,  chemical c o n t r o l f o r a b r i e f p e r i o d , during of o x i d a t i o n period  proceeds r a p i d l y .  i n which the o x i d a t i o n  p o s s i b l y as a r e s u l t of s o l i d  first,  which the rate  T h i s i s followed  by a longer  rate i s considerably state d i f f u s i o n .  the k i n e t i c s has been presented i n Chapter 8.  slower,  D i s c u s s i o n of The l a c k of a  p r e c i s e understanding of the k i n e t i c s i n the l a t e r stage of  201  transformation makes i t F(t)  virtually  impossible to calculate  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 experimental  the  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 a n a l y s i s , the following relationship was found:  %S = 0.31435 - p.01052t + 0.00011237t  The calculated values of  %S  (10  2  in the calcines using Eq.  (10.4) are plotted in Figure 10.1,  together with the experi-  mental data from the f l u i d i z e d bed reactor for Brenda concentrates.  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 o r i g i n a l l y obtained.  For other temperatures, different  have to be developed, using the appropriate  relationships experimental  data.. A second attempt to f i t a mathematical  the experimental  data to  expression was made by assuming that the slow  202  step in the l a t t e r stages of oxidation was diffusiopal in nature.  It was further  assumed that the mathematics which  ho]d for diffusion of gas through porous s o l i d also applied to the oxidation of MoS at these high conversion l e v e l s . 2  Then for slab-shaped particles of MoS , 2  x  M  0  S  v a p l 2  ' s with e  time in the following way  X  MoS  =  1  2  X  Mo0  ~  3  (10.5)  1  The fraction of solids not t r a n s f o r m e d . , ^ <. , discharged r'  •  2  in the calcines from the reactor i s given, therefore by substituting Eq. (10.5) which is the expression for F ( t ) , into Eq. (10.2) to y i e l d *  t =• x  MoS  -  1 -  V  i t=0  1  s  1  I e-  tn  t  dt  (10.6)  Integrating this expression and neglecting the terms smaller than t "  3 / 2  ,  Eq. (10.6) becomes:  7t  The lower l i m i t of the integral should, s t r i c t l y speaking, be a non-zero value of time corresponding to the onset of the f i n a l stage of oxidation. However, since the f i n a l 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 t e d to the experimental  data of  Figure 9.7. f o r the Brenda concentrates by successive iterations.  The resulting equation, in terms of the residual  sulphur in the c a l c i n e s , is as follows:  = 1.35  1+  - T / t  3  - 1  (10.8)  The total time of reaction, T , c a n 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 r e l a t i o n s h i p :  JL(yO . 7 6 5 %S = 1.35 1 + 1  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 since a successful industrial operate with  times is not serious  scale reactor would have to  t > 20 hr to produce low sulphur c a l c i n e s .  Based pn the comparison outlined above, i t i s clear that either  Eqs. (10.4)  or (10.9) could be used to  scale-up the f l u i d i z e d 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 MoS p a r t i c l e s i z e , d~ , can 2  be included in Eq. (10.9) as follows  s  205  t  Figure 10.1.  , Average  time  of  reaction  (hr).  Experimental and calculated residual sulphur in the calcines as a function of the average retention time and temperature.  206  = 1.35 1 + exp -j  10"?d  2  0.765  T + 520  ; JO  I  2 . 7 6 5.  j  +  .  5  2  g J  (10.10)  However, i n s u f f i c i e n t data was obtained in the present investigation to check the r e l i a b i l i t y  10.1.2  of this equation.  Reactor Design and Simulation of Operating Performance  Based upon the sulphur levels desired in the c a l c i n e s , the dimensions of a reactor can be calculated. optimum ranges of roasting temperature  Since the  and s u p e r f i c i a l gas  velocity are very narrow, the following values have been used in the calculations: T = 550°C u o = 25 cm/sec The average p a r t i c l e size of concentrates is taken to be -325 mesh, d~ = 10 microns. s  The concentrates are assumed  to be 100% MoS , with 0.04% Ca. 2  The most important scale factor used in the c a l c u l a 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  207  dictated by the need to maintain the s t a t i c bed height, L  m  ,  such that the pressure drop through the bed w i l l not be excessive (see Chapter 5).  Consequently the reactor  diameter  required tp achieve a given value of if for a given feed  rate,  F , is given by the expression 0  0.333  (m)  where bed and  F t  (10.11)  is the total amount of calcines in the f l u i d i z e d  0  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 l u i d i z e d bed, should then be roughly 2.5 times the reactor  diameter  R ;=• 2.5 R 1  (m)  d  based on experience with the small p i l o t  (10.12)  reactor.  The % SQ in t h e off gases can be calculated 2  in terms of the feed rate of MoS and the total rate of 2  gas flow through the- bed,. G , as follows f  1 2.775 F.  %S0 =  (10.13)  0  2  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 / R m  = 1 and 0.5.  d  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. shallower bed, that is  /  L m  =  0-5,  the reactor  For a  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 t s industrial the multiple hearth furnace.  equivalent,  A 10 TPD f l u i d i z e d bed reactor  with 1.4 m deep bed that produces calcines with 0.1% S, w i 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 %S0 in the roasting 2  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 %S0 in the off gases i n 2  creases almost l i n e a r l y with increasing reactor capacity; at lower capacities the S0 sharply.  For L / R m  d  higher than for L / R m  2  content in the gases decreases  = 1, the calculated values are about 40% ri  = 0.5,  as would be expected.  The  RI ,  0.051 10  1  1.5  Reactor  length  (m)  L__  —i  2.0 Rd , Reactor  2.5 diameter  i  3.0 (m)  I  3.5 o to  Figure 10.2.  Calculated reactor dimensions as a function of the sulphur level in c a l c i n e s , for different output capacities.  210  calculated va]ues of S0  2  in the off gases of an i n d u s t r i a l  f l u i d i z e d bed of over ~5 TPD is greater than 2.5% under any conditions, and i t a shallow bed. S0  2  is about 4% S0  2  for a 20 TPD reactor with  These compare favourably with the level of  found during standard operations of multiple  furnaces of "1.5% S 0 , or 3-3.5% S0 2  2  hearth  for spray water  operation [127]. In the 12.5 cm f l u i d i z e d 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  Lm/Rd. = 0.5 to 1 . The predicted concentration of MoS at the i n j e c 2  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 initial  stage of oxidation would be completed.  The off gases  from this reactor would contain up to 10% S 0 , depending 2  T  8  = 550  °C iO JUL m  0.1 %  S  in calcines  <> / 6 CD CO O CP  10 C,  Reactor  Figure 1 0 . 3 .  output  15 (ton  Mo S  20 2  / day )  C a l c u l a t e d l e v e l of S 0 i n the o f f gases as a f u n c t i o n of the r e a c t o r c a p a c i t y . 2  25  212  upon the capacity of the reactor.  The calcines would t h e n  be discharged to a second, larger reactor where the transformation would be completed.  It  is important to note, however,  that o n l y 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  Fluidized  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 i n Figure 10.4.  A slurry feed consisting of 50% MoS s o l i d s 2  has been assumed; the same basic layout would apply however to the case of solids feeding.  The calculated dimensions  of t h e reactor and predicted process performance are given in Table 10.1. The throughput of the 10 TPD f l u i d i z e d 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 l u i d i z e d bed is from 3.3 to 5 times larger than the output of the multiple hearth when the l a t t e r is calculated on the basis of f l o o r area, or 33 to 50 times larger based on total hearth area.  Off Gases Gases Elutriated Calcines  •ligh Efficiency Cyclone System  Molybdenite Slurry Feed  V distribution Valve  Water Scrubber System  do  Fluidized Bed  Make-up Water  Reactor  Fluidized Bed  -Fuel-  Pump  Recycling System  Combustion: -Chamber  6  /Pump  Calcines  Air  Feeding System Preheated  Air  *<  Blcwer  Figure 10.4.  Layout of a 10 TPD f l u i d i z e d bed plant for roasting molybdenite concentrates.  214  Table  10.1  Dimensions and Operating Performance of a 10 TPD Plant for the Roasti ng of MoS 2  " ~  Capacity Concentrates p a r t i c l e  size  Calcium in concentrates Calcines Reactor dimensions Recirculation ratio Gas flow rate  1  1  '  — •  :  10 TPD MoS  : :  -325 mesh, d~ = 10 microns < 0.04%  :  0.10% S  : :  2.8 m diam. x 6.3 m height 3.41:1  2  s  19 m  3  Superficial gas velocity Temperature of the bed Off gases  : :  :  Cyclones efficiency  :  Scrubber efficiency  :  Overall collecting efficiency Solids from scrubber to reactor  : :  air/min  25 cm/sec 550°C 4.3% S0 95% 95%  2  99.75% 0.3 5 Kg/min calcines (0.70  Table  concentrate  Kg slurry 50% s o l i d s / min)  10.2  Calculated Output of Reactors r~"  .  