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Fluidized bed gasification of some western Canadian coals Gutierrez Despouy, Luis Alberto 1979

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FLUIDIZED BED GASIFICATION OF SOME WESTERN CANADIAN COALS by L U I S . A L B E R T O . G U T I E R R E Z D E S P O U Y A p . S c . i n C h e m . E n g . , U n i v e r s i d a d d e C h i l e , 1 9 7 3 A T H E S I S S U B M I T T E D I N P A R T I A L F U L F I L L M E N T OF T H E R E Q U I R E M E N T S F O R T H E D E G R E E OF M A S T E R " O F - A P P L I E D S C I E N C E i n T H E F A C U L T Y OF G R A D U A T E S T U D I E S D e p a r t m e n t o f C h e m i c a l E n g i n e e r i n g We a c c e p t t h i s t h e s i s a s c o n f o r m i n g t o t h e r e q u i r e d s t a n d a r d THE UNIVERSITY OF BRITISH COLUMBIA June, 1979 ( c p Luis Alberto Gutierrez Despouy, 1979 In presenting th i s thesis in pa r t i a l f u l f i lment of the requirements for an advanced degree at the Univers i ty of B r i t i s h Columbia, I agree that the L ibrary sha l l make i t f ree l y ava i lab le for reference and study. I further agree that permission for extensive copying of th i s thesis for scholar ly purposes may be granted by the Head of my Department or by his representatives. It i s understood that copying or publ icat ion of th i s thesis fo r f i nanc ia l gain sha l l not be allowed without my written permission. Department of Chemical Engineering The Univers i ty of B r i t i s h Columbia 2075 Wesbrook Place Vancouver, Canada V6T 1W5 Date June 28, 1979 i i ABSTRACT Three d i f f e r e n t Western Canadian coals were g a s i f i e d with a i r and steam i n a f l u i d i z e d bed of 0.73 mm Ottawa sand and c o a l , at atmospheric pressure, and temperatures of 1023 - 1175 K to produce a low Btu gas. The coals •:, tested were of two types: one non-caking and two caking coals. The r e s u l t s were compared with those previously obtained f o r the same three coals when g a s i f i e d i n e s s e n t i a l l y the same equipment, but operated as a spouted bed. The e f f e c t s of temperature, coal feed rate, a i r to coal r a t i o , steam to coal r a t i o , c o al q u a l i t y , coal p a r t i c l e s i z e , and bed depth on the gas composition, gas c a l o r i f i c value and the operating s t a b i l i t y of the gas-i f i e r , were established by running g a s i f i c a t i o n t e s t s over a wide range of operating conditions. T y p i c a l c a l o r i f i c value of the gas obtained f o r a l l three coals was i n the range of 2.0 - 2.6 MJ/m , which i s lower than reported f o r the spouted bed and commercially a v a i l a b l e f l u i d i z e d bed g a s i f i e r s . Analysis of the r e s u l t s suggested that i n the present low temperature g a s i f i e r , the com- . bustion and p y r o l y s i s reactions predominate over the g a s i f i c a t i o n reactions. The a b i l i t y to treat caking coals i n f l u i d i z e d bed and spouted bed reactors i s discussed. I t i s concluded that the dispersion of coal i n a bed of i n e r t s i l i c a and ash, rather than hydrodynamic c h a r a c t e r i s t i c s i s the key-factor i n t h e i r success ,in handling caking coals. i i i TABLE OF CONTENTS PAGE ABSTRACT i i ACKNOWLEDGEMENTS . . . . x i CHAPTER I - INTRODUCTION 1 Major Objective of the Research Program 1 Background 1 Theory of Coal G a s i f i c a t i o n . . 7 Experimental Plan H CHAPTER I I - EXPERIMENTAL APPARATUS 15 Segregation Studies and Minimum F l u i d i z a t i o n V e l o c i t y 15 G a s i f i c a t i o n Experiments . 15 CHAPTER I I I - EXPERIMENTAL METHODS 22 Minimum F l u i d i z a t i o n V e l o c i t y and Segregation Studies 22 Coal G a s i f i c a t i o n Experiments 23 General experimental procedure . . 23 Coal preparation and a n a l y s i s 24 Coal feed r a t e 29 F l u i d flow measurement 29 Temperature measurement 31 Sol i d s e l u t r i a t i o n r a t e 31 So l i d s a n a l y s i s 31 Tar determination 33 Gas a n a l y s i s 33 Continuous CO monitor 33 Gas chromatograph 34 Gas c a l o r i f i c value 35 i v C H A P T E R I V - R E S U L T S 36 M i n i m u m F l u i d i z a t i o n V e l o c i t y a n d S e g r e g a t i o n 36 C o a l G a s i f i c a t i o n . . . . . . . 41 O p e r a t i n g e x p e r i e n c e w i t h t h e f l u i d i z e d b e d 41 E f f e c t o f t e m p e r a t u r e 46 E f f e c t o f c o a l q u a l i t y . . . 46 F o r e s t b u r g c o a l 46 S u k u n k a c o a l 50 C o l e m a n c o a l 55 E f f e c t o f c o a l p a r t i c l e s i z e 57 E f f e c t o f a i r t o c o a l r a t i o 61 E f f e c t o f t h e s t e a m t o c o a l r a t i o 73 E f f e c t o f b e d d e p t h 78 C a r b o n c o n t e n t o f t h e b e d 78 S o l i d s a n d c a r b o n e l u t r i a t i o n 78 T a r p r o d u c t i o n 85 C H A P T E R V - D I S C U S S I O N OF R E S U L T S 87 M a s s B a l a n c e s 87 O v e r a l l m a s s b a l a n c e s 87 W a t e r b a l a n c e 89 H y d r o g e n b a l a n c e 90 O x y g e n b a l a n c e 90 C a r b o n b a l a n c e . . 91 E n e r g y B a l a n c e s 91 T h e r m a l E f f i c i e n c i e s 94 U s e f u l C a r b o n R a t i o s 97 V E q u i l i b r i u m C o n s i d e r a t i o n s 99 C o m p a r i s o n o f G a s i f i c a t i o n R e s u l t s f o r F l u i d i z e d a n d S p o u t e d B e d 104 C H A P T E R V I - C O N C L U S I O N S • 109 R E F E R E N C E S 113 L I S T OF T A B L E S v i L I S T OF F I G U R E S v i i i L I S T OF P H O T O G R A P H S i x L I S T OF A P P E N D I C E S x v i LIST OF TABLES TABLE PAGE 1 Typ i c a l Data f o r Low Btu Gas F l u i d i z e d Bed Coal G a s i f i e r s . . . . 4 2 Heats of Reation and Equilibrium Constants f o r Main G a s i f i c a t i o n Reactions (Adapted from reference 8) 9 3. Approximate Relative Rates of the Gas-Carbon Reactions at 800°C and 10.1 kPa Pressure (29) 12 4. P a r t i c u l a r s of Major Equipment 19 5 Ultimate Analysis of Some Western Canadian Coal Samples 26 6 Proximate Analysis of Some Western Canadian Coal Samples . . . . 27 7 Heating Value of Some Western Canadian Coal Samples 28 8 Agglomerating C h a r a c t e r i s t i c s of Some Western Canadian Coal Samples 28 9 F l u i d Flow Measurement Equipment D e t a i l 30 10 Thermocouple Location i n the F l u i d i z e d Bed 32 1.11 F l u i d i z a t i o n of Mixtures of Coal and S i l i c a 37 12 Ty p i c a l Results f o r the g a s i f i c a t i o n of -4.76 mm Forestburg Coal. 51 13 Typical Results for the G a s i f i c a t i o n of -4.76 mm Caking Coals . . 54 14 E f f e c t of Coal P a r t i c l e Size f o r the G a s i f i c a t i o n of Forestburg Coal . . . . 59 15 E f f e c t of Coal P a r t i c l e Size f o r the G a s i f i c a t i o n of Sukunka Coal 60 16 E f f e c t of the Steam to Coal Ratio i n the G a s i f i c a t i o n of Forestburg Coal 74 17 E f f e c t of the Steam to Coal Ratio i n the G a s i f i c a t i o n of Sukunka Coal 76 18 Average Bed Carbon Content During the G a s i f i c a t i o n of Forestburg Coal 79 19 Average Bed Carbon Content During the G a s i f i c a t i o n of Sukunka Coal 80 L i s t o f T a b l e s ( c o n t ' d ) T A B L E 2 0 T y p i c a l S i z e D i s t r i b u t i o n o f S o l i d s E l u t r i a t e d D u r i n g t h e G a s i f i c a t i o n o f 2 . 1 8 mm F o r e s t b u r g C o a l ( R u n N o . 2 0 ) 8 1 2 1 S o l i d s a n d C a r b o n E l u t r i a t i o n D u r i n g t h e G a s i f i c a t i o n o f F o r e s t b u r g C o a l 8 3 2 2 S o l i d s a n d C a r b o n E l u t r i a t i o n D u r i n g t h e G a s i f i c a t i o n . o f C a k i n g C o a l s 8 4 2 3 G a s T a r C o n t e n t f o r t h e G a s i f i c a t i o n o f S o m e W e s t e r n C a n a d i a n C o a l s 8 6 2 4 T h e r m a l E f f i c i e n c i e s a n d & U s e f u l C a r b o n f o r t h e G a s i f i c a t i o n o f S o m e W e s t e r n C a n a d i a n C o a l s . , 9 5 2 5 P r o d u c t R a t i o s f o r t h e . G a s i f i c a t i o n o f S o m e W e s t e r n C a n a d i a n C o a l s 1 0 0 2 6 C a r b o n i n C a r b o n M o n o x i d e t o C a r b o n F e e d R a t i o f o r t h e G a s i f i c a t i o n o f F o r e s t b u r g C o a l s 1 0 3 27 C o m p a r i s o n o f R e s u l t s f o r t h e G a s i f i c a t i o n o f S o m e W e s t e r n C a n a d i a n C o a l s i n F l u i d i z e d a n d S p o u t e d B e d 1 0 5 2 8 C o m p a r i s o n o f t h e M a i n O b s e r v e d O p e r a t i n g C h a r a c t e r i s t i c s o f a F l u i d i z e d a n d S p o u t e d B e d R e a c t o r s , 1 0 7 v i i i LIST OF FIGURES FIGURE PAGE 1 F l u i d i z e d Bed for Segregation Studies 16 2 Schematic Diagram of F l u i d i z e d Bed.Pilot Plant . 17 3 F l u i d i z e d Bed G a s i f i e r 18 4 F l u i d i z a t i o n of a Mixture of 80% 3.36-1.18 mm S i l i c a and 20% 3.36-1.18 mm Coal 38 5 F l u i d i z a t i o n of Mixtures of 0.73 mm Ottawa Sand and 3.36-1.18 mm Coal . 40 6 Carbon Content of F l u i d i z e d Bed During the G a s i f i c a t i o n of Forestburg Coal 44 7 Carbon Content of F l u i d i z e d Bed During the G a s i f i c a t i o n of Sukunka Coal 45 8 Ty p i c a l A x i a l Temperature P r o f i l e s of F l u i d i z e d Bed 47 9 E f f e c t of Temperature on the G a s i f i c a t i o n of Forestburg Coal .... 48 10 E f f e c t of Temperature on the G a s i f i c a t i o n of Sukunka Coal . . . . 49 11 E f f e c t of A i r to Coal Ratio and P a r t i c l e Size on the C a l o r i f i c Value of Gas Produced From Forestburg Coal 58 12 E f f e c t of A i r to Coal Ratio on the Gross C a l o r i f i c Value of Gas Produced from Sukunka Coal. 62 13 E f f e c t of Coal Feed Rate on the Concentration of Methane Produced from Forestburg Coal 64 14 E f f e c t of Coal Feed Rate on the Production of Methane from Forestburg Coal 65 15 E f f e c t of Coal Feed Rate on the Methane Production from Caking Coals . 66 16 E f f e c t of Coal Feed Rate on the Production of Hydrogen from Forestburg Coal 68 17 E f f e c t of Coal Feed Rate on the Production of Carbon Monoxide from Forestburg Coal 69 18 E f f e c t of A i r Flow on the Gas Composition from Forestburg Coal at a Fixed Coal Feed Rate 70 L i s t o f F i g u r e s ( c o n t ' d ) F I G U R E 1 9 E f f e c t o f A i r t o C o a l R a t i o o n t h e G a s C o m p o s i t i o n f r o m F o r e s t b u r g C o a l . . , 7 1 2 0 E f f e c t o f t h e A i r t o C o a l R a t i o o n t h e G a s C o m p o s i t i o n f r o m S u k u n k a C o a l 7 2 2 1 E f f e c t o f t h e S t e a m t o C o a l R a t i o o n t h e G a s C o m p o s i t i o n f r o m 4 . 0 6 mm D i a m e t e r F o r e s t b u r g C o a l 7 5 2 2 E f f e c t o f t h e S t e a m t o C o a l R a t i o o n t h e G a s C o m p o s i t i o n f r o m S u k u n k a C o a l 77 2 3 C a l c u l a t e d E f f e c t o f E n e r g y L o s s e s o n t h e T e m p e r a t u r e o f t h e G a s L e a v i n g t h e C o a l G a s i f i e r 9 3 2 4 E q u i l i b r i u m P r o d u c t - S t e a m R a t i o s a s a F u n c t i o n o f T e m p e r a t u r e . 1 0 1 L I S T OF P H O T O G R A P H S P H O T O G R A P H P A G E 1 B e d S o l i d s D u r i n g t h e G a s i f i c a t i o n o f 4 . 0 6 mm . S u k u n k a C o a l ( R u n N o . 4 0 ) 5 2 2 C a k e d C o l e m a n C o a l a t E n d o f R u n N o . 4 1 5 6 X LIST OF APPENDICES APPENDIX PAGE I DETAILS OF FLUIDIZATION GRID . 118 II TOTAL. GAS FLOW MEASUREMENT BY ORIFICE PLATE 120 P i t o t T u b e 121 O r i f i c e Plate C a l i b r a t i o n . 127 II I IMPINGER SYSTEM FOR GAS SAMPLE CLEANING 131 IV CALIBRATION OF ROTAMETERS .. • 134 Scrubbing Nitrogen Rotameter 135 G a s i f i e r A i r Rotameter . . 135 Incinerator A i r Rotameter .. 139 Steam Meter 139 Gas Sample Rotameter . . . . . 144 V CALIBRATION CURVES FOR GAS ANALYSIS 147 IV MINIMUM FLUIDIZATION VELOCITY AND SEGREGATION 154 IIV GASIFICATION RESULTS 161 IIIV MASS BALANCES 171 General Procedure . . . . 172 Sample C a l c u l a t i o n Run No. 38 172 IX ENERGY BALANCES . 184 General Procedure . . . . 185 Sample Cal c u l a t i o n , Run No. 38 189 ACKNOWLEDGEMENTS This work was supported by grants from the Department of Energy, Mines and Resources, Ottawa, and the B.C. Department of Mines and Petroleum Resources. The contribution of the Department of Chemical Engineering, U.B.C. to t h i s work i s acknowledged, i n p a r t i c u l a r to Dr. A.P. Watkinson, f o r d i r e c t i n g the research program, and to the Chemical Engineering workshops which b u i l t parts of the equipment. Thanks are also due to several s t a f f members of B.C. Research f o r t h e i r support and f o r making a v a i l a b l e f a c i l i t i e s and some equipment. To Dr. C C . Walden, Dr. J.C. Mueller, Mr. A. Bruynesteyn and others f or t h e i r support and encouragement. To Mr. N. Bundalli f o r h e l p f u l d i s -cussions. To Mrs. P. MacLeod f o r her patience and capable services i n typing the manuscript, and very s p e c i a l l y to Dr. P.C. T r u s s e l l f o r . f i n a n c i a l assistance during these studies. L a s t l y , to my wife, Monica, for her excellent drawings and continued support and encouragement during my graduate studies. CHAPTER I - INTRODUCTION MAJOR OBJECTIVE OF THE RESEARCH PROGRAM The major objective of t h i s research program was to study the g a s i f i -c a t i o n of some Western Canadian coals i n a f l u i d i z e d bed under conditions as close as possible to those under which spouted bed g a s i f i c a t i o n of these coals had been previously investigated i n t h i s Department. A secondary objective was to elucidate whether the r e l a t i v e success i n gasi f y i n g caking coals i n a spouted bed was due to the high gas v e l o c i t i e s i n the spout or to a coal d i l u t i o n e f f e c t . BACKGROUND The impending shortage of o i l and natural gas compounded with the un-r e l i a b i l i t y of continuous supply and the escalating prices of these premium fue l s have prompted a massive research e f f o r t i n coal conversion technology as a means of providing clean burning f u e l and petrochemical feedstocks. The p r i n c i p a l scene of t h i s research development has been Europe and the (2 4") United States, although i t has been suggested v ' that Canada should also be involved i n developing t h i s technology. A very large number of processes have been investigated or proposed f o r the production of low, medium or high Btu gases. Inasmuch as several comprehensive reviews on the subject have (5-9) been published , only a few s a l i e n t f a c t s w i l l be pointed out here. Low Btu gas i s produced by the g a s i f i c a t i o n of c o a l , char or even wood, with steam and a i r , generally at atmospheric pressure, and temperatures of about 1073-1400 K. The r e s u l t i n g gas i s therefore d i l u t e d with nitrogen, 2 which comprises 45 - 60% by volume of the dry gas, the r e s t being carbon monoxide, hydrogen, carbon dioxide and small amounts (up to 3%) of methane. Low Btu gases have a gross c a l o r i f i c value i n the range of 3.73 - 9.31 MJ/std m3 (100-250 Btu/scf). Medium Btu gas i s produced by g a s i f i c a t i o n of coal or coke with steam and oxygen at atmospheric or higher pressure, and temperatures of 755-1755K. The produced gas has l i t t l e nitrogen (2%), carbon monoxide and hydrogen con-centrations of 30-40%, up to 4% methane, and a gross c a l o r i f i c value of 9.31-20.49 MJ/std m3 (250-550 Btu/scf). The chemistry involved i n the pro-duction of medium and low Btu gas i s e s s e n t i a l l y the same and w i l l be analysed l a t e r i n t h i s chapter. High Btu gas or synthetic natural gas i s e s s e n t i a l l y methane and has a gross heating value of 35.40^37.26 MJ/std m3 (950-1000 Btu/scf). High Btu gases can be produced by two routes. The f i r s t s t a r t s from a medium Btu gas and involves s h i f t conversion of water to hydrogen by carbon monoxide f o l -lowed by methanation. The s h i f t conversion (eq. (1)) CO +, H 20 — ^ C0 2 + H 2 AH = -32.55 KJ/at.g C (1) i s generally done at 588 - 700 K and 2750 kPa i n the presence of a c a t a l y s t , i n such a way that the C0:H 2 r a t i o i s adjusted to 1:3 i n preparation f o r the c a t a l y t i c methanation step (eq.(2)). Carbon dioxide i s removed by chemical absorption. N i CO + 3H 2 ;=± H 20 + CH 4 AH = - 232.50 KJ/at.g C (2) The second route i s the d i r e c t hydrogenation of coal or h y d r o g a s i f i c a t i o n (eq. (3)). C + 2H 2 CH 4 AH = - 91.46 KJ/at.g C (3) The reaction i s c a r r i e d out at pressures i n excess of 3435 kPa. From the c o a l g a s i f i c a t i o n l i t e r a t u r e i t i s evident that although there are many processes that have been proposed, or are under development, only a few have been commercially proven. These are, with some v a r i a t i o n s , essent-i a l l y four: The f i x e d bed L u r g i : (lO --*-^ y the entrained bed Koppers-„ fc ,(10,11,15) , ... , . , „ . , , (10,11,14) , , , , T o t z e k v the f l u i d i z e d bed Winkler , and the moving bed Wellman-Galusha g a s i f i e r . A l l of these reactors w i l l y i e l d a low Btu gas when operated with a i r , and a medium Btu gas when operated with oxygen. The Koppers-Totzek i s only operated with oxygen. Comparison of the commerc-i a l l y a v a i l a b l e technology i s provided i n references 7, 8, 14, & 18. From an inspe c t i o n of the coal g a s i f i c a t i o n l i t e r a t u r e , i t also appears that the processes which are e i t h e r a v a i l a b l e now or l i k e l y to become a v a i l -able i n the near future are c o a l g a s i f i e r s producing low or medium Btu gas. Of these, many use a f l u i d i z e d bed reactor e.g. the Winkler g a s i f i e r , the CO2 acceptor process (Conoco Coal Development C o . ) , the Synthane Process (U.S. Bureau of Mines, ERDA) ^>18)^ y _ g a s p r o c e s s ( i n s t i t u t e of Gas T e c h n o l o g y ) t h e Agglomerating Ash Process (Batelle/Union Carbide) ^ \ the Westinghouse Process ^ O ' ^ ^ t h e H V c j r a n e p r o c e s s (U.S. Bureau of Mines, ERDA) ^ < " > ' ^ \ and the Hygas Process ( I n s t i t u t e of Gas (12 13) Technology) ' . A l l of these processes except for the Winkler g a s i -f i e r , are at the development stage, e i t h e r on a p i l o t or demonstration s c a l e . In s p i t e of a wealth of d e s c r i p t i v e l i t e r a t u r e on f l u i d i z e d bed g a s i f i e r s very l i t t l e data on actual operating conditions or q u a l i t y of the gas pro-duced i s a v a i l a b l e . Table 1 summarizes such data f o r f l u i d i z e d bed reactors producing low Btu gas. The a v a i l a b l e data show that f o r a i r blown f l u i d i z e d bed reactors operating at atmospheric pressure, a gas q u a l i t y of 3.71-4.40 TABLE 1: TYPICAL DATA FOR LOW BTU GAS FLUIDIZED BED COAL GASIFIERS Variable (21) Bamag Winkler (22) Davy-Power Gas Winkler Modified ( l 8 ) Synthane (5) U Gas (19) Westinghouse Type of Coal Subbituminous A Montana Subbituminous I l l i n o i s No.6 Pittsburg coal Indiana & Pittsburg bituminous Coal size (mm) 15.9 - 0 9.53 - 0 0.84 - 0 - 0.8 - 1.2 (av.) Coal feed rate (g/s) or (g/s x m2) 135.6 (g/s m 2 ) 1.89 - 2.58 (g/s) - 75.6 - 163.8 (g/s) Pressure (kPa) 101 101 2020 2405 1546 Average temperature (K) 1366 1253 - 1292 811 1200 - 1294 3 Air flow std (m /s) 1.18-2.2 x l O - 3 Steam flow (g/s) N i l 1.26-2.39 Gas velocity at average temp (m/s) 0.07-0.16 0.37 - 0.61 Gas composition % by volume (dry basis) H 2 11.7 13.1 .21.5 13.3 CO 21.7 22.10 10.1 19.4 CHt, 0.7 0.84 5.6 4.7 C0 2 9.8 7.12 17.9 10.1 N 2 55.3 56.82 43.5 51.9 C 2H 6 - 0.7 -H2S 0.8 .0.1 0.7 0.7 Gas Gross heating value ( MJ/std m^ ) 3.91 4.40 6.15 5.71 3.73 - 4.47 5 MJ/std m i s expected, while operating at higher pressures y i e l d s a gas of • 3 . •.. a higher heating value, i n the range of 4.47-6.15 MJ/std m . Advantages of f l u i d bed operation are excellent g a s - s o l i d contact, high heat t r a n s f e r r a t e s , r e l a t i v e l y low residence times, high turndown c a p a b i l i t y , and uniform temperature and bed composition which provides good c o n t r o l . How-ever, one of the main disadvantages of f l u i d bed g a s i f i e r s i s t h e i r i n a b i l i t y to t r e a t caking c o a l s without pretreatment. When heated, caking coals f i r s t soften and flow l i k e thermoplastics and then r e s o l i d i f y (cake) into a swollen porous s o l i d (coke). The r e s o l i d i f i -c a t i o n process i s accompanied by emission of gas and condensable vapours. Two types of models f o r the caking process have been proposed: p h y s i c a l and (23) physiochemical v . In the p h y s i c a l models, caking i s assumed to be a superposition of the p h y s i c a l phenomenom of melting and the chemical phen-omenon of p y r o l y s i s . In the physiochemical models, the softening and r e -s o l i d i f i c a t i o n of c o a l i s considered a consequence of p y r o l y s i s which a l t e r s chemical bonds c r o s s l i n k i n g polynuclear s t r u c t u r e s . Caking properties of (24) coal have been r e l a t e d to t h e i r petrography and rank . The caking properties of c o a l are customarily represented by t h e i r f r e e swelling index (FSI), which i s a measure of the increase i n volume of a c o a l when i t i s heated without r e s t r i c t i o n under s p e c i f i e d c o n d i t i o n s . The (25) standard method of t e s t f o r FSI involves heating several 1 g samples of c o a l to 1090 K w i t h i n a s p e c i f i c time to prepare buttons of coke. The shape or p r o f i l e of the coke buttons i s standardized and numbered from 1.-9. The sketch below i l l u s t r a t e s these p r o f i l e s f o r c o a l samples with FSI at 1,4 and 7. 6 FSI =1 FSI = 4 FSI = 7 For a coal with a FSI of 7 the area of the p r o f i l e i s roughly 5 times that'of a coal with a FSI = 1 When a caking coal i s fed into a f l u i d bed g a s i f i e r , . agglomer-ation of the coal occurs, p a r t i c u l a r l y near" the feed point. As we l l , s i n t e r i n g at the f l u i d i z a t i o n g r i d may occur. These two problems lead to d e f l u i d i z a t i o n and render smooth, continuous operation impossible. The caking problem i s aggravated i n anhydrogen atmosphere and with an increase i n operating pressure . Since most of the coal i n the Eastern. States and- a 'sizeable portion of that i n Western Canada i s of the caking type, e f f o r t s have been directed towards developing processes to accommodate these coals. Proposed solutions for avoiding d e f l u i d i z a t i o n i n g a s i f i e r s have been reviewed i n the (Oft) l i t e r a t u r e . These include oxidative pretreatment of the coal i n a separate (27) f28) f l u i d i z e d v e s s e l , a l k a l i n e pretreatment , entrained pretreatment , char d i l u t i o n , good s o l i d mixing i n a turbulent or fa s t f l u i d i z e d bed coupled with coal dispersion i n dry ash, and mechanical s t i r r i n g of the bed. However, a l l of these propositions imply the use of yet unproven technology, added hardware or added chemical costs. Another recent approach that was reported to have per-(19) mitted the successful g a s i f i c a t i o n of various caking coals consisted of pneumatically introducing the coal into the f l u i d i z e d bed reactor through a single o r i f i c e at the bottom. Foong j^t a l . at the University of B r i t i s h Columbia ^"^noted that t h i s and other proposed modifications tended to make the f l u i d i z e d bed reactor 7 s i m i l a r to a spouted bed which they had been i n v e s t i g a t i n g for the g a s i f i -c a tion of two Western Canadian caking coals. Their r e s u l t s showed that.these coals could s u c c e s s f u l l y be g a s i f i e d i n a spouted bed of s i l i c a p a r t i c l e s of the same s i z e as the coal feed, a l b e i t with some problems i n feeding the coal into the reactor. However, a question arose from t h i s work. Was t h i s success due to the d i l u t i o n e f f e c t of the s i l i c a and ash or to the a b i l i t y of the high v e l o c i t y gases i n the spout to break-up any agglomeration? Or was i t a combination of these ' two phenomena? Furthermore, since gas q u a l i t y de-pends on coal properties, i t i s d i f f u c u l t to assess a novel g a s i f i e r based on t e s t s on a few coals which have not been g a s i f i e d i n more standard equip-ment . The objective of t h i s research was to elucidate these questions by pro-cessing the same coals i n a f l u i d i z e d bed, under operating conditions as s i m i l a r as possible to the ones used i n the spouted bed experiments. This work would also provide basic data on the g a s i f i c a t i o n of Western Canadian coals i n a f l u i d i z e d bed. THEORY OF COAL GASIFICATION Coal g a s i f i c a t i o n i s the reaction of coal with a i r or oxygen and steam, or mixtures of these, to y i e l d a gaseous product s u i t a b l e as a f u e l or as a petrochemical feedstock. This gaseous product i s a mixture of carbon mon-oxide, carbon dioxide, hydrogen, methane, nitrogen, l i g h t hydrocarbons and s u l f u r bearing compounds, p r i n c i p a l l y hydrogen s u l f i d e . The chemical reactions occurring i n a c o a l g a s i f i e r can be divided i n three main groups: combustion reactions, g a s i f i c a t i o n reactions and pyro-l y s i s reactions. 8 The combustion reactions are homogeneous or heterogeneous reactions with oxygen as shown below (Eq. (4) to ('7) (4) (5) (6) (7) A l l of these reactions are exothermic (Table 2) and provide the heat nec---essary»for some of - the endothermic g a s i f i c a t i o n reactions, as well as the heat necessary to bring the reactants to.the r e a c t i o n temperature. Overall the g a s i f i c a t i o n system i s autothermic. The g a s i f i c a t i o n reactions produce combustible gases from heterogeneous reactions between carbon and steam or gaseous products, or from homogeneous reactions among the gaseous products. The p r i n c i p a l g a s i f i c a t i o n . , reactions are as follows: C + 2H 2 ;== CH^ (3) C + H 20 ^ CO + 'H (8) C + 2H 20 ^ C0 2 + 2H 2 (9) C + C0 2 ^ 2 CO (10) These reactions are endothermic (Table 2) except reactions (9) and (3). In addi-t i o n the^homogeneous-water-gas s h i f t r eaction w i l l take place i n a g a s i f i e r CO + H 20 ^ C0 2 + H 2 (11) At higher pressures the following reactions can also occur: CO + 3H 2 ^ CH 4 + H 20 (12) C0 2 + 4H 2 ;==± CH 4 + 2H 20 (13) Of the g a s i f i c a t i o n reactions, the carbon steam r e a c t i o n (eq. 8) i s con-sidered to be the most important and t y p i c a l of coal g a s i f i c a t i o n . Although reactions (3) ,(10) , (12)&'(13) produce a gas of a much greater heating value, •C + V 0 2 ;==^  CO CO + h o 2 ^ c 0 2 C + o 2 ^ co 2 H 2 + h 0 2 H 2 0 TABLE 2: HEATS OF REACTION AND EQUILIBRIUM CONSTANTS FOR MAIN GASIFICATION REACTIONS (Adapted from reference 8) Reaction A H (KJ/mol) Kp 1100 K 1300 K 1100 K 1300 K C + h 0 2 CO -112.614 -113.882 8.8 x 10 9 1.31 x 10 9 CO + h 0 2 ^ C0 2 -282.295 -281.416 7.21 x 10 8 6.29 x 10 6 C + 0 2 ^_ co 2 -394.913 -395.298 6.35 x 1 0 1 8 8.25 x 1 0 1 5 H 2 + h o 2 v = ± H 20 -248.422 -249.685 7.60 x 10 8 1.15 x 10 7 C + H 20 ^ — * CO + H 2 135.807 135.636 2.62 1.14 x 10 2 C + 2H 20 ^—* C0 2 + 2H 2 -146.492 -145.780 1.11 x 10 1 6.24 x 10 1 C + C0 2 ^ ^ 2 CO 169.685 167.534 1.22 x 10 1 2.08 x 10 2 CO + H 20 C0 2 + H 2 -33.878 -31.898 1.10 0.55 C + 2H 2 ^ — - CH^ -90.605 -91.735 3.68 x 10" 2 7.93 x 10~ 3 t h e i r i m p o r t a n c e i s l i m i t e d f o r g a s i f i c a t i o n a t a t m o s p h e r i c p r e s s u r e . P y r o l y s i s r e f e r s t o t h e e n d o t h e r m i c t h e r m a l d e c o m p o s i t i o n o f c o a l i n t o c h a r a n d v o l a t i l e c o m p o u n d s . T h e l a t t e r a r e r e p r e s e n t e d b y t a r , l i g h t o i l , w a t e r , h y d r o g e n , c a r b o n m o n o x i d e , c a r b o n d i o x i d e a n d l i g h t h y d r o c a r b o n s ( E q . ( 1 4 ) . C o a l t h e r m a l ^ > C Q + + H + H „ 0 + C » H •> + t a r + l i g h t o i l + c h a r ( 1 4 ) d e c o m p o s i t x o n 2 2 2 m n T h e v o l a t i l e c o m p o n e n t s c a n f u r t h e r r e a c t w i t h w a t e r o r h y d r o g e n . F r o m t h e n u m b e r o f r e a c t i o n s i n v o l v e d i t c a n b e a p p r e c i a t e d t h a t a g a s -i f i e r i s a v e r y c o m p l e x c h e m i c a l s y s t e m w i t h a l l r e a c t i o n s o c c u r r i n g s i m -u l t a n e o u s l y t h r o u g h o u t t h e r e a c t o r o r i n l o c a l i z e d a r e a s ( t y p i c a l l y i n a f i x e d b e d ) . S o m e t i m e s i t m a y e v e n b e o f a d v a n t a g e t o c a r r y o u t s o m e o f t h e s e r e a c t i o n s i n s e p a r a t e v e s s e l s . F o r i n s t a n c e , m a n y p r o p o s e d g a s i f i c a t i o n p r o c e s s e s c a r r y o u t d e v o l a t i l i z a t i o n o f c o a l i n o n e r e a c t o r f o l l o w e d b y g a s i f i c a t i o n o f t h e c h a r i n a s e c o n d r e a c t o r . T a b l e 2 l i s t s t h e e q u i l i b r i u m c o n s t a n t s f o r t h e i m p o r t a n t r e a c t i o n s a t t w o t y p i c a l g a s i f i c a t i o n t e m p e r a t u r e s a n d a t m o s p h e r i c p r e s s u r e . T h i s d a t a s h o w s t h a t f r o m t h e t h e r m o d y n a m i c p o i n t o f v i e w t h e c o m b u s t i o n o f c a r b o n t o c a r b o n d i o x i d e i s b y f a r t h e m o s t f a v o r a b l e r e a c t i o n . T h e r e s t o f t h e c o m b u s t i o n r e a c t i o n s f o l l o w , w h i l e t h e d i r e c t h y d r o g e n a t i o n o f c a r b o n ( r e a c t i o n ( 3 ) ) i s t h e l e a s t f a v o r a b l e r e a c t i o n . A t e m p e r a t u r e i n c r e a s e f a v o r s t h e , c a r b o n - s t e a m r e a c t i o n ( E q . ( 8 ) ) a n d t h e B o u d o i i a r d r e a c t i o n ( E q . ( 1 0 ) ) t h e m o s t , , w h i l e t h e . i m p o r t a n c e o f t h e r e s t o f t h e c o m b u s t i o n r e a c t i o n s d e c r e a s e s . T h e r e f o r e , a n i n c r e a s e i n g a s i f i c a t i o n t e m p e r -a t u r e s h o u l d d o u b l y i n c r e a s e t h e c o n c e n t r a t i o n o f c o m b u s t i b l e g a s e s . Although thermodynamics i s a very u s e f u l t o o l i n p r e d i c t i n g the maximum concentrations obtainable, the over-riding consideration i n a p r a c t i c a l system i s the k i n e t i c of the reactions involved. Unfortunately, the k i n e t i c -of a coal g a s i f i c a t i o n system are extremely d i f f i c u l t to analyse due to the number of reactions superimposed, e f f e c t s of the coal and ash structure and complex f l u i d c i r c u l a t i o n patterns. However k i n e t i c studies of the i n d i v i d u a l reactions involved can at l e a s t provide a comparison of the rates of the d i f f e r e n t reactions. These studies i n d i c a t e that the f a s t e s t reaction i n the g a s i f i c a t i o n system i s the oxidation of hydrogen to water. H 2 + h 0 2 ^ H 20 (7) The rest of the oxidation reactions are also f a s t when compared with the g a s i f i c a t i o n reactions (8) and (10). C + H 20 ;=± CO + H 2 (8) C + C0 2 ^ = 2 CO (10) while the slowest r e a c t i o n i s the d i r e c t hydrogeneration of carbon (Eq.(3), Table 3). C + 2H 2 CH 4 (3) In summary, both thermodynamic and k i n e t i c considerations i n d i c a t e that i n a coal g a s i f i c a t i o n system at atmospheric pressure as long as any oxygen i s present the combustion reactions are favoured over the g a s i f i c a t i o n r e -actions of C-H20 and C-C0 2 while the d i r e c t hydrogenation of carbon i s the lea s t favorable r e a c t i o n . EXPERIMENTAL PLAN Given the objectives of t h i s research, the experimental /apparatus used ' for spouted bed g a s i f i c a t i o n ^ " ^ was e s s e n t i a l l y conserved with only a few modifications and improvements. The spouted bed was transformed into a 12 TABLE 3: APPROXIMATE RELATIVE RATES OF THE GAS-CARBON REACTIONS AT 800°C AND 10.1 kPa PRESSURE (29) Reaction Relative Rates C - 0 2 1 x 10 5 C - H 20 3 C - C0 2 1 C - H 2 3 x 10 - 3 f l u i d bed by simply introducing a f l u i d i z a t i o n g r i d above the spouting o r i f i c e . Other mo d i f i c a t i o n s were r e l a t e d to so l v i n g the p r e v i o u s l y en-countered problems i n feeding caking c o a l into the bed, improving the method of obtaining mass balances by metering the t o t a l gas flow out of the reactor, and making a rough measurement of the amount of t a r s produced. Since c o a l was to be fed i n t o a f l u i d i z e d bed of i n e r t s , some room temper-ature experiments were c a r r i e d out to assess the extent of segregation under f l u i d i z a t i o n conditions of d i f f e r e n t mixtures of c o a l and i n e r t p a r t i c l e s . The g a s i f i c a t i o n experiments i n the f l u i d i z e d bed were to u t i l i z e the same three Western Canadian coals used i n the spouted bed experiments: One non-caking coal (Forestburg) and two caking coals (Sukunka and Cole-man) . The non-caking c o a l was used i n a s e r i e s of experiments to estab-l i s h the e f f e c t of the various operating parameters on both the q u a l i t y of operation and of the gas obtained. These parameters included: Bed temperature Coal feed rate A i r to coal r a t i o - F l u i d i z i n g v e l o c i t y - Steam to coal r a t i o Bed depth - P a r t i c l e s i z e Presence of f i n e s i n feed. Having established the e f f e c t of the d i f f e r e n t operating parameters, and the conditions under which best r e s u l t s could be obtained with non-caking c o a l , the g a s i f i c a t i o n of caking coals i n the f l u i d i z e d bed was attempted under selected operating cond i t i o n s . During a l l g a s i f i c a t i o n runs the following raw data were obtained addition to the values of the operating parameters: - Gas composition - Total gas flow - Ash production and c h a r a c t e r i s t i c s - Carbon bed content - Tar production (selected runs). CHAPTER I I - EXPERIMENTAL APPARATUS SEGREGATION STUDIES AND MINIMUM FLUIDIZATION VELOCITY P r i o r to s t a r t i n g any g a s i f i c a t i o n experiments, the minimum f l u i d i z a t i o n v e l o c i t y and mixing patterns of c o a l / s i l i c a mixtures were studied at room temperature i n the simple experimental set-up shown i n Figure 1. This con-si s t e d of a 0.15 m (6") I.D. by 0.79 m (31") long glass column with a 60° brass c o n i c a l base. The column was f i t t e d with a water manometer f o r measur-ing the pressure drop across the bed and a c a l i b r a t e d rotameter f o r measuring the a i r flow. The f l u i d i z a t i o n grid" was a perforated a c r y l i c p late designed according to the procedure outlined by Kunii and Levenspiel (-^^. i t con-si s t e d of a c i r c u l a r flange with an inner 0.15 m (6") diameter c i r c l e per-forated with 61 holes of 2 mm diameter arranged i n a tr i a n g u l a r p i t c h (Appendix I ) . GASIFICATION EXPERIMENTS The coal g a s i f i c a t i o n work was conducted i n the small p i l o t plant schem-.:.. a t i c a l l y shown i n Figure 2. P a r t i c u l a r s of the main items of equipment are l i s t e d i n Table 4. The f l u i d bed reactor and coal feeding arrangement are shown with more d e t a i l i n Figure 3. The p r i n c i p a l components of the p i l o t plant are described below. Numbers i n brackets r e f e r to legend i n Figure 2. The c o a l , crushed and screened to the desired p a r t i c l e s i z e was loaded into the storage b i n (1)' maintained under a small nitrogen pressure. Coal flow was aided by a small nitrogen current. When working with caking coals a small scraper was f i t f l u s h with the feeder's r o t a t i n g d i s c . Coal flowed then into a slanted 0.13 m (5") I.D. pipe and then into the f l u i d i z e d bed (3), at a point near i t s top. The slanted feeding pipe was cooled by water flowing through a copper c o i l wrapped around i t (Figure 3). When working WATER MANOMETER FIGURE 1 : FLUIDIZED BED FOR SEGREGATION STUDIES ROTAMETER AIR R R O T A M E T E R TO E X H A U S T S Y S T E M STIRRER I 2 3 4 5 6 7 8 9 10 I I 12 13 14 15 C O A L S T O R A G E BIN C O A L F E E D E R FLUID IZED B E D R E A C T O R P R O P A N E B U R N E R C Y C L O N E A S H C O L L E C T O R O R I F I C E P L A T E I N C I N E R A T O R S C R U B B E R I M P I N G E R S ICE BOX D E H Y D R A T O R G L A S S WOOL G A S S A M P L I N G CONTINUOUS CO C O L U M N P O R T MONITOR N I T R O G E N DRAIN E M E R G E N C Y Q U E N C H I N G NITROGEN FIGURE 2: SCHEMATIC DIAGRAM OF FLUIDIZED BED PILOT PLANT 18 a b c d e f g h i N . S I G H T G L A S S t G A S TO C Y C L O N E S C R A P E R V I E W I N G P O R T L O A D I N G F U N N E L C O A L S T O R A G E BIN C O A L T A B L E F E E D E R S T I R R E R C O O L I N G C O I L E X P A N S I O N Z O N E F L U I D I Z E D B E D S O L I D S D I S C H A R G E P R O P A N E B U R N E R W A T E R h P R E S S U R E G A U G E — — 0 R U P T U R E DISC V I E W I N G A N D L O A D I N G P O R T S ^ / T H E R M O C O U P L E S F L U I D I Z A T I O N GRID P R O P A N E — FIGURE 3: FLUIDIZED BED GASIFIER TABLE 4: PARTICULARS OF MAJOR E Q U I P M E N T F l u i d Bed: Diameter: . 1 5 m Height: . 8 6 m Cone angle: 6 0 ° F l u i d i z a t i o n g r i d : perforated plate (Appendix I) Mat e r i a l : column type 3 1 6 L S.S. Coal Feeder: Type: Table feeder, v a r i a b l e speed Capacity: 2 5 k g/h G a s B u r n e r f o r s t a r t - u p : P r o p a n e f i r e d , P r e m i x L i n i n g : r e f r a c t o r y c e m e n t M a n u f a c t u r e r : E c l i p s e F u e l E n g i n e e r i n g C o . G a s / a s h S e p a r a t i o n c y c l o n e : D i a m e t e r : 1 5 0 mm C y l i n d e r h e i g h t : 5 0 0 mm C o n e h e i g h t : 3 0 0 mm Off-gas combustor: Type: E l e c t r i c furnace Scrubber: Type: water spray, counter current to .gas flow Diameter: 1 5 0 mm Height: 1 . 8 m Packing: 1 3 mm p o r c e l a i n saddle or m e t a l l i c lathe shavings with caking coal a small v a r i a b l e speed s t i r r e r was used to break up any agglomerations forming i n the feeding pipe. The feeding pipe was also f i t t e d with a viewing port kept, clean by a small nitrogen flow. A l l n i t r o -gen fed into the reactor was metered by a rotameter. The f l u i d i z e d bed (Figure 3) consisted of a main c y l i n d r i c a l section of 0.15 m (6") I.D. by 0.61 m (24") long type 316 L s t a i n l e s s s t e e l with a wall thickness of 6.4 mm (V). A s i m i l a r section of 0.25 m (10") long could.be added to lengthen the reactor to a t o t a l f l u i d bed depth of 0.86 m (34"). The main section was equipped with two viewing ports which could also be used as loading ports, a valved 25 mm (1") s o l i d discharge pipe, a rupture d i s c for safety, and chromelSalumel thermocouples spaced at 0.15 m. The bottom of the reactor was f i t t e d ; with a f l u i d i z a t i o n g r i d . This consisted of a 6.4 mm (V) t h i c k s t a i n l e s s s t e e l perforated plate of i d e n t i c a l design to the a c r y l i c f l u i d i z a t i o n g r i d used during the segregation studies (see Section 1 t h i s Chapter). Even a i r d i s t r i b u t i o n into the f l u i d i z a t i o n grid-, ' was achieved by a 60° c o n i c a l section j u s t below i t . The g a s i f i c a t i o n reactor was also equipped with a premix propane burner for start-up (4). A i r and propane flows into the refactory l i n e d combustion chamber were metered by c a l i b r a t e d rotameters. Just above the burner and below the c o n i c a l section a r e f r a c t o r y l i n e d mixing chamber was f i t t e d . This mixing chamber was a short c y l i n d r i c a l section f i t t e d with entrances for reacting steam and nitrogen f o r emergency quenching of the reactor. Gas produced i n the f l u i d i z e d bed reactor flowed into an expansion zone fo r s o l i d disengagement con s i s t i n g of two cones joined at t h e i r base. The gas then flowed into a cyclone (5) (see Table 4) equipped with an ash c o l l e c -t i o n chamber (6). Cleaned hot gas was metered by a c a l i b r a t e d o r i f i c e p l ate (7) (see Appendix II). before being incinerated i n an e l e c t r i c furnace (8). A small gas stream of approximately 80 cm~/s by-passed the i n c i n e r a t o r and was pumped by a vacuum pump through an ice-cooled t r a i n of four impingers (10) (31) s i m i l a r to the ones used for high volume stack sampling (Appendix I I I ) . The impingers were used to cool the gas sample, to capture remaining s o l i d p a r t i c l e s and to condense any t a r s . The cleaned gas sample was then passed through a d r i e r i t e desiccating column (12) , a glass wool f i l l e d column (13), and pumped into a continuous CO analyser (15) which monitored the g a s i f i e r performance. A c a l i b r a t e d rotameter metered the gas. A sampling port (14) j u s t before the entrance of the continuous CO monitor enabled a gas sample to be taken for complete analysis i n a gas chromatograph (see Chapter I I I ) . CHAPTER I I I - EXPERIMENTAL METHODS 22 MINIMUM FLUIDIZATION VELOCITY AND SEGREGATION STUDIES The minimum f l u i d i z a t i o n v e l o c i t y and s o l i d s mixing pattern.of mixtures containing 0, 5, 15 and 20% coal i n s i l i c a were studied at room temperature using the experimental set-up described i n Chapter I I . The s i l i c a p a r t i c l e s used were either crushed and screened quartz of 2.3 mm nominal average s i z e (1.18 mm - 3.36 mm) or commercially a v a i l a b l e 90% 0.73 mm (0.84-0.59 mm) Ottawa sand. The coal was crushed and screened Forestburg coal of nominal average si z e of 2.3 mm (1.18 mm - 3.36 mm). The experimental procedure for each sample was as follows. A 4.54 kg (10 lbs) s o l i d s sample was prepared by thoroughly mixing ap-proximate amounts of s i l i c a and coal p a r t i c l e s . The mixture was then poured into the glass f l u i d i z a t i o n column. Random loose packing was obtained by gently f l u i d i z i n g the s o l i d s i n a i r and then turning the a i r o f f . The height of the column was measured. Starting from t h i s point, the a i r flow into the column was increased stepwise. For each a i r flow (measured i n the c a l i b r a t e d rotameter), pressure drop, column height and v i s u a l observations r e l a t i n g to f l u i d i z a t i o n state and mixing of coal and s i l i c a were recorded. P l o t s of pressure drop versus s u p e r f i c i a l a i r v e l o c i t y were prepared to es t a b l i s h the minimum f l u i d i z a t i o n v e l o c i t y of the d i f f e r e n t mixtures, according to the method most generally accepted i n the l i t e r a t u r e . xhe v i s u a l observations, on the other hand, were u s e f u l to e s t a b l i s h the nature of f l u i d i z a t i o n , the extent of segregation, and the minimum s u p e r f i c i a l v e l o c i t y at which segregation of coal and s i l i c a seemed to disappear. 23 COAL GASIFICATION EXPERIMENTS: General Experimental Procedure Coal crushed and screened to the desired p a r t i c l e s i z e was loaded into Che coal storage b i n . A i r was then turned on to a low flow to pre-vent s i l i c a p a r t i c l e s from dropping into the combustion chamber of the propane burner while loading the i n e r t s i n t o , reactor. One of the view-ing ports of, the reactor was used for charging the reactor with 0.73 mm Ottawa sand i n amounts varying between 4.5 kg and 7.9 kg according to the desired bed depth. Next, water flow to the feed pipe cooling c o i l , the scrubber nitrogen to various parts of the coal feeding system and a i r to the e l e c t r i c furnace were turned on. A i r flow into the g a s i f i e r was increased to support propane combustion and increase heat t r a n s f e r to the bed of i n e r t s . Once the bed had reached a uniform temperature of 773 - 873 K (about one hour) the coal feeder was turned on at a small feed rate while the propane feed rate was decreased u n t i l shut-o f f . The coal feed rate was then slowly increased u n t i l the reactor reached a temperature of approximately 1073 K. At t h i s point the ex-perimental conditions were set. F i r s t , steam was turned on to the desired l e v e l . Then, since i t was found that coal feed rate and a i r flow were not independent 1 v a r i a b l e s , one of these two operating v a r i a b l e s was f i x e d at a desired l e v e l and the other adjusted u n t i l a steady operation was reached. It. was considered that the reactor had reached steady operation when both the bed temperature (measured at one point approxi-mately i n the middle of the bed) and the carbon monoxide content of the produced gas gave steady continuous readings. At t h i s point, gas samples were taken approximately every 15-30 min and analyzed i n the gas chroma-24 tograph. When s u f f i c i e n t data had been c o l l e c t e d , the operating conditions were changed to obtain a new steady operation or the reactor was shut down. Each g a s i f i c a t i o n run consisted of 1-4 hours of steady operation. Bed s o l i d s samples of approximately 100 g were taken at i n t e r v a l s of 0.5 - 1 hour using the s o l i d s sampling port, while the ash c o l l e c t e d i n the cyclone was sampled only at the end of the run. The shut-down procedure was as follows. Coal feeder and a i r were shut o ff while the quenching nitrogen was turned on and the steam feed rate increased. When the bed reached a temperature of 573-673 K, the bed s o l i d s were discharged through the s o l i d s sampling port. The a i r was then turned back on to complete the cooling of the reactor. Coal Preparation and Analysis: A l l coal samples were crushed and screened to the desired s i z e range by B.C. Research. The coal preparation procedure was as follows. The coal samples (nominal - 12.7 mm) were f i r s t a i r dried for approximately a week and then screened i n a vibr a t o r y screen through a set of screens of 4.76 mm, 3.36 mm and 1.0 mm. The oversize (+ 4.76 mm) was crushed by an adjustable jaw crusher i n such a way as to obtain a maximum amount of sample i n the 1.0 mm - 3.36 mm s i z e range. Representative portions of the d i f f e r e n t coal samples and p a r t i c l e sizes were then sent to General Testing Co., Vancouver, f or an a l y s i s . For one of the coal samples (Forestburg coal) a -4.76 mm + 0.00 mm composite sample was obtained by combining the i n d i v i d u a l s i z e f r a c t i o n s as shown below: 25 Size range (mm) Nominal Average Size (mm) % by weight 4.76 - 3.36 4.06 15.0 3.36 - 1.00 2.18 50.0 1.00 - 0.0 0.50 35.0 -4.76 + 0.0 dp =^ V— = 1.03 100.0 Z_ (-JL_) i dp'i This coal sample had a nominal average s i z e of 1.03 mm and i t s compos-i t i o n was judged as t y p i c a l f o r a -4.76 + 0.0 mm sample of t h i s coal crushed following the procedure outlined above. Samples of nominal average s i z e of 0.95 mm and 0.53 mm were also obtained from t h i s coal by using a set of screens of 1.18 mm, 0.710 mm and 0.35 mm, and c o l l e c t i n g the 1.18 - 0.710 mm and 0.710 - 0.355 fr a c t i o n s r e s p e c t i v e l y . Three types of Western Canadian coals were g a s i f i e d i n the small p i l o t plant. These had the same o r i g i n as the coal samples g a s i f i e d i n the spouted bed by Foong et^ al^. Ultimate and proximate analysis of the d i f f e r e n t coal samples are presented i n Tables 5 and 6 r e s p e c t i v e l y . Heating values are presented i n Table 7 and agglomerating c h a r a c t e r i s t i c s i n Table 8. Forestburg coal from the Luscar operation i n Alberta i s a sub-bituminous coal with zero free swelling index. Since t h i s i s a non-caking c o a l , i t was used,in-the.-main part of t h i s research i n e s t a b l i s h i n g the e f f e c t s of the d i f f e r e n t operating parameters and most adequate conditions f o r the f l u i d i z e d g a s i f i e r . Coleman coal from Coleman C o l l i e r i e s on the B.C. - Alberta border i s described as a medium v o l a t i l e bituminous coal of moderate caking properties with a free swelling index of 4. Sukunka coal from the Chamberlain seam of TABLE 5 ULTIMATE ANALYSIS OF SOME WESTERN CANADIAN COAL SAMPLES Forestburg Coleman Sukunka Analysis (%) _-4.76 mm + 0 . 0 mm 3.36 mm - 1 .00 mm- 4.76 mm - 3. 36 mm 3.36 mm - 1. 00 mm as received dry as received dry as received dry as received dry Carbon 50.96 66.79 77.24 78.29 78.31 79.06 79.65 80.81 Hydrogen 5.89 4.25 4.50 4.41 4.50 4.44 4.50 4.45 Sulphur 0.46 0.60 0.30 0.30 0.54 0.55 0.50 0.50 Ash 8.0 10.48 10.66 10.80 12.30 12.42 10.72 10.31 Oxygen 33.47 16.28 6.04 4.92 3.10 2.27 3.27 2.56 Nitrogen 1.22 1.60 1.26 1.28 1.25 1.26 1.36 1.37 T A B L E 6 P R O X I M A T E A N A L Y S I S O F S O M E W E S T E R N C A N A D I A N C O A L S A M P L E S Forestburg Coleman Sukunka Analysis (%) -4.76 mm - 0.0 mm 4.76 mm - 3 36 mm 3.36 mm - 1 00 mm 3.36 mm - 1'. 00 mm 4.76 mm - 3. 36 mm 3.36 mm - 1. 00 mm as received dry as received dry as received dry as received dry as received dry as received dry Total moisture 23.7 - 24.2 - 24.4 - 1.34 - 0.95 '- 0.82 -Ash 8.00 10.48 6.65 8.7 7 7.21 9.58 10.66 10.80 12.30 12.42 10.72 10.81 Vo l a t i l e Matter 27.38 35.88 27.76 36.60 27.59 36.47 25.61 25.96 21.77 21.98 22.32 22.51 Fixed carbon 40.92 53.64 41.39 54.63 40.80 54.00 63.39 63.24 69.98 65.60 66.14 66.68 TABLE 7 HEATING VALUE OF SOME WESTERN CANADIAN COAL SAMPLES Sample d e s c r i p t i o n Heat value (KJ/kg) as i s dry Forestburg(-4.76 mm +0.0 mm) Coleman (3.36 mm - 1.00 mm) Sukunka (4.76 mm - 3.36 mm) Sukunka (3.36 mm - 1.00 mm) 20694 31424 31612 31577 26215 31852 31914 31838 TABLE 8' AGGLOMERATING CHARACTERISTICS OF SOME WESTERN CANADIAN COAL SAMPLES Sample (a) Free Swelling Index Fluidity (by Gieseler Fluidity Test) Initial Softening Maximum Fluidity Solidification Temperature (K) DDPM(b) Temperature (K) DDPM Temperature (K) DDPM Forestburg 0 - _ _ _ - -Coleman 4 725 1 739 3 765 0 Sukunka 7 723 1 761 140 795 0 (a) Coal size = 2.18 mm for a l l samples (b) 1 DDPM = 1 dial division per minute = 3.6 are degree per minute in Gieseler fluidity testing aparatus (25). 29 the B.P. Canada property in Northeastern B.C. is a bituminous highly caking coal with a free swelling index of 7. Coal Feed Rate: The coal feed rate into the reactor was measured by the rate of dis-placement of the gas-solid interface in the coal storage bin. This could be followed through the viewing ports on the side of the bin (See .. Figure 3). Before loading the coal bin, the coal was weighed and i t s bulk density determined. By taking the time in which the coal level in the bin dropped a known distance, the volume of coal could be calculated. Since the dimensions of the bin and the coal bulk density were known, the coal mass feed rate could be calculated. " This figure was checked against the feed rate calculated by dividing the amount of coal utilized during the entire run over the run duration. These two measurements were found to be in good agreement. Fluid Flow Measurement Except for the total wet gas flow out of the reactor, which was measured by a calibrated o r i f i c e plate, a l l f l u i d flows in and out of the reactor were measured by calibrated f u l l view rotameters. The methods of calibration, calibration data and curves are given in Appendices II and IV respectively. Particulars of the flow metering equipment are presented in Table 9. A l l gas flows are reported at the standard conditions of 294 K and 101.3 kPa. TABLE 9 FLUID FLOW MEASUREMENT EQUIPMENT DETAIL Stream Flow Measurement Equipment Used Scrubbing nitrogen Gilmont rotameter s i z e 14 G a s i f i e r a i r Brooks rotameter. Tube s i z e R-10M-25-1, Floa t s i z e 10-RS-64 Incinerator a i r Brooks rotameter. Tube s i z e R-10M-25-3. Floa t s i z e 10-RV-64 Steam meter Brooks rotameter. Tube s i z e R-8M-25-2 Floa t s i z e 8-RV-3 Gas sample meter Gilmont rotameter . Size F622 To t a l wet gas O r i f i c e p l a t e , own design. (Appendix I I ) . Temperature Measurement: A l l temperatures were measured with c a l i b r a t e d Chromel-alumel thermo-couples connected to a multiple channel switch and a d i g i t a l d i s p l a y . Temperature on the reactor was measured at four or f i v e d i f f e r e n t depths (See Table 10), at 25 mm from the reactor w a l l . Thermocouples were also used to measure the temperature below the f l u i d i z a t i o n g r i d , ambient temperature, a i r temperature at f l u i d i z a t i o n a i r rotameter, gas'tempera-•ture upstream of the o r i f i c e p l a t e , at the entrance of the gas sampling l i n e , and at the ou t l e t of the scrubber. Solids E l u t r i a t i o n Rate: Solids e l u t r i a t e d ' from the f l u i d i z e d bed reactor were measured a f t e r each g a s i f i c a t i o n experiment simply by emptying and weighing the content of the receptacle below the cyclone. By analysing the cyclone catch f o r ash content, both the ash and carbon e l u t r i a t i o n rates could be c a l c u -lated (see next paragraph). Solids Analysis: Both the s o l i d s samples withdrawn from the f l u i d i z e d bed and a rep-resentative sample of the cyclone catch were analysed for ash content by i n c i n e r a t i o n . The 50 - 100 g samples were weighed, put into tared c r u c i b l e s • and incinerated f o r 10 hours at 1173 K i n an e l e c t r i c furnace. A f t e r cooling, the samples were weighed again to determine the r e s i d u a l ash. The weight l o s s was taken as carbon due to the f a c t that the high reactor temperatures would have driven off water and v o l a t i l e matter from the s o l i d samples. TABLE 10 THERMOCOUPLE LOCATION IN THE FLUIDIZED BED Distance from Fluidization Grid (m)./ 0.70 m Fluid bed 0.86 m Fluid bed 0.05' 0.21 0.36 0.53 . 