1  Kg/m day per individual hearth 2  Multiple  hearth  480  Kg/m day total hearth area 2  48  Fluidized Bed 2400 -7  1260  10.2.1  Capital Costs  The capital cost of i n s t a l l i n g a 10 TPD f l u i d i z e d bed roaster and a multiple hearth furnace with a similar capacity are detailed in Tables 10.3 and 10.4 r e s p e c t i v e l y . The equipment l i s t e d for the multiple hearth process are standard for this type of plant [128], whereas the equipment required for the f l u i d i z e d bed process has been estimated based o n information obtained in the p i l o t f l u i d i z e d bed reactor and t h e previous c a l c u l a t i o n s .  The equipment costs  f o r 1 9 7 4 have been taken to be 1.3 times the 1971 prices N29J.  Type 3 1 6 stainless steel was assumed to cost three  times the p r i c e o f plain carbon s t e e l . A comparison of the capital costs in Tables 10.3 a n d 1C.4 indicates c l e a r l y that the f l u i d i z e d bed process c a n be expected to enjoy a considerable advantage over the multiple hearth furnace. bed  process  process.  The capital cost of the f l u i d i z e d  is 60% less than that of the multiple  hearth  This reduction in cost is due to two f a c t o r s :  less  d i v e r s i f i e d equipment, and a much smaller reactor.  1C . 2 .2  Operating Costs  A comparison of operating costs for the f l u i d i z e d bed a n d multiple hearth processes is given in Table 10.5. In arriving a t many of these f i g u r e s , values from a r e l i a b l e  216  Table 10.3 Capital Cost for a Molybedenite Roasting Plant using a  Number of Units:  Multiple  Hearth Furnace  Equipment  Description  Total Cost  1  Filter,  1  Multiple  1  Multiple hearth furnace, 16' , 30' height by 10 hearths, complete, i n stalled and with controls [128]  1  Rotoclone c o l l e c t o r CFM  2  Bins, 90 %, MoS  1  Bin, 10 %, MoS  8  Screw conveyors, 3 to 5 inches. <> f  60,000  5  Bucket elevators, 6 to 35 feet  40,000  3  Bins, 25 T, Mo0  35,000  2  Multiclone, bank of 4 each (SS 316)  40,000  1  E l e c t r o s t a t i c precipitator CFM (SS 316)  1  Stack, 4'6" x 148' (SS 316)  1  Impact Mi 1.1  5,000  1  Packer Unit  10,000  6' x 4' discs  $  hearth dryer, 10' < p x 4 hearth  (SS 316), 45,000  6,500 200,000 1 ,050,000 30,000 46,000  2  6,500  2  3  16,000  200,000 35,000  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  Filter,  1  Blower, lobular,  1  C y l i n d r i c a l reactor tank (SS 316, 1/2" t h i c k ) , 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)  1  Discharge system (SS 316)  1  Slurry feeding system  1  Agitation  1  Bin, IQ %, MoS  3  Bins, 25 %, Mo0  35,000  1  Scrubber system  55,000  1  Stack (SS 316, ? > x 150')  25,000  1  Scraper system for reactor  -  Piping, insulation for reactor  1  Screw conveyors  -  Instrumentation  25,000  1  Packer Unit  10,000  6' x 4' discs  $  500 CFM  6,500 5,000  8,000 10,000 5,000  tank for MoS  2  slurry  25,000 6,500  2  3  (SS 316)  15,000 20,000  [3"$)  5,000  CAPITAL COST ' Installation TOTAL CAPITAL COSTS 1  •~  382,000 440,000 ~ $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  Power  0.1  tf/lb  Gas cleaning  1.5  tt/lb  Total Operating Costs  6.7  tf/lb  industrial  source [128]  0.04  <t/1 b  0.1  tf/lb 1 tf/lb  -4.2  have been used.  tf/lb  Labour has been taken  to be the same for both processes, even though lower values of maintenance and gas cleaning costs for the f l u i d i z e d bed are probable. attention,  This is because the reactor requires less  e . g . continuous cleaning of the discharge ports,  than the multiple hearth furnace.  In addition there is no  expense for the operation of e l e c t r o s t a t i c The current prices for the  precipitators.  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 l u i d i z e d 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 l u i d i z e d bed reactor  using slurry feed is given in Figure 1 0 . 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 f l u i d i z e d bed reactor, no fuel  is required, and indeed a  large surplus of heat is generated From this a n a l y s i s , i t lower capital  (4.31  x 10  5  Kcal/hr).  seems clear that substantially  investment and operating costs can be achieved  using the f l u i d i z e d bed process rather than a multiple furnace for molybdenite roasting.  hearth  220  vaporized water from scrubber slurry 2.0 X \tf kcal/ hr  [preheat air discharged calcines 2.5 X IO kcal/hr 2  heat losses in reactor and lines  vaporized water from slurry feed  io  1.8 x kcal/hr  /4.26 X io  3.28 X |0  4  kcal/hr  kcal/hr  total heat to the reactor calcines 2.5 XIO kcal/hr  6.60 X |0  kcal/hr  5  3  heat of reaction 6.58 X I O  Figure 10.5.  5  kcal/hr  fuel 2.5 X io  3  kcal/ hr  Thermal balance for the f l u i d i z e d 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  recirculati  f l u i d i z e d 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) inert material  To avoid s t r a t i f i c a t i o n (sand)  of the calcines and  inside the reactor, the l a t t e r should  have a wide size d i s t r i b u t i o n , such as -40/+140 mesh; a l s o , the f l u i d i z e d bed reactor should be operated with a superficial  gas velocity in excess of 18 cm/sec at the roasting  temperature. 3)  For the bimodal system of particles of coarse  sand and fine c a l c i n e s , the concentration of the calcines 221  222  in the bed has a s i g n i f i c a n t l y larger effect on the f l u i d i z a tion properties than the sand for mixtures containing over 20 wt-% c a l c i n e s .  For the range 20-60 wtr% c a l c i n e s , an  empirical relationship was obtained to estimate the size of the gas bubbles along the f l u i d i z e d bed. 4)  Using the pneumatic injection system for  the  feed and a long average residence time for the s o l i d s , e . g . over 15 hours, the reactor is v i r t u a l l y  backmixed for the  solids with a large dispersion of gas along the f l u i d i z e d bed.  However, a small concentration p r o f i l e seems to exist  for the  MQS  2  from the feed point to the top of the bed at  steady state conditiops.  For the rate of e l u t r i a t i o n  fines from ^he bed, an exponential  of  expression of the super-  f i c i a l gas velocity was found to hold. 5)  The transformation of MoS to 2  M0O3  i s a complex  process which involves an i n i t i a l fast stage which seems to be chemical reaction c o n t r o l l e d . final  This is followed by a  slow, and apparently s o l i d diffusion controlled regime.  Vaporization of  M0O3  and subsequent condensation seems to  play an important role in the kinetics of oxidation. 6)  The optimum operating conditions of the  f l u i d i z e d bed reactor l i e within the following ranges:  223  Superficial  gas v e l o c i t y  :  18  Temperature  o f t h e bed  :  520  Average  7)  residence  time  20  of solids:  t o 35  cm/sec  t o 550°C t o AO h r s  To obtain calcines with a low sulphur concentra-  t i o n , the molybdenite concentrates must not contain more than about 0.04% calcium, with an average p a r t i c l e size of MoS less than 15 microns and 100%-325 mesh. 2  8)  For scale-up purposes, a semi-empirical r e l a t i o n -  ship based on the average size of molybdenite p a r t i c l e s and the temperature of roasting has been used.  The scale-up  indicates that industrial size f l u i d i z e d 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 l u i d i z e d beds (5 to 20 TPD) should range from 3 to 4%. 9)  The capital cost and operating costs of a  f l u i d i z e d bed process appear to be about 50% and 60%, respect i v e l y of the capital and operating costs of a conventional multiple hearth process of the same capacity.  224  SUGGESTIONS FOR FURTHER RESEARCH  1)  A full  size p i l o t plant of 0.3-0.5 ton/day is  required to assess on a larger scale the f e a s i b i l i t y of this process for possible application in industry.  2) stratification  More information is required concerning the process for different sizes and densities of  p a r t i c l e s , as well as the influence of external  variables on  phenomenon such as the f l u i d c h a r a c t e r i s t i c s , s u p e r f i c i a l gas velocity and bubble size and d i s t r i b u t i o n .  3)  Additional research on the solid and momentum  transfer of a pneumatic injection system discharged into a f l u i d i z e d bed reactor is needed.  4)  Further studies on the kinetics of the MoS  2  roasting in a f l u i d i z e d bed seems d e s i r a b l e , as well as a study of the f i n a l  5)  stage of sulphur elimination.  More research on the e l u t r i a t i o n  phenomenon  is needed to predict from given c h a r a c t e r i s t i c s of the and solids in a f l u i d i z e d bed reactor the rate of  fluid  elutriation  and the size d i s t r i b u t i o n of the elutriated material.  NOMENCLATURE  b  adsorption factor of gas  ( - )  C  capacity of reactor  (ton/day)  C  concentration of tracer gas  (cm /lt)  concentration of oxygen in the reaction gas  (gr-mole/cm )  C  Q 2  C  p,s'  C  C  p s s  s  C°  e  a  t  c a  P  a c i t  y ° f s o l i d gas  3  (cal/gr-mole)  concentration of tracer solid  (%)  concentration of MoS at feeding point  {%)  concentration of MoS in calcines  (%)  2  C*  C(t),  h  3  2  internal age d i s t r i b u t i o n function of tracer as a function of time t and dimensionless time 6  ( -')  dg  bubble diameter  (cm)  d  cloud diameter  (cm)  £  C(e)  elutriated  225  226  axial dispersion c o e f f i c i e n t of solid  (cm /sec)  axial dispersion c o e f f i c i e n t of gas  (cm /sec)  radial dispersion c o e f f i c i e n t of solid  (cm /sec)  radial dispersion c o e f f i c i e n t of gas  (cm /sec)  average diameter of calcine parti cles  (cm)  average diameter of sand p a r t i c l e s  (cm)  a v e r a g e diameter of MoS particles  (cm)  diameter of MoS particles  (cm)  diffusion coefficient  (cm /sec)  exit age d i s t r i b u t i o n function of tracer in the bed as a function of time t and dimensionl ess time 9  ( - )  f  expansion factor of f l u i d i z e d bed  ( - )  f.  volumetric factor of f l u i d i z e d bed  ( - )  f  voidage factor of  (  D a  °  q , y  -  r  D r  q  d c a 1  d S 7  -;  2  d s  2  D E(t),  r s  F(t)  E(6)  2  2  2  2  - ) • :  volumetric fraction of downflow phase  (cm /cm )  transformation function of an individual MoS p a r t i c l e  (  2  F*  fluidization  2  elutriation  flux  3  -  3  )  (gr/cm sec) 2  227  volumetric fraction of upflow phase  (cm /cm )  feed rate of molybdenite concern trates  (gr/min), (Kg/hr) (Ton/day)  gas flow rate at 25°C  3  3  (1/sec)  (m /min) 3  gas flow rate at roasting temperature  (1/sec) (m /min)  free energy change  (Kcal/mol )  number of holes in the gas d i s t r i butor gri d  (  heat transfer c o e f f i c i e n t between gas and particles  (cal/cm sec °C)  heat of reaction  (Kcal/mole)  elutriation  (gr/min) (Kg/hr)  rate  3  - )  2  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 c o e f f i c i e n t between >ubble and cloud  (sec )  gas transfer c o e f f i c i e n t between cloud and emulsion  (sec )  - 1  - 1  228  overall transfer c o e f f i c i e n t of gas between the gas phase and the emulsion phase  '  (sec ) - 1  f l u i d i z e d bed height  (cm),  (m)  s t a t i c bed height  (cm),  (m)  minimum f l u i d i z i n g bed height  (cm),  (m)  s t a t i c bed height of sand  (cm),  (m)  adsorption c o e f f i c i e n t  ( - )  number of holes per unit area of the gas distributor  ( - )  Nusselt number  ( - )  partial  (atm)  pressure of oxygen  total pressure  (atm)  Peclet number for mass transfer  ( - )  heat flow  (cal/sec cm )  heat losses by convection  (cal/sec cm )  heat losses by radiation  ( c a . / s e c cm )  r e c i r c u l a t i o n ratio  (.-•  2  2  2  )  229  reactor  diameter  (cm),  (mt)  reactor  height  (cm),  (mt)  Reynolds number with respect to the gas Reynolds number with respect to the sol id  ( - )  ( - )  reactor cross sectional area  (cm ),  s p e c i f i c surface area of mplybde ni te concentrates  (cm /mole-gr)  surface area of bubble  (cm )  temperature of the bed, average  (°C)  temperature of the gas  (°C)  temperature parti cles  (m )  2  2  2  2  of reacting MoS  2  (°C)  time  (sec),  (min),  average residence time of solids in the bed  (min),  (hr)  s u p e r f i c i a l gas velocity of gas^ at minimum f l u i d i z i n g conditions  (cm/sec)  s u p e r f i c i a l gas velocity  (cm/sec)  (hr)  230  bubble r i s i n g  velocity  emulsion downflow  (cm/sec)  velocity  (qm/sec)  bubble volume  (cm )  superficial of solids  (cm/sec)  3  velocity of downflow  s u p e r f i c i a l velocity of upward flow of solid  (cm/sec)  total weight of solids in the bed  (gr),  (Kg),  (Ton)  (gr),  (Kg),  (Ton)  (gr),  (Kg),  (Ton)  weight fraction  of calcines in  the bed  weight fraction  of sand in the  bed  '  Fraction of MoS in the discharged ca1ci nes 2  fraction of M o 0 calcines volumetric bubbles  3  ( - )  in the discharged  fraction  of wake in the  (-)  fraction of the bed consisting of gas bubbles  ( - )  voidage of s t a t i c bed  (  voidage of b e d at minimum f l u i d i z i n g conditions  ( - )  emmissivity  ( ^ )  coefficient  -  )  231  e  voidage of f l u i d i z e d bed  ( - )  n  collection efficiency  {%)  6  dimensionless time  (  f  T  )  y  g  v i s c o s i t y of gas  (poise)  P  s  density of s o l i d  (gr/cm )  P  g  density of gas  (gr/cm )  P  s  average bulk density of solids  (gr/cm )  P  c  average bulk density of calcines  (gr/cm )  P 1  average bulk density of sand  (gr/cm )  a  Stefan-Boltzmann constant  (cal/cm sec °K*)  variance of d i s t r i b u t i o n  (sec )  i  total time of transformation  (min),  ip  centroid of d i s t r i b u t i o n  ( - )  Y  stoichiometric factor M0S2/M0O3  ^ ( - )  a  s i  a  2  l  3  3  of  3  3  3  2  2  (hr)  REFERENCES  Zelikman, A . 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N . , Nienow, A.W. and Agbim, A . J . , Trans. Chem. E n g . , 50 (1972), 310-323.  Inst.  G i b i l a r o , L.G. and Rowe, P . N . , Chem. Eng. S c . , 29, No. 6 (1974), 1402-1412. Davidson, J . F . and Harrison , D., "Fluidization op. cit., p. 26-45. K u n i i , D. and Levenspiel op. c i t . , p. 152.  ?  0. , "Fluidization  K u n i i , D. and Levenspiel , 0. , Ibid.,  Particles,  Engineering,"  p. 76.  Lewis, W.K. , G / i l l i l a n d , E.R. and Lang, P-M., Ch. Eng. Progress Symp, S e r i e s , 5_8, No. 38 (1 962), 65.  239  99.  Levenspiel ,. 0. and Bishoff, K. , Ind. Eng. Chem., 53 (1961 ), 313. ' -~  100.  Levenspiel, 0 . , "Chemical Reaction Engineering," 2nd E d . , J . Wiley and Sons, New York (1972), 277.  101.  Levenspiel, 0 . , Ibid.,  102.  K u n i i , D. and Levenspiel, 0 . , I. 8, No. 3 (1 969), 402-406.  103.  K u n i i , D. and Levenspiel, 0 . , pp. cit., pp. 186-187.  104.  Chiba, T. and Kobayashi, H., Ch. Eng. S o . , 25 (1970), 1375-1385.  105.  Drinkenburg, A.A.H. and Rietema, (1972), 1765-1774.  1 06 .  p. 277.  Kunii D. and Levenspiel , 0 . , op. c i t . , p. 178,  and EC Fundamentals,  "Fluidization  Engineering,"  K., Ch. Eng. S c . , 27  "Fluidization  Engineering ,'  107.  Kobayashi, H. , A r a i , F. and Sunakawa, T . , Kogaku-Kagaku, 31_ (1 967), 239.  108.  K u n i i , D., and Levenspiel, 0 . , op. cit., 140-161.  109.  "Fluidization  !  Engineering,"  Van Deemter, J . J . , Proc. of an Int. Symp. on F l u i d i z a t i o n , Netherland University Press (1967), 334-347.  110.  K u n i i , D. and Levenspiel, 0 . , op. c i t . , p. 158.  "Fluidization  Engineering,"  111.  Bart, R., Ph.D. Thesis, MIT, Cambridge (1950), C i t . in K u n i i , D, and Levenspiel, 0 . , "Fluidization Engineeri n g , " op. c i t . , p. 147.  112.  Ammann, P.R. and Loose, T . A . , op. cit., p. 891.  240 113.  Ammann, P.R. and Loose, T . A . , Ibid.,  114.  Cardoen, C . R . , op. oit., p. 66.  115.  Ong, J . , op. oit., p. 71-73.  116.  Zelikman, A . N . and Belaevskaya, L . V . , op. oit., p. 18.  117.  Cardoen, C.R. and Sepulveda , J . , Intermet. B u l l . , 2, No. 3 (1973), 37.  118.  Ammann, P.R. and Loose, T . A . , op. oit., p. 892.  119.  Cardoen, C . R . , op. oit., p. 69.  120.  Ammann , P.R. and Loose, e . A . , op. oit., p. 892.  121.  Kel1ey, K . K . , op. oit., p. 72.  122.  Grimes, G.R and Witkamp, G . , op., oit., p. 7.  123.  Sidgwick, N V . , "The Chemical Elements and Their Compounds," V o l . II, Oxford Press (1952), 1046.  124.  Sidgwick, N.V., Ibid.,  125.  Grimes, G.R. and Witkamp, G . , op.' oit., p.. 2.  126.  Belton, G.R. and Jordan, A . S . , J . Phys. C h e m 6 9 , (1965), 2065-71.  127.  Grimes, G.R. and Wilkamp, G . , op. oit., pp. 2-8.  128.  I.A.  129.  Popper, H . , E d . , "Modern Cost Engineering Techniques," McGraw-Hill ,•• New Y o r k (1970), 1 09-192. Baerns, M., Fetting F. and Shurgerl , K . , Chem. Ing. Tech. (1 963), 35, 699.  130. 131.  p. 892.  T  T  Warren, personal  p. 1046.  No. 6  communication.  Schurgerl, K. , Chem. Ing. T e c h . , 38 (1966), 1 1 69.  241  APPENDIX 1  VAPOUR PRESSURE OF Mo0 AS A FUNCTION OF TEMPERATURE 3  Temperature (°C)  243  APPENDIX 2  ENTHALPY AND FREE ENERGY OF REACTION  * S(g) s  °2 2 ° (g) S  +  298-2000°K  2  AH° = 0172,630 - 1.49% + 0.712 x 1 O + 0.336 x 10 T -  -3  T  2  cal/mol  1  cal/mol  AG-J = -173,240 + 34.62T MoS  + 7/2 0 ^ M o 0  2 ( s )  2  3 ( s )  + 2S0  298-1068°K 2  A H " =, -297,590 - 4.49T + 7.862 x 10~ T + 3.191 x 10 T 3  5  2  1  cal  A G j = -266,840 + 12.75T x logT - 1.35 x IO  -3  T  2  + 36.44T + 1.54 x 10~ T " 5  1  cal M o S 2  (s)  +  3  °  2  2  M o  ° ( ) 2  s  + 2 S0  AG? = -224,340 + 9.82T log T + 0.2 x IO" T + 2.56T 2 t  2  (s)  +  0  2  +  M  ° ° 3 (  s  )  298-1100°K  2  3 c  a  l  298-1068°K  244  A G = -45,500 + 2.93TlogT _ 1.55 x 10~ T + 1.54 x 10 T 3  2  T  5  6 Mo0  3 ( s )  1  + MoS  2 ( s )  I  7 Mo0  2 ( s )  + 2 S0  A G = 30,660 - 7.76TlogT + 9.5 x 10" 9.24 x 10" T~ - 200.84T T  5  1  3  298-1068°K  :  T  cal  2  -  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 p l a s t i c reactor Ottawa sand (-40/+140 mesh) (-325 mesh)  M0O3  f (It/min) (air 25°C) G  1)  10  AP  (cm/sec)  (cm H 0) 2  (cm)  2.2 4.9 7.0 7.7 8.4 9.0 9.1  8.5 8.7 8.8 8.9 8.9 9 9.3  5.1 3.7 8.9 8.4 8.4 9.0 9.5 10.3  8.7 8.7 8.8 9.0 9.3 9.5 9.9 10.3  L  f1  M0O3  wt-%  Lm = 8.5 cm  L  m  .5 .8.6 12.0 16.1 24.0 32.2 40.8 2)  Uo  mf =  8  '  8  2.0 3.4 4.6 6.2 9.0 12.0 15.0  20 wt-% Mo0 m -  3  L  =  8  5.0 7.0 8.6 12.0 16.1 24.0 32.2 40.8  7  c m  L  mf  2.0 2.7 3.4 4.6 '6.2 9.0 12.0 15.1  =  9  '°  c m  248  u (It/min)  (air 3)  AP  0  (cm/sec)  25°C)  40 w t - % Mo0 = 11.7 cm  (cm  H 0) 2  (cm)  3  l  m  5.0 7.0 8.6 12.0 16.1 24.0 32.2 40.8  L  mf  =  12 .2 cm  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  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  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  1  4)  50 w t - % MoO ' L  m  =  1  5  -  2  c  ™  L  5.0 7.0 8.6 12.0 16.1 24.0 32.2 40.8 5)  mf  =  15 .3 cm  2.0 2.7 3.4 4.6 6.2 9.0 12.0 15.1  60 w t - % MoO L = 19.2 cm 3  m  5.0 7.0 8.6 12.0 16.1 24.0 32.2 40.8  L  mf  2.0 2.7 3.4 4.6 6.2 9.0 12.0 15.1  -  20 .4  249  Uo ( I t / m i n) ( a i r 25°C) 6)  0% M o 0 m  L  =  8  c  (cm/sec) (100  3 m  wt-%  100  L  m  wt-%  H 0) 2  f  (Cm)  m f  2.0 2.7 3.4 4.6 6.2 8.0 9.0 11.0 12.0 15.1 22.0 27.5 M0O3  = 2 7 . 