0.06 0.31 0.46 0.61 0.76 33 Tar Determination: On selected runs, the approximate tar content of the gas produced was determined by the following method. After the reactor had reached steady operation, the gas sample stream was diverted from the cleaning set of impingers i n use into a second set of impingers f o r a f i x e d period of time (1 - 2 hours). Here, the t a r condensed i n the cold water and on the walls pf the impingers. The tar was then removed by d i s s o l v i n g i t with acetone. The acetone was then evaporated under vacuum and the r e s i d u a l tar weighed. The tar content of the gas was calculated by d i v i d i n g the t o t a l tar thus obtained over the t o t a l gas volume passing through the impingers during the tar sampling period. Gas Analysis: A l l gas analyses were reported on a dry.basis. The carbon monoxide content of the gas was continuously monitored by an i n f r a r e d analyser while the remaining gases were analysed from samples taken at convenient time i n t e r v a l s (15 - 30 min) i n a gas chromatograph. Both instruments were connected to a Watanabe chart recorder. Continuous CO monitor: A Beckman continuous i n f r a r e d analyser, model No 864-13-4, which operated i n the range of 0 - 25% CO was employed. The instrument gave a continuous s i g n a l on the recorder chart. Carbon monoxide content of the gas was obtained from a factory provided c a l i b r a t i o n chart (Figure V - l , Appendix V). The instrument was c a l i b r a t e d f o r every run by adjusting the readings of two points of the c a l i b r a t i o n curve. The zero reading was adjusted by passing through the i n s t r u -ment a standard gas stream containing 10% and 90% nitrogen. The second c a l i b r a t i o n point was obtained by adjusting the i n s t r u -ment's reading on a standard sample containing 10% Cp^, 21.3% CO and nitrogen as balance, to the c a l i b r a t i o n curve reading (91% deflec-:. t i o n ) . The continuous CO analyser provided an excellent means of moni-toring the g a s i f i e r performance as well as providing an i n d i c a t i o n of whether the reactor was operating under steady conditions. Gas chromatograph: 3 At convenient time i n t e r v a l s , 5 cm gas samples were injec t e d into a Hewlett-Packard gas chromatograph model 5710 A. The gas chromatograph was equipped with a molecular sieve column, a poro-pack column and a thermal conductivity detector which resolved hydrogen, carbon dioxide, nitrogen, oxygen, methane and carbon monoxide. Ty p i c a l gas chromatograph tracings are shown as Figure V-2, Appendix V. The percentage content of each gas i n the sample was determined from the peak height and the corresponding c a l i -b ration curves. (Figures V-3 - V-5, Appendix V). In order.to avoid possible v a r i a t i o n s of the c a l i b r a t i o n curves, these were obtained for each run by analysing d i l u t i o n s i n a i r of a standard gas sample (Table V - l ; Appendix V). Experience showed that carbon monoxide was best obtained from the i n f r a r e d analyser and carbon dioxide by d i f -ference. Therefore, only H^, and CH^ were obtained from the gas chromatograph (since there was no 0^ under g a s i f i c a t i o n operation conditions). 35 Gas C a l o r i f i c Value: The c a l o r i f i c value of the gas was calculated from the gross (high) heat of combustion of the three combustible components of the gas pro-duced (hydrogen, carbon monoxide and methane), and the dry gas compos-i t i o n . In order to enable comparison with values reported i n the l i t e r -ature, the gross c a l o r i f i c value of a unit volume of gas (h ) was eg calculated at the North-American standard conditions (288.6 K = 60°F and 101.6 kPa = 30 i n Hg dry). Then: v/v) + h CH 4 where: h = 12.109 (MJ/m3) H2 h C Q = 11.997 (MJ/m3) h_„ = 37.743 (MJ/m3) (*) M Fuel Flue Gases I I American Gas Association CHAPTER IV - RESULTS MINIMUM FLUIDIZATION VELOCITY AND SEGREGATION The minimum f l u i d i z a t i o n v e l o c i t y of 4.45 kg mixtures of coal.and s i l i c a each of a s i z e range 1.18 - 3.36 mm were determined at room temperature i n a glass column of the same diameter as the coal g a s i f i e r (Figure 2). At the same time the mixing patterns were v i s u a l l y observed to determine conditions where segregation or slugging affected f l u i d i z a t i o n . Pressure drop across the bed, bed height and v i s u a l observations made at d i f f e r e n t a i r flows are presented i n Tables VI-1 to VI-6, Appendix VI, f o r the various mixtures studied. This data i s summarized i n Table 11. For a l l mixtures of coal and s i l i c a tested, the curves showing pressure drop across the bed as a function of the s u p e r f i c i a l a i r v e l o c i t y (Figures 4 & 5) had a s i m i l a r shape and were t y p i c a l of f l u i d flow through a bed of so l i d s For a fixed bed, the pressure drop increased almost l i n e a r l y with the f l u i d v e l o c i t y . Upon reaching the minimum f l u i d i z a t i o n v e l o c i t y the pressure drop ceased to increase and remained almost constant or de-creased s l i g h t l y as the f l u i d v e l o c i t y was further increased. The minimum f l u i d i z a i t o n v e l o c i t y of a mixture of 80% 3.36 - 1.18 mm s i l i c a and 20% coal of the same s i z e was approximately 0.78 m/s (Figure 4). Once the bed became f l u i d i z e d , coal segregated from.the s i l i c a and f l o a t e d at the top of the bed. As the s u p e r f i c i a l a i r v e l o c i t y was increased, seg-regation started to disappear, and at a s u p e r f i c i a l v e l o c i t y of approximately 1.10 m/s (1.41 Uj^ £?') the bed was completely mixed. However, the disappearance of segregation corresponded to the commencement of slugging (Figure 4). TABLE 11 FLUIDIZATION OF MIXTURES OF COAL AND SILICA 80% 3.36-1.18mm S i l i c a 20% Coal 80% Ottawa Sand 20% Coal 85% Ottawa Sand 15% Goal -90% Ottawa Sand 10% Coal 95% Ottawa Sand 5% Coal 100% Ottawa Sand Superficial a i r velocity (m/s) Pressure drop (kPa) Superficial a i r velocity (m/s) Pressure drop (kPa) Superficial a i r v e l o c i t y (m/s) Pressure drop (kPa) Superficial a i r velocity (m/s) Pressure drop (kPa) Superficial a i r velocity (m/s) Pressure drop (kPa) Superficial a i r velocity (m/s) Pressure drop (kPa) 0 0 0 0 0 0 0 0 0 0 0 0 0.26 0.32 0.21 1.35 0.16 1.02 0.16 1.00 0.16 1.02 0.16 1.17 0.39 0.75 0.26 1.64 0.21 1.30 0.21 1.32 0.21 1.27 0.21 1.50 0.52 1.22 0.31 2.19 0.26 1.77 0.23 1.54 0.23 1.50 0.23 1.74 0.65 1.87 0.34 2.19 0.28 2.02 0.26 1.67 0.26 1.69 0.26 2.02 0.70 1.89 0.36 2.17 0.31 2.17 0.28 1.97 0.28 1.97 0.28 2.24 0.78 2.19 0.41 2.14 0.34 2.12 0.31 2.12 0.31 2.17 0.31 2.27 0.83 1.99 0.47 2.12 0.36 2.12 0.34 2.09 0.34 2.17 0.34 2.24 0.88 1.99 0.52<a) 2.12 0.41 2.12 0.36 2.09 0.36 2.14 0.36 2.27 0.93 1.94 0.56 2.07 0.47 2.07 0.39 2.07 0.39 2.14 0.41 2.22 0.98 1.89 0.62<b) 2.07 0.52 ( a ) 2.09 0.41 2.07 0.41 2.12 0.47 2.22 1.04<b> 1.87 0.78 1.99 0.57 2.04 0.44 2.07 0.44 2.12 0.52 2.19 1.09<a) 1.87 0.62 2.02 0.47 2.04 0.47 2.07 0.65<b> 2.14 1.14 1.84 0.67 1.97 0.52<a) 2.04 0.52 ( a ) 2.07 0.73 2.12 0.72<b> 1.94 0.57 2.02 0.57 2.02 0.91 2.07 0.91 1.87 0.62<b> 0.70 0.78 0.91 1.97 1.94 1.92 1.84 0.65<b) 0.70 0.78 0.91 2.02 1.99 1.94 1.87 Coal and s i l i c a in size range 1.18 - 3.36 mm (a) Minimum s u p e r f i c i a l velocity at which segregation disappears (b) Onset of slugging. 38 3.0 2.0 o Q_ 1.0 0.8 0.6 cn o S 0.4 CO CO LU cr CL 0.2 — » - i r—b ! 1 i / f t ' j oo j °°°Oooo FIXED BED BUBBLING BED SLUGGING BED SEGREGATED BED WELL MIXED BED 0.2 0.4 0.6 1.0 SUPERFICIAL AIR VELOCITY, m/s 2.0 FIGURE 4: FLUIDIZATION OF A MIXTURE OF 80% 3.36-1.18 mm SILICA AND 20% 3.36-1.18 mm COAL Substitution of the coarse crushed s i l i c a with uniformly sized 0.73 mm Ottawa sand resulted i n an improvement of the mixing of coal and i n e r t s i n the f l u i d i z e d bed. The minumum f l u i d i z a t i o n v e l o c i t y of four d i f f e r e n t mix-tures of Ottawa sand and varying amounts of coal (5% to 20%) was very s i m i l a r , at approximately 0.32 m/s (Figure 5). For these mixtures very l i t t l e segre-gation was observed, even f o r f l u i d v e l o c i t i e s s l i g h t l y above the minimum f l u i d i z a t i o n v e l o c i t y , ( U m f ) • V i s u a l observation indicated that in;general, for a f l u i d v e l o c i t y larger' than. 0.5 m/s (1.56 U- •) excellent dispersion of m i coal i n s i l i c a was obtained. At t h i s v e l o c i t y , the bed was vigorously bub-b l i n g , while at f l u i d v e l o c i t i e s l arger than 0.7 m/s for a l l cases the bed was slugging. Ov e r a l l , ttiese experiments suggested that s o l i d s segregation was l e s s intense and f l u i d i z a t i o n much smoother f o r mixtures of 3.36 - 1.18 mm coal and Ottawa sand than f o r mixtures of coal and s i l i c a of the same s i z e . The minimum s u p e r f i c i a l v e l o c i t y required f o r achieving good dis p e r s i o n of coal i n s i l i c a was i n the range: 1.41 U , < U < 1.56 U c mf — — mf (33) which i s i n l i n e with values c i t e d i n the l i t e r a t u r e . As expected f o r beds of the same weight, pressure drop across the bed was s i m i l a r i n a l l cases. As well the minimum f l u i d i z a t i o n v e l o c i t y f o r mixtures of 3.36 -1.18 mm coal i n Ottawa sand were much smaller than f o r mixtures of coal; and s i l i c a of the same (3.36 - 1.18 mm) s i z e . This could be an important consideration i n an i n d u s t r i a l a p p l i c a t i o n . As a r e s u l t of these t e s t s , i t was decided to operate the f l u i d bed gas-i f i e r using 0.73 mm Ottawa sand as an i n e r t dispersing medium for coa l . 3.0 2.5 2.0 1.5 .0 OTTAWA 3.36- 1.18 mm SAND COAL % % O 80 20 A 85 15 a 90 10 v 95 5 • 100 0 -PACKED-BED BUBBLING BED-1 -SLUGGING-BED SLIGHT- •NO-SEGREGATION SEGREGATION 0.1 0.2 0.4 0.6 SUPERFICIAL AIR VELOCITY , m/s 0.8 1.0 FIGURE 5: FLUIDIZATION OF MIXTURES OF 0.73 mm OTTAWA SAND AND 3.36-1.18 mm COAL COAL GASIFICATION The three coals described i n Chapter I I I were g a s i f i e d i n a f l u i d i z e d bed of 0.73 mm Ottawa sand. As outlined before, most of the g a s i f i c a t i o n experi-ments u t i l i z e d the non-caking Forestburg c o a l , and the data derived was used i n e s t a b l i s h i n g the e f f e c t of the d i f f e r e n t operating parameters on the qu a l i t y of the operation and of the gas obtained. G a s i f i c a t i o n of caking coals was attempted only under selected conditions so as to e s t a b l i s h the f e a s i b i l i t y of g a s i f y i n g caking coals i n the f l u i d i z e d bed and as a means of confirming some of the experimental findings on the e f f e c t of the operating v a r i a b l e s derived i n the f i r s t part of t h i s research. The experimental r e s u l t s are summarized i n Tables VII-1 and VII-2, Appendix VII. The tabulated values are averages of 2 to 10 d i f f e r e n t measure-ments, and i n general, observed v a r i a t i o n s from the reported averages are within 5%, except f o r gas analysis where v a r i a t i o n s of up to 10% may have occurred. Except when otherwise noted, a l l reported values correspond to steady operation of the g a s i f i e r . The g a s i f i e r was considered to be oper-ating under steady conditions when there was neither a continuous decrease or increase of temperature at a rate exceeding 0.8 K/min at a t y p i c a l oper-ating temperature of 1080 K nor a continuous change i n the carbon monoxide content of the gas. The e f f e c t s of the d i f f e r e n t operating parameters on the g a s i f i e r oper-ati o n and on the gas q u a l i t y are discussed i n the following sections. Operating Experience with the F l u i d i z e d Bed The s t a b i l i t y of the f l u i d i z e d bed operation was found to be a d e l i -cate balance between mass and heat t r a n s f e r , and both the gas composition and bed temperature were very s e n s i t i v e to changes i n the operating parameters. These parameters were i n t e r r e l a t e d and i t was found impos-s i b l e to change one operating v a r i a b l e and keep the others constant with-..; out introducing i n s t a b i l i t y i n the operation. The two main operating v a r i a b l e s were the coal feed rate and the a i r flow rate. Once one of these parameters was f i x e d , the other as well as the bed temperature, was f i x e d within a r e l a t i v e l y narrow range. A higher degree of freedom to change the steam i n j e c t i o n rate was experienced, and t h i s v a r i a b l e could be changed within a wider range without d i s e q u i l i b r a t i n g the system. However, an increase i n steam i n j e c t i o n rate i n v a r i a b l y caused a decrease of the bed temperature, as well as a change i n the gas com-p o s i t i o n . The a i r flow into the reactor was l i m i t e d at i t s lower end by the flow corresponding to a s u p e r f i c i a l a i r v e l o c i t y (at the average reactor temp-erature) equal or larger than the previously determined minimum non-segregating v e l o c i t y (0.5 m/s) (see Section 1, t h i s Chapter). The upper 3 l i m i t f o r the a i r flow was approximately 0.007 - 0.008 std m /s or a corresponding s u p e r f i c i a l v e l o c i t y of about 1.6 m/s. For values larger than t h i s the bed began to slug excessively and overflowed into the f r e e -board. In p r a c t i c e , i t was found necessary to operate at a i r flows r e -s u l t i n g i n s u p e r f i c i a l -velocities i n the range of 1.0 - 1.4 m/s; 1.2 m/s being a t y p i c a l value. Under these conditions the bed was v i o l e n t l y agitated or slugging. On the other hand, coal feed rate was l i m i t e d at i t s lower end by the minimum feed rate which at the given a i r flow would r e s u l t i n an operating temperature below 1175 K, the maximum safe temperature f o r the unlined s t a i n l e s s s t e e l reactor. This.corresponded to a feed rate of about 0.27 g/s of Forestburg coal (see run 10, Table VII-1, Appendix VII) at the m i n i m u m a i r f l o w f o r t h e a b s e n c e o f s e g r e g a t i o n i n t h e b e d . T h e m a x i m u m c o a l f e e d r a t e w a s d e t e r m i n e d b y t h e c o a l r e a c t i v i t y . I f c o a l i s f e d t o t h e b e d a t a h i g h e r r a t e t h a n i t c a n r e a c t , i t w i l l a c c u m u l a t e , c a u s i n g a n i n c r e a s e o f t h e b e d ' s s o l i d s c o n t e n t , a n d e v e n t u a l l y , d e f l u i d i z a t i o n . T h e c a r b o n c o n t e n t o f t h e b e d w a s m o n i t o r e d a s a f u n c t i o n o f t i m e f o r v a r i o u s g a s i f i c a t i o n r u n s u s i n g F o r e s t b u r g a n d S u k u n k a c o a l s ( T a b l e s V I I - 3 a n d V I I - 4 , A p p e n d i x V I I ) a n d t h e r e s u l t s p l o t t e d i n F i g u r e s 6 a n d 7 r e -s p e c t i v e l y . D u r i n g t h e g a s i f i c a t i o n o f F o r e s t b u r g c o a l , c a r b o n d i d n o t a c c u m u l a t e i n t h e b e d w h e n t h e a i r t o c o a l r a t i o w a s >_ 4 . 9 3 ( F i g u r e 6 ) . H o w e v e r , f o r a n a i r t o c o a l r a t i o o f 4 . 0 1 t h e b e d c a r b o n c o n t e n t i n - . . c r e a s e d r a p i d l y , t o o v e r 3 0 % , i n l e s s t h a n 1 h o u r . T h e r e f o r e , t h e m i n i m u m a i r t o c o a l r a t i o t h a t a l l o w e d o p e r a t i o n w i t h o u t c a r b o n a c c u m u l a t i o n i n . t h e b e d w a s i n t h e r a n g e o f 4 . 0 1 - 4 . 9 3 , a n d p r o b a b l y s l i g h t l y - a b o v e 4 . 0 j u d g i n g f r o m o p e r a t i n g e x p e r i e n c e . I f t h e p r e v i o u s l y s t a t e d m a x i m u m a i r f l o w i s u s e d , t h e m a x i m u m c o a l f e e d r a t e w i t h o u t c a r b o n a c c u m u l a t i o n i s e x p e c t e d t o b e 1 . 9 5 - 2 . 3 9 g / s . T h e e x p e r i m e n t a l . . r e s u l t s c o n f i r m e d ' t h i s s i n c e a l l r u n s w i t h a i r t o c o a l r a t i o s b e l o w 4 . 0 w e r e u n s t a b l e ( s e e T a b l e V I I - 1 , A p p e n d i x V I I ) . S u k u n k a c o a l a p p e a r e d t o b e s o m e w h a t l e s s r e a c t i v e t h a n F o r e s t b u r g c o a l . F i g u r e 7 s h o w s t h a t a t a n a i r t o c o a l r a t i o o f 4 . 6 3 , i n i t i a l l y c a r b o n a c c u m u l a t e s i n t h e b e d q u i t e r a p i d l y , w i t h t h e r a t e o f a c c u m u -l a t i o n d e c r e a s i n g a f t e r a b o u t a n h o u r o f o p e r a t i o n . F o r a s l i g h t l y h i g h e r a i r t o c o a l r a t i o ( 4 . 6 9 ) t h e b e d c a r b o n c o n t e n t o s c i l l a t e d b e t w e e n 2 1 % a n d 1 0 % , w h i l e o p e r a t i o n a t a n a i r t o c o a l r a t i o o f 4 . 1 1 r e s u l t e d i n a s o m e w h a t u n s t a b l e o p e r a t i o n w i t h t e m p e r a t u r e d e c r e a s i n g a t a r a t e o f a p p r o x i m a t e l y 0 . 8 K / m i n ( R u n N o . 3 8 ) . 44 35 50 100 150 200 250 300 TIME , min FIGURE 6: CARBON CONTENT OF FLUIDIZED BED DURING THE GASIFICATION OF FORESTBURG COAL 45 0 50 100 150 200 250 TIME , min FIGURE 7: CARBON CONTENT OF FLUIDIZED BED DURING THE GASIFICATION OF SUKUNKA COAL E f f e c t of Temperature A x i a l temperature p r o f i l e s i n the reactor were t y p i c a l of a f l u i d -ized bed (Figure 8) i . e . f a i r l y f l a t with maximum temperature d i f f e r - .. ences of approximately 50 K. The maximum temperature i n the bed occurred at a point approximately 13 mm below the feed point. As mentioned e a r l i e r , one of the basic l i m i t a t i o n s of the present experimental set-up was the i n a b i l i t y of operating at average temper-atures over about 1175 K because the reactor was not r e f r a c t o r y l i n e d . This severely affected the q u a l i t y of the gas produced since most of the data a v a i l a b l e i n the l i t e r a t u r e suggests that the carbon-steam re a c t i o n does not proceed at considerable rates at temperatures below about 1273 K (see Chapter I ) . A l l of the experimental data was obtained i n the 1023 -1173 K temperature range. Within t h i s range, the r e s u l t s suggest that temperature had no e f f e c t on the q u a l i t y of the gas obtained from either Forestburg coal (Figure 9) or Sukunka coal (Figure 10). E f f e c t of Coal Quality Forestburg coal A t o t a l of 33 g a s i f i c a t i o n runs were conducted with Forestburg coal of f i v e d i f f e r e n t p a r t i c l e s i z e s , a l l below 4.76 mm (Table VII-1, Appendix VII). As expected with a non-caking coal the f l u i d -ized bed g a s i f i e r operated smoothly, and no p a r t i c u l a r problems were encountered under the wide range of operating conditions tested. For coal feed rates of 0.27 - 2.37 g/s, a i r flows of 2.7.5 x 10" 3 - 7.92 x 10~ 3 m3/s and steam feed rates of 0.0 - 1.165 g/s the best gas obtained had a c a l o r i f i c value of 2.94 MJ/m (78.9 Btu/cf) 4 7 200 I 1 0 0 UJ rr ZD < or UJ CL UJ I -1000 900 FEED POINT FEED POINT O BED A BED DEPTH DEPTH = 0.61 m = 0.86 m (RUN (RUN 25) 17) 1 1 1 1 1 i i 0.2 0.4 0.6 0.8 D I S T A N C E F R O M F L U I D I Z A T I O N GRID , m FIGURE 8: TYPICAL AXIAL TEMPERATURE PROFILES OF FLUIDIZED BED 3.0 ro E ^ 2.5 2.0 UJ _ J < > O u. 1.5 cc o _J < o 1.0 CO CO o g 0 5 A O ° o 0 o ° • P A R T I C L E D I A M E T E R • 0 . 5 8 mm • 0 . 9 5 mm O 2 . 18 mm A 4 0 6 mm 0 - 4 0 6 + - 0 . 0 mm A A om 1000 1 0 4 0 1080 1120 A V E R A G E R E A C T O R T E M P E R A T U R E , K 160 FIGURE 9 : EFFECT OF TEMPERATURE ON THE GASIFICATION OF FORESTBURG COAL oo 3.0 to UJ _ J < > O LL. E o _l < o CO CO o Q: CD 2.5 2.0 1.5 1.0 0.5 0 C O A L P A R T I C L E D I A M E T E R O 2 .18 mm S U K U N K A A 4 0 6 mm S U K U N K A V 2 . 1 8 mm C O L E M A N _L 1020 1060 1100 1140 AVERAGE REACTOR TEMPERATURE , K FIGURE 10: EFFECT OF TEMPERATURE ON THE GASIFICATION OF SUKUNKA COAL 100 50 and the poorest 0.41 MJ/m 3(ll Btu/cf) (Table 12). T y p i c a l l y the g a s i f i c a t i o n of Forestburg c o a l produced a gas with a c a l o r i f i c value of 2.0 - 2.5 MJ/m3 (53.7 - 67.1 Btu/cf) T y p i c a l dry gas composition was as f o l l o w s : H 2 : 5.9 - 10.2% CO : 5.9 - 12% CH.: 0.5 - 0.9% 4 C0 2: 8.2 - 14.1% • N 2 : 68.1 - 75.5% The abnormally high nitrogen content of the gas was p a r t l y due to the use of nitrogen i n the feeding system (Figure 3). Assuming that t h i s nitrogen has only a d i l u t i o n e f f e c t , the gas composition and the gross c a l o r i f i c value of the gas have been corrected f o r the i n t r o d u c t i o n of nitrogen (Table VII-5, Appendix V I I ) , which would not be necessary i n an improved design of the feed system. This c o r r e c t i o n increases the concentrations of a l l gases but nitrogen, and the gross c a l o r i f i c value by percentages between 9.6 and 2.6%, depending on the a i r input, and are higher f o r lower a i r input. Under most operating conditions, the c o r r e c t i o n f a c t o r was between 3 and 4%. A l l compositions reported are uncorrected, except f o r those i n Tables VII-5-6, Appendix VII. Sukunka c o a l A t o t a l of seven g a s i f i c a t i o n runs using e i t h e r 2.18 mm or 4.06 mm Sukunka c o a l were c a r r i e d out. (Tables VII-2, Appendix V I I ) . This coal which has a swelling index of 7, could be g a s i f i e d at feed rates of up to 1.52 g/s without agglomeration i n the f l u i d bed (See Photograph 1) or s i n t e r i n g at the f l u i d i z a t i o n g r i d . Higher TABLE 12 TYPICAL RESULTS FOR THE GASIFICATION 0F-4.76 mm FORESTBURG COAL Type of r e s u l t Run No. (a) Average c o a l p a r t i c l e s i z e (mm) Coal feed r a t e (dry basis) (g/s) A i r flow (b) (m3/s) Steam feed rate (g/s) Gas Composition, % (dry Basis) (v/v) Gross gas c a l o r i f i c value (MJ/m 3)( c) H 2 CO CH^ C0 2 N 2 Best 26 4 06 1 .259 6.11xl0~ 3 0. 763 11. 6 10 9 0.6 9 .1 67. 8 2 .94 2 0 53 1 .060 4.68xl0~ 3 0. 377 7. 8 8 0 0.5 8 .2 75. 5 2 .09 3 0 .53 1 .060 3 .03 0. 395 8. 7 5 9 0.7 11 .9 72. 8 2 03 5 0 .95 0 .983 4 .19 0. 542 9. 0 7 9 0.7 11 .4 71. 0 2 30 7 0 .95 0 .983 5 .44 0. 433 7. 7 8 9 0.5 12 .2 70. 7 2 19 13 2 .18 0 .798 5 .48 0. 283 7. 3 8 2 0.5 10 .8 73. 2 2 06 14 2 18 1 .185 6 .51 0. 542 8. 6 7 2 0.7 10 .9 72. 6 2 17 T y p i c a l 15 2 18 1 .185 6 .52 0. 250 7. 2 7 2 0.7 10 .7 74. 2 2 00 16 2 18 1 .273 4 .74 0. 00 5. 9 12 0 0.7 9 .8 71. 6 2 42 19 2 18 1 .673 7 .00 0. 167 6. 4 8 0 0.7 12 .3 72. 5 2 00 22 2 18 2 .369 7 .92 0. 526 10. 2 8 2 0.9 12 .6 68. 1 2 56 25 4 06 1 .223 6 .36 0. 885 10. 1 8 9 0.6 13 .4 67. 0 2 52 27 4 06 1 .516 6 .67 0. 00 7. 7 10 2 0.6 12 .2 69. 3 2 38 28 4 06 1 .516 6 .74 0. 542 8. 5 8 3 0.6 13 .8 68. 8 2 25 29 4 06 1 .516 6 .75 0. 748 8. 7 7. 6 0.7 13 .9 69. 1 2 23 30 4 06 1 .516 6 .76 1. 165 10. 0 6 9 0.7 14 .1 68. 3 2 30 32 -4 76+0.0 1 .059 4 .27 0. 00 6. 3 10 1 0.6 10 .6 72. 4 2 20 33 -4 76+0.0 1 .059 4 .42 0. 433 9. 9 7. 4 0.7 12 .2 69. 8 2 35 Worst 10 2 .18 0.271 2.75xl0~ 3 0. 283 1. 0 1. 8 0.2 11 .5 85. 5 0. 41 (a) See Table VTI-1, Appendix VII f o r d e t a i l s (b) At 294 K and 101.3 kPa (c) At 288.6 K and 101.6 kPa (North-American Gas Standard). PHOTOGRAPH 1: B ed s o l i d s during the g a s i f i c a t i o n of 4.06 mm Sukunka coal (Run No. 40) feed rates were not t r i e d . Agglomeration of the coal i n the feed pipe was experienced i n i t i a l l y , but t h i s problem was solved by the introduction of a small v a r i a b l e speed a g i t a t o r into the feed pipe (Figure 3). For coal feed rates i n the range of 1.46 - 1.58 g/s, a i r flow of 5.43 x 1 0 - 3 std m-Vs a n < i steam feed rates of 0.00 - 0.455 g/s, the Sukunka coal produced a gas with a gross c a l o r i f i c value i n the range of 1.85 - 2.43 MJ/m3 (49.7 - 65.2 Btu/cf) (Table 13). These values are s l i g h t l y lower than those obtained with Forestburg coal i n the same range of operating conditions. Under the. operating conditions described above, t y p i c a l dry com-p o s i t i o n of the gases obtained f o r the g a s i f i c a t i o n of Sunkunka coal was: 4.3 - 6.8% 6.6 - 7.1% 1.0-2.0% 12.1 - 15.7% 70.0 - 74.8% As can be seen from the gas composition, methane contributed an im-portant part of the gas c a l o r i f i c value; and since methane i s mostly produced by p y r o l y s i s , coal feed rate played an important r o l e i n determining the gas q u a l i t y (see Table VII-2, Appendix VII). Gas composition and c a l o r i f i c value ' of the gas produced . from g a s i f i c a t i o n of Sukunka coal corrected f o r the introduction of nitrogen i n the feed,system are presented i n Table VII-6, Appendix VII. Corrections factor were very s i m i l a r , and averaged 6.10%, H2 : CO : CH, : C0 2: N„ : T A B L E 1 3 T Y P I C A L R E S U L T S F O R T H E G A S I F I C A T I O N OF - A . 