5 cm 2 3 5 6 9 11 16.1 21 24  (cm  L  Ottawa Sand -40/+140) L = 8 . 1 cm  5.0 7.0 8.6 12.0 16.1 21.0 24.0 30.0 32.2 40.8 49.6 75.0 7)  AP  1.6 2.8 3.5 5.6 7.3 9.0 9.4 9.8 10.0 10.0 10.1 10.1  8.0 8.0 8.0 8.0 8.1 8.1 8.1 8.2 8.4 8.8 9.3 10.8  4.5 9.0 5.2 5.6 7.6 6.0 6.6 10.2 8.6  27.5 27.5 27.6 27.7 28 28.3 29.5 31 38  (-325) L  mf  1.0 1.4 2.0 2.7 3.5 4.2 6.2 8.0 9.0  =  2  7  -  8  cm  250  APPENDIX 5  AVERAGE MEASURED DIAMETER OF BUBBLES FROM HIGH SPEED PICTURES 1  ,  Test No  wt-%  u  0  Calci nes cm/sec  B.106-A.1 A.2 A.3  0  B.106-B.1 B.2 B.3 B.4  20  B.106-C.1 C.2 C.3 C.4  40  B.T06-D.1 D.2 D.3. D.4 D.5  60  B.106-E.1 E. 2 E.3 E.4 E.5  80  II  II II M  II  n n  II II II II  II  n II  20 30 40  Distance  0-5 0.5  from  0.9 1.3  0.6 1.4* - e ! 1 ,4 2.3*  15 20 30 40  1.5 1.2 1 .0* 1.6  15 20 30 40 10  1 .9 2.6* 0.9 1 .2 1.2 1.6 0.8 1 .4  1  1.2  c m  the  gas distributor,  cm  5-10 10-15 15-20 20-25 25-30 3Q-35  15 20 30 40  10 15 20 39 40  ^b'  1.9 2.1 1 .8* 1.9  0.9 1 .8 1.9 2.0 2.6  1 .2 1 .8 2.3  1.0 2.6 2.8  1 .4 1 .6 6.1  1 .6  2.4 3.2 4.6*  "°* 2.5 4.4*  3.3 5.6  1.9 1 .7 2.5* 3.7  3.4 - * 3.7 3.6  4.8 1.8 2.6 2.8 0.5  2.4 3.6 3.0 1.1  2  1 .2 2.4 2.2 2.4 2.7  2  2  1.6 3.7 1 .8 2.8  T  Average from two different picture measurements.  3.0 2.6 4.4 4.0 4.4 2.1 3.9 4.5 1.9 2.9 3.7 4.2  6.3  2.2  2.4 4.5 5.2 2.0 2.9 4.2 2.8 3.2  -  3.1 5.8  4.3  APPENDIX 6  STRATIFICATION OF SOLIDS TEST RESULTS RUN B.I 20 wt-% Mo0  3  - 325 mesh  80 wt-% sand -40/+70 mesh L = 25 cm m G  f  = 28.2 lt/min  U  0  = 30 cm/sec  time of test = 180 sec time of delta input = 3 sec Sample No  Di stance cm  S i l i c a sand gr  Calcines gr wt-%  B.l-1 2  0-2  35.5  4.45  2-4  28.7  0.55  11.1 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  6  10-12  7  0.15  0.0 0.2 0.1  8  12-14 14-16  48.5 73.1  0.07 0.17  43.8  0.12  0.1  9 10  16-18 18-20  53.5 49.2  0.15 0.08  0.1 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-% Mo0  -325 mesh  3  80 wt-% Ottawa s i l i c a sand -40/+70 mesh 1m = 25 cm G  f  = 28.2 It/min  U  0  = 30 cm/sec  time of test:  Sample . No  30 sec  Di stance,, cm  S i l i c a sand gr  Calcines  MoS  2  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 6-8  47.2 57.8  0.45 0.83  0.9  11 .40  1 .4  6.75  50.0 .  6.16  4 5 6  '  8-10 10-12  33.8  1 .05  10.9 21 .9 31 .1  0.26 0.12  7 8  12-14  36.4  9.45 13.45  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  11  20-22  38.8  17.37  31 .3 30.8  0.0  12  22-24  13  24-26  26.3 23.1  12.96 1 5.30  33.0 39.8  0.0 0.0  CONTINUED  —  253 APPENDIX 6 (Continued) RUN B.4 20 wt-% Mo0  3  -325 mesh  80 wt-% Ottawa s i l i c a sand -40/+70 mesh 1m = 25 cm G  f  = 28.2 It/min  U  0  = 30 cm/sec  time of test - 15 sec time of delta input = 3 sec  Sample No  Calcines  Distance cm  S i l i c a sand gr  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  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  5 :  MoS  2  gr  CONTINUED  wt-%  %  APPENDIX 6 (Continued) RUN B.64 60 wt-% Mo0  3  -325 mesh  4.0 wt-% Ottawa sand -40/+140 mesh G  f  = 28.2 It/min  U  0  = 30 cm/sec  time of test = 180 sec  Sample No  Distance cm  S i l i c a sand gr  Calcines gr  , wt-%  B.64-1  0-2  27.76  19.05  40.7  2 3  2-4  21 .82  16.60  43.2  4-6  16.89  18.26  51 .9  4  6-8  18.69  54.3  5  8-10 10-12  15.73 15.33  56.8  14.11  20.15 20.63  59.4  17.00 17.47  60.2 57.5  24. 52  61 .6 62.3 63.8  6 7 8  12-14  11.26  14-16  9 10  16-18 18-20  12.93 15.32  11  20-22  13.13  12  22-24  12.12  24.33 23.19 21.51  13 14  24-26 26-28 28-31  12.82 13.82  23.43 25.52  64.6 64.8  14.18  28.82  67.0  15  14.72  CONTINUED  63.9  255 APPENDIX 6 (Continued) RUN B.65 60 wt-% Mo0  3  -325 mesh  40 wt-% Ottawa sand -40/+140 mesh 1 = 31 cm m G  f  = 10 It/mi n  U  0  = cm/sec  time of test = 180 sec time of delta input = 3 sec Sample No B.65-1 2 3  Distance cm  S i l i c a sand gr  Calcines  MoS  2  gr  Wt-%  %  17.57  35.8  2-4 4-6  31 .56 40.05  6.63  42.72  13.5 12.1  0.08 0.08  0-2  4 5  6-8  34.60  5.91 . 8.25  8-10  16.39  21 .29  6  10-12  10.50  7  12-14 14-16 16-18 18-20 20-22  0.05  19.3  0.08 0.53  25.23  56. 5 70.6  8.93 7.55  22.95 21 .56  72.0 74.1  1 .60 1.21  9.32  26.82  74.2  0.80  9.74 8.17  28.43 24.91  74.5  0.73 0.60  8.52 8.65  26.93  76.0  13  22-24 24-26  28.72  76.9  0.53 0.43  14  26-28  12.62  54.84  81.3  0.25  15  28-31  9.95  51 .63  83.8  0.05  8 9 10 11 12  75.3  CONTINUED  1 .83  APPENDIX 6 (Continued) RUN B.93 60 wt-%  -325 mesh  .M0O.3  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 0-2  S i l i c a sand gr  Calcines gr  wt-%  17.69  39.7  12.56 17.44  31 .0 45.6 57.7  B.93-1 2 3  2-4 4-6  4  6-8  5  8-10  20.84 16.73 14.11  6 7  10-12  11 .44  12.-14  10.26  21 .86 22.22  8 9  14-16 16-18  10.51  20.00  68.4 65.6  10  26.35 24.71  66.4 66.1  11  18-20 20-22  13.33 12.67  25.71  12 13  22-24 24-26  13.25 1 2.37 9.46  26.70 22.53  66.0 68.3 67.7  14  26-28  5.76  13.25  69.7  15  28-31  3.07  8.01  72.3  26.84 28.00  22.84 23.82  CONTINUED  62.8 65.6  257  APPENDIX 6 (Continued) RUN  B.94  40 wt-% Mo0  -325 mesh  3  60 wt-% Ottawa S i l i c a  sand -40/+140 mesh  1 = 31 cm m G =28.9 It/min f  U  0  = 30 cm/sec  time of t e s t = 180 sec time of d e l t a  input = 3 sec  Sample No  Distance cm  B.94-1  0-2  2  Silica gr  sand  Calcines  MoS  2  gr  wt-%  %.  39.48  12.74  24.4  0.08  2-4  35.39  9.57  21 .9  0.35  .3 4  4-6  31.32  11 .72  27.2  0.15  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 8  0.55  14  26-28  4.07  3.83  48.5  0.40  15  28-31  1.92  2.05  51.6  0.35  T  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  U  0  = 30 cm/sec  time of test = 180 sec  Sample No B.95-1  Dis tance cm  S i l i c a sand gr  Calcines gr  wt-% 13.8 14.3 16.0 17.9 18.1  0-2 2-4  42.46 37.72  6.80 4. 21  3 4  4-6  39.03 39.62  6.50 7.57  5 , 6 7  8-10  41.42 39.85 43.96 35.12  9.00 9.13 10.72 9.48 11 .00 11.11  20.7  2  6-8 10-12 12-14 14-16  10.0  19.6 21 .3 20.1  8 9  16-18  10  18-20  43.86 42.59  11  20-22  43.97  1 2.09  21 .6  12 13  22-24 24-26  42.17 43.41  11 .83 12.00  21 .9 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 MoS CONCENTRATES, AND 2  CYCLOSIZER ANALYSIS OF MoS CONCENTRATES 2  SIZE DISTRIBUTION OF STANDARD OTTAWA SILICA SAND  ~  mesh, US Std.  T  wt-%  -  d , mm s  -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  1—:  d  s  = 0.311 mm  1  CC-1?»J<  PARTICLE SI2 E DISTRIBUTION  ^y .^.ilL5^.i iriIi b  S A M P L E IDENTIFICATION Density Preparation  **-  g/cc  8 0  0 . 35 g for  n  LIQUID: Density  0  ,  8  2  1  (HoS  9  LiH  in  30 ml  +rtr  i 1 j t~ • c b i —— ' 1—-t-1*0  Tt.t.r  '  Viscosity.  ir  Hi:  tta  L f lie in: i^:: LHi  /tiiiiti"  U  iii  •if: Liif •Hii.  it Li Li5i Hi? ii±i  ANALYSIS NO. 1^86  :.r.tt-  : rv£  is 5§  •§3r  :  ± f S r  Iii!  :: j  ai;  30 C  Lii  r  : i : ; Hi i  i j: i :; l:  iii:  : ;  I Tin !"'"tM" *  - iiir.|: t '  ..-tii-i  • i , 60  rt'.: i i i i i :  Hift  lliiLi  .- :l4-t:. • : • so  tftf  ;:::  4  •!•::'  :  p. ;40  mi:. • t-' -. : : t l . t 51 ,  if iii  i - i i i - i:i i l • ;-:::t.-:;.virxL: • » • M it t i  1  •Li: i : Li i  -T-HH 80 Hi  r: 3rrr  iiii  iii:  ii »  iiii  at  ::+. :.'t: riLHlHi  iiii  i pi i i i i * i i i ' •*-tlS5S jr_g:  tit "till'-  m\  i f f i :i~iiiL-  5 u  :rii  iiii il}: Liu  ::f:t-:T-i:.t. r. j .  D  112  .! i •.:•.: i.: : 9o H i : " :;M-:p-r  Lii'L  : : :V  > <  Holl  R.  BY  ;  SI  U:il  l-i  LM-:  :::i  Hr? ilHT  40  cps  0  liH :. _i_:..  : 1 • H i - \: •V 50  50  8  i-rt-j- Tt"ii£-  M •;ii"f  :'t:i:  :1  •-.'•l-f 60 LLiiBt  60  6  iiii  iiU  V  ss.  iiii  LlL-i H ! is o •rr=t4r #LtfLS t4:r.r.  80  ,  DATE  RATE  :  100  7  9-26-197 * 4  •  enda  .g/cc  !"T:; 9b : ;. t.:.  C/5 tn <  B  SEOISPERSE-E  Hi';  2  )  sampl e were, . d i s p e r s e d ..u. I t r a s o n i x a J I y . _ w . i t h j s i m u l t a n e o u s . _ s t i r r ing  5 minutes  UJ  ?  H_i"  •iv  :ii: ::i: LH; 'ti' iiii 113  r  :  11-iH  : : : :  ' T- f H i : .fRljl IHitT li-Jt::. i i ii i i i i ; "t- " : ;  m  Jtrr  >A___L____  tri r  1  0.8  0.6  ni'- •  0.5  0.4  0.3  EQUIVALENT SPHERICAL DIAMETER, MICRONS. ro  M I C R O N DATA LABORATORIES, INC. 22 Secord Dr., St. Cat larines, Ont., C a n a d a L2N IK8 Phone (116) 934-0629  PARTICLE SIZE S A M P L E IDENTIFICATION Density  *».80  Preparation  Molybdenite  g/cc  5._"»nutes  (HoS ) g /  r-rrrrriv  ~—"Pico  Sffjif .-,-4-1  iff: I i Iril  ::!:  j f f i r \ t 70 :.;4_!. ::.: :. ; \ i i 1 _;V._4—  Hill';::;  iff!  i ...-4-...-.  iff':  i f\*o i-fii^ : i i i V :  •fffH-: :ru  :  :  ::n  if::  mi..::::  tffrj  ::::  iiiibii  i'fff.  tO  in  in-;  io  • .i-i 70:.i§rrhl: itfrjr.r. : : : 60  i : i f  iff:  tflfi  : i i :•  ifff i f t f  tiM fffri  : i:; f f f 30 :r£::::;._:  rcr-l-:-.  •.rr;r:f:i:. f.fH- " 1 " i o ::::  rfrFEif:  Hf:i  *4-Li.iff: f t f f " f f - i i . i> rCrrrtn.:-. ^ii i .  TO f iffi"  a t 30 C  :;i;  f-ffiT  :::>  .:;i4  i i f i —A  .1--j ffffffi  \ ::\  112  ;f n liii  inf. i f i i i  :  li»78  I W  : : ... I-.J 50  iH  tiff  1.  • ••  fff.T.'  i -i i'i'.ff; r j r r r r r *(j  R. H o l t  RATE  iff;  - Jr4-  ii. '•'•'• 30  M  iff  Mm  fiff-  fff-.fr ffffr  U ifff  BY  ifff  filii"  :.j.::::  9-19-197^  :;t:  ~t-i~r-f" :  DATE  ANALYSIS NO.  ffffi ;fff  .tiff  muii •Hi  i::i  ifff  k|3 f f e  •; -;  mm  '• so>  iEffifi  cps  stirring  M  it-n  iii: iff ::::  r r  m-  •ifi 1 R 1  i f i f iiii ;|t:::!:: iii: s m :::: It4i-::il  iiifHif;'  :  simultaneous  ;.  iff.ff  /  .7..680  Viscosity  c  i n 30 ml S E O I S P E R S E - S  :::|:::  :  c  d i s p e r s e d ul t r a s o n i c a l 1 y w i t h  ;  •:  II.C. Moly  2  LIQUID: Density .Q.-.821?  0 - 3 5 9 s a m p l e were  for  concentrate  GC-19924  DISTRIBUTION  iii: '•|* ff§  P+3  lit- • r t  : : T :  ifff.  ::::  fHfrF 1  S O . 40  0.8  0.6 O.S 0.4  EQUIVALENT SPHERICAL DIAMETER. MICRONS. MICRON DATA LABORATORIES, INC. 22 Secord Dr., St. Catharines, Ont., Phone (416) 934-0629  Canada L2N IK8  PO  264  APPENDIX 8  ELUTRIATION TESTS RESULTS  A)  B.C. Moly Calcines  Temp. °C  f,25°C 1t/mi n  540 540 540 540 540 480 460 523 560 505 539 560 525 550 550 550 550 550 550 550 574 574 574 550 550 550 560 550 550 540 550  81 90 76 70 66 77 77 81 81 62 76 80 81 77 77 77 77 69 62 88 68 70 73 68 72 82 41 50 77 56 52  G  f ,T°C lt/min  G  226 253 214 198 187 204 202 224 232 170 214 231 223 218 218 218 218 198 177 247 199 206 213 195 205 233 114 138 218 153 143  Uo , »i  cm/sec 29.7 33.2 28.0 26.0 24.6 26.4 26.5 29.4 30.6 22.3 28.2 30.2 29.4 28.5 28.5 28.5 28.5 26.1 23.3 32.3 26.0 27.0 28.1 25.4 26.8 30.7 15.0 18.0 28.5 20.0 18.6  k* gr/mi n 98.2 105.2 74.6 61 .0 52.8 53.0 87.8 113.6 98.9 46.7 84.0 101 .0 96.7 51.0 104.2 106.1 106.4 95.2 68.6 136.9 56.4 68.4 73.2 82.8 90.0 111.8 11.1 14.1 73.2 51,0 36.0  CONTINUED  265  APPENDIX 8 (Continued) Temp.  B)  °C  f ,25°C 1t/mi n  560 534 450 486 342 540 550 550 540 510 560 555 580 567 556 556 500 546 546 518 446 430 504 540 538 330 540 550 550 518 500 480 550  55 51 77 61 60 77 62 68 72 76 63 58 72 69 66 64 72 70 74 63 62 43 65 65 65 66 66 66 77 50 70 86 67  G  .  Uo,  k*  f ,T°C 1t/mi n  cm/sec  gr/min  154 137 198 165 141 215 177 195 203 207 184 168 214 198 191 184 195 199 212 175 163 103 178 186 183 1 56 187 191 218 132 190 228 193  20.3 17.9 25.9 22.4 18.7 27.8 23.3 25.4 26.5 27.4 24.0 22.0 27.9 25.9 25.1 24.2 25.4 26.1 27.6 23.5 21 .4 13.8 23.7 24.3 24.0 20.6 24.6 25.0 28.0 17.5 25.0 30.0 25.3  38.3 23.0 92.4 60.5 29.9 56.6 60.9 76.7 84.4 85.5 42.2 56.9 57.4 79.0 54.6 49.2 55.4 86.8 91 .2 96.3 48.0 17.4 82.8 81 .5 83.6 81 .2 107.5 88.9 115.2 40.4 72.5 129.5 90.5  193 192 177 119 188  25.3 25.2 23.2 15.6 24.4  85.4 79.0 69.5 26.6 75.6  G  T  Brenda Calcines 550 551 540 530 546  67 66 62 44 65  CONTINUED  266  APPENDIX 8 (Continued) Temp.  C)  °c  f,25°C 1t/mi n  520 480 470 490 490 490 490 360 360 340 330 550 555 555  64 64 72 48 60 52 72 62 57 72 83 63 65 66  f ,T°C 1t/mi n  G  Uo,  T  k*  cm/sec  gr/min  177 171 189 125 154 134 193 154 142 171 192 179 187 193  23.2 22.4 24.8 16.7 20,2 17.6 25.3 20.2 18.7 22.7 25.3 23.7 24.7 25.3  54.7 53.8 48.8 47.1 56.0 50.8 84.1 19.5 19.3 33.4 42.5 56.7 65.0 59.8  184 2Q8 208 225 180 190 205 235 188  24.2 27.5 27.5 29.6 23.6 24.9 26.8 31 .0 24.6  39.8 61 .0 74.0 73.2 49.6 56.1 68.1 100.0 67.1  109 116 123 145 161 173 173 174 154 147 160 175 179  14.4 15.4 16.1 19.1 21 .2 22.7 22.7 22.9 20.1 19.4 21 .1 23.0 23.6  26.5 35.2 28.0 56.0 77.4 50.3 89.8 44.7 44.8 63.3 113.4 105.8 120.4  Kennecott Calcines 570 550 525 520 544 550 550 550 544  D)  G  63 72 77 82 63 65 72 83 65  Endako Calcines 235 300 350 400 425 450 495 503 410 430 470 51 5 530  64 64 64 64 64 67 64 64 67 62 64 64 64  267  APPENDIX 9  • •  RECORDED SIGNAL FROM THE INFRARED ANALYSER TO A S0 PULSE INPUT 2  268  9912M-S ' ° N IHVHJ  KOrmoHiimii  IT  o T =. 7  U.L-:  -  77 37  -  r>  ///.-...^  <<  i  PJTC  t ^ r p ? / / A * - * j.r/c9ic::  :  -rh  --0S-  -08:-  mis (Mi*-™ --14  B 0 -  269  70/ £ 1 trzr  7v;  -1  m m  -0C—H  -0f-  -—-Of-  •:r-  rpr-h  ' : > '  :• ' •  uniiti  -r/H-  -06-  ' trzr::t.rr '-•••••^4 = ^-t-^i; R  !  •^Ift  #r  7^7 —  '&?<>•:&#.: fit  ;  en  1332!  jrr  '  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 MICHIGAN  TF.F MI MAL  SYSTEM  C C  UF'TRAN  SYMPCLS  USED  0(41336) IN  '  MAIN  10rl6-74  CO MPUT AT I ON  c — V  ?  c C c C C c C c c.  AALPH AC A DCA ADI S ADS AOS I ADT AFF AF I  V c cc  .  c  .c c c c c. c c c c  r  c c c c c c c c c c c c c c c c c c c •• c c c c c •c c c  ^  •  VPICAGE COEFFICIENT C F C l . C U C - W A K E AND F M L L S I LIN ABSORPTION COEFFICIENT C F THE C A S I N THE SOLID 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 IN R E A C T O R , G R / C U . C M AV F E A G E S U E D I A M E T E R OF S I L I C A S A N D P A R T I C L E S , CM DIMtfTIONLFSS IME E X P A N S I O N F A C T O R C F F L U C U E C BED R E S P E C T TO THE S T A T I C BED Vt.LUMETRIC F A C T O R C F I N C R E A S E C F B E D R E S P E C T OF S I L I C A SAND BED ALM S T A T I C B E D HC IGHT , CM AKCE 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 C L O U D AND E M U L S I O N , CM/SEC AKPC MASS T R A N S F E R C O E F F I C I E N T OF C A S B E T W E E N B U B f i L F AND C L C U D , CM/SEC AKBE 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 AKEEB OVERALL COEFFICIENT OF G A S I N T E R C H A N G E B E T W E E N G A S B U B B L E S AND E M U L S I O N , I/SEC ' AKCEB 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 B E T W E E N C L O U D A N D E M U L S I O N , I/SEC A K B C E 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 A N D rt n u n . i / S F C A K h E S 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 FLUICIZED FED H E I G H T , CM AL F E E D I N G P O I N T , CM ALO AL F TOTAL H E I G H T GF F L U I D I Z E D B E D , CM A L P H A F R A C T I O N OF WAKE IN GAS B U B B L E S STATIC B E D H E I G H T C F S I L I C A S A N D , CM AL S ANH N C " 3 F - K R O L E S PER U N I T A R E A I N C A S D I S T R I B U T O R , I/SO.CM A P D C A 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 APDSI AT". 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 I N BED,MIN BN T O T A L NUMBER CF B U B B L E S AT L E V E L L IN T H E F L U I D I Z E B E D , CU.C.M BNT NUMBER O F 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 OF' R E A C T O R S , I./SEC 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 CG CGB C C N C E N T RAT I O N O F R E A C T IN G G A S I N S I D E B U B R L E S . M O L . G R / C U . C M CGE C O N C E N T R A T I O N OF R F A C T I N G C A S I N E M U L S 1 0 N , M O L . G R / C U . C M • C GM C O N C E N T R A T I O N O F R EAC T I N G G A S , M C L . G R / C U . C M CK B COEFFICIENT GF R I S I N G V E L O C I T Y OF B U B B L E S , (0.90) C P " 0 S 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 CPMOX H E A T C A P A C I T Y OF M 0 0 3 , C A L / M O L - D E G . K C O N C E N T R A T I O N O F R E A C T I N G S O L I D I N GAS P H A S E AT A N Y P O I N T IN CS9 F L U I D I Z E D BED CDNCFNTRATIONOF REACTING SCLID IN G A S P H A S E AT FEEDING CS EH 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 G A S P H A S E AT T H I S PCINT C O N C E N T R A T I O N OF R E A C T I N G S H L . I D IN GAS P H A S E A T T O P OF B E D CSBE T  11 2 6 : 3 3 6.000 7. 000 7. 000 7.000 P. 000 9. 000 10.000 1 1. 0 0 0 12. 000 13.000 14. 000 15. 000 16.000 1 7. 0 0 0 • 18.000 19. 000 20. 000 21.000 2 2 . 000 2 3 . 000 24.000 2 5 . 000 2 6 . 000 27.000 28. 000 29.000 30. 000 3 1 . 000 32.000 3 3 . 000 2 4 . 000 35.000 3 6 . 000 3 7. 0 0 0 38.000 . 3 9 . 000 40.000 41.000 4 2 . 000 •43.000 44.000 45. 000 46.000 4 7. 000 48. 000 49.000 5 0 . 000 51.000 52.000 53.000 54.000 55. 000 56. 000 57.000 58. 000  RAGE  PU  271  MICHIGAN TERMINAL  MAI N  SYSTEM ' -OF TP. AN G( 4 12 2C) CSE CSEE  csro  "C C C  CT G  ~t  CT5  C C ~C C  DAG —DAS DB  CCNCENTPAT ION OF REACTING SOLID IN EMULSION PHASE AT ANY PCINT IN FLU 1CIZED BEC TOP BED CONCENTRATION OF REACTING SOLID E MIL SI ON PHASF FEEDING CCNCENTPATION OF REACTING SOLID EMULSIUN PHASE PCn>rTT~AS"F R A C T 1 ON 0~F C A L C 1 N E S AT THIS PCINT CONL ENTRAT ION CF TRACE AT M EA SLR I NG POINT, CU.CM TRACER/LT GAS _ t m ^ T T ' R T r T l W T j F SOLID TpACEft -'AT SAHPTlNd POINT, K T . K E R CfcNT AXIAL DISPERSION COEFFICIENT OF GAS IN FLUIDIZED BED, SCO/SEC A X U U D I S P C P " S 1 C ' N I COEFFICIENT OF SCLIDS I N iHt b. t U , SU . CM/ b te. DIAMETER OF G A S BUBBLE,- EQUATORIAL, CM 0  D B O !iini_ __ L?iL -  c  11:26:31-  10-16-74  L  BL  L  R  C  -L^—- l  T 9  Hf0  *_ J* C  DC ' OIAMETTk " O F BUBBLE"' S~" CL'Cub , ""EO U A T pp. I A L , CM DC A DENSITY C F CALCINES, CR/CL.CM C D DELTA EF F FFX IACT T O ION R FES1LTNCE 0 F BEL) CONSISTING TIME. DISTRIBUTION OF CAS BUBBLES O F TRACER 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 E D I FPUS ION CO E F FI CIE N T ' OF GAS TN THE FLUIDIZED FEDi SC.CM/SFC DMMOS MOLAR DENSITY OF MOLYBDENUM CI SUL P H I DE . GR-MO L/CU .CM ~D>MOX MOLAR DENSITY CF MOLYBDENUM TR 10 X ID E , GR - MOL/CU . C M Mint 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 C ORG RADIAL DfSPERSIONXOEFF1CIENT OF GAS IN FLUIDIZED BED, C SO.CM/SEC C DS RACIAL DISPERSION COFFFICIFNT CF SOLIDS IN THE BED,SO.CH/SEC T~~ ~DS"' AVERAGE SIZE DIAMETER CF M:CS 2 PARTICLES IN FEED, CM • 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 _ _ _ 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 C EOCA STATIC VOIDAGE O F CALCINES BED C EOSI STAT IC VO I DA GE OF. SILICA SAND BED C ER ELUTRI AT ICN RATE, GR/SEC ~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 C REACTOR., TfJN/SU.MT.»DAY C FV FLUIDIZATION FACTOR CF VOICAOE, CONSTANT FOR 10 TO 60 C WT.PERCENT OF CALCINES ~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 C GK THERMAL CONDUCTIVITY OF AIR . ~< 4.32*10**-£l ,CAL7CM*2*SEC + CEGC C GL TOTAL GAS AT LEv EL L IN THE FLUIDIZED EEC, BU.CM C GM TOTAL GAS'FLPV, RATE AT 25 01 ;G.C FOR ThE SIMULATED FLUIDIZED 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  ~C" C  -  C  r  ;  r  59. 