7 6 mm C A K I N G C O A L S T y p e o f c o a l R u n N o . ( a ) A v e r a g e c o a l p a r t i c l e C o a l f e e d A i r f l o w s t d ( m 3 / s ) S t e a m f e e d G a s c o m p o s i t i o n % v / v ( d r y b a s i s ) G r o s s g a s c a l o r i f i c s i z e (mm) r a t e ( g / s ) r a t e ( g / s ) H 2 CO CHt, C 0 2 N 2 v a l u e ( M J / m 3 ) S u k u n k a 3 6 2 . 1 8 1 . 4 6 2 5 . 5 8 x l 0 - 3 0 . 0 0 4 . 6 7 . 0 1 . 2 1 2 . 4 7 4 . 8 1 . 8 5 S u k u n k a 3 7 2 . 1 8 1 . 4 6 2 5 . 6 9 0 . 4 5 5 5 . 6 6 . 6 1 . 2 1 2 . 3 7 4 . 4 1 . 9 2 S u k u n k a 3 8 _ 2 . 1 8 1 . 5 8 0 5 . 4 3 0 . 0 0 6 . 8 7 . 1 2 . 0 1 2 . 1 7 2 . 0 2 . 4 3 S u k u n k a 3 9 4 . 0 6 1 . 4 6 2 5 . 6 8 0 . 0 0 4 . 3 7 . 1 1 . 0 1 5 . 3 7 2 . 4 1 . 7 5 S u k u n k a 4 0 4 . 0 6 1 . 4 6 2 5 . 7 3 x l 0 - 3 0 . 3 5 8 6 . 2 7 . 0 1 . 0 1 5 . 7 7 0 . 0 1 . 9 7 C o l e m a n 4 1 2 . 1 8 1 . 9 6 8 5 . 7 2 x l 0 - 3 0 . 0 0 7 . 1 7 . 8 1 . 9 1 0 . 9 7 2 . 5 2 . 5 1 ( a ) S e e T a b l e V I I - 2 , A p p e n d i x V I I f o r d e t a i l s higher than f o r the Forestburg coal since introduction of more n i t r o -gen was necessary when operating with caking c o a l . As a consequence, the corrected c a l o r i f i c values of gas produced from Forestburg and Sukunka coal were s i m i l a r . Coleman coal Only one g a s i f i c a t i o n run was c a r r i e d out with Coleman coa l , a caking coal of free swelling index 4. Im.this experiment, 2.18 mm coal was fed to the reactor at a rate of 1.97 g/s. A i r flow was 5.72 x 1 0 - 3 m3/s and steam feed rate was zero. (See Table VII-2, Appendix VII). Under these conditions, steady g a s i f i c a t i o n pro-ceeded without apparent problems producing a gas of 2.51 MJ/m3 (67.37 Btu/cf) and the composition shown i n Table 13.. However, a f t e r approximately 2 hours of operation, the pressure i n s i d e the reactor started to b u i l d up, and the operation was shut down. When a i r was used to cool down the reactor, a f i r e started i n the upper d i s -engaging section which resulted i n the destruction of that section. Dismantling of the reactor showed that the coal had caked (See Photograph 2) and bridged across the bed near the feed point, causing almost complete blockage of the upper part of the reactor and the consequent pressure build-up. Analysis of the data obtained under these conditions reveals that the gas c a l o r i f i c value was i n the range obtained f o r s i m i l a r oper-ating conditions for the two previously tested coals (see runs 21 and 38, Appendix V I I ) , and that an important portion (28.5%) of the gas c a l o r i f i c value was produced by methane. Since i n the present experimental set-up the only p o s s i b i l i t y of achieving trouble-free 56 PHOTOGRAPH 2: Caked Coleman coal at end of Run No. 41 operation with Coleman coal would require decreasing the coal feed rate with the consequent decrease of methane production and gas qu a l i t y , i t was decided not to pursue any further experiments with t h i s type of caking c o a l . E f f e c t of Coal P a r t i c l e Size The gross c a l o r i f i c value of the gas obtained from narrowly sized Forestburg coal under a wide range of operating conditions i s plotted as a function of the a i r to coal r a t i o - (dry basis) i n Figure 11. Com-parison of the curves obtained f o r each coal s i z e suggests that i n . general, an improvement on the gas qu a l i t y may be expected as the coal p a r t i c l e s i z e increases. However, the data i s somewhat scattered and for c e r t a i n operating conditions i t was possible to obtain gas of si m i l a r c a l o r i f i c values i r r e s p e c t i v e of the coal s i z e . Table 14 i l -l u s t r a t e s that g a s i f i c a t i o n of Forestburg coal of four d i f f e r e n t s i z e s at a i r to coal r a t i o s of 5.0 - 5.3 produced gas of a c a l o r i f i c value i n the range of 1.94 - 2.30 MJ/m3 and that there was no c o r r e l a t i o n with the coal s i z e . When coal, of a wider s i z e d i s t r i b u t i o n (-4.76 mm + 0.00 mm) was g a s i f i e d under the above conditions, a s i m i l a r gas was obtained (Run 33, Table 14). Note that although not a l l the experi-ments compared i n Table 14 were performed at the same steam to coal r a t i o , t h i s operating parameter did not have a substantial e f f e c t ori • the gas heating value as. discussed l a t e r . L i t t l e e f f e c t of coal s i z e on gas q u a l i t y was observed for Sukunka coal (Table 15). At a i r to coal r a t i o s of 4.6 r 4.7 g a s i f i c a t i o n of 2.18 mm and 4.06 mm Sukunka coal gave s i m i l a r r e s u l t s . 58 3.0 ro E 2.5 -. 2 .0 UJ 13 _ J < > O U _ E o < o CO CO o CD 1 . 5 1 . 0 0.5 0 FORESTBURG COAL STEAM/COAL - 0.30-1.8 w/w T Q V - I 013 - I 165 K BED DEPTH = 0.51-0.86 m COAL PARTICLE DIAMETER A 0. 53 mm _ • 0.95 mm O 2.18 mm . r? 4.06 mm • -4.76+-0.0 mm 2 4 6 8 10 12 14 A I R TO C O A L R A T I O , w / w ( D R Y B A S I S ) FIGURE 11: EFFECT OF AIR TO COAL RATIO AND PARTICLE SIZE ON THE CALORIFIC VALUE .OF GAS PRODUCED FROM FORESTBURG COAL TABLE 14 EFFECT OF COAL PARTICLE SIZE FOR THE GASIFICATION OF FORESTBURG COAL Run // (a) Nominal average s i z e (mm) A i r to coal r a t i o (dry basis) (w/w) Steam to dry coal r a t i o (w/w) Gas Composition (% v'/v) (dry basis) Gas Gross heating value (MJ/m3) H2 CO CH, 4 co 2 N2 2 ( b ) 0.53 5.29 0.436 7.8 8.0 0.5 8.2 75.5 2.09 5 ( b ) 0.95 5.10 0.632 9.0 7.9 0.7 11.4 71.0 2.30 19 2.18 5.01 0.436 6.4 8.0 0.7 12.3 72.5 2.00 12 2.18 5.25 0.698 7.3 6.9 0.6 10.2 75.0 1.94 28 4.06 5.33 0.690 8.5 8.3 0.6 13.8 68.8 2.25 33 -4.76+0.0 5.01 0.732 9.9 7.4 0.7 12.2 69.8 2.35 (a) See Table VII-1, Appendix VII for d e t a i l s (b) Unstable operation. TABLE 15 EFFECT OF COAL PARTICLE SIZE FOR THE GASIFICATION OF SUKUNKA COAL Run #.. (a) Nominal average s i z e (mm) A i r to coal r a t i o (dry basis) (w/ w) Steam to dry coal r a t i o (w.w) Gas Composition (T (dry basis) v/v) Gas Gross c a l o r i f i c Value (MJ/m3) H2 CO CH, 4 co 2 N2 36 2.18 4.58 0.020 4.6 7.0 1.2 12.4 74.8 1.85 39 4.06 4.66 0.021 4.3 7.1 1.0 15.3 72.4 1.75 37 2.18 4.67 0.331 5.6 6.6 1.2 12.3 74.4 1.92 40 4.06 4.71 0.261 6.2 7.0 1.0 15.7 70.0 1.97 (a) see Table VII-2, Appendix VII for d e t a i l s . A l t h o u g h c o a l p a r t i c l e s i z e d i d n o t h a v e a p a r t i c u l a r l y s t r o n g e f f e c t o n t h e g a s q u a l i t y , i t h a d a n i m p a c t i n t h e o p e r a t i o n o f t h e g a s i f i e r . W h e n g a s i f y i n g 0 . 5 3 mm a n d ' 0 . , 9 5 mm F o r e s t b u r g c o a l , t h e r e -a c t o r w a s d i f f i c u l t t o o p e r a t e u n d e r s t a b l e c o n d i t i o n s . I n s p i t e o f a d j u s t m e n t o f t h e o p e r a t i n g v a r i a b l e s , i n m o s t r u n s ( S e e T a b l e V I I - 1 , A p p e n d i x V I I ) t h e t e m p e r a t u r e o f t h e r e a c t o r d r i f t e d u p w a r d s o r d o w n -w a r d s a t a r a t e i n e x c e s s o f 1 K / m i n . A s t h e c o a l s i z e i n c r e a s e d t h e o p e r a t i o n b e c a m e m o r e s t a b l e , a n d s t e a d y g a s i f i c a t i o n w i t h 2 . 1 8 mm a n d 4 . 0 6 mm c o a l w a s p o s s i b l e . G a s i f i c a t i o n o f - 4 . 7 6 mm + 0 mm F o r e s t b u r g c o a l a l s o w e n t s m o o t h l y , w h i c h s u g g e s t s t h a t l a r g e r c o a l p a r t i c l e s h a v e a b u f f e r i n g c a p a c i t y t o a b s o r b s o m e v a r i a t i o n s i n t h e o p e r a t i n g c o n -d i t i o n s . E f f e c t o f A i r t o C o a l R a t i o T h e m o s t i m p o r t a n t p a r a m e t e r i n d e t e r m i n i n g t h e g a s c a l o r i f i c v a l u e a p p e a r e d t o b e t h e a i r t o c o a l r a t i o . F o r f o u r d i f f e r e n t s i z e s o f F o r e s t b u r g c o a l a l i n e a r c o r r e l a t i o n b e t w e e n t h e g a s c a l o r i f i c v a l u e a n d t h e a i r t o c o a l r a t i o w a s f o u n d ( F i g u r e 1 1 ) . I n s p i t e o f s o m e s c a t t e r a n d t h e f a c t t h a t t h e d a t a c o v e r d i f f e r e n t t e m p e r a t u r e s a n d s t e a m t o c o a l r a t i o s , a d e f i n i t e t r e n d e x i s t s t o w a r d s i n c r e a s e d g a s c a l o r i f i c v a l u e w i t h d e c r e a s i n g a i r t o c o a l r a t i o s . S i m i l a r c o r r e -( 3 4 ) l a t i o n s h a v e b e e n f o u n d b y o t h e r a u t h o r s i n t h e g a s i f i c a t i o n o f c o a l a n d w o o d . I n s u f f i c i e n t d a t a w a s g e n e r a t e d t o d e f i n e s u c h a c o r r e l a t i o n f o r t h e g a s i f i c a t i o n o f S u k u n k a c o a l , b u t t h e s a m e g e n e r a l t r e n d c a n b e o b s e r v e d i n F i g u r e 1 2 . 62 3.0 2.5 ro " 2.0 UJ ZD -J o |.5  LL. cr o _i < o cn CO O cr 0.5 C O A L P A R T I C L E D I A M E T E R O 2 . 1 8 mm 4 . 0 6 mm 2.18 mm F O R E S T B U R G C O A L S U K U N K A C Q A L S T E A M / C O A L = 0 . 0 2 0 - 0 . 9 5 w/w T E M P E R A T U R E = 1 081 - 1 171 K B E D D E P T H = 0 . 6 8 - 0 . 7 6 m 1 1 1 l 2 4 6 8 10 12 AIR TO COAL RATIO, w/w (DRY BASIS) 14 FIGURE 12: EFFECT OF AIR TO COAL RATIO ON THE GROSS CALORIFIC VALUE OF GAS PRODUCED FROM SUKUNKA COAL The e f f e c t of the a i r to coal r a t i o on the gas composition i s also of i n t e r e s t . Since the a i r to coal r a t i o can be changed by changing the coal feed rate, the a i r flow or both parameters, these e f f e c t s are considered i n d i v i d u a l l y . Since i n coal g a s i f i c a t i o n at atmospheric pressure the amount of methane formed by re a c t i o n of C and n i s i n s i g n i f i c a n t at present oper-ating temperatures and pressure, most of the methane produced i s expected to be a product of p y r o l y s i s , and therefore a d i r e c t function of the coal feed rate. Indeed, g a s i f i c a t i o n of Forestburg coal showed that methane concentration i n the gas increased with coal feed rate (Figure 13). The data i n Figure 13 are scattered because not a l l other operating conditions are constant. However, i f the methane production i s calculated from the t o t a l volume of gas produced (see Table VII-7, Appendix VII) and plotted as a function of the coal feed rate, a good l i n e a r c o r r e l a t i o n i s obtained '(Figure 14). The slope of t h i s l i n e suggests that the methane y i e l d of Forestburg coal i s 2.5%. A r e l i a b l e c o r r e l a t i o n could not be obtained f o r the g a s i f i c a t i o n of caking coals due to the r e s t r i c t e d number of experi-ments performed, but the a v a i l a b l e data (Table VII-8, Appendix VII) suggests a s i m i l a r trend of increased methane production with coal feed rate, and also a higher methane y i e l d f o r the Sukunka and Coleman coals (Figure 15). Since steam feed rate may have an e f f e c t on the production of hydro-gen and carbon monoxide, the e f f e c t of coal feed rate on the production of these gases should be studied at a constant steam to coal r a t i o (Table VII-9, Appendix VII). Figures 16.and.17 r e s p e c t i v e l y show that the hydrogen and carbon monoxide production rates from Forestburg coal g a s i f i c a t i o n at a steam to coal r a t i o of 0.32 - 0.34 are c l e a r l y depend-64 FIGURE 13: EFFECT OF COAL FEED RATE ON THE CONCENTRATION OF METHANE/PRODUCED. FROM.FORESTBURG. COAL FIGURE 14: EFFECT OF COAL FEED RATE ON THE PRODUCTION OF METHANE FROM FORESTBURG COAL 66 100 £ < x i -LU 80 O 60 I— o a o £ 40 UJ 20 1 1 1 1 PARTICLE DIAMETER - O 2.18 mm SUKUNKA V 4.06 mm SUKUNKA 0 2.18 mm COLEMAN ° 0 — 8 — O V 1 1 1 1 0.4 0.8 1.2 1.6 DRY COAL F E E D R A T E , g/s 2.0 FIGURE 15: EFFECT OF COAL FEED RATE ON THE METHANE PRODUCTION FROM CAKING COALS ent on the coal feed rate. From the slope of these curves, the hydrogen y i e l d from Forestburg coal i s calculated as 3.37% (Figure 16) which com-pares with a hydrogen content of 4.25% i n the coal (Table 5). By a sim-i l a r a n a l y s i s , the carbon, monoxide y i e l d under these conditions i s c a l -culated at 52.6% (Figure 17). Operating experience indicated that the e f f e c t of increasing the a i r flow into the reactor while keeping the coal and steam feed rates constant was to decrease both the hydrogen and carbon monoxide content of the gas while the carbon dioxide concentration increased. These fin d i n g s are i n l i n e with r e s u l t s reported elsewhere and are i l l u s t r a t e d i n Figure 18 f o r the g a s i f i c a t i o n of Forestburg coal at a feed rate of 1.27 g/s and steam to coal r a t i o of 0.33. However, as discussed e a r l i e r i n t h i s chapter, a l t e r i n g the a i r flow rate while keeping the coal feed rate constant introduces a c e r t a i n measure of i n s t a b i l i t y into the operation of the reactor. The combined e f f e c t on the gas composition of changing both the coal feed rate and the a i r flow rate at a constant steam to coal r a t i o i s shown i n Figures 19 and 20,for Forestburg and Sukunka coal r e s p e c t i v e l y . Increasing i n such a manner the a i r to coal r a t i o resulted i n a decrease of the hydrogen and methane and an increase i n the carbon dioxide content of the gas from both coals while the carbon monoxide content did not change appreciably. The o v e r a l l r e s u l t , as discussed e a r l i e r , (Figures 11 and 12) was then a decrease of the gas c a l o r i f i c value.-68 0 I 1 1 1 1 1 0 0.5 1.0 1.5 2.0 2.5 C O A L F E E D R A T E , g / s FIGURE 16: EFFECT OF COAL FEED RATE ON THE PRODUCTION OF HYDROGEN FROM FORESTBURG COAL 69 FIGURE 17: EFFECT OF COAL FEED RATE ON THE PRODUCTION OF CARBON MONOXIDE FROM FORESTBURG COAL 70 1.4.0 12.0 10.0 co CO < CD > o r o > > 8.0 o P 6.0 CO o CL O ° 4.0 co < RUN No. • 16 • 17 COAL FEED RATE = 1.27 g/s STEAM /COAL RATIO = 0. 33 2.0 CH. 0.005 0.006 AIR FLOW , std ( m 3 / s ) 0.007 FIGURE 18: EFFECT OF AIR FLOW ON THE GAS COMPOSITION FROM FORESTBURG COAL AT A FIXED COAL FEED RATE 71 CO CO < m >-cn o > > o CO o CL o o CO < 14.0 12.0 h 10.0 \-8.0 6.0 4.0 2.0 3.5 CO' H. RUN No. • 21 • 17 STEAM/COAL RATIO =0.33 CH, 4.0 4.5 5.0 5.5 AIR TO COAL RATIO , w / w 6.0 FIGURE 19: EFFECT OF AIR TO COAL RATIO ON THE GAS COMPOSITION FROM FORESTBURG COAL 14 10 8 0 1 1 7 • co2 • - • — -RUN N2 0 38 (2.18 mm COAL) • 36 (2.18 mm COAL) • 39 ( 4 0 6 mm COAL) STEAM/COAL RATIO =0.02 CO _ n • - -• - • -. CH4 i 1 i 4.0 4.4 4.8 AIR TO COAL RATIO , w/w FIGURE 20: EFFECT OF THE AIR TO COAL RATIO ON THE GAS COMPOSITION FROM SUKUNKA COAL 7 3 E f f e c t o f t h e S t e a m t o C o a l R a t i o T h e e f f e c t o f i n c r e a s i n g t h e s t e a m t o c o a l r a t i o w a s s t u d i e d i n R u n s 27 t o 3 0 u s i n g 4 . 0 6 mm d i a m e t e r F o r e s t b u r g c o a l . O t h e r o p e r a t i n g p a r a m e t e r s w e r e , k e p t c o n s t a n t . T h e r e a c t o r t e m p e r a t u r e d e c r e a s e d a s a c o n s e q u e n c e o f t h e i n c r e a s e d s t e a m i n p u t f r o m 1 1 1 5 K a t a s t e a m t o c o a l r a t i o o f 0 . 3 3 t o 1 0 5 9 K a t a s t e a m t o c o a l r a t i o o f 1 . 1 0 ( T a b l e 1 6 ) . . A s d i s c u s s e d e a r l i e r , a t e m p e r a t u r e c h a n g e i t s e l f d o e s n o t a f f e c t t h e g a s c o m p o s i t i o n . A l t h o u g h i n c r e a s i n g t h e s t e a m t o c o a l r a t i o d i d n o t a p p r e c i a b l y c h a n g e t h e g a s c a l o r i f i c v a l u e w h i c h r e m a i n e d i n t h e r a n g e o f 2 . 2 3 - 2 . 3 8 M J / m 3 ( T a b l e 1 6 ) , t h e r e w a s a m a r k e d e f f e c t o n t h e g a s c o m p o s i t i o n . F i g u r e 2 1 s h o w s t h a t t h e h y d r o g e n a n d c a r b o n d i o x i d e c o n -t e n t o f t h e g a s i n c r e a s e d w h i l e t h e c a r b o n m o n o x i d e c o n t e n t d e c r e a s e d . M e t h a n e c o n c e n t r a t i o n r e m a i n e d f a i r l y c o n s t a n t . T h e s a m e e f f e c t c o u l d b e o b s e r v e d f o r o t h e r o p e r a t i n g c o n d i t i o n s w i t h 2 . 1 8 mm F o r e s t b u r g c o a l ( c o m p a r e f o r e x a m p l e r u n s 1 4 a n d 1 5 a n d r u n s 1 7 a n d 1 8 , T a b l e V I I - 1 , A p p e n d i x V I I ) , a n d f o r t w o d i f f e r e n t s i z e s o f S u k u n k a c o a l ( T a b l e 1 7 , F i g u r e 2 2 ) . T h e s e r e s u l t s s u g g e s t t h a t u n d e r t h e p r e v a i l i n g o p e r a t i n g c o n d i t i o n s i n c r e a s e d c o n c e n t r a t i o n o f w a t e r i n t h e r e a c t o r f a v o u r s t h e r e a c t i o n CO + H 2 0 C 0 2 + H 2 r a t h e r t h a n t h e c a r b o n s t e a m r e a c t i o n * S t e a m r a t i o = r a t i o o f t o t a l w a t e r f e d : . : ( . i ; > e . s t e a m + H 2 0 i n c o a l + H O i n a i r ) t o c o a l f e d . TABLE 16 EFFECT OF THE STEAM TO COAL RATIO IN THE GASIFICATION OF FORESTBURG COAL Run # Steam to Coal Ratio (w/w) Average Reactor Temperature (K) Gas Composition (dry basis) % (v/v) Gross Calorific Value (MJ/m3) H 2 CO CH„ C02 N 2 27 0.332 1115 7.7 10.2 0.6 12.2 69.3 2.38 28 0.690 1085 8.5 8.3 0.6 13.8 68.8 2.25 29 0.826 1079 8.7 7.6 0.7 13.9 •69.1 2.23 30 1.101 1059 10.0 6.9 0.7 14.1 68.3 2.30 Coal particle size: Coal feed rate: Air flow: Expanded bed depth 4.06 mm. 1.516 g/s (6.67-6.71) x l O - 3 std. m 3/i 0.86 m 75 F I G U R E 2 1 : E F F E C T OF T H E S T E A M TO C O A L R A T I O ON. T H E G A S C O M P O S I T I O N F R O M 4 . 0 6 * ; . i r a n ' ' D I A M E T E R F O R E S T B U R G C O A L TABLE 17 EFFECT OF THE STEAM TO COAL RATIO IN THE GASIFICATION OF SUKUNKA COAL Run # Steam to Coal Ratio Average Reactor Gas Composition % (v/v) (dry basis) Gross C a l o r i f i c Value (w/w) Temperature (K) H 2 CO CH 4 C0 2 N 2 (MJ/m3) 36 0.020 1132 4.6 7.0 1.2 12.4 74.8 1.85 37 0.331 1081 5.6 6.6 1.2 12.3 74.4 1.92 39 0.021 1160 4.3 7.1 1.0 15.3 72.4 1.75 40 0.261 1107 6.2 7.0 1.0 15.7 70.0 1.97 Coal p a r t i c l e s i z e : Runs 36 & 37 = 2 18 mm, Runs 39 & 40 = 4.06 mm Coal feed rate: 1.46 g/s A i r flow: (5.58-5.73) x 10* 3 std m3/s Expanded bed depth: 0. 76 m 77 _ 14 C O C O < m 12 >-CC O > > C O o CL o o CO < 10 8 » — CO-co 2 • S U K U N K A C O A L 2 .18 mm 4 . 0 6 mm A I R / C O A L = 4 . 6 - 4 . 7 w/w T E M P E R A T U R E = 1 0 8 1 - 1 160 K C H , C H , 0.0 0.2 0.4 STEAM TO COAL RATIO, w/w FIGURE 22: EFFECT OF THE STEAM TO COAL RATIO ON THE GAS COMPOSITION FROM SUKUNKA COAL E f f e c t of Bed Depth E f f e c t s of the bed depth on the gas q u a l i t y were masked by changes i n other operating v a r i a b l e s . However, inspection of the data (Table VII-1, Appendix VII) indicates that no noticeable improvement on the gas q u a l i t y was obtained when the bed height varied from 0.5 to 0.86 m. This i s i n agreement with experimental r e s u l t s reported for the g a s i f i c a t i o n of coal and chars at 1273 K . However, a deeper bed made the operation smoother and easier to c o n t r o l . Carbon Content of the Bed The carbon content of the bed i s not an operating v a r i a b l e i n i t s e l f , but rather a r e s u l t of other operating conditions. The e f f e c t of coal r e a c t i v i t y on the bed carbon content was discussed e a r l i e r i n t h i s Chapter. In general, the average carbon bed content remained below 5% fo r stable g a s i f i c a t i o n of Forestburg co a l , (Table 18) and was not cor-rel a t e d with either the a i r to coal r a t i o (for conditions not exceeding the coal r e a c t i v i t y l i m i t ) or the temperature of the reactor. Also, there appears to be no c o r r e l a t i o n between the q u a l i t y of the gas pro-duced and the bed carbon content. Under equivalent operating conditions, the carbon content of a bed ga s i f y i n g Sukunka coal was higher (Table.19) than f o r Forestburg coal (Table 18). This i s probably due to a di f f e r e n c e i n coal r e a c t i v i t y . Solids and Carbon E l u t r i a t i o n Table 20 shows the s i z e d i s t r i b u t i o n of the s o l i d s e l u t r i a t e d from the bed for t y p i c a l g a s i f i c a t i o n conditions of 2.18 mm diameter Forestburg c o a l . Approximately 95% by weight of the s o l i d s had a s i z e below 355 :V®[' \ -79 TABLE 18 AVERAGE BED CARBON CONTENT DURING THE GASIFICATION OF FORESTBURG COAL R u n # A v e r a g e C o a l P a r t i c l e S i z e (mm) A i r t o C o a l . R a t i o ( w / w ) A v e r a g e R e a c t o r T e m p e r a t u r e ( K ) A v e r a g e B e d C a r b o n C o n t e n t (%) 1 0 . 5 3 5 . 3 9 1 0 3 3 0 . 4 4 5 0 . 9 5 5 . 1 0 1 1 2 3 - 1 0 8 8 2 . 2 9 8 0 . 9 5 6 . 6 5 1 1 5 6 0 . 5 0 9 0 . 9 5 9 . 0 4 1 1 7 2 0 . 2 0 1 0 2 . 1 8 1 2 . 1 6 1 0 8 8 2 . 7 1 1 1 2 . 1 8 7 . 6 2 1 0 5 0 2 . 9 0 1 2 2 . 1 8 5 . 2 5 1 0 4 0 4 . 72 1 3 2 . 1 8 8 . 2 3 1 0 6 1 3 . 9 6 .14 2 . 1 8 6 . 5 9 1 0 3 1 1 1 . 7 1 1 5 2 . 1 8 6 . 5 9 1 0 3 1 9 . 8 9 1 6 2 . 1 8 4 . 4 6 1 1 0 7 2 . 4 7 17 2 . 1 8 5 . 7 0 1 1 6 5 0 . 9 7 1 8 2 . 1 8 5 . 9 4 1 1 4 5 0 . 7 4 1 9 2 . 1 8 5 . 0 1 1 0 4 9 1 0 . 3 6 2 1 2 . 1 8 3 . 6 6 1 0 8 0 2 . 6 9 2 2 2 . 1 8 4 . 0 1 1 0 2 6 1 6 . 6 - 3 1 . 3 2 3 4 . 0 6 1 0 . -36 1 1 0 8 1 . 9 5 2 4 4 . 0 6 6 . 9 7 1 0 9 6 4 . 4 4 2 5 4 . 0 6 6 . 2 3 1 0 7 4 4 . 4 3 2 6 4 . 0 6 6 . 1 0 1 0 6 5 1 0 . 6 7 27 4 . 0 6 5 . 2 8 1 1 1 5 4 . 2 1 2 8 4 . 0 6 5 . 3 3 1 0 8 5 4 . 3 1 2 9 4 . 0 6 5 . 3 2 1 0 7 9 3 . 6 1 3 0 4 . 0 6 5 . 3 4 1 1 5 3 4 . 0 1 3 1 4 . 0 6 5 . 6 8 1 0 7 8 8 . 4 5 3 2 4 . 7 6 - 0 . 0 4 . 8 7 1 1 2 5 1 . 5 2 3 3 4 . 7 6 - 0 . 0 5 . 0 1 . 1 0 5 5 1 . 9 1 8 0 TABLE 19 AVERAGE BED CARBON CONTENT DURING THE GASIFICATION OF SUKUNKA COAL R u n # A v e r a g e C o a l P a r t i c l e S i z e (nm) A i r t o C o a l R a t i o ( w / w ) A v e r a g e R e a c t o r T e m p e r a t u r e ( K ) A v e r a g e B e d C a r b o n C o n t e n t (%) 3 4 2 . 1 8 8 . 1 2 1 1 5 3 2 . 6 8 3 5 2 . 1 8 1 0 . 8 6 1 1 7 1 2 1 . 4 3 6 2 . 1 8 4 . 5 8 1 1 3 2 2 1 . 7 37 2 . 1 8 4 . 6 7 1 0 8 1 2 7 . 2 38 2 . 1 8 - 4 . 1 1 1 1 6 5 1 8 . 3 4 3 9 4 . 0 6 4 . 6 6 1 1 6 0 1 6 . 2 4 0 4 . 0 6 4 . 7 1 1 1 0 7 1 6 . 2 TABLE 20 TYPICAL SIZE DISTRIBUTION OF SOLIDS ELUTRIATED DURING THE GASIFICATION OF 2.18 mm FORESTBURG COAL (RUN NO. 20) P a r t i c l e Size (ym) Weight Percentage (%) + 841 0.53 - 841--. 425 3.62 - 425-+ 355 1.09 - 355 + 177 8.36 - 177 + 150 6.43 - 150 + 125 3.35 - 125 + 72 8.99 - 72 + 63 6.98 - 63 + 44 9.14 - 44 51.49 82 This is in good agreement with theoretical predictions where the maximum size of a solid particle elutriated is calculated from the particle terminal velocity at the prevailing operating conditions ( 3P). From these calculations, the maximum diameter of an elutriated particle would be 727 ym' for a charcoal particle and 130 y m • for a s i l i c a particle. Solids elutriation rates were expressed in terms of g of solids per g of dry coal fed into the reactor. Carbon elutriation:: rates were calcu-lated from the solids elutriation rate and the carbon content of the solids elutriated. Although the data i s somewhat scattered due to the fact that in some cases elutriation rates were measured as averages of several runs, in general, elutriation rates for narrowly sized forest-burg coal were in the range of 0.078-0.108 g/g coal, and seemed to i n -crease with the average superficial velocity in the reactor and were largest for the smallest coal particle size (Table 21). When Forestburg coal sample of a wider size distribution which included fines (-4.76 +0.0 mm) was gasified, elutriation of solids sharply increased to 0.187 g/g coal (runs 32 and 33, Table 21). The same trends with particle size and fluidization velocity were followed by carbon content of the solids elutriated as with total solids elutriated. As a consequence the carbon elutriation - rates (Table 21) increased with superficial velocity and decreased with increasing particle size. Solids and carbon elutriation rates measured for the gasification of Sukunka coal were higher by a factor of 3 than for Forestburg coal, while for the run with Coleman coal they were intermediate between the two (Table 22). These differences existed in spite of the fact that coal ash content of a l l samples was similar (see Table 6) and may be an indication of the differences in coal reactivity. Indeed, during the TABLE 21 SOLIDS AND CARBON ELUTRIATION DURING THE GASIFICATION OF FORESTBURG COAL Run 9 Average Coal P a r t i c l e Size (mm) Weighted Average A i r V e l o c i t y a t Reactor's Temp.(a) (m/s) Solids E l u t r i a t i o n Rate (g/g coal) E l u t r i a t e d Solids Carbon Content (%) Carbon E l u t r i a t i o n Rate (g/g coal) 1 & 4 0.53 1.10 0.167 62.31 0.104 2 & 3 0.53 0.80 0.102 55.14 0.056 5-9 0.95 1.15 0.078 29.50 0.023 10,12 & 13 2.18 0.79 0.084 36.79 0.031 11 & 19 2.18 1.13 0.086 50.44 0.043 14 & 15 2.18 1.25 0.093 57.99 0.054 16,17,18 & 21 2.18 1.22 0.091 29.66 0.027 22 2.18 1.52 0.093 58.48 0.054 23,25 & 31 4.06 1.21 0.083 38.14 0.032 24 4.06 0.93 0.093 30.01 0.028 26 4.06 1.28 0.091 43.35 0.039 27-30 4.06 1.35 0.108 40.96 0.044 32 & 33 -4.76+0.0 0.88 0.187 61.45 0.115 When s o l i d s e l u t r i a t i o n rate corresponds to the average of more than one run the average a i r v e l o c i t y i s the average of the a i r v e l o c i t i e s weighted by the duration of the run. TABLE 22 SOLIDS AND CARBON ELUTRIATION DURING THE GASIFICATION OF CAKING COALS Run # Average Coal P a r t i c l e Size (mm) Weighted Average A i r V e l o c i t y At Reactor's Temp.( a) (m/s) Solids E l u t r i a t i o n Rate (g/g coal) E l u t r i a t e d Solids Carbon Content (%) Carbon E l u t r i a t i o n Rate (g/g coal) 34 2.18 0.80 0.227 81.66 0.185 35 2.18 1.16 0.266 85.17 0.227 36-38 2.18 1.17 0.272 80.33 0.219 30-40 4.06 1.21 0.276 61.88 0.171 41 2.18 1.17 0.139 71.86 0.100 Runs 34-40 used Sukunka Coal Run 41 used Coleman Coal (a) When s o l i d s e l u t r i a t i o n rate corresponds to the average of more than one run, the average a i r v e l o c i t y i s the average of the a i r v e l o c i t i e s weighted by the duration of the run. g a s i f i c a t i o n of Sukunka coa l , the bed had a much higher carbon icontent (and thus probably a higher content, of fine, and l i g h t p a r t i c l e s ) than ' f o r Forestburg c o a l . Therefore, the e l u t r i a t i o n rates w i l l be also higher with the former c o a l . Tar Production Tar content measurements i n the gas on some selected runs (Table 23) showed that the tar produced during g a s i f i c a t i o n was notably higher f o r the caking coals. Tar production was highest from Sukunka coa l , followed by Coleman coal while Forestburg.coal yielded the lowest amount of t a r . These measurements also showed a strong e f f e c t of the p a r t i c l e s i z e with the smaller 2.18 mm coal producing as much as three times more tar than 4.06 mm c o a l . This l a s t e f f e c t was v e r i f i e d f o r both the Forestburg and Sukunka c o a l . Unexpectedly, a Forestburg coal sample including f i n e s . (-4.76 mm+ 0.0 mm) (see runs 31 and 32) and having a nominal average s i z e of -1.03 mm yielded an amount of tar intermediate between the .2.18 mm and 4.06 mm Forestburg c o a l . Tar l e v e l s are also dependent on the coal feed l o c a t i o n . If the coal had been fed into the bottom of the bed, lower tar l e v e l s would be expected f o r both coals. TABLE 23 GAS TAR CONTENT FOR THE GASIFICATION OF SOME WESTERN CANADIAN COALS Run # Coal Type Coal P a r t i c l e Size . (mm) /.Dry Coal Feed Rate (g/s) A i r Flow Std. (m3/s) Steam Feed Rate (g/s) Average Reactor Temperature (K) Gas Tar Content (g/m3) Tar Production (mg/g coal). 