000 60. 000 61.000 62. OOJ bl,000 64.000 65. 000 66.000 67. 000 68. OOP 69.000 70. 000 71. 000 72.000 73. 000 74. 000 75.000 76. OOJ 77.000 78.000 7?. 000 8 0.000 81. 00) 82.000 84. 000 85.000 86.000 87. 000 88.000 8'9. 000 90. 000 91.000 92.000 93. 000 94.000 95. 000 96.000 97.000 98. 000 99.000 100. 000 101. 000 102.000 103.000 104.000 105.000 106. 000 107.000 108.000 109. 000 110.000 111. 000 112.000 113.000  PAGE P002  272  :  i  ! •  1  c c  SYSTEM  FORTRAN  GT  TOTAL  G  GI41326) GAS IN  RAIN  R E D AT  INITIAL  ANYTIME,  10-16-74  CU.CM.  HEIGHT OF HEAT OF RFACTIUN  HP P  HE A T T R A N S F E R COEFFICIENT TLTAL PRESSURE, ATM.  c  PG PN  PARTIAL PARTIAL  c c r c c c c c r  PO PR  R AR I 1 AL PRtSSUPR Of- C X I GEN,  PSO  PARTIAL  PTK  PARTICLE TRANSFER COEFFICIENT CLOUD A N D WAKE B A S E D IN UNIT APEA  c  C  c c c c c c c c c c c c c c c c c c c c c c c c c c c  c  c c  c  r c c c c c c  PRESSLRE PRESSURE  PRANDL  GAS-PARTICLE,  OF REACTING G A S , OF N I T R O G E N , A T M  115. M003,  CAL/MOL.GR  C4L/SO.CM*SEC*DEC.C '  ATM  AT M  PRESSURE  FLUX  OF SULPHUR  CIOXIDE, ATM  12 3 .  SOLID OF BUBBLE  IN B U B B L E S SURFACE,  AND  REG REP P.R  R.S RTD  Sb SBK  P E A C TOR FEYNULOS RFYNOLCS  P E RUNIT  PER UNIT  AREA  A P E G F B E DFEP. UNIT  C F B EC  PER. U N I T  T  I MF,  TIME, GR  NUMBER  RECIRCULATION REACTOR CROSS  / SC . CM * S E C  FUR SCLICS SO.CM CF  GAS,  CIMENTIONLESS  SULPHUR  SGK SH  VOLUME OF SCLI'J SHERWOOD NUMBER  SP SPF SR  VOL  UR  CALCINES  F R CM  L  INSIDE  O F GAS  THE REACTOF CAS  TE  3 G  TF T TH  H S  BASED  IN  »EACTION UNIT  SUPERFICIAL  MFAN  THE  TE MPFR A TLRE OT TEMPERATURE OF T H E GAS TIME, SEC  BASED  VCLUME  ON U N I T  AR F A O F  144.000  G A S AT  OF  GAS VELOCITY,  CONSTANT,  GP/CL.CM  TK  TEMPERATURE BUBBLE RISING  UF UF  U"F UN  TEMPERATURE  PURREE  SURFACE  OF SOLID  OF- THE RE'.CTCR VELOCITY,  VELOCITY  RESPECT  B E D , DEG.C  THE  CAS  C  150.  000  151.000 152.000 153.000  159.000 160.JOO  DEC.C  161.  TC EMULSION,  000  157.000 158. 0 0 0  CM/SEC  000  162.000  CM/SEC  EMULSION 'VELOCITY, SUPERFICIAL GAS VELOCITY IN t-HJLSICN,  SOLID  FLUIDIZED  WALLS, CM/SEC  148.  149.000  155.000 156. 0 0 0  F L U I D I Z E D B E D AT ANY T I M E OF THE REACTING M0S2, DEG.C  IN  000  146.003 147. 000  154.000  FLUICIZEC B E D , DEG.C IN F L U I D I Z E D B E D , O E G . C  BUBrtLES IN  TOTAL NUMBER CF TEMPERATURE A T THE  UPR  140.000 141.OOJ  145.  RATE  TN 7R TS  UB  138.000 139. 0 0 0  142.OOJ  REP CENT SULPHUR UIOXIL'E IN C F F GASES SPECIFIC SURFACE AREA OF MOLYBDENITE C O N C E N 1 R A T E , S 0 . C M,/ G R SPHERICITY FACTOR OF SOLID  MEAN  000  BURPLES  SLURRY R E C Y C L E T C REACTOR A S 5 0 5 S C L I OS.KG/MIN SPECIFIC 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 TEMPERATURE OF EMULSION PHASE IN F L U I C I Z F G B E D , DEG.C  SS  000  134.  143.JOO  FDR CHEMICAL  RECIRCULATION S02  OF DISCHARGED  UNIT  133.  13 7. 0 0 0  SCAL  PER  000  135.000 1 3 6 . JOO  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 ) NUMBER OF B U R B L E S P E RU N I T AREA C F B E D . l / S C . C M  ELUTRIATION  129.000 131.  SB N  SKS  128.000  132.000  EXIT A G E D I S T R I B U T I O N FUf>CTICN SURFACE OF BURBLE, SC.CM  RATE CONSTANT SCLID, CM/SEC  000  127.000  CM ,MT  RATIO SECTION,  CONTENT  JOO  125.000  DISPERSION NUMBER NUMBER FOR GAS  SK  000  130.000  riAMETEP,  R DN  ooo  124.000 126.  GAS CONSTANT REACTOR  RD  lie.  122.000  BEJWFFN  SOLID  117.000  121.  MUMPER  Vl.LUMETRIC GAS FLUX CU.CM/SO.CM*SEC  CS P.  000  116.000  119.000 12 0. 0 0 0  CR/SC.CM*SEC OB  PAGE  114.000  BURRLES FORMAT I O N , CM F OR CX1 C A T I C N • C F M C S 2 T C  HR HEAT  c C C  11: 2 6 : 33  163.000 THROUGH M/ S E C  EMULSION  TO  THE  164.  000  165.000  REFFREu  SUPERFICIAL VELOCITY MINIMUM F L U I D I Z I N G C C N C I T I C N S , MUSSELT NUMBER  RELATIVE  TO THE EMPTY CM/SEC  REACTOR  AT  166.000 167.  000  168.OOO  P003  273  ' IC HI GAN T E RMIrAL SYSTEM FOPTFJN GI41336)  MAIN  SUPERFICIAL GAS VELOCITY FEFFP.PED WOPK I N G T E M P E R A T U R E CM/SFC  UR  AVERAGE SUPERFICIAL GAS FMJBBLES, CM/SEC  ur  L INEAR  V E L O C I T Y "OF"  FMULSION,  C C  J  C C _C  VAPJANCE  OF  DISTRIBUTION,  RFACTOR  OF RECIRCULATION  11:26:2; AT  INSIDE  RELATIVE TO THE  S E C . **2 , MI N . **2  -  1  1  f  {UN «FTN .SPl,'NCh- = C C N P P f i F I L C "~."UT ICN""a CIMS  i  T H E EMPTY  l O T""TTHT EMULSION  SCL  T  -  c  TO  169, 170. 171. 172.  T7T  000 000 000 000 "5oar 000  174 0 0 0 175 0 0 0 'VCluFE 0-" B U ? OLE",* CU.CM ""' ' " ' " — "T76". VOLUME OF HUBBLE S C L C U D , CU.CM 177, 0 0 0 PATE CONSTANT F O R C H F M IC A L R E A C T I O N B A S E O IN U N I T VOLUME OF 178. 0 0 0 SOLID, MCL/CU.CM.*SEC 179. 0 0 0 VMAIP MOLAR VOLUME CF AIR, (APPARENT), CC/GR-MCL 180. 0 0 0 YJ?'__ MQ LAP VOLUME CF NITRCGEN, CC /GP.-MOL 181. 0 0 0 VKCX" "MOLAR VOLUME "CF 6YlCTt'NTTr/GP-"Mclt 182. 0 0 0 V«SO Mf LAP VOLUME CF SULPHUR DIOXIDE, CC/GR-MCL 183. 0 0 0 ' VP VOLUME OF A SINGLE PAR ICL E, CU.CM 184. 0 0 0 VS SPECIFIC VOLUME OF SCLID PER UNIT MASS O F PART I CL E S , 105. 0 0 0 CU .CM/GR 186. 0 0 0 TO TAL VOL UM E CF_ FLUlrlZFC PFC, CU. .CM VT 18 7.0 0 0 t C TA l" B "0" t'B L CSV OTU W I ( C T C U ID I Z EC" BED, CU.CM"" VTB 188, 0 0 0 TOTAL CLCUD VOLUME IN FLUIDIZED 3E0, CU.CM 18 9.0 0 0 VOLUME OF BUBBLE WAKE, CU.CM VW 190. OOP V T O T A L WEIGHT Cr SOLID IN "THE FLUIDIZED" B E D , G R , KG~" 191, 0 0 0 WCA WEIGHT O F CALCINES IN FED, G R 192, 0 0 0 J^'''L F _!:"-' ! i: E ' c u t A i: h ^ i ? t l T _ - 0 f L _ / L I i ' ' _ ( A P P A R E N T ) , G R / ' H O L ' 193. 0 0 0 _ W-'"OS MOLECULAR"WE I GUT OF M 6 L V R~D E N L M DI MJOMTT b~E , OPT""" 0L"~ 194 0 * 0 0 WMMO MOLECULAR WEIGHT O F MOLYBDENUM, G R / M O L 195. 0 0 0 WM*"TX MOLECULAR WEIGHT CF MCLYP.nRKLM. TRICXIDE, G R / M O L 196, 0 0 0 Mi. L E'CUL AF IGHT OF UX IGLN-" GR TFol '. ' 197, O O P v. !•• r x MOLECULAR OF SULPHUR CICX ICE, GR/MQL 198, 0 0 0 WMSO WEIGHT OF WEIGHT 199, 0 0 0 :'i_SLfiGJ-I__?^T!.£-l!:_' GR _WF SPFCfFIC WEIGHT OF SOLID PER*UNI T V O I T J M E~P A R TIC L E S, GR/CU.TM" 200. 000 VP V 201. OOP WMCHT O F SILICA SA.NC IN BED, GR WS I 202. OQi) IGHT REACTION OF CALC I N F S IN BED _XCA F R A C 1 1UN OF BED "203. POO CLOUDS X LL WEIGHT FRACTICN OF AT LEVEL L FN THE FLUIDIZED 204. 000 XP WE IGHT FRACTION OF SCLID TRANSFORMED 205. 00 0_ SILICA SAND IN BED XS I 206. OOP VAC VB VC VK  C C _C C C C  CM/SEC  GAS V E L O C I T Y  "10-16-74  PAGE P004  274 C  TESTS  R f SPONSfc AND  GAS  'Of-  AND'Sntin  GAS TRACER INPUTS IN FIVE INCHES REACTOR O I S P R K S I ON_ COEFFICIENT CALCULATIONS IN THE BCD  \\7 238 24 0 24 1  C  I.L.L.P.FRFORMf 0 IN WURKING CONDITIONS^ TYPICAL, OF HIGH CON VER T I ON LEVELS OF MOLYBDENITE TI I 7-ICL I b OENUN T R I>JX I UV ~' ' MOLYBDENITE CONCENTRATE FROM "BRENDA MINES".-32b MF SH PAR TI C LE S.  _Q  s  C  C  •c  ~2A~  "  2^3 24 4  C  '  — ,  C C C P  — _  ._  TEMPERATURE CF FLUIDIZED BED,510 TO 520 CEG C TRACER SUIPHUR DIOXIDE INPUT TIME FIVE SECONDS VOLUMF. TRACER 16.5 LITERS  •  J^246 247 248 249 250  ;  . . .  252 53 254" 255 2*6 25T" 258 g |g-gj 262  C  INTEGER TI . • 01 MENS I CN TM (20) ,CTG( 20) ,ADT( 18) ,OEF( 20) ,RTD( 18) ,AT( ICOi REALH5.10) TIME,EFAM,ALPHA,ALS,RS,G,GF,UMF,WCA,WSI,ERK,AFI,AFE,DSI 1 .DCA.ADCA.AOSI .APOSI , APP CA , F V , DH , WM SC , WM OX , VM SO , VMCIX , VMA I P , RD , CK R , 2HSP0.P.AC — R E AD ( 5 j 20) (CTG( Jl ,J = 1,18) PEAD(5,20) ( TM( J), J=l. 18 ) f 0RMAT(8F":i0.4) ~~~ FORMAT!RF10.2) CALCULATION CF GAS RESIDENCE TIME DISTRIBUTION FUNCTION IRTD) . . SUMC = 0  ? • 3  2  1  !  4 _5 6 7  10 20 C v.  8 9  •  10 11 12 13 14  SUMTC-0  16  2  . WP IT El  20  65  '.  6, 6 5 )  ""  ""'  266 267 6B 269 2  "  270 .  271 272 273  ( A P T ! J ) ,J-1  60  FORMAT ( I X ,1 8 F 6 . 3 )  66  FORMAT! IX, ' F X I T  70 Q  WPITF(6,7'O)  23  2 7 5 "  TIME /)  276  1  ,18)  277 2.78  >IPITE(6,66)  24  25 26  „ AGE  D I S TP..  C  279  F UNC T. ' / I  -r  r k T [ ) i j i , J = I , la) ; ~ FORMAT!IX,18F6.3) CALCULATION OF REACTCR DISPERSION NUMBER OF GAS I RDN ) T " " — " " R-  27  274  ;  1  WRITE16.60)  '  .  ^  FORMAT(IX, DIMENTIONLESS  21 22  ;  :  28 0  "  ; .  R-0  8  0  29 30  31  •  800  32 ' C. 33 34 35 36 38 37  1  =  0  •  '  ' DO 800 J = l , 1 8 ' 0=0+ACT(J)**2.*RTD(J)  :  !  IST~ 282 283 284" 2  2  1  1  FORMAT(IX, ' Q F A C T O R . D I S P • N O . ' / )  40 '41 42  80 90  WRITE(6,90) RDN F0RMATIF6.4) WP IT EJ 6 , 200 I  43  C  CALCULATION wVi'fEffiVfi'srOF THE FLUIDIZING PROPERTIES OF THE BED  8  5  286  287"  !  F=R+RT0!J)  1  C  ~i  8  3  289 290 291  VAC=0/R-1.  ITERATION LOOP FOR CALCULATION THE DISPERSION NUMBER R DN= VAD/ 2. ' DO 170 M = l,100 ~ ' i VA0D = 2.*RDN-2.*RDN«*2.*(1.-EXP(-1./RCN)) IF ( (_V AO- VA D D K L F.J.I) GO TO 8 0 170 V.PITEI6.75) CONTINUE " ~ ~  75  3  265  R T O I J ) = T ME AN *D EF t J )  50  6  2 b l >  DEFlJi=CTG(J)/A  18 19  3?  6  AOT! J )=TM( J ) / T M F A N  17  2  2  '  DO 40 J=l,l( SUMC=SUMC+CTG!J) 40 SUMTC = SUMTC-«C.TG( J.)*TMU ) A*TIME*SUMC "" TMEAN=SUMTC/SUMC DO 5 0 J = l,18  _J5  ?5  :  2  ~— '  9  2  29T" 294 29 5 ~~296~ 29 7 298  29 9 300 301 302 303 30-H  275  C  C C C  44  C Al C U L AH O N S  MADE  TICNS  ASSUMING OF  GFT=GF*(  OF  THE FLUIDIZING  AND  PRESSURE  ALMf=1.01"ALS*AFI  46  ALF=ALMF*AFE  55 56 57 58 59 60 61 62 63 64 65 66  P/ 68 69 70 71 72 J_3  C C C  C C  c  78 79 80 81 _82  83 84  AND TEMPERATURE  GASES  OF  OF AT  T i iT  HT fj A R E  THF  WORKING  '  1  '  ^oTT  306  CONDI  T H E RED.  307  308 3  •  XCA= XSI = WSI/ IW'CA + WSI ) ECSI = l.-APD SI/OSI EpCA-J.-APDCA/DCA E0 = EDS I*XSI + FCCA*XCA EMF=FV*'-n EA-1.20*EMF . . . . . . AALPH = 11 . -E A) / (1 .-t.MF ) * ( EME/ EA )**2. THE AVERAGE RECIRCULATION VELOCITY PF GAS INSIDE BUtiKLES WAS CAL CU LATED USING THE EXPRESSION OF LEUNG AMD SANDRORD  0  312 313 314 31 5 316 31 I 31 3 319 320 321 32 2 323 324 325 326 32 7 32 8 '32 9 330 . 33 1 332 333 334 335 1  >  3  3  •  3  g 5  A  .  2  3  bL = b H / I S U U , » U l l l * R S  3  , n  BN=Gl/VP VOLUMETR'IC GAS FLOW ' R'ATE 'OF RECIRCULATION INSIDE BUBBLES WAS CAICU LATED EY MEANS CF THE EXPRESSION CF CAVICSON AND HARRI SON GR=3.*UF*EMF*3.14*(DB/2. ) * * 2 T  •'  "  1  '  1  346 3vT -348 34 9 3  5  0  SB = 4. *3. 14*<DP/2. ) * * 2 . VC_M0.6/D8**0.5 )*V9 • US = ALPHA*DE.LTA*UB/ I 1 . - D E LT A- AL PH A* C EL T A ) ~ : J 5 3 DC-I 6.*VC/3. 141**0. 333 X C L = 2 . / 3 . * T B / U P * ( 8 N « U B / A L »_*_( DC/T) B ) *_*3_. /J_0 0_. 355 DIFFUSION C C E F F n c I E N t OF' GAS" C AL CULAT E D E Y MEANS CF THE RFLATIONSH : ^ K T IP OF ANCRUSSOW \ r : ., ' 3 5 8 r 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) DIFFE-DIFF»EMF 36 1 3  3  1  :  :  5  -  5  1  2  :  3  C C C  0  u ^ 344 2  ...  37  syg-  :  UF=UM'F/ETMF '  C C C  9  1  il I  1  UH-0 .00 2b?«- AL" ( (UU-UNF)/ANH) UB = GfJ-UMF«-0. 711*IG*OR)**O 5 '. DELTA=(U0-UMr)/U6 UE = UMF/EMF-1 ALPHA».UO/ I 1. -CELT A-ALPHA*CELTA >r-ALPHA»UMF ) EF = UE /R S V P = ( ( 4 . / 3 . )*3.14*(UB/2.1**3,)*(1.-ALPHA) _UBR = 0. 7U*(G*PB)**'0.5  0  3  UR= IUMr*EA)/( AAIPH*FMF ) MO-ALF/0.5 TI=TIO. UO=GFT/R$ ANHJ-CH/RS 00 100 K=1,TI Y = FLOAT I K ) SUUM--0 ' — . . AL = SUUM + 0. 5*Y AT (K>AL/0.5 AOSj-DCA^XCA + IVS I*XS I ADIS-APCA*XCA + ADS! *XSI BUBBLE DIAMETER CALCULATOR USING THE EXPRFSSION DEVEIOP-'D I N THI S PRESENT WORK FROM TWO-DIMENSIONAL EPOS EXPERIMENTS ' ' f  13  77  FOR  1  74  76  CCNDITICNS  REHAVOIR  550.+?73. )/27?.*( 1000./60.)  45  47 48 49 50 51 52 53 54  IDEAL  3  5  5  4  7  :  :  3  6  0  3fe2  CALCULATION OF GAS I NT PR C HAN GE COEFFICIENTS BETWEEN GAS BUBBLE, CLOUDS AND EMULSION  363 3  6  4  276  85 86 87  f "  C  88 89  V  r  1  .. 90 91 92 93 94 95  96 97 ' .99 99 100  101 102 103 104 106 107  c c 200 115 300 500 C C  c c 116  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 AKBC=0.975*DI FF**0 .5 » < C,/DP ) **0 . 25 . 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) CALCULATION OF OISPEPSIUN COEFFICIENTS OF GAS ' DRG=0.2*DP**2.*AKBEB/DELTA 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 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, SOL ID INTEPCM.U-WAKE' ,5X,'SOL ID AXIAL DISP.COEFF.•,5X, 2'SOLIO RADIAL 0 ISP.COEFF.•/) 1  WPITC(6 ,300)  FORMAT! IX, 12IP 7. 4, 2X>/I 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 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 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). ' KPIIFtt.Ufcl AKEBS.DAS.DRS FORMAT! ix, 3F10.4) STOP ENO  '  365 36 6 36 7 368 369 370 371 372 373 374 375 376 377 373 379 380 381 362 383 384 385 386 387 388 389 390 39 1 392 393 394 395 396 59 7 39 8 399  APPENDIX 11  COMPUTER PROGRAM TO ESTIMATE THE TEMPERATURE REACTION PARTICLES OF MoS  2  '  ~'  J.CCMPILE C . C G  ^  '  ; i  .  ''  ~  ^C C  ' ' ' . '  C • C  ; '  .1 2 '3 4 5 • 6 10 7 • . 20 B 30 • 9  'n II  210  12 13  15 16 17 18 19 20 21 22 23 24 25 26 27 28 29  THE T CMPE RA TUP.E AT THE R E A C T I N G SURFACE CP MGLYBP6A l THE I N I T I A L CHEMICAL REGIME OF TRANSFORMATION.  S I M U L A T I O N CH ROASTING C C M i l T I U N S IN AIR FOP PART I C L F S IN EMULSION MO L Y CPE NI TE CONCENTRATE FROM "BROKCA M I N E S " . - 3 2 5 MESH P A R T I C L E S . : — . . ' ' . ' • . DIMENSION I 1 20) , S M 2C) P F f D 1 5 , ; j ) AC, F G i U N . V 5. PO. EM, I, SBO.CF , A , S . READ ( 5 , 2 0 ) N A READ ( 5, 10) ( TI J ) , J = l",NA ) PE AD ( 5,10) (SK ( J I, J = l ,Na ) FORMAT ( 5 F 1 0 . 5) FORMAT I 1 I 1 C I F C RM AT (B Fl 0. £ ) 00 2C0 J = I ,NA ' U R I T FIft.210 ) ECRM AT 1 5 X , ' T E M P . P E P . ' ,5X,'TEMP.PART.SURF•• , 5 X , ' H f c A I F t A l l . 1.HEATH'/) CPMCS=(19.7+3.15/lO.**:.*((T(J)»27 3 .)-29E.))/16J.09 HE ATP=2 S 75 9 0 * 4 . 4 9* T < •)) - 7 . 86 2 10 •> <• ( - 3 . ) ) » I T I J ) ) * g 2 . - 3 .91 9* 1 1 1 0 . ) * * 1 5 . >MT < J I I T * ( - l . I  ki=L*S  14 . •  C A L C U L A T I O N CP NI TE P A R T I C L E S  300 310 320 2C0  GK=AC*FG SBK=S60*CF HP=GK*LN/DS CGM=PO*ll./224GQ.)*(273./(T(J)+273.)) TPS = T U I + 5 0 . 0 DO 3 0 0 M = l , 5 0 TP A= [-EM I * S h K » [ 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 . )/ 1 0 0 . >««-4. )/ HP I F ( ( 7 R S - T P A ) . L E . 1.0) GG 70 3 1 0 T RS = T R A CONTINUE '. ' W P I T E ( 6 , 3 2 0 ) T ( J ) ,TRS,HtAIR,CPMOS FORMAT (4 F1 5 . 2 ) CUNHNLE STOP ENC  APPENDIX 12  COMPUTER PROGRAM TO CALCULATE THE REACTOR SIZE AND PERFORMANCE  279  f  tCCMP!LF  ""  . - - ••  —  -  - .  -  C C  C ^  c  C C C C c  C  _C C C C  FLUIDIZED BED REACTOR FOR V,H VBDEN ITE ROASTINC  _ I  _  It 17  I !  I j (~. : " j '  9  :  n  t ALCOLATTnN"'.TF~"R^  PRL-DICTICN OF SLLRHUR LEVELS IN DISCHARGED CALCINES AND SIMULATION CF_REACTCR PERFORMANCE •  12 13  MOLYBDENITE CONCENTRATE: PARTICLE S I Z E - -325 MESH.  15 16  - _ _ ^ _ _ _ _ _ _  • -  i  c  ^  r  j  ?  l  r  -  H  T  O  -  •  •  —  18  CYCLONE SYSTEM EFFICIENCY = 9 .?; " "SCRUBBERETFTCI E N'C'Y'"""=""" 9 5"'}I ' ' ' SLIP ERF.GAS VELOC. AT ROASTING TEMPERATURE =25 CM/SEC FOASTING WITH AIR  IP 2"0"~ 21  :  DIMENSION T(20) ,FOl20),DS<20),ATR( 20) Rt AC (5 ,_1 1 J )_XC A ,_X_S I_, DCA,DSI,AFE,EFAM,ERK " R.-.ADI 5, 33.!) NN V NC i'MOTivM ' "" RtAD(3,221) < FC'( I ) , 1= 1,NN ) RE_A_0_(_5 ._220_)_l T ( J ) ,J=1,NC) "Rt AD( 5", 220 ) (OST"K) ,"'K = 1 ,NDI RF AD(5 ,550 ) ( ATR (L l,L= 1,NM ) J I 0 FO R MA JJ_7 F1 0. 2J_ ' ; 3 30 FORMAT! 4110V 220 FORMAT (3F10.2 ) 221 FORMATI5FI0.2) ' _ 550 FORMAT ( 6F10.2 ) DO 200 1 = 1 , M.N WRITEI6.20) F0< n ' - WRITE ( 6 , 40) ~ "' " . DO 300 J=l , NC WRITE I 6 , 25) T<J)  ? 2  • '  o.  24 25 2o  • \  1  —  ~ '  '  '  '  ;  19  WRITE(6,30) OS(K) • DO 500 L = I , N M 2 1 w R . IT! (4, 3") I i r - ( L ) ' ~ 22 20 FORMAT(5 X,'iFED RATE =',F8.2,IX,•GR.MIN•) _ 23_ 25 FORMAT! 5X • TE MP . = • ,F R. 2 _1X 'DEG.C • ) . • ' " 2 4 " "30 " F_*MAT ("5"X,'"'MOi'2^ ' S f/E = "'TFff. 2 7 1 X 7 * M I C P T H ' ' . 2 5 35 • FORMAT <5X, • AV.RFS.T IME=' , FS.2 , IX, • MIN' 1 26 40 FORMAT! 3X, "PER CENT S ' , 3 X , ' CH ARGE ' , 3 X , • R E AC T . DI AM . < ,3X , • FLU I D BED " " " 1HF IGHT ' , 3X , • GAS FLOW RAT E , 3X, 'P ER "CENT 502* 73X7' RECTRCTRATT 0"* ,3X~J 2'SLURRY RECIRC.'/) 27 AD_S-DCA*XCA+DSI *XSI "28 SC AL"= 1.3 5i (1 " +( EXP i-0 .1 *D"_"<K7*<""( TTj ) ) **0 . 765-TTj") +52C.) /ATRTTTJTI 1*(2./((T(J))**0.765-T(Jl+520.1-1,)) 29 W=ATF.(L)*FO( I )/ 1000. C RD:AH = l : l " " " ' 30 RD= ( <4.*F0f I ) *ATR <L )/( ADS*3. 14) 1**0.333)/100. 31 _ P.S = 3.J4* (R_D/2 . ) **2. • 32 " GM = '(64.*RS"/T"4";) " "~ ," 33 GF = 64.*RS/0.0 14 34 ALF=AFE*EF A M * K l i 35 S02 = (FU( I)*0.3o5*22.4*100./64.TA<"GF^ ' 36 ER = ERK*RS*fcO.* 10000. 37 RR = EF/FO(I ) . L  !  27 28 29~" 30 31 32" 33 34 3"5"" 36 37 33"" 39 40 «2 A3 44 45  20  46 ZT" 48  JL  50""  1  ;  38 39 40 41 42 45 3 44 "*5 45 50 0 46 400 47"" 30 0 48 20 0 49  i  y  -j-jr-  r  1 _2____ 3 4 '5 ~6 7 10 11 12 13 14  6 7  .  50  SR= 2 . * (0 .0 50*RR I FCM=FO(I )/60. CSE0=(F0M*100.)/(FUMf ER/60.) WRITEI6.45) SCAL,w,RD,ALF,GM,S02,RR,SR WRITE (6, 55) CSEC FCRMAT13F15.3)  7 D » ATTR G"7"5 T C0NT1NUE CONTINUE "CUNT INUE CONTINUE _STOP " ;•;.}  " '  '  51 52 ~ "5T~ 54 55 Sb~ 57 58 59' 60 61 6_" 63 64  65~" 66 67 68 69 70  •  71~  1  ;  .  1  :  ' •  72 73 TCT 75 76 77"  49  A P P E N D I X 13  EXPERIMENTAL DATA OF OPERATION FLUIDIZED BED REACTOR, 12.5 CM DIAMETER  Molybdenite  coriceniTQhe. source  Run N-  T °C  t  <V  hr  u* Fo 9 7T .K 7. I I3 /min /sec /onset /  Uo  I-  B . C . Moly  B.7.  570  2.0  70  27 A  it  B.8  580±5  1 .4  "7 O  OO  11  B, 1-2  575  3.2  .11  B.13  541  4.3  II  B.14  532  11.7  II  B.15  560  9.0  81  B.I 7  582  3.2  81  ii  B.18  540  9.7  81  u  B.20  511  13.2  81  it  B.21  524  13.2  81  B.22  524  10.4  81  B.23  550  10.4  B.-26  588  3.7  n  II  it II  it  -\0  0.99  Po  %50  2  2  calcines  gale;  dt  %S  solid SCfUbb  1  Cyclone.S a u t t e r  oxidized  mice  atm.  12  0.21  2.22  0.98  i  ?  0.21  3.56  0.91  7.6149 2.6028  0.9507 0.9110  |0-3  2.5  33 1  7  2.5  35  1  0.74  62  2.5  25 1  12  0.21  3.22  1 .21  4.8394  0.9195  72  26.6 . 0.90  72  26.5  i C 64  A  24'. 8  • 1 .07  82  I A Fo  p.  .75  2.5  30 1  12  0.21  1 .74  0.91  3.7353  0.9565  0.90  75  2.5  30 1  1 2 0.21  1 .20  0.91  1.3857  0.9700  30.6  1 .45  121  2.5  48 1  12  0.21  1 .42  0.78  1 .7861 -  31 .4  1 .60  132  2.5  53 1  12  0.21  1 .57  0.78  4.6914  99.4  0.9607  30.0  1 .36  113  2.5  45  1  12  0.21  1 .33  0 .