14 & 15 Forestburg 2.18 1.185 6.52 xd'O"3 0.25-0.54 1031 0.841 5.17 24 Forestburg 4.06 0.782 4.53 0.549 1096 0.274 1.80 32 & 33 Forestburg 4.76-0.0 1.059 4.35 . 0.0433 1090 0.434 2.11 37-38 Sukunka 2.18 1.46-1.58 5.57 0.0455 1126 3.307 14.16 39-40 Sukunka 4.06 1.462 5.71 0.0358 1134 0.913 4.25 41 Coleman 2.18 1.968 5.72 x l 0 ~ 3 0.0 1096 2.122 7.22 CHAPTER V. - DISCUSSION OF RESULTS MASS BALANCES In order to v e r i f y the experimental data and gain some in s i g h t into the g a s i f i c a t i o n reactions, o v e r a l l and d e t a i l e d mass balances were c a r r i e d out for 38 g a s i f i c a t i o n runs. The method followed i n doing such balances and a de t a i l e d sample c a l c u l a t i o n are found on Appendix VIII. Results are summarized i n Tables VIII-1 to VIII-4, Appendix VIII. O v e r a l l Mass Balances The o v e r a l l mass balances (Table VIII-1, Appendix VIII) showed that i n general the outputs were short of the inputs by 2-20%, but i n some cases the outputs were larger than the inputs. O v e r a l l , the average discrepancy i n the mass balances was .-- 4.7 + 9.2%, which i s within acceptable margins of experimental erro r . Inaccuracies introduced i n the mass balances which contributed to the imbalance between outputs and inputs are discussed below i n order of importance. 1. A l l mass balances were c a r r i e d out at average operating con-d i t i o n s (Table VII-1 and VII-2, Appendix VII) and therefore do not consider v a r i a t i o n s of parameters such as a i r flow, reactor tempera-ture, and gas composition, which a c t u a l l y occurred during a run. Calculations done at s p e c i f i c times during a run yielded better balances, but are not included i n t h i s thesis f o r the sake of bre v i t y . 2. The water output of the reactor was measured i n d i r e c t l y , by difference between the measured wet gas flow and the calculated dry gas flow (see Appendix VIII). Some degree of inaccuracy i s introduced here since i n most g a s i f i c a t i o n runs the bed was vigorously agitated or slugging, which caused considerable o s c i l l a t i o n s of the water manometer measuring the pressure drop across the o r i f i c e p l ate used to determine the wet gas flow. 3. Inert ash p a r t i c l e s having a terminal v e l o c i t y equal to or higher than the p r e v a i l i n g gas s u p e r f i c i a l v e l o c i t y w i l l not be e l u t r i a t e d from the reactor u n t i l such time that t h e i r s i z e i s s u f f i c i e n t l y reduced by a t t r i t i o n . Therefore, at any given time there i s l i k e l y to be a net accumulation of ash i n the reactor. This was confirmed by the ash balances (Table VIII-4, Appendix VIII) which i n v a r i a b l y show an ash input larger than the ash output. This accumulation term (which w i l l increase with the coal p a r t i c l e s i z e and decrease with the a i r flow) i s not considered i n the o v e r a l l mass balances shown on Table VIII-1, Appendix VIII, and causes a s h o r t f a l l i n the t o t a l mass output. 4. In most cases, the reported s o l i d s e l u t r i a t i o n rates represent averages of 2 - 4 g a s i f i c a t i o n runs c a r r i e d out under r e l a t i v e l y s i m i l a r operating conditions. Although some inaccuracy i s i n t r o -duced by t h i s procedure, the contribution of the e l u t r i a t e d s o l i d s to the t o t a l mass output i s below 5%. 5. Although tar measurements were made only for a few selected runs, these showed that the tar contribution to the t o t a l mass output i s below 1% and therefore of l i t t l e consequence i n the over-a l l mass balances. Water Balance Water balances (Table VIII-1, Appendix VIII) indicated that i n only 12 of 38 g a s i f i c a t i o n runs the water output was lower.than the water input. Of these, four corresponded to obvious inaccuracies i n the mass balances since the wet gas flow measured was lower than the calculated dry gas flow. The balance of these 12 runs corresponded to a v a r i e t y of operating conditions and i t did not appear to be a s p e c i f i c set of conditions that favored water dissapearance i n the reactor. I t i s therefore safe to say that i n general water was formed i n the reactor. Since the experimental data provides evidence of the reaction CO + H 20 ^ C0 2 + H 2 .occurring (see Chapter'IV, e f f e c t of steam) the'observed excess water can only be explained by assuming that the other water consuming reaction: C + H 20 v ^ CO + H 2 does not proceed to any appreciable extent under the p r e v a i l i n g op-erating conditions, and that water i s formed through the rea c t i o n of oxygen and hydrogen and through the combustion of hydrocarbons at higher rates than i t disappears. This i s supported by both the k i n e t i c s and thermodynamics of the p r i n c i p a l reactions involved i n a g a s i f i c a t i o n system (see Chapter I ) . Hydrogen Balance An Important part of the hydrogen entering and leaving the reactor i s water bound, and therefore t o t a l hydrogen mass balances (Table VIII-2, Appendix VIII) are very s e n s i t i v e to inaccuracies i n the water determination discussed e a r l i e r . Accordingly, an imbalance between t o t a l hydrogen input and output was generally obtained. Note however, that the impact of such imbalances i n the o v e r a l l mass balance i s minor. On the other hand, the amount of hydrogen entering the reactor with the dry coal and the amount of hydrogen leaving the reactor with the dry gas are subject to le s s error since they only depend on coal and gas analyses and measure-ment of the coal feed rate and dry gas flow out of the reactor. Comparison of these two quantities (Table VIII-2, Appendix VIII) shows that i n general, the hydrogen i n the gas i s s i m i l a r to or lower than the hydrogen present i n the dry coal feed. This again suggests that the formation of hydrogen through the carbon steam reaction C + H 2 0 ~ — ^ H 2 + CO i s not very important, and that hydrogen i s a v a i l a b l e f or the formation of water. Oxygen Balance A t o t a l oxygen balance (Table VIII-3, Appendix VIII) i s also very s e n s i t i v e to the inaccuracies i n water input and output de-terminations and i s not of much value i n i n t e r p r e t i n g the present experimental data. However, i n a l l g a s i f i c a t i o n runs, the oxygen gas content was lower than the oxygen entering the reactor with the a i r except i n 8 runs where they were s i m i l a r . In a l l but two runs, the oxygen contained i n the dry gas was lower than the oxygen entering the reactor with the dry coal and the a i r . This provides a d d i t i o n a l evidence that i n general some oxygen goes into the formation of water. Carbon Balance In general, the t o t a l carbon output was equal to or lower than the t o t a l carbon input (Table VIII-4, Appendix VI I I ) . The major inaccuracies introduced i n the carbon balances was i n the deter-mination of the carbon e l u t r i a t i o n rates. ENERGY BALANCES A de t a i l e d energy balance i n the f l u i d i z e d bed reactor would require the knowledge of the amounts of gas produced by each of the p y r o l y s i s , combustion and g a s i f i c a t i o n reactions and consideration of the d i f f e r e n t heats of reaction at the p r e v a i l i n g operating conditions. A d d i t i o n a l l y , a measurement of heat losses from the reactor would be necessary. Since much of t h i s data i s not a v a i l a b l e , and i n order to s i m p l i f y the ana l y s i s , a s i m p l i f i e d energy balance was c a r r i e d out for various operating con-d i t i o n s for the g a s i f i c a t i o n of Forestburg and Sukunka coals. The method considered the reactor as a "black box" and compared the heat content of the inputs (coal, a i r and steam) with the heat content of the outputs (sensible and c a l o r i f i c value of the gas, heat content of e l u t r i a t e d carbon and heat content of steam). The heat losses from the reactor were then calculated by dif f e r e n c e . A detailed'• sample c a l c u l a t i o n i s presented on Appendix IX, while energy balances for selected g a s i f i -c a tion runs are calculated i n Table IX-1, Appendix IX). The energy losses calculated f or 17 g a s i f i c a t i o n runs varied between 10.3% and 37.2% of the t o t a l energy input. Such v a r i a t i o n s are not only due to experimental errors and v a r i a t i o n s i n the mass balances already discussed, but also to the dependence of the percentage heat losses on the temperature of the reactor and on the heat input. The mean energy losses and standard deviation were 25.8% + 8.8%, as would be expected f o r a r e l a t i v e l y small externally insulated f l u i d i z e d bed. The sub-s t a n t i a l energy losses have no doubt an impact i n the q u a l i t y of the gas produced, and on the e f f i c i e n c y of the system since an important portion of the energy generated by the combustion of carbon, hydrocarbons and hydrogen w i l l be l o s t to the environment instead of being a v a i l a b l e f o r the endothermic g a s i f i c a t i o n reactions. In other words, more carbon than otherwise i s necessary goes into the formation of CO2 rather than combustible gases. The e f f e c t that reduced heat losses could have i n the g a s i f i e r operation i s i l l u s t r a t e d on Figure 23. For run No. 38, which has a rather t y p i c a l energy loss of 28.5%, the temperature of the gases leaving the reactor has been calculated assuming that the energy losses could be cut down to d i f f e r e n t percentages to a loss of 10%, t y p i c a l of a larger scale operation. This i s made under the assumption that the increased reactor temperature w i l l not a l t e r the s p e c i f i c heat and the mass of the gases leaving the reactor. Of course, t h i s i s not quite true since the change i n temperature w i l l increase the rates of the d i f f e r e n t reactions involved and thus change the gas composition. However, the point to be made here i s that i f the reactor temperature increases, the endothermic reactions are l i k e l y to proceed at higher CL < UJ X o H Lt_ O to UJ CO CO O < UJ X 30 25 20 15 -10 EXPERIMENTAL POINT ( R U N 38) _L J_ 1000 1200 1400 1600 1800 REACTOR OUTLET GAS TEMPERATURE , K 100 ^ 90 80 70 60 2 000 >-o z UJ o u. UJ cn UJ x 50 S CD 40 o x 30 LLI =5 FIGURE 23: CALCULATED EFFECT OF ENERGY LOSSES ON THE TEMPERATURE OF THE GAS LEAVING THE COAL GASIFIER rates (for instance, the reaction of carbon and steam which i s known to become important only at temperatures over 1300 K). This i n turn w i l l tend to consume the extra heat and bring the reactor temperature down below the calculated value,but the q u a l i t y of the gas produced could improve considerably. This factor i s important when comparing r e s u l t s of t h i s i n v e s t i g a t i o n with those from larger scale equipment where heat losses are low or with small scale studies where stronger gases are pro-duced, by providing the reaction heat f o r example by external e l e c t r i c a l heating. In the l a t t e r case one could gasify coal by steam only with no nitrogen to d i l u t e the off-gas, and consume l i t t l e coal i n CO2 production. THERMAL EFFICIENCIES The thermal e f f i c i e n c y of a g a s i f i c a t i o n system i s a measure of the e f f i c i e n c y with which the energy contained i n the coal i s transformed into a gaseous f u e l . Two kinds of thermal e f f i c i e n c i e s are defined here. The f i r s t i s the thermal e f f i c i e n c y of the clean cool gas, and i s simply the r a t i o between the combustion heat at 288.6 K of the clean (no s o l i d s , no tars) gas produced by a unit mass of coal and the heat value of that unit mass of coal, i . e . The clean gas thermal e f f i c i e n c y for the g a s i f i c a t i o n of various sizes of Forestburg coal (Table 24) was generally between 40 and 50% , which i s s i m i l a r to the e f f i c i e n c y obtained for g a s i f i c a t i o n of the same coal towards improved e f f i c i e n c y with increasing coal p a r t i c l e s i z e . This appears to be related with a decrease of the amounts of carbon e l u t r i a t e d x 100 i n a spouted bed (1) The data for t h i s coal shows a general trend TABLE 24 THERMAL EFFICIENCIES AND % USEFUL CARBON FOR THE GASIFICATION OF SOME WESTERN CANADIAN COALS Run a Thermal e f f i c i ency (%) Useful carbon n < cc 7 ' raw 1 25.7 55.8 2 39.0 78.6 3 26.3 70.8 4 24.5 54.2 . Mean + S.D. 28.9 + 6.8 64.9 + 11.8 5 44.4 90.8 6 73.3 94.1 7 54.2 92.4 8 35.8 88.8 9 11.9 72.1 Mean + S.D. 43.9 + 22.7 87.6 + 8.9 10 16.2 73.6 11 41.6 83.4 12 36.8 60.0 85.9 , 13 61.0 91.0 14 51.4 83.2 83.1 15 46.4 77.9 81.6 16 40.1 73.0 88.6 17 35.2 73.8 84.6 18 34.5 84.3 86.9 19 36.0 57.5 81.2 21 38.1 64.0 88.1 22 39.1 • 58.5 78.8 Mean +_ S.D. 39.7 + 10.7 70.2 + 10.5 83 .9+4 .8 23 30.2 61.1 82.6 24 62.9 91.8 25 57.9 70.6 90.5 26 66.1 89.9 27 47.3 84.8 28 45.5 84.3 29 44.9 84.0 30 46.9 84.6 31 61.1 80.2 90.8 Mean + S.D. 50.9 + 10.9 70.6 + 9.6 87.0 + 3.6 32 39.3 64.0 33 45.0 66.9 Mean + S.D. 42.2 + 4.0 65.5 + 2..1 36 25.2 53.8 40.8 37 26.8 58.3 42.4 38 31.1 49.0 45.3 Mean + S.D. 27.7 + 3.1 53.7 + 4.7 42.8 + 2.3 39 25.8 48.2 47.0 40 29.0 52.8 51.2 Mean + S.D. 27.4 + 2.3 50.5 + 3.3. 49.1 + 3.0 41 26.8 61.3 S u n s 1- i: 0 . S 3 ran F c r e s t b u r g c o a l " R u n s ' 3 2 - 3 3 : " 4 . 7 6 ' m m O.'C OT F o r e s t b u r g c o a l Runs 5- 9 : 0 . 9 5 m F o r e s t b u r g c o a l Runs 3 6 - 3 8 : 2 . 1 8 mm S u k u n k a c o a l Runs 1 0 - 2 2 : 2 . 1 8 mm F o r e s t b u r g c o a l Runs 3 9 - 4 0 : 4 . 0 6 mm S u k u n k a c o a l Runs 2 3 - 3 1 : 4 . 0 6 mm F o r e s t b u r g c o a l Runs 4 1 : 2 . 1 8 am C o l e m a n c o a l from the bed as the coal s i z e increases, leaving more carbon i n the bed f o r g a s i f i c a t i o n . The clean gas thermal e f f i c i e n c y f o r g a s i f i c a t i o n of Sukunka and Coleman coal was approximately 27% (Table 24) i . e . s u b s t a n t i a l l y lower than for Forestburg co a l . This difference i s larger than the one expected from the differences i n gas q u a l i t y f o r these two coals (Tables 12 & 13), and seems to confirm the importance of carbon e l u t r i a t i o n rates, which were s u b s t a n t i a l l y higher for the caking coals (see Tables 21 & 22). A second thermal e f f i c i e n c y f o r the system, that of the wet hot gas can be defined as the r a t i o between the t o t a l heat value of the gas (combustion heat + sensible heat + sensible heat of steam) over the energy input to the system (heat content of c o a l , a i r and steam fed). Such e f f i c i e n c i e s are p a r t i c u l a r l y meaningful for i n - s i t u i n d u s t r i a l a p p lications where the " d i r t y " gas can be used d i r e c t l y . Computation of the raw gas thermal e f f i c i e n c y requires carrying out a complete energy balance, and values for selected runs are presented i n Table 24. Note that since i n t h i s work the tar content of the gas was determined only on selected runs, and that the tars were not analysed, the thermal e f f i c i e n c y of the raw gas as defined here does not include the heat value of the t a r . The average raw gas thermal e f f i c i e n c i e s were 70 - 71% for the g a s i f i c a t i o n of Forestburg coal and 50 - 54% for Sukunka coal. Again, the difference i s due to the energy losses to the system i n the form of e l u t r i a t e d carbon which were s u b s t a n t i a l l y higher for Sukunka coal (see Table IX-1, Appendix IX) heat content of hot wet gas t o t a l heat input F u l l scale g a s l f l e r design c a l c u l a t i o n s suggest clean gas e f f i c i e n c i e s of 70 - 80% and hot gas e f f i c i e n c i e s of 85 - 90%. These higher e f f i c i e n c i e s are due to the production of a better q u a l i t y gas i n the absencedof large heat losses, and also to the recycle of carbon e l u t r i a t e d from the bed. USEFUL CARBON RATIOS Yet another way of defining the e f f i c i e n c y of a g a s i f i c a t i o n system (3, i s computing the useful carbon r a t i o or % us e f u l carbon. This i s defined as the r a t i o between the t o t a l u s e f u l carbon, C , over the carbon a v a i l a b l e u f o r g a s i f i c a t i o n , C^. The useful carbon i s defined as the carbon converted to synthesis gas, C^, (CO + H^) plus the carbon converted to gaseous hydrocarbons, C^. then, C = C + C, u s h On the other hand the carbon a v a i l a b l e f or g a s i f i c a t i o n i s defined as C = C + C , + O. + C . a s h 1 e; where C = unreacted carbon = e l u t r i a t e d carbon, e and C^ = carbon i n l i q u i d hydrocarbons, but from the stoichiometry of the g a s i f i c a t i o n reactions: C + H 20 _ CO + H 2 and C + 2H20 _ -» C0 2 + 2H 2 C + r.n^ _ -» 2 CO i t i s evident that a l l of these reactions produce two moles of synthesis gas per mole of carbon reacted. Accordingly: C s = S / 2 where S = moles of CO + H 2 produced. Therefore: C C , C. • u s + h yc = = C C + C, + C, + C a s h -. 1 e [h (moles H„ + moles CO) + moles CH.] x 12 yc = 2 4j [^(moles H 2 + moles CO) + moles CH^] x 12 + e l u t r i a t e d carbon If the amount of carbon i n the l i q u i d hydrocarbons i s neglected and the sole gaseous hydrocarbon i s methane, the u s e f u l carbon r a t i o i s therefore a means of separating the combustion from the g a s i f i c a t i o n reactions and assessing how much of the carbon that does not go into the formation of C0 2 goes into the formation of combustible gases. Consequently, the u s e f u l carbon r a t i o i s a strong inverse function of the amount of e l u t r i a t e d carbon, and points out the importance of heat losses to the environment through e l u t r i a t e d carbon previously discussed i n t h i s chapter. This i s i l l u s t r a t e d by Table 24 showing that the u s e f u l carbon r a t i o was higher for Forestburg coal than for the caking coals where carbon e l u t r i a t i o n r a t i o s are higher. For the Forestburg c o a l , the u s e f u l carbon r a t i o s were lowest for the smallest p a r t i c l e s i z e (0.53 mm) and the coal containing fines (-4.72 + 0.0 mm) which exhibited the highest carbon e l u t r i a t i o n rates. I t should be noted that no attempt was made to recycle the e l u t r i a t e d s o l i d s i n t h i s work.• In p r a c t i c e these f i n e s would ei t h e r be recycled to the g a s i f i e r or burned to r a i s e steam f or g a s i f i c a t i o n , and would not contribute so s i g n i f i c a n t l y to an o v e r a l l plant e f f i c i e n c y . EQUILIBRIUM CONSIDERATIONS An attempt to c l a r i f y which are the preponderant reactions i n the present coal g a s i f i c a t i o n system i s presented below by consideration of some of the equilibrium data a v a i l a b l e i n the l i t e r a t u r e . Experimental evidence has been provided i n Chapter IV (see Figure 14) that p y r o l y s i s reactions are the main i f not the sole source of methane. The observed e f f e c t of the steam r a t i o i n the gas composition (see Chapter 4) and the mass balances points;to the fac t that i n the present system, hydrogen..is...also mainly a product of coal rather than steam de-composition. On the other hand, the energy balances showed that i n a g a s i f i e r characterized by large heat losses, the combustion of carbon to CO2 i s a predominant factor i n determining the gas composition. Consider now the following reactions i n a carbon steam system: C +-H20 _ r.n + H 2 (1) CO + H 20 „ C0 2 + H 2 (2) C + C0 2 ^ 2C0 (3) C + ? H - **- r H (4) and compare the experimental product to steam r a t i o s obtained (Table 25)) with the corresponding equilibrium r a t i o s (Figure 24, a f t e r r e f e r - , ence 29) TABLE 25 PRODUCT RATIOS FOR THE GASIFICATION OF SOME WESTERN CANADIAN COALS Run Reactor Product, ra t io s # average Temperature (K) CO; ,/H20 CH 4/H 20 H 2/H 20 C0/H20 10 1088 0..445 0. 018 0. 039 0.070 11 1050 0.786 0. 028 0 372 0.441 12 1040 0 539 0 032 0 386 0.364 13 1061 0 690 0 032 0 466 0.524 14 1031 0 614 0 039 0 484 0.406 15 1031 0 708 0 046 0 477 0.477 16 1107 0 289 0 021 0 175 0.355 17 1165 0 576 0 025 0 171 0.346 18 1145 0 360 0 014 0 115 0.143 19 1049 1 343 0 076 0 699 0.873 21 1080 0 527 0 028 0.341 0.458 22 1026 0 976 0 070 0 790 0.635 23 1108 1 984 0 044 0 415 0.503 24 1096 0 301 0 023 0 442 0.433. 25 1074 1 335 0 060 1 006 0.886 26 1065 1 404 0 093 1 789 1.681 27 1115 0 503 0 025 0 318 0.421 28 1085 0 411 0 018 0 253 0.277 29 1079 0 443 0 022 0 278 0.242 30 1053 0 354 0 018 0 25 T 0.173 32 1125 0 233 0 013 0 139 . 0.223 33 1055 0 208 0 012 0 169 0.126 36 1132 0 397 0 038 0 147 0.224 37 1081 0 286 0 028 0 130 0.154 38 1165 1 427 0 236 0 802 0.837 39 1160 1 033 0 068 0.290 0.479 . 40 1107 0.847 0 054 0 334 0.378 41 1096 0 596 0 104 0 388 0.427 101 FIGURE 24: EQUILIBRIUM PRODUCT-STEAM RATIOS AS A FUNCTION OF .TEMPERATURE FOR THE REACTIONS :. C + H 20 —=- CO + H 2; CO + H 20 ~ C 0 2 + H 2 C + C0 2 ^  2C0 and C + 2H 2 ^ *»• CH^ PERFECT GAS LAW ASSUMED 102 An inspection of the data shows that only the C^/H^O product r a t i o f o r reaction (2) i s near equilibrium. The remainder of the experi-mental product r a t i o s are one or two orders of magnitude away from the equilibrium. Discounting the production of methane and hydrogen by p y r o l y s i s , and bearing i n mind the fa c t that the presented equilibrium data do not consider the presence of oxygen i n the system, makes i t even more evident that the contribution of the above reactions, perhaps with the exception of reaction (2), to the production of combustible gases i s n e g l i g i b l e . The extent to which reaction (2) occurs i s d i f f i c u l t to assess from the a v a i l a b l e data since one of the products of the reaction, i s mainly produced from carbon and oxygen. Since reactions (1) and (3) do not contribute s i g n i f i c a n t l y to the production of carbon monoxide, t h i s gas must be produced mainly by reaction(5) C + h o 2 CO (5) From Figure 17, at a constant steam r a t i o the production of carbon monoxide was a l i n e a r function of the coal feed rate, with a carbon monoxide y i e l d of 52.6%. This implies that under t h i s s p e c i f i c operating conditions, the r a t i o of carbon i n carbon monoxide to carbon i n the coal i s : 5 2 ' 6 x 1 2 = 33.8% 28 x 0.6679 If t h i s r a t i o i s calculated for a l l g a s i f i c a t i o n runs with Forestburg coal (Table 26), the average values obtained for a l l p a r t i c l e sizes but the smaller ones are i n the range of 30.7 to 36.5%. This suggests that TABLE 26 CARBON IN CARBON MONOXIDE TO CARBON FEED RATIO FOR THE GASIFICATION OF FORESTBURG COALS Run Carbon in CO I w /w 1 # Carbon in coal \ ri / W l 1 14.7 2 29.1 3 14.9 4 14.3 Mean + S .D. 18.3 + 7.2 5 29.7 6 48.2 7 42.6 8 26.5 9 10.7 Mean + S .D. 31.5 + 14.7 10 13.9 n 33.1 12 25.3 13 47.4 14 33.3 15 32.7 16 38.8 17 33.1 18 26.5 19 28.1 21 31.8 22 24.4 Mean + S .D. 30.7 + 8.2 23 23.3 24 46.5 25 39.9 26 47.8 27 39.5 28 32.7 29 29.9 30 27.5 31 41.4 Mean + S .0. 36.5 + 8.6 32 35.3 33 27.7 Mean + S .D. 31.5 + 5.4 carbon monoxide production i s a l i n e a r function of the coal feed rate, quite independently of other operating v a r i a b l e s , i n p a r t i c u l a r the steam to coal r a t i o . Therefore, carbon monoxide appears to be produced mainly from reaction (5). This i s i n agreement with other r e s u l t s ( O (. \ reported i n the l i t e r a t u r e f o r a s i m i l a r system L a s t l y , the preceding discussion on the r e l a t i v e importance of the d i f f e r e n t reactions occuring i n a coal g a s i f i c a t i o n system i s supported by the thermodynamic and k i n e t i c data of a coal g a s i f i c a t i o n system pre-sented i n Chapter I. COMPARISON OF GASIFICATION RESULTS FOR FLUIDIZED AND SPOUTED BED One of the major objectives of t h i s research was to compare the performance of a g a s i f i e r operated as a f l u i d i z e d bed and as a spouted bed with the three d i f f e r e n t Western Canadian coals. This comparison i s d i f f i c u l t because t y p i c a l operating conditions for the two types of reactor were d i f f e r e n t . In some cases i t was possible.to operate under s i m i l a r conditions, and i n other cases i t was not. Bearing i n mind the l i m i t a t i o n s noted above, a comparison of the performance of the f l u i d i z e d bed reactor as obtained i n t h i s work with the experimental r e s u l t s of Foong et a l . for a spouted bed ^ " ^ i s presented i n Table 27 for " t y p i c a l " and " s i m i l a r " operating conditions with the three d i f f e r e n t coals. In general, the data show that for t y p i c a l operating conditions the gas q u a l i t y obtained from the spouted bed was 40 - 56% better than from the f l u i d i z e d bed g a s i f i e r . On the other hand, at s i m i l a r a i r to coal TABLE 27 COMPARISON OF RESULTS FOR THE GASIFICATION OF SOME WESTERN CANADIAN COALS IN A FLUIDIZED AND SPOUTED BED Type of r e s u l t F l u i d i z e d bed (this work) Spouted bed C 1 ) Coal feed r a t e (g/s) i.Air to coal r a t i o (w/w) Gas gross c a l -o r i f i c value (MJ/m3) Coal feed rate (g/s) A i r to coal r a t i o (w/w) Gas gross c a l -o r i f i c value (MJ/m3) Typical of 2.18 mm Forestburg coal 1.516 5.3 2.25 2.528 2.3 3.51 Forestburg c o a l under s i m i l a r op.• conditions 1.673 5.01 2.00 1.211 5.1 1.35-1.79 Typi c a l of 2.18 mm Sukunka coal 1.462 4.7 1.92 1.889 2.7 2.69 Sukunka coal under s i m i l a r operating conditions 1.580 4.1 2.43 1.500 4.4 1.87 Coleman coal under s i m i l a r operating conditions* 1.968 3.48 2.51* 1.750 2.90 1.79 Under t h i s condition coal caked i n the f l u i d i z e d bed r a t i o s , which i n both cases was the most important operating parameter, the gas q u a l i t y of the gas produced f o r a f l u i d i z e d bed was 12 - 48% better than from the spouted bed. Although the a i r to coal r a t i o may have an i n d u s t r i a l s i g n i f i c a n c e , the r a t i o does not nec e s s a r i l y mean the same thing from a process point of view f o r both types of reactor. In a f l u i d i z e d bed, as long as there i s not extensive bubbling the a i r to coa l r a t i o can be representative of the e f f e c t i v e oxygen to carbon r a t i o a v a i l a b l e f o r reaction. In a spouted bed on the contrary, the a i r to coa l r a t i o i s s u b s t a n t i a l l y higher f o r the spout region than f o r the annulus region, and both i n turn are d i f f e r e n t from the o v e r a l l a i r to coal r a t i o . This stretches the meaningfulness of comparing the performance of both reactors under " s i m i l a r " operating conditions. A basic difference between the operation of both reactors was the a b i l i t y to obtain lower a i r to coal r a t i o s and operate at higher coal feed rates i n the spouted bed. As pointed out e a r l i e r (see Chapter 2) the coal r e a c t i v i t y l i m i t e d the minimum a i r to coal r a t i o to approxi-mately 4.0 for Forestburg coal and to ^  5.0 for Sukunka coa l . Since on the other hand the maximum a i r flow was l i m i t e d by the s i z e of the equipment and the p r a c t i c a l i t y of operation under f l u i d i z i n g mode to -3 3 approximately 8 x 10 std'm /s;: the maximum coal feed rate was l i m i t e d to 2.00 - 2.40 g/s while t y p i c a l feed rates were 1.2 - 2.0 g/s. By contrast, t y p i c a l feed rates f o r the g a s i f i c a t i o n at Forestburg coal i n a spouted bed were 2.50 - 3.33 g/s i . e . s u b s t a n t i a l l y higher. The preceding f a c t s i n d i c a t e that f o r a given equipment s i z e , l a rger coal throughputs can be achieved with a spouted bed (Table 28). This allows the production of a gas of better q u a l i t y . T A B L E 2 8 C O M P A R I S O N OF T H E M A I N O B S E R V E D O P E R A T I N G C H A R A C T E R I S T I C S OF A F L U I D I Z E D AND S P O U T E D B E D R E A C T O R S O p e r a t i n g v a r i a b l e F l u i d i z e d b e d S p o u t e d b e d T y p i c a l t h r o u g h p u t ( g / s m 2 ) f o r n o n - c a k i n g c o a l ( F o r e s t b u r g ) 6 5 . 8 - 1 0 9 . 7 1 3 7 . 1 - 1 8 2 . 6 E f f e c t o f c o a l f e e d r a t e o n g a s q u a l i t y c r i t i c a l c r i t i c a l E f f e c t o f s t e a m r a t i o o n g a s q u a l i t y l i t t l e n o n e E f f e c t o f b e d d e p t h o n g a s q u a l i t y n o n e n o n e O b s e r v e d m a x i m u m t h r o u g h p u t ( g / s m 2 ) f o r S u k u n k a ( c a k i n g ) c o a l 8 6 . 6 1 2 4 . 9 O b s e r v e d m a x i m u m t h r o u g h p u t ( g / s m 2 ) f o r C o l e m a n ( c a k i n g ) c o a l 1 0 7 . 9 ^ 9 5 . 9 ( a ) U n d e r t h i s c o n d i t i o n c o a l s e v e r e l y c a k e d i n t h e b e d . In s p i t e of these d i s i m i l a r i t i e s , the spouted and f l u i d i z e d bed have several common c h a r a c t e r i s t i c s (Table 28). In both systems, the coal feed rate and a i r to coal r a t i o were c r i t i c a l i n determining the gas q u a l i t y while t h e i r performances were l a r g e l y i n s e n s i t i v e (other than gas composition for the f l u i d i z e d bed) to the steam to coal r a t i o . This i s no doubt a c h a r a c t e r i s t i c of a system operating at r e l a t i v e l y low temperature where p y r o l y s i s and p a r t i a l combustion of carbon are the preponderant reactions i n determining the gas q u a l i t y . Both systems were also i n s e n s i t i v e to changes i n bed depth, probably for the same reasons mentioned above. F i n a l l y , the maximum throughput of caking coals f or both systems were s i m i l a r . The maximum throughput of Sukunka 2 coal i n the spouted bed was 124.9 (g/s m ), at which point caking 2 problems arose, while a throughput of 86.6 (g/s m ) of t h i s c oal was processed without any problems by the f l u i d i z e d bed. Higher through-puts could not be achieved i n the l a t t e r system because of coal r e a c t i v -i t y and equipment l i m i t a t i o n s rather than caking. A comparison of f l u i d bed versus spouted bed performance i n a larger r e f r a c t o r y l i n e d reactor would be valuable. The maximum throughput with Coleman coal f or the spouted bed was 2 95.9 (g/s m ). In the f l u i d i z e d bed Coleman coal could be g a s i f i e d at a 2 throughput of 107.9 (g/s m ) for approximately 2 hours a f t e r which time the coal was found to have caked. G a s i f i c a t i o n of t h i s coal at lower throughputs was. not .attempted. O v e r a l l , i t appears that ;the . a b i l i t y of gasify i n g caking coals i n a spouted bed i s mainly due to a d i l u t i o n e f f e c t of the i n e r t s i l i c a p a r t i c l e s rather than because of the p o t e n t i a l of the high v e l o c i t y spout to break up agglomerates. 109 CHAPTER VI - CONCLUSIONS Three d i f f e r e n t Western Canadian coals, one non-caking and two of the caking type of coal, were g a s i f i e d under a wide range of operating conditions i n a small p i l o t plant capable of steady operation at through-2 puts of up to 109.6 g/s m (2.0 g/s) of dry non-caking co a l . The gas-i f i e r was an a i r f l u i d i z e d bed of 0.73 mm Ottawa sand and coal operated at atmospheric pressure. Discounting the d i l u t i o n e f f e c t of some n i t r o -gen used i n the coal feeding system, a l l three coal tested t y p i c a l l y 3 produced a gas of a c a l o r i f i c value i n the range of 2.0 — 2.6 MJ/m . These values are lower than for commercially a v a i l a b l e f l u i d i z e d bed reactors which produce gas of a c a l o r i f i c value i n the range of 3.91 -3 4.40 MJ/m . The difference i n performance i s a t t r i b u t e d to the i n -a b i l i t y to operate at average temperatures over about 1175 K because the reactor was not re f r a c t o r y l i n e d , and to the important impact that large energy losses from the reactor (about 25% of energy input) had on the q u a l i t y of the gas produced. Analysis of the e f f e c t of the d i f f e r e n t operating v a r i a b l e s i n -dicated that the sing l e most important operating parameter i n determin-ing the gas q u a l i t y was the coal feed rate, while an inverse c o r r e l a t i o n between the a i r to coal r a t i o and gas q u a l i t y was generally observed. Minimum a i r to coal r a t i o s were l i m i t e d by coal r e a c t i v i t y to 4.0 and about 5.0 for Forestburg and Sukunka coal r e s p e c t i v e l y . Temperature i n the range of 1023 - 1175 K, p a r t i c l e s i z e , bed depth and steam to coal r a t i o did not have important e f f e c t s on the gas c a l o r i f i c value. However, increasing the steam to coal r a t i o resulted i n an increase of the hydrogen concentration and a decrease of the carbon monoxide concentration. 110 Even though the coal p a r t i c l e s i z e had only minor e f f e c t s on the q u a l i t y of the gas produced i t had an impact on the q u a l i t y of the oper-a t i o n . Increasing the coal p a r t i c l e s i z e made the operation more stable, decreased the amount of tar produced and decreased the amounts of s o l i d s and carbon e l u t r i a t e d from the bed thereby increasing the raw gas thermal e f f i c i e n c y and the useful carbon r a t i o . G a s i f i c a t i o n of coal with a wider s i z e d i s t r i b u t i o n which included the f i n e s (-4.76 + 0.0 mm), as opposed to.operation-with narrowly sized coal, did not have a detrimental e f f e c t either on the gas q u a l i t y or the s t a b i l i t y of the operation a l -though an increase i n s o l i d s and carbon e l u t r i a t i o n rates was observed. Tar production from t h i s coal was le s s than expected from i t s nominal average s i z e . The coal c h a r a c t e r i s t i c s also had an impact on the q u a l i t y of the operation. The non-caking coal (Forestburg) and one caking coal of free swelling index' of 7 (Sukunka) could be g a s i f i e d without any operating 2 2 problems at throughputs of 109.6 g/s m and 86.62 g/s m r e s p e c t i v e l y . A second caking coal, with a free swelling index of 4 (Coleman) caked 2 when g a s i f i e d at a throughput of 107.9 g/s m . Solids and carbon e l u t r i a t i o n rates for the g a s i f i c a t i o n of Sukunka coal were higher by a factor of 3 than for Forestburg coa l , while for the one run with Coleman coal they were intermediate between the two. Tar production from these three coals were i n s i m i l a r proportions to the e l u t r i a t i o n rates. Analysis of the data indicated that i n the present system com-bustion and p y r o l y s i s reaction prevailed over g a s i f i c a t i o n reactions. Most of the hydrogen and e s s e n t i a l l y a l l of the methane produced arose from coal thermal decomposition while most of the carbon monoxide Ill appeared to be produced from p a r t i a l combustion of carbon. In most g a s i f i c a t i o n runs, water was produced i n the reactor from combustion of hydrogen and hydrocarbons. Carbon dioxide appeared to be mostly produced by the combustion of carbon, although there was experimental evidence of the reaction of carbon monoxide with water to form carbon dioxide and hydrogen. The main g a s i f i c a t i o n reaction, that of carbon and steam to produce carbon monoxide and hydrogen did not appear to occur to any important extent. These conclusions were i n agreement with r e s u l t s reported i n the l i t e r a t u r e and thermodynamic and k i n e t i c p r edictions f o r a g a s i f i c a t i o n system operated at temperatures below 1175 K which allowed f o r r e l a t i v e l y large energy losses from the reactor. Comparison of the performance of the f l u i d i z e d bed and a spouted bed for g a s i f i c a t i o n of the three coals was d i f f i c u l t because i t was not always possible to operate under s i m i l a r conditions. A basic d i f f e r e n c e between the two g a s i f i e r s was the a b i l i t y to operate at higher coal feed rates and lower a i r to coal r a t i o with the spouted bed. Since i n both cases these operating parameters, were the most important factors i n determin-ing the gas q u a l i t y , the data showed that f or " t y p i c a l " operating con-d i t i o n s the gas q u a l i t y obtained from the spouted bed was 40 - 56% better than for the f l u i d i z e d bed g a s i f i e r . On the other hand, at " s i m i l a r " a i r to coal r a t i o s , the gas q u a l i t y of the gas produced i n the l a t t e r reactor was 12 - 48% better than from the spouted bed. This indicated that for a given equipment s i z e , l arger coal throughputs can probably be achieved with a spouted bed, allowing the production of a gas of better q u a l i t y . In s p i t e of t h i s d i s s i m i l a r i t y , both g a s i f i e r s had several common c h a r a c t e r i s t i c s . In both systems the coal feed rate and a i r to coal r a t i o were c r i t i c a l i n determining the gas c a l o r i f i c value, while t h e i r performance was -~ l a r g e l y i n s e n s i t i v e to the steam to coal r a t i o , temper-ature, and bed depth. This i s no doubt a c h a r a c t e r i s t i c of a system operating at r e l a t i v e l y low temperature where combustion and p y r o l y s i s reactions dominate. The a b i l i t y of both systems to process caking coal was also s i m i l a r . The maximum throughput of Sukunka coal i n the spouted bed was 124.9 g/s n 2 while 86.62 g/s m of t h i s coal was processed without any problems i n the f l u i d i z e d bed. Larger throughputs could not be achieved i n the l a t t e r because of coal r e a c t i v i t y and equipment l i m i t a t i o n s rather than caking. The maximum throughput of Coleman coal i n the spouted bed was 2 95.9 g/s m while i n the f l u i d i z e d bed t h i s coal caked a f t e r 2 hours of 2 operation at a throughput of 107.9 g/s m . In conclusion, i t appears that the previously noted a b i l i t y of a spouted bed to trea t l i m i t e d throughputs of caking coals i s due to the dispersion of the coal i n a bed of i n e r t s i l i c a and ash rather than to the a b i l i t y of the high v e l o c i t y gas spout to break-up agglomerates. REFERENCES F o o n g , J . , W a t k i n s o n , A . P . , a n d K . B . M a t h u r . 1 9 7 8 . " G a s i f i c a t i o n o f C a k i n g C o a l s i n a S p o u t e d B e d " , Y e a r E n d R e p o r t t o t h e D e p a r t m e n t o f E n e r g y , M i n e s a n d R e s o u r c e s , O t t a w a . D e p a r t m e n t o f C h e m i c a l E n g i n e e r i n g , U n i v e r s i t y o f B r i t i s h C o l u m b i a , B . C . N . B e r k o w i t z . 1 9 7 8 . " W e s t e r n C a n a d i a n I n t e r e s t s i n C o a l C o n v e r s i o n " . P a p e r p r e s e n t e d a t t h e C o a l R e f i n i n g S y m p o s i u m a t E d m o n t o n , A l b e r t a . A p r i l 2 0 - 2 1 . J . M . T a y l o r . 1 9 7 8 . " C o a l G a s i f i c a t i o n f o r P r o d u c t i o n o f F u e l G a s a n d P e t r o c h e m i c a l F e e d s t o c k s " , P r e s e n t e d a t t h e C o a l R e f i n i n g S y m p o s i u m a t E d m o n t o n , A l b e r t a . A p r i l 2 0 - 2 1 . A l b e r t a E n e r g y R e s o u r c e s C o n s e r v a t i o n B o a r d . 1 9 7 5 . P r o s p e c t s f o r  C o a l G a s i f i c a t i o n i n A l b e r t a . 6 t h A v e . S . W . , C a l g a r y , A l b e r t a , C a n a d a . A . V e r m a . 1 9 7 8 . " F r o m C o a l t o G a s " . P a r t I . C h e m t e c h V o l . 8 , N o . 6 p p . 3 7 2 - 3 8 1 . P a r t I I . C h e m t e c h V o l . 8 , N o . 1 0 , p p . 6 2 6 - 6 3 8 . J . Y e r u s h a l m i , 1 9 7 7 . R e p o r t o n E P R I ' s W o r k s h o p o n C l e a n G a s e o u s F u e l f r o m C o a l . H e l d a t B u r l i n g a m e , C a l i f o r n i a , M a y 1 6 , 1 9 7 6 . M c C a l e b , T . L . , a n d C L . C h e n . 1 9 7 7 . " L o w B . t . u . G a s a s I n d u s t r i a l F u e l " . C h e m . E n g . P r o g r . 2 3 , J u n e 1 9 7 7 . p p . 8 2 - 8 8 . 114 8. V.P. Baboudjian, 1976. " G a s i f i c a t i o n of Coal to Low B.t.u. Gas and P o t e n t i a l Use of a Spouted Bed", Essay for the Degree of Master Engineering. Chemical Engineering Department. U n i v e r s i t y of B r i t i s h Columbia, Vancouver, Canada. 9. H. Perry, 1974. "Coal Conversion Technology". Chem. Eng. July 22, 1974. pp. 88-102. 10. I n s t i t u t e of Gas Technology. 1973. Clean Fuels from Coal. Symposium I Papers. 11. I n s t i t u t e of Gas Technology. 1975. Clean Fuels from Coal. Symposium II Papers. 12. American Gas Association Inc. 1975. Proceedings of the Seventh Synthetic P i p e l i n e Gas Symposium, Chicago. 13. Lurgi Canada Ltd. 1976. Coal G a s i f i c a t i o n Seminar. Calagary, Canada. 14. Kasturirangan, V.N. et a l . 1973. "Comparative Study of Commercial Coal G a s i f i c a t i o n Processes". Kopper Totzek, L u r g i and Winkler". Koppers Co. Inc. 15. Coal Processing Technology. 1974. A CEP Technical Manual. A; I. Ch. E. 115 16. Coal G a s i f i c a t i o n and L i q u e f a c t i o n . 1974. Best Prospects f o r Commercialization Symposium. U n i v e r s i t y of P i t t s b u r g h . 17. Watson, W.B. and C P . Curran. 1977. "Comparison of Coal G a s i -f i c a t i o n Processes". Proceedings of 42nd Mid-year Meeting of  the American Petroleum I n s t i t u t e , Process R e f i n i n g Department. May 9-12, 1977. Chicago, 111. 18. C a s i o r , S.J. et a l . 1975. " F l u i d i z e d Bed G a s i f i c a t i o n of Various Coals", Chem. Eng. Prog. V o l . 71, No. 4, p. 147. 19. J.B. Yasinsky. 1978. "Advances i n F l u i d i z e d Bed G a s i f i c a t i o n Process Development" i n Coal Technology '78. V o l . I I . Proceedings of the I n t e r n a t i o n a l Coal U t i l i z a t i o n Conference. October 17-19, Houston, Texas, pp. 22-35. 20. Symposium on Coal G a s i f i c a t i o n . ACS. 1973. D a l l a s , Texas. 21. T.A. Hendrickson, 1975. S y n t h e t i c Fuels Data Handbook, Cameron Inc., Denver. 22. R.A. Ashworth, 1978. "Co-Generation " P l u s " f o r I n d u s t r i a l Parks". i n Coal Technology '78 Proceedings of the I n t e r n a t i o n a l Coal U t i l i z a t i o n Conference. Oct. 1978. Houston, Texas, pp. 215-234. 116 23. Kam, A.Y. , Hixson, A.N., and D.D. Perlmutter. 1976. "The Oxi-dation of Bituminous Coal. 3. Effect or Caking Properties". Ind. Eng. Chem. Proc. Des. Dev. Vol. 15, No. 3, pp. 416-427. 24. Davis, A., Spackman, W., and P.H. Given. 1976. "The Influence of the Properties of Coals on their Conversion into Clean Fuels". Energy Sources. Vol. 3, No. 1. pp. 55-81. 25. Annual Book of ASTM Standards. Part 19. Gaseous Fuels and Coke. 1972. American Society for Testing and Materials. 26. J. Yerushalmi, 1977. "Fluid Bed Processing of Agglomerating Coals". CEP Manual of Coal Processing. Vol. III. pp.156-165. 27. Crewe, G.F., Gat, U., and V.K. Dhir. 1975. "Decaking of Bituminous Coals by Alkaline Solutions". Fuel, Vol. 54. January, pp. 20-23. 28. Saroff, L. et a l . 1977. "Entrained Pretreatment and Coal Trans-port". CEP Manual of Coal Processing Technology. Vol. III. pp. 172-179. 29. Walker, P.L. Jr., Rusiko, F. Jr., and L.C. Austin. 1959. "Gas Reactions of Carbon" in Advances in Catalysis'. Vol. XI. Academic Press, New York. pp. 133-221. 117 30. Kunii, D. and 0. Levenspiel. 1968. Fluidization Engineering. J. Wiley, New York. 31. B.C. Research. 1978. Design, Development and Evaluation of a Source Particulate Sampler. Prepared for the Council of Forest Industries of B.C., Vancouver, B.C. 32. Perry, J.H. 1963. Chemical Engineers Handbook. 4th Edition, McGraw H i l l Co. New York. 33. J. Highley et a l . 1975. "Application of Fluidized Bed Combustion to Industrial Boilers and Furnaces". Proc. Inst. Fuel, 1(1). Fluidized Combustion Conf., London, Sept. 1975. 34. E l l i o t , M.A. et a l . 1952. "Gasification of Pulverized Coal with Oxygen and Steam i n a Vortex Reactor". Ind. Eng. Chem. Vol. 44, pp. 1074-1082. 35. H.A. Simons Ltd. 1978. Engineering Feasibility Study of the British Columbia Research Hog Fuel Gasification System. Project Report No. 4122A. 36. Shaw, J.T., and N.P. McC. Paterson. 1978. "Studies of the Gasification of Solid Fuels in a Fluidized Bed at Atmospheric Pressure". Fluidization: Proc. 2nd. Eng. Foundation Conference on  Fluidization, J.F. Davidson and D.L. Keairns (Ed.) Cambridge Press pp.229-234. APPENDIX I DETAILS OF FLUIDIZATION GRID 119 FIGURE 1-1: DETAIL OF FLUIDIZATION GRID APPENDIX II TOTAL GAS FLOW MEASUREMENT BY ORIFICE PLATE 121 TOTAL GAS FLOW MEASUREMENT BY ORIFICE PLATE The t o t a l gas flow out of the g a s i f i e r was measured with a 19.1 mm (3/4") o r i f i c e plate of our own design (Figures I I - l and II-2) at a point between the cyclone and the gas in c i n e r a t o r . Pressure and temperature upstreame of the o r i f i c e p l ate were measured.by a mercury manometer and a c a l i b r a t e d chromel-alumel thermocouple r e s p e c t i v e l y . The o r i f i c e plate was i n s t a l l e d and then c a l i b r a t e d with a i r by using a Straushibe (S-type) p i l o t tube. PITOT TUBE: The p i t o t tube was supplied by B.C. Research and c a l i b r a t e d i n the wind tunnel of the Department of Mechanical Engineering of U.B.C. A i r v e l o c i t i e s i n the wind tunnel were obtained by measuring the pressure drop i n the tunnel with a Betz manometer. Since then where V = a i r v e l o c i t y (m/s) g = ac c e l e r a t i o n due to gra v i t y = 9.80665 m/s H = pressure d i f f e r e n t i a l i n m of flowing gas h i s given by: h = h. where: h.. = Betz manometer pressure drop, mm H O ;under c a l i b r a t i o n conditions : = 28.91 x 101.592 x 10 3 a i r 8.31439 x 296.3 1.1922 x 102 g/m3 WELDED 2 LENGTH OF 1/4" S.S. T U B E \ ^ WELDED 1/4 PRESSURE GAUGE FITTING 1/8 KLINGERITE GASKET ^ ORIFICE PLATE (SEE DETAIL) I (/> = 1/16" I 7/8' 14' 2 PIPE THREAD WELDED 1/8' THERMOCOUPLE FITTING 4 7/8 c/)=27/8" 2" PIPE THREAD \ J 31/2 1/16 -12' 1 ^ T '(£ = 1/4" CLEARANCE (SIX BOLTS) FIGURE I I - l : ORIFICE PLATE CONSTRUCTION DETAIL FIGURE II-2: ORIFICE PLATE DETAIL Then: 2 x 9.80665 x h± x 10 3 x 1.0 x 10 6 1.1922 x 10 V = 4.056 •}) h x' (m/s) (1) S i m i l a r l y , from B e r n o u i l l i ' s equation the v e l o c i t y measured by the p i t o t tube i s related with the d i f f e r e n t i a l pressure A P by V = C ^ 2 g A P where C = discharge c o e f f i c i e n t of p i t o t tube under the c a l i b r a t i o n conditions V = 4.067 C V V 7 or C = (2) 4.056 -\/A P The discharge c o e f f i c i e n t of the p i t o t tube was calculated f o r each c a l i b r a t i o n condition (Table I I - l ) . The average discharge c o e f f i c i e n t was found.to be : C = 0.839 + 0.006 Therefore, the c a l i b r a t i o n equation f o r the p i t o t tube i s : ' ' V 2g R V = C 2 g-J AJL = c i A P x T V = 338.81 or, f o r a i r : V = 63.013 A P x T P x M.W (m/s) P x M.W. (3) A P x T (m/s ) (4) Where: A P = P i t o t tube pressure drop, mm H^ O T = upstream temperature K P = upstream pressure, Pa The c a l i b r a t i o n curve of the p i t o t tube f o r a i r i s presented i n Figure II-3 TABLE I I - l PITOT TUBE CALIBRATION DATA AP P i t o t Tube mm H2O Betz Monometer Pressure D i f f e r e n t i a l (mm H 20) V e l o c i t y (E.q. (D) (m/s) Discharge C o e f f i c i e n t C (Eq. (2)) 3.30 2.30 6.15 0.835 5.598 3.85 7.96 0.830 7.87 5.50 9.51 0.836 10.41 7.30 10.96 0.838 13.46 9.50 12.50 0.840 19.81 14.14 15.25 0.845 22.86 16.35 16.40 0.846 A i r upstream pressure = 101.592 kPa (30 i n Hg) A i r upstream temperature = 296.3 K T E M P E R A T U R E = 2 9 6 . 3 K P R E S S U R E = 101.59 kPa GAS M O L E C U L A R WEIGHT (AIR) =28.91 g/mol 1.0 2 4 6 8 10 2 4 6 8 IC PRESSURE DIFFERENTIAL AT PITOT TUBE AP , mm H, FIGURE II-3: CALIBRATION CURVE FOR THE PITOT TUBE 127 ORIFICE PLATE CALIBRATION The procedure to calibrate the or i f i c e plate was as follows. For different air flows the pressure drop across the or i f i c e plate as well as the upstream pressure and temperature were recorded. Downstream (approximately 1 m) of the or i f i c e plate, the air velocity at 3 different points of a cross section of the air pipe (Table II-2) were measured using the previously calibrated pitot tube. These three points were chosen according to the principle of proportional areas; i.e. each point represented the middle point over a pipe diameter of three sections of the pipe having the same area. The ai r flow was then calculated from the average a i r velocity and pipe cross sectional area (21.646 x 10 ^  2 (31) m ). Since the flow through an ori f i c e plate i s given by: where: -4 2 A^ = or i f i c e area ='2.852 x 10 m -4 2 A^ = pipe cross area section = 21.646 x 10 m M.W. = Molecular weight of gas = 28.91 g/mol for air T = upstream temperature = 290.36 K A P = pressure drop across o r i f i c e plate, mm R^ O P = upstream pressure, Pa C = discharge coefficient, dimentionless from which the or i f i c e plate discharge coeficient can be calculated by: Then Q = 0.368C P C = 128 TABLE II-2  ORIFICE PLATE CALIBRATION DATA: ORIFICE PLATE PITOT TUBE Flow &P Temperature Relative inner point middle point outer point Average (mm H20) • <°C) Pressure (kPa) (mm H 20) Velocity (m/s) <iP (mm H2O) Velocity (m/s) i P (mm H 20) Velocity . (m/s) Velocity (m/s) (m3/s) 39.37 17 4 0 0 381 2 07 0.508 2.39 0 584 2 57 2 34 5.07 x 10-3 48.26 17 3 0 0 508 2 39 0.635 2.68 0 711 2 83 2 63 5 69 59.69 17 2 0 0 686 2 78 0.813 3.03 0 813 3 03 2 95 6 39 76.20 17 2 0. 32 0 889 3 17 1.016 3.38 1 067 3 47 3 34 7 23 85.09 17 2 0. 37 1 016 3 38 1.016 3.38 1 219 3 71 3 49 7 55 95.25 17 4 0 50 1 219 3 71 1.321 3.86 1 321 3 86 3 81 8 25 107.95 17 3 0 62 1 270 3 78 1.473 4.08 1 499 4 11 3 99 8 64 118.4 17 3 0 75 1 524 4 14 1.575 4.21 1 575 4 21 4 19 9 07 127.00 17 4 0 87 1 575 4 21 1.727 4.41 1 778 4 48 4 37 9 46 147.32 17 4 0 97 1 778 4 48 2.032 4.79 2 032 4 79 4 69 10 15 177.80 17 3 1 32 2 159 4 93 2.337 5.13 2 413 5 22 5 09 11 02 201.93 17 4 1 49 2 540 5 35 2.667 5.48 2 667 5 48 5 44 11 78 231.14 17 4 1 74 2 743 5 56 3.048 5.86 3 048 5 86 5 76 12 47 254.00 17 4 1 99 3 048 5 86 3.302 6.10 3 302 6 10 6 02 13 03 280.6? 17 4 2 24 3 302 6 10 3.683 6.44 3 683 6 44 .6 33 13 70 300.99 17 4 2 37 3 683 6 44 3.937 6.66 3 937 6 66 6 59 14 26 328.93 17 4 2 61 3 937 6 66 4.267 6.94 4 318 6 98 . 6 86 14 85 360.68 17 4 2 94 4 369' 7 02 4.699 7.28 4 699 7 28 7 19 15 56 388.62 17 4 3 19 4 699 7 28 5.080 7.57 5 207 7 66 7 50 16 23 415.29 17 .4 3 44 4 953 7 47 5.334 7.75 5 461 7 85 7 69 16 65 454.66 17 .4 3 74 5 334 7 75 5.842 8.12 6 096 8 29 8 05 17 43 467.36 17 .5 3 91 5 .461 7 85 5.842 8.12 6 .096 8 29 8 09 17 51 495.30 17 .5 4 .23 5 .842 8 .12 6.350 8.46 6 350 8 46 • 8 35 18.07 x 10" Atmospheric pressure: 102.269 kPa; Average temperature: 290.36 K T h e a v e r a g e d i s c h a r g e c o e f i c i e n t w a s ( T a b l e I I - 3 ) : C = 0 . 7 2 2 + 0 . 0 0 7 T h e r e f o r e , t h e f o l l o w i n g e q u a t i o n w a s u s e d t o m e a s u r e t o t a l g a s f l o w o u t o f t h e r e a c t o r : Q = 4 0 3 . 8 2 3 A 2 C A P x T M . W . x P 1 Q = 0 . 