-78  1 .6668-  99.2  0.9667  25.4  0.7-9  67  2.5  27 1  12  0.2:1  0.93  0.78  1 .2363 :  29.4  1 .25  104  2.5  42 1  12  0.21  0.87  0.78  1 .2088  29.4  1 .25  104  2.5  42 1  12  0.21  0.90  0.78  1.5640  81  30.2  1 .38  115  2.5  46 1  12  0.21  0.78  0.78  1.5686  0.9805  77  30.0  1 .36  113  2.5  45  12 . 0.21  1 .74  0.86  4.9050  0.9665  1  0.9545  0.9767 0.9782 . . .0.9775  0.5  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  .II  B.32  570  10.7  77  29.4  1 .25  104 . 2.5  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  0.21  0.48  0.86  0.7278  0.33 97.3  0.9880  0.28  0.9867  II  ii ti it n II  it II  n  550  22.6  77  28.5  1.14  95  2.5  38 1  21  B.42 . 550  15.3  77  28.5  1 .14  95  2.5  38 1  21  0.21  0.53  0.86  1.0725  95  2.5  38 1  21  0.21  0.56  0.86  0.7888  0.8760  59  2.5  24 1  21  0.21  0.60  1 .04  0.7880  0.9850  B.41  it  II  1  B.43  550  20.8  77  28.5  1.14  B.44  550  20.8  65  24.5  0.71  B.45  550  20.8  83  31 .0  1 .53  128  2.5  51 1  21  0.21  0.56  0.76  0.7888  B.46  550  26.0  77  28.5  1.14  95  2.5  38 1  21  0.21  0.42  0.86  0.6342  B.47  550  26.5  77  28.5  1.14  95  2.5  38 1  21  0.21  0.39  0.86  0.6189  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  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  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  0,9860 . 97.8  0.9895 0.9902 Oo  ccocentKihe, Source Moly  B.C.  "B.51  B..52  II  B.53  11  B.54  . II  B.55  n B  .56  II  B.57 '  ti  B.-60  II  B.61  II  B.62  H  3.63  u  B.68  Cchloride leached) Moly  B.C. u  t  u  11  B.C.  T t •c hr  Run N-  Moly  11  11 11 II 11 . II  11 II II  550 550 550 573 524 573 425 574 550 550 549 550  Uo =  I- -10  K  cry 9<y. '/mii /secV/t /onset 1  28  62 28.0 58 28.0 87 -1 5.3 77 8.0 77 22.7 77 23.3 77 22.0 77 23.3 77 23.3 77 23.7 77 25.0 77  23.6 0.64 5-4 21.0 0.46 39 32.5 1 .88 158 •2-9.4 :1 .24 104 27.5 1.01 84 29.4 1.24 104 ~27.5 1.01 "84 29.6 1.28 105 28.5 1.14 95 28.5 : 1.14 95 28.5' 1.14 95 28.5 1.14 95  Fo /•run  2.5 2.5 2.5 2.5 2.5 2.5 2.5 2.5 2.5 2.5 2.5 2.5  •Tiler:  1 21 16 1 .21 63 1 21 42 1 21 34 1 21 42 1 21 34 1 21 42 1 21 38 1 21 38 1 21 38. 1 21 38. 1 21  %  atm.  <7oS calctrcc.  cit solid gates nnole/hr x scrubb. |0-3 er  % 1 Sautter oxidized  0.44 0.48 0.44 0.49 1.04 0.38 0.41 0.39 0.41 0.57  1  0.044 0.46  .08 0.5886 1.17 0.5880 0.71 0.5886 99.4 0.86 1.0736 0.86 2.0291 0.86 0.7283 0.86 0.7069 0.33 99.8 0.86 0.7428 0.46 99.9 0.86 0.6996 0.37 97.9 0.80 0.7041 0.30 98.5 1.55 0.61 0.6866 0.52 98.7 0.47 0.86 0.6588  21 21 21 21 21 21 21 21 12 12 12 12 12  0.21 0.21 0.21 0.52 0.83 0.055 0.21 0.21 0.21 0.21 0.21  0.52 0.64 0.42 0.41 0.38 0.67 0.57 0.66 0.62 0.62 0.82  0.21 0.21  1.13  0.86 1.2337 99.3 0.86 2.0500 99.6 0.86 0.8315 0.86 0.8317 0.86 0.6903 97.1 0.86 0.6924 94.2 1 .36 0.6940 1.18 0.6926 95.8 0.86 0.7341 0.86 0.7343 0.83 1.0532 0.80 1.6122 1.48 0.84 1.6588 2.18  12  0.21  22  0.21  0r21 0.21 0.21 0.21 0.21 0.21 0.21 0.52 0.11  0.9890 ^0.9880 0.9890 0.9877 0.-9740 0.9905 0.9897 - 0.9902 0.9897 0.9857 0.9612 0.9882  -  530 7 0 550 B.71 550 B.72 550 B.73 550 B . 7 4 550 B.75 550 B.76 550 B.77 524 B.78 550 B.79 550 B.81 500 B.82 550 B.83 524 B. 69 :  B.  13.4 77 28.1 8 77 28.5 19.8 77 28.5 19.8 77 28.5 23.7 77 28.5 23.7 77 28.5  23.7 50 23.7 22.4 77 22.4 77 15.5 77 9.9 77 9.9 77 8 7 7  17.8 22.0  27.8 28.5 28.5  1.08 1.14 1:14 1.14 1.14 1.14 0.27 0.53 1.05 1.14  90 95 95 95 95 95 22  1  1.14  27.4 0.99 28.5 27.8  1.14  1.05  2.5 2.5 2.5 2.5 2.5 2.5 2.5 2.5 2.5 2.5  36 1 38 1 38 1 38 1 38 1 38 1 11 1 18 1 35 1  45 88 95 38 1 95 2.5 38 1 83 2.5 33 .1 95 2.5 38 1 88 2.5 35 .1  0.85  1.28  0.7.6  2.0166  99.5  0.9870 0.9840 0.9895 0.9897 0.9905 0.9832 0.9857 0.9835 0.9845 0.9845 0.9795 0.9717 0.9787 0.9680  Molybdenite coooinrrar-e Run source B . C . Moly II ll li tl it  B.84  T  t hr  "C 502  Uo —-—i  <?oS  Fo  /sec  mice _tm.  /mir  calo'nej  gates  cit  %$  1 solid saubb Cyclone. er  ^M„S  oxidized  23  77  27.5  1 .00  83  2.5  33-. 1  12  0.21  0.42  0.86  0.7170  98.3  0.9895  23  77  26.8  0.92  78  2.5  31 :1  12  0.21  5.37  0.86  0.7179  97.8  0.9907  22.7  77  29.2  1.21  101  5.0  2 0 :1  12  t).21  4), 3 7  1.71  0.7284  97.7  0.9907  11.3  77  29.2  1.21  101  10.0  1 0 :1  12  0.21  0.56  3.45  1.4294  98.6  0.9860  15.2  77  29.2  1.21  101  7.5  14  :.l  12  0.21  0.44  2.55  1.0868  97.9  0.9890  24  77  29.5  1.27  106  2.5  43.1  12  0.21  0-45  0.86  0.6866  98.8  0.9887  550  15.4  64  24.2  0.89 -  2.5  3 0 :1  10  0.21  0.218  1.04  1.0810  B.91 5 7 5 B.92 5 5 0 B.96 _ 5 5 0 B.98 5 5 0 B.99 5 5 0 B . l 00. .524 • B.-101-526  15.4  64  25.9  1.08  .•91  2.5  36. :1  10  0.21  0.282  1 . 0 4 . 1.0792  21 .0  64  24.2  0.39  75  2.5  3 0 :1  10  0.21  0.14.1 1 . 0 4  0.7971  17.2  64  24.2  0.89  75  2.5  3 0 :1  10  0.21  0.182  1 .04  0.9664  B.85 5 1 1 B.86 5 6 0 ± 3 B.87 5 5 0 ± 3 B.88 . 5 6 0 ± 3 B.-89 ~ 5 7 5  BRENDA-  B.90 II 11 II ii H II II  •  75-  0.13  0.3.6  98.4  0.9945  94.5  0.9929  98.4  0.9964  96.7  0.9954  19.4  64  24.2  0.89  75  2.5  3 0 •1  10  0.21  0.131  1 .04  0.8592  0.25  96.5  0.9967  19.4  64  _4.2  0.89  75^  2.5  30 1  10  0.21  0.131  1 .04  0.8592  0.27  95.1  0.9965  34.5  64  23.2  0.8O  68  2-5  27  1  IQ  0.21  0.115  1 .04  0.4793  97.0  0.9971  27  64  23.4  0.81  69  2.5  27  1  10  0.21  0.119  1.04  0.6154  97.2  0.9970  0.14  tl  B. 1 02- 5 3 6  27  64  23.7  0.85  71  2.5  28  1  10  0.21  0.105  1 .04  0.6156  0.14  97.5  it  0.9973  B.-103  549  27  64  24.1  0.88  74  2.5  30.. 1  10  0.21  0.108  1 .04  II  0.6162  0.21  99.2  B . l 05  97.6  0.9973  550  20.6  64  24.2  0.89  75  2.5  30  1  10  0.21  0.123  1 .04  if  0.8039  0.38  98.8  98.2  0.9969  B..110  5.50  20.6  64  24.2  0.89  75  2.5  30  1  10  0.60  0.130: 1 .04  0.8038  0.9967  BRENDA  Cdou.led  1eached) II  <  B.ni  560  17.6  64  24.2  0.89  74  2.5  30  1  10  0.21  0.130  1 .04  0.9494  0.9967  8.112  550  17.6  64  24.2  0.89  75  2.5  30  1  10  0.60  0.118  1 .04  0.9495  0.9970  0.4376  BRENDA II  B.113  550  38.0  64  24.2  0.89  75  1 .25 60: 1  10  0.21  0.085  0.52  «  B.114  550  20.7  64  24.2  0.89  75  2.50  30: 1  10  0.21  0.132  1.04  KENNECOTT  B.115  550  17.1  64  24.2  0.51  42  2.50  17: 1  28  0.21  0.65  1.04  ll  B.116  570  17.1  64  24.5  0.53  45  2.50  18: 1  28  0.21  0.56  T.04  0.9650  2  98.2  97.9  0.7973  98.0  99.0  0.9645  95.3  0.9837  97.0  0.9860  0.9978 0.9967  Molybdenite source  KENNECOTT II 11  II  ENDAKO 11 11 11  T  t  N°-  8.117  ••hr 524  1- -10  Uo Cry  .  Fo V I . /mil  /sec  0 atm.  %so _ d r i o S  ?  Z  /min  mice  calcines  at  z  %S  01  solid gases m c l e / h r _ SCrufcb er |0"3  \  Sautter  ^ M o S  oxidized  17.1  64  23.2  0.44  37  2.5  li.rl  28  0.21  0.73  1 .04  0.9624  92.9  0.9817  20.0  64 " 2 4 . 2  -0.51  42  2.-50- 1 7 : 1  28  0.21  0.54  1.04  0.8220  94.6  0 79865  B . l 1 9 . 550 •  .2-5,8  64  24.2  0.51  42  2.50  1"/ :1  28 - 0 . 2 1  0.48  1 .04  0.6415  96.3  0.9880  B.120  25..8 . 64  '24.2  0.51  42  2.50  17:1  28  0.21  0.42  1 .04  0.6425  B . l 18' 550  54-9  :  :  B.126  520  118  64  23.2  1.12  -96  2.5  3-7:1  8  0.21  0.32  1 .04  1 .389  B.127  520  15.4  64  23.2  1 .12  96  2.5  37:1  8  0.21  0.30  1 .04  1 .0679  8.128  550  15.4  64  24.2 -  1.1-5  96  2.5  38:1  8  0.21  0.28  1 .04  1 .0793  B . 1 2-9 550  24.0  64  24.2  1 .15  96  2.5  3*.: 1  8  0.21  0.148  1.04  2.6918  2  0.9895:  99  99.5  0.9920  99.5  0.9925  99.5  0.9930  99.5  0.9962  -  ro oo -P=.  A P P E N D I X 14  CALIBRATION CURVE OF THE SCREW FEEDER FOR MoS  2  APPENDIX 15  SULPHUR ANALYSIS OF  Samples of 0.1  CALCINES  M0O3  gr of calcines for %S < 0.5 or  gr for %S > 0.5 were used to determine content of the c a l c i n e s .  0.01  the total sulphur  The sulphur was analyzed by iodimetry  of the gaseous products of samples fused in a Leco induction furnace.  The analysis proceeds according to the  KI0  S0  2  + 5KI + 6H6 * 6KC1 +  3  + I  2  + 2H 0 * H S0 2  2  4  3I  reactions:  2  + 2HI  where the free iodine is used as an indicator by the coloring of a starch s o l u t i o n .  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  iron and tin accelerators for the "Leco" analyzer. of confidence of the standard was ±0.02%S. were also analyzed independently 287  standard The range  -  Samples of calcines  at Can Test laboratories  and  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 0.03  0.88  + 0.06  0.05  + 0.02  0.40  0.55  + 0.15  1.10  0.96  -0.14  0.54  0.54  0.78  0.82  0.00 + 0.04  0.78 0.82 1.41  0.82 0.90 1.78  0.56 0.46 0.42  0.58 0.39 0.39  0.59 0.54  0.49  Endako  This Thesis  0.28  0.280  0.19  0.148  0.42  + Q.04 + 0.08 +0.37 + 0.02 -0.07 -0 07 -0.10 -0.12 ?  L\%S  0.00 -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 S0 AT STEADY STATE 2  289  —  ^  PUBLICATIONS A. Sutulov and I. Wilkomirsky, " P u r i f i c a t i o n of • molybdenite concentrates by c h l o r i nation i n 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 e l e c t r o l y t e , Part I, Thermodynamic properties of the Pb-Ag system", Transactions of the AIME S o c , 242, No. 2 (1968). I. Wilkomirsky and G. Morizot, "Theoretical and p r a c t i c a l 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 v o l a t i l i z a t i o n i n a multiple hearth furnace", I n s t i t u t i o n of Mining and Metallurgy S o c , S e c C, V o l . 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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