0 8 4 - \ | A P x T f ( m 3 / s ) M . W . x P w h e r e A P , T , P w e r e m e a s u r e d a n d M . W . i s k n o w n o n c e t h e g a s c o m p o s i t i o n i s k n o w n . TABLE II-3 DETERMINATION OF ORIFICE PLATE DISCHARGE COEFFICIENT AP Absolute Flow Discharge Pressure C o e f f i c i e n t (mm H 20) (kPa) (m3/s) C 39.37 102.27 5.07 x 10" 3 0.702 48.26 102.27 5.69 0.712 59.69 102.27 6.39 0.719 76.20 102.59 7.23 0.721 85.09 102.64 7.55 0.713 95.25 102.37 8.25 0.736 107.95 102.89 8.64 0.725 118.11 103.02 9.07 0.728 127.00 103.14 9.46 0.733 147.32 103.24 10.15 0.730 177.80 103.59 11.02 0.723 201.93 103.76 11.78 0.726 231.14 104.01 12.47 0.719 254.00 104.26 13.03 0.717 280.69 104.51 13.70 0.718 300.99 104.64 14.26 0.723 328.93 104.88 14.85 0.721 360.68 105.21 15.56 0.722 388.62 105.46 16.23 0.727 415.29 105.71 16.65 0.722 454.66 106.01 17.43 0.723 467.36 106.18 17.51 0.717 495.30 106.50 18.07 x 10-3 0.720 APPENDIX I I I IMPINGER SYSTEM FOR GAS SAMPLE CLEANING 132 IMPINGER SYSTEM FOR GAS SAMPLE CLEANING The basic gas sample cleaning u n i t was an impinger as shown i n Figure I I I - l . This was b a s i c a l l y an hermetic (except for gas i n l e t and outlet) 3 s t a i n l e s s s t e e l scrubber f i l l e d with approximately 150 cm of water. The impinger t r a i n consisted of two p a r a l l e l sets of four impingers i n s e r i e s . In each set, the second impinger had a perforated d i f f u s e r section as shown i n Figure I I I - l , while i n the remaining three impingers the d i f f u s e r section was simply an open end tube. The two sets of impingers were f i t t e d with a system of valves which allowed operation of one set while the other was by-passed. This permitted continuous operation of one set f o r gas cleaning while the second set was only connected when measuring tar content of the gas. The whole system was kept cool by immersion i n a bath of cracked i c e . GAS INLET RUBBER -GASKET 0 RING DIFFUSER SECTION TUBES DEFLECTION PLATE 50.8 mm STAINLESS * STEEL TUBE FIGURE I I I - l : IMPINGER DETAIL APPENDIX IV CALIBRATION OF ROTAMETERS 135 CALIBRATION OF ROTAMETERS SCRUBBING NITROGEN ROTAMETER This rotameter was factory c a l i b r a t e d i n a d i r e c t reading scale. GASIFIER AIR ROTAMETER The g a s i f i e r rotameter was c a l i b r a t e d by using a dry gas meter (CHE 2856). This gas meter was i n turn checked against a high p r e c i s i o n wet test meter made av a i l a b l e by the Water Resources Laboratory i n Vancouver. The agreement between both gas meters was excellent, the U.B.C. gas meter reading (x) being related to the high p r e c i s i o n meter reading (y) by the equation: y = 0.98 x + 0.1 (1) with a c o r r e l a t i o n factor of 0.98 at standard conditions (101.3 Pa = 1 Atm. and 294 K). Just before the entrance to the rotameter, the a i r pressure and temperature were measured with a 170.3 kPa (10 p s i g ) . Matheson pressure gauge and a c a l i b r a t e d chromel-alumel thermocouple r e s p e c t i v e l y . The data taken during the c a l i b r a t i o n of the f l u i d i z a t i o n a i r rotameter against the UBC gas meter are shown i n Table IV-1. The gas meter readings (x).on the l a s t column of Table IV-1 were then corrected according to Eq. (1) and are shown i n Table IV-2. F i n a l l y the c a l i b r a t i o n curve (Figure IV-1) was drawn at standard con-d i t i o n s . This required c a l c u l a t i n g the a i r flow under standard conditions that w i l l give the same rotameter reading as the flow under the actual measur-(32) ing conditions. This i s given by the following equation: where F = volumetric flow f J = f l u i d density P = absolute pressure T = absolute Cemperature 136 TABLE IV-1 GASIFICATION AIR ROTAMETER CALIBRATION DATA ROTAMETER UBC GAS-METER Reading Relative Pressure kPa Temperature o c -Relative Pressure kPa Temperature oc (a) Flow (m3/s) (b) Flow (m3/s) 20 - 15.5 0.20 21.0 0.94xl0 - 3 0.94xl0 - 3 40 - 14.9 0.22 20.0 . 1.57 1.58 60 - 15.1 0.22 20.0 2.27 2.27 80 - 15.2 0.25 20.5 2.87 2.88 100 - 15.5 0.25 21.0 3.57 3.58 120 - 15.9 0.25 21.0 4.27 4.28 140 - 16.1 0.25 21.0 4.93 4.95 160 - 16.0 0.25 21.0 5.66 5.68 180 - 16.0 0.25' 21.0 6.36 6.38 200 0.55 16.1 0.27 21.0 7.08 7.10 210 1.10 16.2 0.27 21.0 7.50 7.52 220 • 1.38 16.3 0.27 21.0 7.84 7.86 230 2.00 16.3 0.27 21.0 8.27 8.28 240 2.28 16.3 0.27 21.0 8.61 8.63 250 2.62 16.2 0.27 21.0 9.02xl0 - 3 9.04x10"3 (a) At gas-meter pressure and temperature (b) At standard conditions (101.3 kPa = 1 At; 294 K) TABLE IV-2 CALIBRATION DATA OF GASIFICATION AIR ROTAMETER AT STANDARD CONDITIONS Rotameter Reading Corrected gas meter Flow (a) (m3/s) Rotameter Flow (b) (m3/s) 20 1.00 x 10" 3 0.99 x 10" 3 40 1.61 1.59 60 2.28 2.26 80 2.86 2.83 100 3.54 3.51 120 4.22 4.18 140 4.87 4.83 160 5.57 5.52 180 6.24 6.19 200 6.90 6.82 210 7.27 7.17 220 7.57 7.46 230 7.93 7.79 240:. 8.24 8.08 250 8.59 x 10" 3 8.41 x 10" 3 (a) at rotameter temperature and pressure (see table IV-1) (b) at standard conditions i n the rotameter 4 0 8 0 1 2 0 1 6 0 2 0 0 2 4 0 R O T A M E T E R R E A D I N G FIGURE IV-1: CALIBRATION CURVE OF GASIFICATION AIR ROTAMETER u> CO In t h i s case subscripts 1 and 2 denote standard conditions and rotameter actual conditions r e s p e c t i v e l y . Results of these c a l c u l a t i o n s are shown i n Table IV-2: Since i n most of the g a s i f i c a t i o n experiments the a i r flow into the g a s i f i e r was not metered under standard conditions, the reverse procedure as indicated here was followed to c a l c u l a t e actual a i r flow from the c a l -i b r a t i o n curve. In other words, the a i r flow derived from the c a l i b r a t i o n curve (Figure IV-1) was corrected by equation (2). INCINERATOR AIR ROTAMETER This rotameter was c a l i b r a t e d using the same gas meter used for c a l i -b ration of the g a s i f i e r a i r rotameter. Since i t i s not necessary to meter the a i r into the e l e c t r i c furnace with extreme accuracy, and the a i r i s metered at conditions very close to standard, no corrections allowing f o r temperature and pressure d i f f e r e n c e from standard conditions were made. The i n c i n e r a t o r a i r rotameter c a l i -bration data i s presented i n Table IV-3 and c a l i b r a t i o n curve i n Figure IV-2. STEAM METER Due to the small steam flows involved, the steam fed into the reactor was measured by a rotameter (Table 9). The rotameter was c a l i b r a t e d by completely condensing the steam a f t e r passing i t through the rotameter. The condensed water was c o l l e c t e d over a period of 10 min. f o r each r o t a -meter reading and i t s volume measured i n a graduated c y l i n d e r . C a l i b r a t i o n data and curve are shown i n Table IV-4 and Figure IV-3 r e s p e c t i v e l y . 140 TABLE IV-3 INCINERATOR AIR ROTAMETER - CALIBRATION DATA Gas Meter Rotameter Reading Relative Pressure kPa Temperature °C Flow (a) (m3/s) .Flow (b) (m3/s) 20 0.12 25 1.10 x l O - 3 1.11 x 10" 3 40 0.12 24 2.02 2.03 60 0.17 24 2.90 2.87 80 0.20 24 3.76 3.70 100 0.25 24 4.72 4.64 120 0.25 24 5.63 5.52 140 0.25 24 6.67 6.53 160 0.25 24 7.66 7.50 180 1.25 24 8.69 8.59 200 1.49 24 9.71 9.61 220 1.67 24 10.85 10.75 240 2.17 24 12.07 x 10" 3 12.00 x l O " 3 (a) At Gas meter temperature and pressure. (b) At standard conditions (101.3 kPa, 294K) and corrected for gas meter deviation. (Eq . (D) 141 FIGURE IV-2: CALIBRATION CURVE OF INCINERATOR AIR ROTAMETER T A B L E I V - 4 S T E A M M E T E R C A L I B R A T I O N D A T A R o t a m e t e r R e a d i n g . M a s s o f w a t e r c o n d e n s e d i n 1 0 m i n . ( g ) F l o w ( g / s ) 0 . 5 2 4 3 0 . 4 1 0 . 7 5 3 2 5 0 . 5 4 1 . 0 0 3 6 0 0 . 6 0 1 . 2 5 4 6 0 0 . 7 7 1 . 5 0 6 6 0 1 . 1 0 1 . 7 5 7 5 0 1 . 2 5 2 . 0 0 8 8 0 1 . 4 7 2 . 2 5 1 0 0 6 1 . 6 8 2 . 5 0 1 1 3 6 1 . 8 9 GAS SAMPLE ROTAMETER The continuous gas sample withdrawn from the main gas l i n e was metered with a small rotameter a f t e r i t had been cleaned of s o l i d s and t a r s , cooled and dri e d (Figure 3). The rotameter was c a l i b r a t e d with a i r using a small -3 3 wet test meter (0.118 x 10 ' (m /s) per r e v o l u t i o n ) . The c a l i b r a t i o n data i s shown i n Table IV-5 and the c a l i b r a t i o n curve i n Figure IV-4. T A B L E I V - 5 GAS S A M P L E R O T A M E T E R - C A L I B R A T I O N D A T A : GAS M E T E R R o t a m e t e r R e a d i n g T e m p e r a t u r e ° C A b s o l u t e : ' . P r e s s u r e k P a . . . F l o w ( a ) ( m m 3 / s ) ' F l o w ( b ) ( m m 3 / s ) 1 7 2 2 . 8 1 0 1 . 3 2 2 7 . 3 8 x l 0 3 2 7 . 2 1 x 1 0 3 2 5 2 2 . 8 1 0 1 . 3 3 4 6 . 2 6 4 5 . 9 8 3 2 2 3 . 1 1 0 1 . 3 3 5 6 . 6 4 5 6 . 2 4 4 7 2 3 . 1 1 0 1 . 3 5 9 0 . 6 2 9 0 . 0 0 5 5 2 3 . 1 1 0 1 . 3 5 1 1 2 . 3 4 1 1 1 . 5 7 6 9 2 3 . 1 1 0 1 . 3 7 1 5 1 . 0 4 x l 0 3 1 5 0 . 0 3 x l 0 3 ( a ) A t g a s m e t e r c o n d i t i o n s . ( b ) A t s t a n d a r d c o n d i t i o n s . 146 160 0 20 40 60 80 ROTAMETER READING FIGURE IV-4: CALIBRATION CURVE OF GAS SAMPLE ROTAMETER APPENDIX V CALIBRATION CURVES FOR GAS ANALYSIS 148 PERCENT CO IN N 2 BY VOLUME FIGURE V - l : CALIBRATION CURVE OF CONTINUOUS CO ANALYSER PEAK HEIGHT , mm 6 v l 150 152 160 PERCENTAGE M E T H A N E , % (v/v) FIGURE V-5: GAS CHR0MAT0GRAPH METHANE CALIBRATION CURVE TABLE V - l COMPOSITION OF STANDARD GAS FOR GAS CHROMATOGRAPH CALIBRATION Component Analysis (% v/v) H2 13.35 CO 19.98 co 2 9.93 2.95 N2 53.79 APPENDIX VI MINIMUM FLUIDIZATION VELOCITY AND SEGREGATION 155 MINIMUM FLUIDIZATION VELOCITY AND SEGREGATION TABLE VI-1 SEGREGATION DATA FOR FLUIDIZATION OF A MIXTURE OF 20% 3.36mm. 1.18mm COAL AND 80% 3.36mm. 1.18 mm SILICA Air std. Flow (m3/s) Pressure Drop, across bed (Pa) Expanded Bed Depth (m) Observations 0 0 0. 229 Packed bed. Well mixed solids. 4.72 x l O - 3 323.8 0. 229 Packed bed. Well mixed s o l i d s . 7.08 747.2 0. 229 Packed bed. Well mixed so l i d s . 9 44 1220.5 0. 229 Packed bed. Well mixed so l i d s . 11 80 1868.1 0 229 Packed bed. at top. Very fine solids f l u i d i z e 12 74 1893.0 0. 235 Packed bed. f l u i d i z e at Finer coal p a r t i c l e s top. 14 16 . 2191.9 0 241 Packed bed. 25mm of coal f l u i d i z e s at top. Segregation starts. 15 10 1992.6 0 269 0.18m Packed with strong Some bubbles bed and 0.089m of f l u i d bed segregation of coal at top. 16 05 1992.6 0 279 Similar to above but bed almost complete-ly segregated with coal i n upper 0.1 m section. 16 99 1942.8 0 292 Mildly bubbling bed. Strongly segregated. 17 94 1893.1 0 330 Vigorous bubbling. Strong segregation with some mixing at top. 18 88 1868.1 0 330 Similar to above with increased mixing. Bed starts to slug. 19 .82 1868.1 0 406 Slugging bed No segregation. 20.77 x l O - 3 1843.2 0 432 Vigorous slugging. Excellent mixing. No segregation. 156 TABLE VI-2: SEGREGATION DATA FOR FLUIDIZATION OF A MIXTURE 20% OF 3.36mm. 1.18mm COAL AND 80% 0.73mm OTTAWA SAND. Air std. Flow (m3/s ) Pressure Across (Pa) Drop. Bed Expanded Bed Depth (tn ) Observations 0 0 0. 203 Packed bed. Well mixed. 3.78 x 10" 3 1345 0 0. 203 Packed bed. Well mixed. 4. 72 1643 9 0. 203 Packed bed. Well mixed. 5 66 2191 9 0. 203 Packed bed. Well mixed. 6. 14 2191 9 0 203 Bubbles r i s e through packed bed. Some coal floats at surface. 6. 61 2167 0 0. 203 Similar. More frequent bubbles. 7. 55 2142 9 0. 229 More bubbles. Fair amount of agitation. 25 mm at top of bed enriched i n coal. 8. 50 2117 2 0. 229 Bubbling bed. Good mixing except for some coal at bed's top. 9 44 2117 2 0 241 Vigorous bubbling. Good mixing. Slight coal enrichment at bed's top. 10 38 2067 4 0 292 Vigorous bubbling. No segregation. 11 33 2067 4 0 292 Similar. Sluggish bed. 14.16 x 10" 3 1992 6 0 373 Slugging bed. No segregation. 157 TABLE VI-3: SEGREGATION DATA FOR FLUIDIZATION OF A MIXTURE OF 15% 3.36 ram. 1.18 mm COAL AND 85% OTTAWA SAND Air Flow std. (m3/s) Pressure Drop Across Bed (Pa) Expanded Bed Depth (m) Observations 0 0 0.178 Packed bed. Well mixed. 2.83 x 10""3 1021.2 0.178 Packed bed. Well mixed. 3.78 1295.2 0.178 Packed bed. Well mixed. A.72 1768.5 0.178 Packed bed. Well mixed. 5.19 2017.6 0.178 Packed bed. Small a i r stream breaks through. 5.66 2167.0 0.178 Packed bed. Some channeling. Some coal floats at bed's top. 6.14 2117.2 0.184 Gently bubbling bed. Some 6.4 mm of coal segregates at top. 6.61 2117.2 0.191 Increased bubbling. 63.5 mm of coal segregates at top. 7.55 2117.2 0.203 Bubbling bed. 63.5 mm of coal segregates at top. 8.50 2067.4 0.216 Vigorous bubbling. Enriched top layer of coal starts disappearing. 9.44 2092.3 0.241 Vigorous bubbling. Good mixing. No noticeable segregation. 10.38 2042.5 0.241 Excellent mixing i n bubbling bed. 11.33 2017.6 0.241 Similar to above. But, bubble coalescense. 12.27 1967.7 0.318 Similar to above. 13.22 1942.8 0.330 Sluggish bed. No segregation. 16.52 x l O " 3 1868.0 0.419 Slugging bed. 158 TABLE VI-4: SEGREGATION DATA FOR FLUIDIZATION OF A MIXTURE OF 10% 3.36 mm. 1.18mmm COAL AND 90% 0.73 mm OTTAWA SAND Air Flow Pressure Drop Expanded Observations Across Bed Bed Depth std. (m3/s) (Pa) ( <n) 0 0 0. 171 Packed bed. No segregation. 2.83 x 10~ 3 996. 3 0. 171 Packed bed. No segregation. 3.78 1320. 1 0. 171 Packed bed. No segregation. 4.25 1544. 3 0. 171 Packed bed. No segregation. 4.72 1668 8 0 171 Packed bed. No segregation. 5.19 1967 7 0 171 Packed bed. No segregation. 5.66 2117. 2 0. 171 Packed bed. Small bubbles r i s e up. 6.14 2092 3 0 178 Gently bubbling bed. 12.7 mm of coal at top. 6.61 2092 3 0 178 Similar to above. 7.08 2067 4 0 191 More bubbling. Some segregation. 7.55 2067 4 0 197 S t i l l some segregation. 8.02 2067 4 0 203 Enriched coal layer at top of bed starts disappearing. 8.50 . 2042 5 0 216 Bubbling bed. Hardly any segregation. 9.44 2042 5 0 229 Vigorouse bubbling. No segregation. 10.38 2017 6 0 241 Same as above. 11.33 1967 7 0 254 Bubble coalescence. Slugging starts. 12.74 1942 8 0 279 Slugging bed. 14.16 1917 9 0 318 Similar to above. 16.52x 10" 3 1843.2 0 356 Slugging bed. Some spouting. 159 TABLE VI-5: SEGREGATION DATA FOR FLUIDIZATION OF A MIXTURE OF 5% 3.36 mm. 1.18 mm COAL AND 95% 0.73 mm OTTAWA SAND Air std. Flow (m3/s) Pressure Across (Pa) Drop Bed Expanded Bed Depth (m) Observations 1 0 0 0. 170 Packed bed. Well mixed. 2.83 x 10" 3 1021. 2 0. 170 Packed bed. Well mixed. 3 78 1270. 3 0. 170 Packed bed. Well mixed. 4 25 1494 5 0. 170 Packed bed. Well mixed. 4 72 1693. 7 . o. 170 Packed bed. Well mixed. 5 19 1967. 7 0. 170 Packed bed. Well mixed. 5 66 2167 0 0 170 Packed bed. Well mixed. 6 14 2167 0 0 170 Small bubbles r i s e through bed. Fine coal floats at surface. 6 61 2142 0 0 178 Gently bubbling bed. Some segregation. 7 08 2142 0 0 178 Similar to above. 7 55 2117 2 0 178 Bubbling bed L i t t l e segregation. 8 02 2117 2 0 191 Bubbling bed Mixing starts. 8 50 2067 4 0 191 Vigorous bubbling. Hardly any segregation. 9 44 2067 4 0 203 Vigorous bubbling. No segregation. 10 38 2017 5 0 216 Similar to above. 11 .80 2017 5 0 216 Some slugging and spouting. 12 .74 1992 6 0 279 More slugging. 14 .16 1942 8 0 343 Fair amount of slugging. 16.52 x 1 0 - 3 1868 0 0 .419 Slugging bed No segregation. 160 TABLE VI-6: FLUIDIZATION OF 7.3 mm OTTAWA SAND Air Flow Pressure Drop Expanded Observations Across Bed Bed Depth std. (m-3/s) (Pa) (m) 0 0 0 178 Packed bed. 2.83 x 10- 3 1170. 7 0 178 Packed bed. 3.78 1494. 5 0 178 Packed bed. 4.25 1743 6 0 178 Packed bed. 4.72 2017 6 0 178 Packed bed. 5.19 2241 7 0 178 Packed bed. 5;66 2266 7 0 178 Bubbles r i s e through bed. 6.14 2241 7 0 .178 More bubbles. 6.61 2266 7 0 .178 Bubbling bed. 7.55 2216 8 0 .191 Bubbling bed. 8.50 2216 8 0 .191 Vigorous bubbling bed. 9.44 2191 9 0 .203 Vigorous bubbling bed. 11.80 2142 1 0 .216 Some slugging. 14.16 2117.2 0 .343 More slugging. 16.52 x 10" 3 2067 4 0 .419 Slugging bed. APPENDIX VII GASIFICATION RESULTS TABLE V I I - 1 - EXPERIMENTAL RESULTS FOR FORESTBURG COAI, GASIFICATION Bun 1 Run Ducat Ion (h) Average Coal Particle Size Coal Feed Rate Dry Ranis (»/-) Air Flovfal Steam Feed Rate U/s) Expanded Bed Depth ( •> Average Reactor Temperature (k> Gas Composition IX, Dry Basis v/V) Orosa Gas Heating Valuetb) ( M J / B 3 ) Air Velocity At Average Reactor Temperature*^  (»/s> Air To Coal Feed Stcan ' To Coal Feed Ratio'"1' -To Coal Ratio'e* Average Bed Carbon Cydone Catch Dry Basis (g/R coal) Cyclone Catch Carbon Carbon Elutriation Rate Gas Tar Content H2 CO co 2 " V Ratio Dry Baals (v/«> Dry Basts («/-) Dry Basis Content I Content I Dry Baals (R/B coal) 1 1 0.53 1.059 4.77x10"° 0.472 0.61 1033 5 . 5 4.0 0.6 13.S 76.4 1.37 0.92 5-39 0.446 0.526 0.44 0.167 62.31 0.104 _ 2* 1.5 0.53 1.060 4.68 0.377 0.61 1054-1127 7.8 8.0 0.5 8.1 75.5 2.09 0.95 5.29 0.356 0.436 - 0.102 55.14 0.056 -J* 1 0.53 1.060 ' 3.03 0.395 0.51 1031-1013 8.7 5.9 0.7 11.9 72.8 2.03 0.58 3.42 0.373 0.498 - 0.102 55.14 0.056 -w 1 0.53 2.077 6.57 0.472 0.61 1055-1034 6. 7 5.4 0.9 14.1 72. S 1.80 1.28 3.80 0.227 0.340 0.167 62.31 0.104 -5* l.S 0.95 0.983 4.19 0.542 0.51 1123-1088 9.0 7.9 0.7 11.4 71.0 .2.30 0.86 5.10 0.551 0.632 2.29 0.078 29.50 0.023 _ 6* o.s 0.95 0.983 5.42 0.322 0.36 1093-1059 10.9 9.8 1.1 9.1 69.1 2.91 0.55 3.31 0.328 0.411 - 0.078 79.50 0.023 -1.0 0.95 0.983 5 . 4 * 0.4 33 0.6L 1128 7. 7 8.9 0.5 12.2 70. r Z.19 1.15 6.63 0.4*0 0.524 - 0.078 29.50 0.023 -8* 1.0 0.95 0.983 5.46 0.729 0.61 1156 5 . 4 5.9 0.5 12.2 76.C 1.55 1.18 6.65 0. 742 0.825 0.50 0.078 29.50 0.023 -' 1.5 0.95 0. 983 7.42 1.66 7 0.61 1172 0.9 1.9 0.2 16.0 81.0 0.41 1.62 9.04 1.696 1.785 0.26 0.078 29-50 0.023 -LO* L.7 2.18 0. 271 2.75 0.283 0.43 1088 1.0 1.8 0.2 11.5 85.5 0.41 0.56 12.16 1.044 1.399 2.71 0.OB4 36.79 0.031 -IL 2.0 2.18 0. 715 4.54 0.217 0.61 1050 5 . 4 6.4 0.4 11.4 76.4 1.57 0.39 7.62 0.303 0.645 2.90 0.086 50.44 0.043 ' -12 1.7 2.18 0.798 3.49 0.283 0.51 1040 7.3 6.9 0.6 10.2 75.C 1.94 0.67 5.25 0.355 0.698 4.72 0.084 36. 79 0.031 -L3 2.0 2.18 0.798 5.48 0.283 0.51 1061 7.3 8.2 0.5 L0.8 73.2 2.06 1.08 8.23 0-355 0.698 3.96 0.084 36.79 0.031 -Li 1.0 2.18 1.185 6.51 0.542 0.61 1031 8.6 7.2 0.7 10.9 72.fi 2.17 1.25 6.59 0.457 0.797 11.71 0.093 57.99 0.05'. 0.841 L5 L.O 2.18 1.185 6.52 0.250 0.61 1031 7.2 7.2 0.7 10.7 74.2 2.00 1.25 6.59 0.211 0.550 9.89 0.093 57.99 0.054 0.841 16 L.5 2.18 1.273 4 . 74 0.0 0.69 1107 5.9 12.0 0.7 9.8 71.6 2.42 0.98 4.4<S 0.0 0.334 2.47 0.091 29.96 0.027 17 L.O 2.18 1.273 6.06 0.0 0.86 1165 4 . 1 8.3 0.6 13.8 73.2 1.72 1.32 5.70 0.0 0.337 0.97 0.091 29.96 0.027 -13 I.J 2.18 1.273 6.32 0.600 0.86 1145 5.1 6.3 0.6 15.9 72.1 1.60 1.35 5.94 0.628 0.966 0.74 0.091 29.96 0.027 -19 2.0 2.18 1.673 7.00 0.167 0.61 1049 6.4 8.0 0.7 12.3 72.5 2.00 1.37 5.01 0.100 0-4 36 10.36 0.086 50.44 0.043 -20 1.0 2.18 ' 1.699 5.89 0.258 0.50 1108 11.0 8.8 1.0 21.4 57.8 2.77 1.22 4.15 0.211 0.544 - - - -21* l.P 2.18 2.046 6.25 0.0 0.86 1080 8.4 11.3 0.7 13.0 66.6 2.64 1.26 3.66 0.0 0.332 2.69 0.091 29.96 0.027 -22 2.0 2.18 2.369 1.92 0.526 0.66 1026 L0.2 8.2 0.9 12.6 68.1 2.56 1.52 4.01 0.222 0.555 16.57-31. •7 0.093 58.48 0.0S4 23 2.0 4.06 0. 397 3.43 0.5 70 0.40 1108 2.8 3.4 0.3 13.4 BO.l 0.86 0.71 10.36 1.4)6 1.781 1.95 0.083 38.14 0.032 -21 5.0 6.06 0. 783 4.53 0.549 0.61 1096 9. 7 9.5 0.5 6.6 73.7 2.50 0.93 6.97 0. 702 1.038 4.44 0.093 30.01 0.029 0.274 25 1.3 4.06 1.223 6.36 0.885 0.61 1074 L0.1 8.9 0.6 13.4 67.0 2.52 1.27 6.23 0.724 1.059 4.43 0.083 38.14 0.032 -26 3.0 4.06 1.259 6.11 0.763 0.61 1065 LI.6 10.9 0.6 9.1 67.8 2.94 1.28 6.10 0.606 0.940 10.67 0.091 43.35 0.039 -27 0.7 4.06 1.516 6.67 0.0 0.86 1115 7.7 10.2 0.6 12.2 69.3 2.38 1.39 5.28 0-0 . 0.332 6-21 0.108 40.96 0.044 -23 1.0 4.06 1.516 6. 74 0.542 0.86 1085 8.5 8.3 0.6 13.8 68.8 2.25 1.36 5.33 0.358 0.690 4.31 0.108 40.96 0.044 -29 2.0 4.06 1.516 6.75 0.748 0.86 1079 8.7 7.6 0.7 13-9 69.1 2.23 1.35 5.32 0.493 0.826 3.51 0.108 40-96 0.044 -30 t.5 4.06 1.516 6.76 1.L65 0.36 1053 10.0 6.9 0.7 14.1 68. 3 2.30 1.33 S.34 0. 768 1.101 4.01 0.108 40.96 0.044 -31 2.6 4.06 1.642 7.75 0.950 0.61 1078 10.9 9.4 0.7 14.0 65.0 2.71 1.56 5.68 0.579 0.912 8.45 0.083 38.14 0.032 32 2.0 -4.76t0.0 1.059 4.27 0.0 0.6b 1125 6.3 10.1 0.6 10.6 72.4 2.20 0.89 4.34 0.0 0.323 1.52 0.187 ' 61.45 0.115 0'. 134 33 2.0 —i.76*0.0 1.059 4.42xl0"3 0.433 0.69 1055 9.9 7 . 4 0. 7 12.2 69.8 2.35 0.87 5.01 0.409 0.732 1.91 0.187 61.45 0.115 0.434 (*> Unstable condition (a) All gas flews at 294 K and 101.3 KPa (b) At Morth American standard conditions: 288.fi K and 101.6 R?a fc) Superficial -'elocity of the air and r.ot rhe gas (d) Ratio of iteam fed to dry con! fed (e) Ratio of total water fed (I.e. steam *• H-0 In coal +• HO ln air) to coal fed (see Appendix VIII) TABLE VII-2 EXPERIMENTAL RESULTS FOR FORESTBURG COAL GASIFICATION Run No. Run Durst ton Average C o i l P a r t i c l e S i r e (mm) Coal Feed Rate Dry Basis (g/s) Air Flow<»' Steam Teed la t e ' (g/s) Expanded Red Depth 0») Average Reactor Temp. (k) Gas Composition (Z) Dry Basis v/v) Cross Cas Heat log Value' 1'' (M.l/m3) Air Velocity at Average Reactor Temp>> (m/s) Air To Coal Feed' Ratio Dry Basis (w/w) Steam To Coal R a t i o " ' Dry Basin (w/w) Steam To Coal I . t l o ( e ' Dry Basis (w/w) Average Bed Carbon Content (I) Cyclone Catch Dry Basis (g/g coal) Cyclone Catch Carbon Content in Carbon E l t i t r i a t i o n Rate Dry Rasls ( R / R coal) Cas Tar Content (ll) (m3/*) " 2 CO CH, co 2 N 2 (g/m-1) 3'. 2.0 2.18 0.53'. 3.69x10"' 0.50S 0.68 1153 4.2 6.9 0.7 13.7 76.6 1.36 0.B0 R.12 0.917 0.045 2.68 0.227 8 1.66 0.1B5 -1 5 1.0 2.IB 0.5(18 5.33 0.35B 0.68 1171 5.5 6.6 0.7 11.7 75.5 1.72 1.16 10.86 0.609 0.645 21.6 0.266 85.17 0.227 -.16 1.4 2 . 1 8 1 .'.62 5.58 0.0 0.76 1132 4.6 7.0 1 . 2 12.4 74 . 8 1.85 1 . 1 8 4.58 0 . 0 0.020 21 . 7 0.272 8 0 . 3 3 0.219 3.307 37 1. . 5 2. IB 1 . ' i 6 2 5 . 6 9 0.455 0.76 1081 5.6 6.6 1.2 12.3 74.4 1.92 1.15 4.67 0.310 0 . 3 3 1 27 . 2 0.272 8 0 . 3 3 0.219 3.307 38* 1.1 2 . 1 8 1 .580 5.43 0.0 0.76 1165 6 . 8 7.1 2.0 12.1 72 . 0 2.43 1 . 1 8 4 . 1 1 0 . 0 0.01B 18.34 0.272 8 0 . 3 3 0.219 3.307 39 2.0 4.06 .1.462 5.68 0.0 0.76 1160 4.3 7.1 1.0 15 . 3 72.4 1 . 7 5 1 . 2 3 4.66 0 . 0 0.021 .16.2 0.276 61 . 8 8 0 . 171 0.913 '.0 1.5 1.06 1 .462 5.73 0.358 0.76 1107 6.2 7 . 0 1.0 15 . 7 70 . 0 1.97 1.18 4.71 0.245 0.261 16 . 2 0.276 61 . 8 8 0 . 1 7 1 0.913 COl.EhV W.COAl. 4 1 1.0 2. in 1.968 5.72xlO - 3 0.0 0.76 1096 7.1 7.8 1.9 1 0 . 9 7 2 , 2.51 1.17 3.48 0 . 0 0.022 26 . 8 0.139 71.86 0.100 2.122 (*) Unstable condition <n) A l l gas Flows at 294 K and 101.3 kPa (b) At North American Standard conditions: 2R8.6 K and 101-6 kPa (c) S u p e r f i c i a l v e l o c i t y of the a i r and not the gas (d) Ratio of steam fed to dry coal fed (e) Ratio of coal water fed ( i . e . steam + II 0 In coal + H_0 Ih a i r ) to coal fed (See Appendix VIJI). TABLE VII-3 BED CARBON CONTENT AS A FUNCTION OF TIME FOR THE GASIFICATION OF FORESTBURG COAL Run Average Coal A i r to Coal Time Bed Carbon # P a r t i c l e Size Ratio (Dry Basis) Content (mm) (w/w) (min) (% w/w) 22 6.44 14 & 15 2.18 6.59 46 8.56 84 12.86 109 11.71 30 1.06 17 & 18 2.18 5.82 65 0.87 225 0.74 22 16.57 22 2.18 4.01 52 20.54 82 25.16 107 31.27 24 8.54 56 3.15 84 2.01 24 4.06 6.97 114 4.10 144 4.54 189 2.28 279 6.48 6. . 37 8.72 67 8.24 26 4.06 6.10 97 9.36 127 12.61 182 8.94 219 10.45 62 3.08 92 3.94 137 4.15 27-30 4.06 5.30 167 3.86 227 4.63 287 3.99 287 4.21 50 0.94 32 & 33 -4.76 + 0.0 4.93 110 2.09 170 1.39 215 2.42 TABLE VII-4 BED CARBON CONTENT AS A FUNCTION OF TIME FOR THE GASIFICATION OF SUKUNKA COAL Run # Average Coal P a r t i c l e Size (mm) A i r to Coal Ratio (Dry Basis) (w/w) Time (min) Bed Carbon Content (%, v/v) 35 19.66 36 & 37 2.18 4.63 70 23.71 125 27.21 39 & 40 4.06 4.69 52 97 157 187 21.79 10.20 15.56 16.87 TABLE VII-5 . GAS COMPOSITION AND GROSS CALORIFIC VALUES FROM FORESTBURG COAL GASIFICATION CORRECTED FOR THE INTRODUCTION OF PURGING NITROGEN Run Gas Composition (%; v/v) Gross # (Dry Basis) C a l o r i f i c Value H2 CO CH4. co 2 N2 (MJ/m3) 1 5.77 4.20 0.63 14.17 75.23 1.44 2* 8.19 8.40 0.53 8.61 74.27 2.19 3* 9.34 6.33 0.75 12.77 70.81 2.18 4* 6.93 5.59 0.93 14.58 71.97 1.86 5* 9.47 8.31 0.74 11. 99 69.49 2.42 6* 11.33 10.19 1.14 9.46 67.88 3.02 7* 8.01 9.26 0.52 12.69 69.52 2.28 8* 5.63 6.15 0.52 12.73 74.97 1.62 9 0.93 1.96 0.21 16.54 80.36 0.42 10 1.10 1.97 0.22 12.61 84.1 0.45 11 5.68 6.73 0.42 11.99 75.18 1.65 12 7.78 7.35 0.64 10.87 73.36 2.07 13 7.60 8.54 0.52 11.25 72.09 2.15 14 8.90 7.45 0.72 11.28 71.65 2.24 15 7.45 7.45 0.72 11.08 73.30 2.07 16 6.17 12.56 0.73 10.26 70.28 2.53 17 4.25 8.61 • 0.62 14.32 72.20 1.78 18 5.28 6.52 0.62 16.46 71.12 1.66 19 6.61 8.26 0.72 12.69 71.72 2.06 21* 8.68 11.67 0.72 13.43 65.50 2.73 22 10.47 8.42 0.92 12.93 67.26 2.63 23 3.05 3.65 , 0.32 14.37 78.66 0.92 24 10.19 9.98 0.53 6.93 72.37 2.63 25 10.45 9.21 0.62 13.86 65.86 2.61 26 12.00 11.27 0.62 9.41 66.70 3.04 27 7.95 10.53 0.62 12.59 68.31 2.46 28 8.77 8.56 0.62 14.24 67.81 2.32 29 8.98 7.84 0.72 14.34 68.12 2.30 30 10.31 7.12 0.72 14.54 67.31 2.37 31 11.18 9.64 0.72 14.36 64.10 2.78 32 6.63 10.63 0.63 11.15 70.96 2.31 33 10.42 7.79 0.74 12.84 68.21 - 2.47 * Unstable condition TABLE VI1-6 GAS COMPOSITION AND GROSS CALORIFIC VALUES FROM CAKING COAL GASIFICATION CORRECTED FOR THE INTRODUCTION OF PURGING NITROGEN Run # Gas Composition (Dry Bas (%, v/v) i s ) Gross C a l o r i f i c Value H 2 CO CH4. c o 2 N 2 (MJ/m3) 35 5.85 7.02 0.74 12.44 73.95 1.83 36 4.89 7.44 1.28 13.18 73.21 1.97 37 5.94 7.00 1.27 13.05 72.74 2.04 38* 7.22 7.54 2.12 12.85 70.27 2.58 39 4.56 7.52 1.06 16.21 70.65 1.85 40 6.55 7.40 1.06 16.60 68.39 2.08 41 7.52 8.26 2.01 11.55 70.66 2.66 * Unstable condition Runs 35-40: Sukunka Coal Run 41: Coleman Coal 168 TABLE VII-7 METHANE PRODUCTION FROM THE GASIFICATION OF FORESTBURG COAL Run Coal Dry Coal Volume of Gas CH 4 CH 4 # P a r t i c l e Feed Rate Produced ( a) Concentration Production Size In Gas (mm) (g/s) (m3/s) % (v/v) (mg/s) 10 2.18 0.271 2.80x10-3 0.2 3.71 11 2.18 0.715 4.97 0.4 13.19 12 2.18 0.798 3.97 0.6 15.80 13 2.18 0.798 6.19 0.5 20.53 14 2.18 1.185 7.35 0.7 34.13 15 2.18 1.185 7.21 0. 7 33.48 16 2.18 1.273 5.53 0.7 25.68 17 2.18 1.273 6.82 0.6 27.14 18 2.18 1.273 7.20 0.6 28.66 19 2.18 1.673 7.90 0.7 36.68 21 2.18 2.046 7.73 0.7 35.89 22 2.18 2.369 9.48 0.9 56.60 23 4.06 0.397 3.65 0.3 7.26 24 4.06 0.782 5.14 0.5 17.05 25 4.06 1.223 7.37 0.6 29.33 26 4.06 1.259 7.42 0.6 29.13 27 4.06 1.516 7.89 0.6 31.40 28 4.06 1.516 8.03 0.6 31.96 29 4.06 1.516 8.00 0.7 37.15 30 4.06 1.516 8.11 0.7 37.66 31 4.06 1.642 9.71 0.7 45.09 32 -4.76+0.0 1.059 4.96 0.6 19.74 33 -4.76+0.0 1.059 5.31xl0 - 3 0.7 24.66 (a) Calculated from mass balance,see Table VIII-1 TABLE VII-8 METHANE PRODUCTION FROM THE GASIFICATION OF CAKING COALS Run P a r t i c l e Size (mm) Dry Coal Feed Rate (g/s) Volume of Gas Produced ( a) (nvVs) CH, 4 Concentration In Gas % (v/v) CH. Production (mg/s) 34 2.18 0.544 - ' 0.7 -"35 2.18 0.588 6.25 x l 0 ~ 3 0.7 29.16 36 2.18 1.462 6.34 1.2 50.47 37 2.18 1.462 6.49 1.2 51.66 38 2.18 1.580 6.43 2.0 85.31 39 4.06 1.462 6.66 1.0 44.18 40 4.06 1.462 6.94 1,0 46.04 41 2.18 1.968 6.70xl0" 3 1.9 84.4 (a) Cal c u l a t e d from mass balance, see Table VIII-1 Runs 34-40: Sukunka Coal R-n 41: Coleman Coal TABLE VII-9 HYDROGEN AND CARBON MONOXIDE PRODUCTION FROM THE GASIFICATION OF FORESTBURG COAL AT A STEAM TO COAL RATIO OF 0.32-0.34 Run Coal P a r t i c l e Size (mm) Dry Coal Feed Rate (g/s) Volume of Gas Produced( a) (m 3/s) H^ Cone. i n Gas % (v/v) H 2 Production (mg/s) CO Cone. In Gas % (v/v) CO Production (g/s) 16 2.18 1.273 5.53xl0 - 3 5.9 27.05 12.0 0.77 17 2.18 1.273 6.82 4.1 23.19 8.3 0.66 21 2.18 2.046 7.73 8.4 53.84 11.3 1.01 27 4.06 1.516 7.89 7.7 50.38 10.2 0.93 32 4.76-0.0 1.059 4.96xl0~ 3 6.3 25.91 10.1 . 0.58 (a) Calculated from mass balance, see Table VIII-1 APPENDIX VIII MASS BALANCES 172 MASS BALANCES GENERAL PROCEDURE: The procedure consists i n c a l c u l a t i n g the dry volume of gas produced during g a s i f i c a t i o n through a nitrogen balance, and determining the volume of water vapour coming out of the reactor by comparing the dry volume of gas with the measured volume of wet gas. Over a l l mass balances and mass balance for the d i f f e r e n t elements can then be c a r r i e d out. The calcu-l a t i o n procedure i s deta i l e d below with s p e c i f i c a p p l i c a t i o n to Run No. 38. Mass balances f o r a l l g a s i f i c a t i o n runs are tabulated i n Tables VIII-1- VII-4, Appendix VIII. SAMPLE CALCULATION RUN NO. 38 Data For Run No. 38 ' Basis : 1 second Inputs:: Dry coal feed rate.'vr^ = 1.580 (g/s) of 2.18 mm Sukunka coal Coal a n a l y s i s : See Table 2 A i r flow: F= 5.43 1 0 _ 3 std (m3/s) Steam feed rate: S= 0.0 (g/s) Purging Nitrogen: N= 0.435 (g/s) Outputs: Wet Gas: O r i f i c e upstream pressure: P= 102.83 kPa O r i f i c e upstream temperature: T= 637 K Pressure drop across o r i f i c e p l a t e : iSP= 138.01 mm H^ O Dry Gas composition: H 2 = 6.8% (v/v) CO = 7.1% CH. = 2.0% 4 C0 2 = 12.1% N 2 = 72.0% 173 Molecular weight dry gas :M.W.^ r^  = 27.83 (g/mol) Gas composition by elements, dry weight basi s : 0 = 17.98% N = 72.19% H = 0.77% C = 9.11% 3 Gas tar content: 3.307 (g/std m ) dry gas Cyclone catch: 0.272 g/g dry coal fed Cyclone catch carbon content: 80.33% Calculations: 1.- Mass balance: Nitrogen input = m = A i r flow x J a i r x (% N„) . (w/w)/100 + coal feed rate 2 a i r x (%N 2in coal)(w/w)/100 + purging N 2 input = 5.43 x 10~ 3 x 28.91 x 0.7553 + 1.580 x 0.0126 + 0.435 24.12 x 10 3 m = 5.372 (g/s) 2.- T o t a l dry gas flow (V) out of the reactor, based on nitrogen input: v = m x 100.0 5.372 x 24.12 x 1 0 _ 3 (%N„v/v), x j 3 N „ 0 .72x28 2 dry gas 2 = 6.43 x 10~ 3 std (m3/s) 3.- T o t a l water input (W): W = water i n a i r + steam + water i n coal = F x j * . x (% H o0 w/w) . /100 + S + a i r 2 a i r r (%H o0)coal c z (1- %H 20)coal - 5.43 x 10 3 x 28.91 x (0.0025) + 0 + -1.580 x 0.0082 24.12 x 10 W = 0.029 (g/s) _3 (1-0-.0082 Wet gas flow: Since c a l c u l a t i n g the wet gas flow from the o r i f i c e p l ate pressure drop requires knowing the molecular weight of the wet gas (see Appendix II) and t h i s i s not known u n t i l the water content of the gas i s known, an i t e r a t i v e method i s required. Assume f i r s t that a l l water entering the reactor i s unreacted, and c a l c u l a t e the steam flow (W) at the conditions p r e v a i l i n g at the o r i f i c e p l a t e . 3 Let v g = s p e c i f i c volume of steam (m /g) = 2.898 x 10" 3 (m3/g) at temperature and pressure measured at the o r i f i c e p l a t e , then: W = W x v (m3/s) s = 0.029 x 2.898 x 1 0 - 3 = 0.084 x 10~ 3 (m3/s) Calculate now molecular weight of gas ( M' W- w e t) o n a w e t basis. F i r s t c a l c u l a t e volume f r a c t i o n of steam (s) i n wet gas. Since at the o r i f i c e plate conditions the dry volume of gas ( V ) i s given by: , = V x T x 101320 294 P 6.43 x 10" 3 x 637 101320 ,„ . n3 , 3 , * = — x = 13.73 x 10 (m /s) 294 102834 then s = -K- = ^084 = 0.0061 W + V 0.084 + 13.73 M.W. = M.W. , (1-s) + 18 s (g/mol) wet dry ° = 27.93 (1-0.0061) + 18 x 0.006 = 27.87 (g/mol) 175 The flow (Q) measured by the orifice plate is given by (Appendix II) Q = 0 . 0 8 4 \ j A P x T ( m ? / s ) P . M . . : x P wet if. = 0 . 0 8 4 \ | 1 3 8 . 0 1 x 6 3 7 8 7 x 1 0 2 8 3 4 = 1 4 . 7 1 x 1 0 3 ( m 3 / s ) Q f - W + V = 1 3 . 8 l 7 x l 0 _ 3 ( m 3 / s ) S i n c e t h e m e a s u r e d v a l u e o f t h e w e t g a s (Q) i s d i f f e r e n t f r o m t h e c a l c u l a t e d o n e u n d e r t h e a s s u m p t i o n t h a t t h e w a t e r d o e s n o t r e a c t u s e t h e n e w v a l u e o f W' = Q ' - V a n d i t e r a t e u n t i l Q = V + W' - 3 3 i n t h i s c a s e , Q = 1 4 . 9 1 x 1 0 (m / s ) T h e r e f o r e , t h e m a s s o f w a t e r c o m i n g o u t o f t h e r e a c t o r i s i i n Q - V 1 4 . 9 1 - 1 3 . 7 3 . / n _ , . . H O = = = 0 . 4 0 7 ( g / s ) v s 2 . 8 9 8 i . e w a t e r i s p r o d u c e d i n t h e r e a c t o r ( w a t e r e n t e r i n g t h e r e a c t o r i s W = 0 . 0 2 9 ( g / s ) ) . 176 5.- O v e r a l l mass balance: Coal = r Inputs <  A i r = 24 F x Water = W Purging N 2 a i r q /•» m' 3 28.91 5.43 x 10 x 24.12 x 10 -3 = 1.580 (g/s) = 6.508 (g/s) = 0.029 (g/s) = 0.435 (g/s T o t a l inputs = 8.552 (g/s) Dry gas = V x. 6.43 x 10 x 27.93 dry gas 24.12 x 10 -3 Outputs Water = H 20 E l u t r i a t e d s o l i d s = r x cyclone catch c = 1.58 x 0.272 Tar = gas t a r content x V o l . dry gas = 3.307 x 6.43 x 10 -3 = 7.446.(g/s) = 0.407:.(g/s) = 0.430 (g/s) = 0.021 (g/s) T o t a l outputs = 8.304 (g/s) Hydrogen Balance: Inputs 1 Water bound = W x 2/18 = 0.029/9 In dry coal r x (% H„ i n coal w/w)/100 c I = 1.58 x 0.0445 To t a l inputs = 0.003 (g/s) = 0.070 (g/s) = 0.073 (g/s) Outputs i. Water bound = H 20 /9 = 0.407/9 In gas = V x - x (% H„in gas w/w)/100 gas z 6.43 x 10" 3 x 27.93 x 0.0077 24.12 x 10 -3 = 0.045 (g/s) = ' 0.057 (g/s) To t a l outputs = 0.102 (g/s) 178 (^Balance: Water bound = W x 16/18 = 0.029 x 16 18 = 0.026 (g/s) Inputs / In dry c o a l = r £ ( % 0^ i n c o a l w/w)/100 = 1.58 x 0.0256 In a i r = F x F . x (% 0 o i n a i r (w/w)/100 a i r 2 5.43 x 10" 3x 28.91 x 0.2314 24.12 x 10 -3 = 0.040 (g/s) = 1.506 (g/s) T o t a l inputs = 1.572 (g/s) Water bound = H 20 x 16/10 Outputs, < 0.407 x 16 18 = 0.362 (g/s) In gas: V x f , x (% 0 o w/w)/100 ° dry gas 2 6.43 x 10" 3 x 27.93 x 0.1793 24.12 x 10 -3 = 1.335 (g/s) T o t a l outputs = 1.697 (g/s) 179 8.- Carbon Balance: Input: = r c x (%C i n coal w/w)/100 = 1.58 x 0.8031 1.269 (g/s) i n gas Output < = V x f x (% C i n gas, w/w)/100 gas b 6.43 x 10" 3 x 27.93 x 0.0911 24.12 x 10 -3 = 0.678 (g/s) In e l u t r i a t e d s o l i d s = r x (carbon e l u t r i a t i o n rate) c = 1.58 x 0.219 = 0.346 (g/s) To t a l outputs 1.024 (g/s) 9.- Ash Balances: Input = r c x (% Ash i n coal w/w)/100 = 1.580 x_0.1081 = 0.171 (g/s) Output = r c x Ash e l u t r i a t i o n rate = = 1.580 x (0.272-0.219) = 0.085 (g/s) TABLE VIII-1 - OVERALL MASS BALANCE I n p u t s O r i f i c e P l a t e Measurements Dry Gas Flow Outputs D i f f e r e n c e Run Dry c o a l A i r Water P u r g i n g N2 T o t a l A P P T Wet gas Std a t T,P Water Dry gas S o l i d s Tar T o t a l o u t p u t - i n p u t No. (g/s) (g/s) (g/s) (g/s) (g/s) (mm H 20) (Pa) (K) f l o w ( T , P ) (m 3/s) (m 3/s) (g/s) (g/s) (g/s) (g/s) (g/s) i n p u t (m 3/s). (%) 1 1 059 5 717 0 557 0 285 7 618 81 28 102168 527 l l . l O x l O - 3 . 5 .21x]0~ 3 11.02x10 3 0.027 6 190 0 177 - 6 394 - 16 1 2 1 060 5 609 0 462 0 285 7 416 93 98 103150 638 12 53 5 18 11 04 0.512 5 847 0 108 - 6 467 - 12 8 3 1 060 3 632 0 475 0 285 5 452 36 83 102325 575 7 35 3 60 6 98 0.141 4 113 0 108 - 4 362 - 20 0 4 2 077 7 875 0 631 0 285 10 868 177 80 104233 703 17 27 7 40 17 21 0.019 8 720 0 347 - 9 086 - 16 4 5 0 983 5 022 0 621 0 285 6 911 22 86 104104 675 3 12 83 4 97 11 11 0.559 5 646 0 077 - 6 282 - 9 1 6 0 983 6 496 0 404 0 285 8 168 48 26 98680 760 8 67 6 49 17 23 0? 7 128 0 077 - 7 205 - 11 8 7 0 983 6 520 0 515 0 285 8 303 93 98 104934 719 12 28 6 37 15 04 . 07 7 366 0 077 - 7 443 - 10 4 8 0 983 6 544 0 811 0 285 8 623 86 36 104934 716 12 30 5 94 13 97 0? 7 016 0 077 - 7 093 - 17 7 9 0 983 8 894 1 755 0 285 11 917 372 75 101728 815 8 27 90 7 46 20 62 1.962 9 371 0 077 11 410 - 4 3 10 0 271 3 296 0 379 0 285 4 231 26 25 103240 496 3 5 70 2 80 4 64 0.470 3 432 0 023 - 3 925 - 7 2 11 0 715 5 442 0 461 0 285 6 903 83 82 104080 588 11 12 4 97 9 68 0.538 5 848 0 062 - 6 448 - 6 6 12 0 798 4 183 0 55.1 0 285 5 817 50 55 104037 533 8 37 3 97 7 01 0.561 4 553 0 067 - 5 181 - 10 9 13 0 798 6 568 0 557 0 285 8 208 140 97 105690 647 15 19 6 19 13 06 0.723 7 127 0 067 - 7 917 - 3 5 14 1 185 7 803 0 944 0 285 10 217 218 44 104843 692 7 19 81 7 35 16 74 0.974 8 356 0 110 0.006 9 446 - 7 5 15 1 185 7 815 0 652 0 285 9 937 199 81 104572 676 18 57 7 21 16 06 0.813 8 298 0 .110 0.006 9 227 - 7 I 16 1 273 5 681 0 425 0 285 7 664 159 17 105136 702 7 17 20 5 53 12 74 1.395 6 41 0 116 - 7 921 + 3 4 17 1 273 7 264 0 429 0 285 9 251 232 75 106942 755 7 20 79 6 82 16 60 1.220 8 222 0 116 - 9 558 + 3 3 18 1 273 7 575 1 230 0 285 10 363 355 6 107394 793 3 26 91 7 20 18 33 2.376 8 701 0 116 - 11 193 + 8 0 19 1 673 8 390 0 730 0 285 11 078 266 70 106866 703 19 64 7 90 17 91 0.540 9 233 0 144 - 9 917 - 10 5 21 2 046 7 491 0 679 0 285 10 501 294 0 106942 768 23 97 7 73 19 02 1.423 8 913 0 186 - 10 522 + 0 2 22 2 369 9 493 1 314 0 285 13 461 312 13 107044 724 7 25 12 9 48 22 11 0.913 10 710 0 220 - 11 843 - 12 0 23 0 397 4 111 0 707 0 285 5 500 40 01 102151 558 3 7 34 3 65 6 87 0.184 4 446 0 033 - 4 663 - 15 2 24 0 782 5 43 0 812 0 285 7 309 100 33 102913 623 3 13 12 5 14 10 73 0.841 5 640 0 073 - 6 555 - 10 3 25 1 223 7 623 1 295 0 285 10 426 198 12 105227 704 3 18 77 7 37 17 00 0.552 8 386 0 102 - 9 040 - 13 3 26 1 259 7 323 1 183 0 285 10 050 176 35 104253 674 6 17 65 7 42 16 55 0.359 8 112 0 115 - 8 586 - 14 6 27 1 516 7 995 0 504 0 285 10 300 306 25 105791 761 5 24 52 7 89 19 57 1.427 9 •117 0 164 - 10 708 + 4 0 28 1 516 8 079 1 046 0 285 10 926 355 60 107009 766 3 26 65 8 03 19 62 2.014 9 295 0 164 - 11 473 + 5 0 29 1 516 8 091 1 252 0 285 11 144 355 60 106694 790 3 27 01 8 00 20 42 1.872 9 250 0 164 - 11 286 + 1 3 30 1 516 8 103 1 669 0 285 11 573 406 4 107146 787 5 29 19 8 11 20 54 2.410 9 278 0 164 - 11 347 - 2 0 31 1 642 9 289 1 497 0 285 12 713 314 96 106672 758 2 23 79 9 71 24 70 0? 10 999 0 136 - 11 135 - 12 4 32 1 059 5 118 0 342 0 285 6 804 141 39 103669 665 3 16 06 4 96 10 97 1.682 5 755 0 122 0.002 7 561 + 11 1 33 1 059 5 298 0 775 0 285 7 417 188 65 106028 682 0 18 98 5 31 11 77 2.323 6 009 0 122 0.002 8 456 + 14 0 36 1 462 6 688 0 029 0 435 8 614 209 55 103607 706 5 19 68 6 34 14 92 1,479 7 529 0 398 0.021 9 429 + 9 5 37 1 462 6 820 0 484 0 435 9 201 260 35 105622 735 5 22 54 6 49 15 58 2.080 7 641 0 398 0.021 10 140 + 10 2 38 1 580 6 508 0 029 .0 435 8 552 138 01 102834 637 14 91 6 43 13 73 0.407 7 445 0 430 0.021 8 303 - 2 9 39 1 462 6 808 0 031 0 435 8 736 188 81 104945 719 3 18 14 6 66 15 73 0.736 8 073 0 404 0.006 9 219 + 5 5 40 1 462 6 868 0 389 0 435 9 154 226 91 104775 764 3 20 79 6 94 17 45 0.960 8 273 0 404 0.006 9 643 + 5 3 41 1 968 6 856 0 044 0 435 9 303 190 50 104775 714 0 18.70xlO~ 3 6 . 7 0 x l 0 ~ 3 1 5 . 7 4 x l 0 - 3 0.914 7 703 0 274 0.008 8 899 - 4 3 TABLE VIII-2 - HYDROGEN MASS BALANCES Inputs Outputs Run Water In dry Total Water In dry Total No. bound coal bound gases (g/s) (g/s): (g/s) (g/s) (g/s) (g/s) 1 0.062 0.045 0.107 0.03 0.029 0.032 2 0.051. 0.045 0.096 0.057 0.038 0.095 3 0.053 0.045 0.098 0.016 0.030 0.046 4 0.070 0.088 0.158 0.002 0.054 0.561 5 0.069 0.042 0.111 0.062 0.043 0.105 6 0.036 0.042 0.078 - 0.071 0.071 7 0.057 0.042 0.099 - 0.046 0.046 8 0.090 0.042 0.132 - 0.032 0.032 9 0.195 0.042 0.237 0.218 0.008 0.226 10 0.042 0.012 0.054 0.052 0.003 0.055 11 0.051 0.030 0.081 0.060 0.026 0.086 12 0.061 0.034 0.095 0.062 0.028 0.090 13 0.062 0.034 0.096 0.080 0.043 0.123 14 0.105 0.050 0.155 0.108 0.061 0.169 15 0.072 0.050 0.122 0.090 0.051 0.141 16 0.047 0.054 0.101 0.155 0.033 0.188 17 0.048 0.054 0.102 0.136 0.030 0.166 18 0.137 0.054 0.191 0.264 0.037 0.301 19 0.081 0.071 0.152 0.060 0.051 0.111 21 0.075 0.087 0.162 0.158 0.063 0.221 22 0.146 0.101 0.247 0.101 0.094 0.195 23 0.079 0.017 0.096 0.020 0.010 0.030 24 0.090 0.033 0.123 0.093 0.046 0.139 25 0.144 0.052 0.196 0.061 0.069 0.130 26 0.131 0.054 0.185 0.040 0.079 0.119 27 0.056 0.064 0.120 0.159 0.058 0.217 28 0.116 0.064 0.180 0.224 0.065 0.289 29 0.139 0.064 0.203 0.208 0.053 .0.261 30 0.185 0.064 0.249 0.268 0.076 0.344 31 0.166 0.070 0.236 0.0 0.099 0.099 32 0.038 0.045 0.083 0.187 0.031 0.218 33 0.086 0.045 0.131 0.258 0.050 0.308 36 0.003 0.065 0.068 0.164 0.037 0.201 37 0.054 0.065 0.119 0.231 0.043 0.274 38 0.003 0.070 0.073 0.045 0.057 0.102 39 0.003 0.065 0.068 0.082 0.035 0.117 40 0.043 0.065 0.108 0.107 0.047 0.154 41 0.005 0.087 0.092 0.102 0.061 0.163 TABLE VIII-3 - OXYGEN MASS BALANCES Inputs . ;0utputs Run Water .': . In dry In a i r Total Water In dry Tota l No. bound.. . . coal bound gas (g/s) (g/s) (g/s) (g/s) (g/s) (g/s) (g/s) 1 0.495 0.172 1.323 1.990 0.024 1.072 1.096 2 0.411 0.173 1.298 1.882 0.455 0.838 1.293 3 0.422 0.173 0.841 1.436 0.125 0.709 0.834 4 0.561 0.338 1.823 2.722 0.017 1.651 1.668 5 0.552 0.160 1.162 1.874 0.497 1.012 1.509 6 0.359 0.160 1.504 2.023 - 1.205 1.205 7 0.458 0.160 1.767 2.385 1.407 1.407 8 0.721 0.160 1.648 2.529 - 1.194 1.194-9 1.560 0.160 2.058 3.778 1.744 1.677 3.421 10 0.337 0.044 0.763 1.144 0.418 0.461 0.879 11 0.410 0.116 1.259 1.786 0.478 0.963 1.442 12 0.490 0.130 0.068 1.588 0.499 0.719 1.218 13 0.495 0.130 1.520 2.145 0.643 1.224 1.867 14 0.839 0.193 1.806 2.838 0.866 1.414 2.280 15 0.580 1.193 1.809 2.582 0.723 . 1.368 2.091 16 0.378 0.207 1.315 1.900 1.240 1.161 2.401 17 0.381 2.207 1.681 2.269 1.084 1.628 2.712 18 1.093 0.207 1.753 3.053 2.112 1.819 3.931 19 0.647 0.272 1.942 2.861 0.480 1.708 2.188 21 0.604 0.333 1.734 2.671 1.265 1.912 3.177 22 1.168 0.386 2.197 3.751 0.812 2.100 2.912 23 0.628 0.065 0.951 1.644 0.164 0.731 0.265 24 0.722 0.127 1.257 2.106 0.748 0.774 1.522 25 1.151. 0.199 1.764 3.114 0.491 1.745 2.236 26 1.052 0.205 1.695 2.952 0.319 1.432 1.751 27 0.448 0.247 1.850 2.545 1.268 1.807 3.075 28 0.930 0.247 1.870 3.047 1.790 1.909 3.699 29 1.113 0.247 1.872 3.232 1.664 1.882 3.546 30 1.484 0.247 1.875 3.606 2.142 1.892 4.034 31 1.331 0.267 2.150 3.748 0.0 2.409 2.409 32 0.304 0.172 1.184 1.660 1.495 1.030 2.525 33 0.689 0.172 1.226 2.087 2.065 1.120 3.185 36 0.026 0.037 1.548 1.611 1.315 1.337 2.652 37 0.430 0.037 1.578 2.045 1.849 1.343 3.192 38 0.026 0.040 1.506 1.572 0.367 1.335 1.697 39 0.028 0.033 1.576 1.637 0.654 1.665 2.319 40 0.346 0.033 1.590 1.969 0.853 1.768 2.621 41 0.039 0.097 1.584 1.720 0.812 1.316 2.128 TABLE V I I I - 4 - CARBON AND ASH MASS BALANCES ;Carbon Ash Run Inputs Outputs Input Output No. Coal Gas Solids Total Coal Solids (g/s) (g/s) (g/s) (g/s) . (g/s) (g/s) 1 0 . 7 0 7 0 . 4 6 9 0 . 1 1 0 0 . 5 7 9 0 . 1 5 5 0 . 0 6 7 2 0 . 7 0 8 0 . 4 3 0 0 . 0 5 9 0 . 4 8 9 0 . 1 5 5 0 . 0 4 9 3 0 . 7 0 8 0 . 3 3 2 0 . 0 5 9 0 . 3 9 1 0 . 1 5 5 0 . 0 4 9 4 1 . 3 8 7 0 . 7 5 2 0 . 2 1 6 0 . 9 6 8 0 . 3 0 3 0 . 1 3 1 5 0 . 6 5 7 0 . 4 9 5 0 . 0 2 3 0 . 5 1 8 0 . 1 2 1 0 . 0 5 4 6 0 . 6 5 7 0 . 6 4 6 0 . 0 2 3 0 . 6 6 9 0 . 1 2 1 0 . 0 5 4 7 0 . 6 5 7 0 . 6 8 4 0 . 0 2 3 0 . 7 0 7 0 . 1 2 1 0 . 0 5 4 8 0 . 6 5 7 0 . 5 5 0 0 . 0 2 3 0 . 5 7 3 0 . 1 2 1 0 . 0 5 4 9 0 . 6 5 7 0 . 6 7 2 0 . 0 2 3 0 . 6 9 5 0 . 1 2 1 0 . 0 5 4 1 0 0 . 1 8 1 0 . 1 8 8 0 . 0 0 8 0 . 1 9 6 0 . 0 2 6 0 . 0 1 4 1 1 0 . 4 7 8 0 . 4 5 0 0 . 0 3 1 0 . 4 8 1 0 . 0 6 i Q 0 . 0 3 0 1 2 0 . 5 3 3 0 . 3 5 0 0 . 0 2 5 0 . 3 7 5 0 . 0 7 6 0 . 0 4 2 1 3 0 . 5 3 3 0 . 6 0 1 0 . 0 2 5 0 . 6 2 6 0 . 0 7 6 0 . 0 4 2 14 0 . 7 9 1 0 . 6 8 8 0 . 0 6 4 0 . 7 5 2 0 . 1 1 3 0 . 0 4 6 15 0 . 7 9 1 0 . 6 6 7 0 . 0 6 4 0 . 7 3 1 0 . 1 1 3 0 . 0 4 6 1 6 0 . 8 5 0 0 . 6 1 9 0 . 0 3 5 0 . 6 5 4 0 . 1 2 1 0 . 0 8 1 17 0 . 8 5 0 0 . 7 7 0 . 0 . 0 3 5 0 . 8 0 5 0 . 1 2 1 0 . 0 8 1 18 0 . 8 5 0 0 . 8 1 7 0 . 0 3 5 0 . 8 5 2 0 . 1 2 1 0 . 0 8 1 1 9 1 . 1 1 7 0 . 8 2 5 0 . 0 7 2 0 . 8 9 7 0 . 1 5 9 0 . 7 1 4 2 1 1 . 3 6 7 0 . 9 6 2 0 . 0 5 5 1 . 0 1 7 0 . 1 9 5 0 . 1 3 0 2 2 1 . 5 8 2 1 . 0 2 3 0 . 1 2 8 1 . 1 5 1 0 . 2 2 6 0 . 0 9 1 23 0 . 2 6 5 0 . 3 1 0 0 . 0 1 3 0 . 3 2 3 0 . 0 3 5 0 . 0 2 0 24 0 . 5 2 2 0 . 4 2 5 0 . 0 2 2 0 . 4 4 7 0 . 0 6 9 0 . 0 5 1 25 0 . 8 1 7 0 . 8 3 9 0 . 0 3 9 0 . 8 7 8 0 . 1 0 7 0 . 0 6 3 . 2 6 0 . 8 4 1 0 . 7 6 1 0.;'049 0 . 8 1 0 0 . 1 1 0 0 . 0 6 5 27 1 . 0 1 3 0 . 9 0 3 0 . 0 6 7 0 . 9 7 0 0 . 1 3 3 0 . 0 9 7 2 8 1 . 0 1 3 0 . 9 0 7 0 . 0 6 7 0 . 9 7 4 0 . 1 3 3 0 . 0 9 7 2 9 1 . 0 1 3 0 . 8 8 3 0 . 0 6 7 0 . 9 5 0 0 . 1 3 3 0 . 0 9 7 3 0 1 . 0 1 3 0 . 8 7 5 0 . 0 6 7 0 . 9 4 2 0 . 1 3 3 0 . 0 9 7 3 1 1 . 0 9 7 1 . 1 6 4 0 . 0 5 3 1 . 2 1 7 0 . 1 4 4 0 . 0 8 4 3 2 0 . 7 0 7 0 . 5 2 5 0 . 1 2 2 0 . 6 4 7 0 . 1 1 1 0 . 0 7 6 3 3 0 . 7 0 7 0 . 5 3 7 0 . 1 2 2 0 . 6 5 9 0 . 1 1 1 0 . 0 7 6 3 6 1 . 1 7 4 0 . 6 5 0 0 . 3 2 0 0 . 9 7 0 0 . 1 5 8 0 . 0 7 7 37 1 . 1 7 4 0 . 6 4 9 0 . 3 2 0 0 . 9 6 9 0 . 1 5 8 0 . 0 7 7 3 8 1 . 2 6 9 0 . 6 7 8 0 . 3 4 6 1 . 0 2 4 0 . 1 7 1 0 . 0 8 4 3 9 1 . 1 5 6 0 . 7 7 5 0 . 2 5 0 1 . 0 2 5 0 . 1 8 2 0 . 1 5 4 4 0 1 . 1 5 6 0 . 8 1 8 0 . 2 5 0 1 . 0 6 8 0 . 1 8 2 0 . 1 5 4 4 1 1 . 5 4 1 0 . 6 8 9 0 . 1 9 7 0 . 8 8 4 0 . 2 1 3 0 . 0 0 7 APPENDIX IX ENERGY BALANCES 185 ENERGY BALANCES GENERAL PROCEDURE Consider the f l u i d i z e d bed reactor as a "black box" i n sketch below: INPUTS COAL AIR STEAM R E A C T O R S Y S T E M OUTPUTS DRY GAS STEAM ELUTRIATED CARBON HEAT LOSSES The enthalpy balance i s car r i e d out i n the basis of one second of operation and with a reference temperature of 288.6 K (60°F = North American standard reference temperature f o r combustion processes). The heat content of inputs and outputs i s calculated as below and the heat losses are calculated by dif f e r e n c e . Inputs: 1) Heat content of c o a l : H .| = Dry coal feed rate (g/s) x heat value of coal (KJ/g) (=) KJ/s coal Heat value of coal obtained from analysis (Table 7) Dry coal feed measured experimentally. 2) Heat content of a i r H . = F a i r a i r J* 1 r e f C p d T = F a i r C p a i r ( T l ~ T r e f } X 4 * 1 8 5 X 1 0 ( = ) K J / s where: F . = molar a i r feed rate a i r Volumetric a i r feed rate (std m/s) 24.12 x 10 3 (std. m3/mol) _3 4.185 x 10 = conversion factor (=) K J / c a l = a i r feed temperature ( K ) , measured experimentally Cp . = mean molal heat capacity of a i r (cal/mol K) axr Cp . i s calculated as the arithmetic mean of Cp . evaluated at a i r r a i r (32) temperatures and ^ T e f t by the expression : Cp . = 6.8085 + 0.0008351 T - 39323.15 T~ 2  r a i r This i s a good approximation since T^ ^ T r e f Heat content of steam: H . = S[ Cp (T, - T ^) + Cp ( T 0 - T, ) + /A ] (=)KJ/S s i rwater b.p. ref steam 2 b.p. s = S.x h . 1 s i where: = mass flow rate of saturated steam (g/s), at 137.89 kPa (20 psia) h g ^ = s p e c i f i c enthalpy of saturated steam ( K J / s ) ( T r e f = 288.6 K) T 2 = 381.9 K = temperature of saturated steam at 137.89 kPa T, = T, . . = 373 K b.p. b o i l i n g point The s p e c i f i c enthalpy of saturated steam h ^ i s given i n the steam (32) tables , but here the reference temperature i s 273 K rather than the reference temperature of 288.6 K used here. Therefore, the value given by the tables h . i s re l a t e d to h . by: gi s i h . = h . - Cp fc (288.6 - 273) s i g i r water since i n t h i s temperature range _3 Cp = 1 (cal/mol K) = 4 , 1 8 5 X 1 0 = 0.23 x 10~ 3 (KJ/gK) *water 18 and h . = 2688.98 x 10" 3 (KJ/g) g i then: H . = S. x 2685.39 x 10~ 3 ( K J / s ) s i 1 187 4) Heat content of purging nitrogen is assumed to be negligible 5) Total inputs = H . + H . + H coal arr steam Outputs: 1) Calorific value (combustion heat) of dry gas: 3 H = Volume of dry gas(m /s) x c a l o r i f i c value of the unit volume of eg gas (KJ/m3) (=) KJ/s Volume of dry gas is obtained from the mass balances (Table VIII-1, Appendix VIII), and the c a l o r i f i c value of the unit volume of gas is % calculated from the gas composition (See Chapter III, and Tables VII-1 and VII-2, Appendix VII). 2) Sensible heat of dry gas: H = G 3 Cp dT = G x Cp (T„ - T .) x 4.185 x 10 3 (=) KJ/s sg J T F g g 3 ref , ref where: G = molar dry gas flow = volume of dry gas ( m o l e s / s ) 24.12 x 10 T3= Reactor outlet gas temperature (K), measured experimentally. _3 4.185 x 10 = conversion factor (KJ/cal) Cp = mean molal specific heat of the gas in the temperature range T . - T_ (cal/mol k) 1 ref 3 The mean molal specific heat of the gas is approximated as the arithmetic mean of the molal specific heats of the gas at the temper-atures T^ and T £. In turn, the specific heat of the gas at these temperatures i s evaluated as the weighted molal average of the individual gases composing i t ; i.e. 188 C p g a s ( T ) = ^ Z c p ^ (% gi) where Cpi = molar s p e c i f i c heat of component i at T and (% gi) = molal f r a c t i o n of component i i n dry gas. The s p e c i f i c heats of the i n d i v i d u a l components at temperature T are obtained from : Cp = 6.62 + 0.00081 T W2 Cp C Q= 6.60 + 0.00120 T C p „ u = 5.34 + 0.0115 T Cn. 4 Cp_. = 10.34 + 0.00274 T - 195500 T~ 2 2 Cp = 6.5 + 0.00100 T. 2 3) Heat of Steam The enthalpy of :the steam i s given by H S O = S Q X h g o (KJ/s) where: S q = water flow out of the reactor (g/s), calculated from the mass balances (Table VIII-1, Appendix V I I I ) . and h g o = s p e c i f i c enthalpy of steam (KJ/g) at o u t l e t of reactor (T^, atmospheric pressure, T ^ ^ = 288.6). h g Q i s obtained from (32) superheated steam tables a f t e r correcting f or the difference i n reference temperature i . e . H = S (h - 3.59 x 10~ 3) so o go where: h ^ = s p e c i f i c enthalpy of steam (KJ/g), from superheated steam tables. 189 4.- Heat value of e l u t r i a t e d carbon: This i s given by H = e x h (KJ/s) where ° J c c c e^ = e l u t r i a t e d carbon (g/s)which i s obtained experimentally. (Tables VII - 1 and VII-2, Appendix VII) and h'^  = heat of combustion of carbon = 32.773 (KJ/g) 5) T o t a l outputs = H + H + H + H eg sg so e Losses: Heat losses are obtained by differ e n c e between the t o t a l energy input and the t o t a l energy output. Sample Ca l c u l a t i o n , Run No. 38: A sample c a l c u l a t i o n following the general procedure outlined above is presented here f o r run No. 38, the same f o r which a sample calculation-of the mass balance was presented i n Appendix VIII. Energy Inputs: 1) Heat content of c o a l : H = 1.586 x 31.34 = 50.305 (KJ/s) coal 2) Heat content of a i r : _3 F . = 5 , 4 3 X 1 0 = 0.225 (moles/s) a i r 24.12 x 10 3 T^ = a i r temperature = 294.8 K Cp . (T,) = 6.8035 + 0.0008351 x 294.8 - 39325.15 x (294.8)~ 2  r a i r 1 = 6.602 (cal/mol K) Cp . (T .) = 6.8085 + 0.0008351 x 288.6 - 39325.15 x (288.6)~ 2 r a x r r e f = 6.577 (cal/mol K) . - 6.602 + 6.577 , crir> , , / , „ x . . Cp . = = 6.590 (cal/mol K) a i r „ H . = 0.225 x 6.590 x (294.8 - 288.6) x 4.185 x 10 3 a i r = 0.038 (KJ/s) 3) Heat content of steam: S ± = 0.0 (g/s) H .= 0.2685.39 x 10~ 3 = 0 (KJ/s) sx 4) T o t a l inputs = 50.305 + 0.038 + 0.0 == 50.343 (KJ/s) Energy Outputs: 1) C a l o r i f i c value of dry gas: C a l o r i f i c value of unit volume of gas: h c g = 0.068 x 12.109 + 0.071 x 11.997 + 0.02 x 37.743 = 2.43 MJ/m3 H = 6.43 x 10" 3 x 2.43 x 10 3 = 15.625 (KJ/s) eg 2) Sensible heat of dry gas: _3 „ 6.43 x 10 _ , T / \ G = = 0.267 (mol/s) 24.12 x 1 0 - 3 T 3 = 1127.7 K S p e c i f i c heat of i n d i v i d u a l dry gas components at temperature T^: Cp„ = 6.62 + 0.00081 x 1127.7 = 7.533 (cal/mol K) H2 Cp = 6.60 + 0.00120 x 1127.7 = 7.953 (cal/mol K) C = 5.34 + 0.0115 x 1127.7 = 18.309 (cal/mol K) PCH. 4 Cp„ n = 10.34 + 0.00274 x 1127.7 - 195500 x (1127.7)~ 2 2 = 13.276(cal/mol K) Cp„ = 6.5 + 0.001 x 1127.7 = 7.628 (cal/mol K) N2 .'. Cp ( T J = 7.533 x 0.068 + 7.953 x 0.071 + 18.309 x 0.02 gas V + 13.276 x 0.121 + 7.628 x 0.72 = 8.542 (cal/mol K) S i m i l a r l y , C p g a g ( T r e f ) = 7.083 (cal/mol K) Then: Cp = 7.083 + 8.542 = 7.813 (cal/mol K) r-gas . .'. H = 0.267 x 7.813 (1127.7 - 288.6) x 4.185 x 10~ 3 sg = 7.314 (KJ/s) 3) Heat of steam H = 0.407 (4280.76 x 10~ 3 - 3.59 x 10 - 3) = 1.741 (KJ/s) so 4) Heat value of e l u t r i a t e d carbon: H = 0.346 x 32.773 = 11.339 (KJ/s) c 5) T o t a l energy output = 15.625 + 7.314 + 1.741 + 11.339 = 36.019 KJ/s Losses: Energy losses = 50.343 - 36.019 = 14.324 (KJ/s) losses _ 14.324 input 50.343 TJ . - i  O Q c „ Percentage losses = = = 28.5% TABLE IX-1 ENERGY BALANCES FOR TYPICAL GASIFICATION OF FORESTBURG AND SUKUNKA COAL Run 9 Air Reactor Energy Inputs (KJ/s) C a l o r i f i c Energy Outputs (KJ/s) Energy Inlet Temperature K Outlet Temperature K a i r coal steam Total C a l o r i f i c Value of Dry Gas Sensible Heat of Gas Elutriated Carbon Heat Content steam Total Losses (KJ/s) Losses Input-Output Input 12 294 5 907 0 024 20 91 0 760 21 694 7 70 3 199 0 819 2 116 13 834 7 860 36 2 14 291 3 1028.3 0 020 31 06 1 455 32 535 15 .95 7 190 2 097, 3 940 29 177 3 358 10 3 15 290 8 1026.3 0 016 31 06 0 671 31 74 7 14 42 7 025 2 097 3 285 26 827 4 92 15 5 16 296 6 - 1025.7 0 043 33.36 0 0 33.403 13.38 5 366 1 147 5 635 25 528 7 875 . 23 6 17 296 1 1124.3 0 052 33 36 0 0 33 412 11 73 7 704 1 147 5 209 25 790 7 622 22 8 18 2'j4 2 1130.0 0 040 33 36 2 148 35 548 11 52 8 272 1 147 10 175 31 114 4 434 12 5 19 293 6 1047.3 0 040 43 85 0 448 44 338 15 30 7 987 2 360 2 208 27 855 16 483 37 2 21 294 6 1064.5 0 043 53 62 0 0 53 663 20 41 8 047 1 803 5 876 36 136 17 527 32 7 22 294 3 1026.7 0 052 62 09 1 413 63.555 24 27 9 345 4 195 3 690 41 500 22 055 34 7 23 294 8 997 0 024 10 41 1 531 11 965 3 14 3 437 0 426 0 731 7 734 4 231 35 4 25 294 3 1068.7 0 041 38 05 2 377 40 468 18 57 7 720 1 278 2 285 29 853 10 635 26 2 31 293 5 1075.4 0 043 43 04 2 551 45 634 26 31 10 303 1 737 0 00 38 350 7 284 16 0 36 295 1 1113.7 0 041 46 54 0 0 46 581 11 73 7 047 10 487 6 278 35 542 11 039 23 7 37 295 7 1070 0 046 46 54 1 222 47 808 12 46 6 807 10 487 8 616 38 37 9 4 38 19 7 38 294 8 1127.7 0 038 50 30 0 0 50 338 15 62 7 314 11 340 1 741 36 015 14 323 28 5 39 294 4 1133 0.038 46 65 0 0 46 688 11 66 7 692 8 193 3 157 30.702 15 986 34 2 . 40 294 9 1080.7 0 041 46 65 0 961 47 652 13 67 7 484 8 193 4 001 33 348 14 304 30 0 Runs 12-31 : Forestburg coal Runs 36-40 : Sukunka coal. 

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