Ind. Eng. Chem. Process Des. Dev. 1983, 22, 563-570 Lawson, 0.B. Ph.D. Thesls. Unlverstty of Manchester, England, 1967. Lewis, I. E. Paper presented at Internatlonal Solvent Extractbn Conference, 1977: p 325. Longsdall, D. H. et al. Trans. Insf. Chem. Eng. 1057, 35, 301. Luhning, R. W.; Sawlstowskl, H. Paper presented at Internatlonal Solvent Extraction Conference, The Hague, 1971. Mizrahl, J.; Barnea, E. Process Eng. Jan 1973, 60. Mumford, C. J. Br. Chem. Eng. 1066, 73,981. Mumford, C. J. Ph.D. Thesis, University of Aston, England, 1970. Nagata. S.; Yamaguchl Mem. Fac. Eng. Kyofo Unlv. 1080, 22, 249. Orjans, C. W.; Godfrey, J. C. Paper presented at Intematlonal Solvent Extraction Conference, 1977; p 340. Qulnn, J. A.; Slgldr, D. B. Can. J . Chem. Eng. 1063, 47, 15. Rcdgar, W. A.; Trice, V. G.: Rushton, J. H. Chem. Eng. Prog. 1056, 52, 515. Rushton, J. H.; Nagata, S.; Rooney, T. B. AIChE J . 1964, 70, 298. Rushton, J. H.: Owshve, J. Y . Chem. Eng. Reg. 19S3, 49,267. Ryon, A. D.; Daley, F. L.; L O W , R. S. Oak Rkige Natlonal Laboratory Rept. ORNL-2951, Oct 5, 1960. Sarkar, S. Ph.D. Thesis. Unlverstty of Aston, England, 1976.
583
Sarkar, S.; Mumford. C. H.; Jeffrays, G. V. Paper presented at 3rd Annual Research Meeting 1976, Inst. Ghem. Engrs., Salford, England. Sarkar, S.; PhHlIps, C. R.; Mumford, C. H.; Jeffreys. G. V. Trans. I . Chem. E . 1080, 58, 43. Selker, A. H.; Sleicher, C. A. Can. J . Chem. Eng. Dec 1065, 298. S t h e r , H. M.: Wahler, F. I . Chem. E . Symp. Ser. 1075, No. 42, Paper 14. Storey, B. L. BSc. Project, Chemical Engineering Dept., Unlverstty of Aston, England, 1971. Thomas, R. J.; Mumford, C. H. Proceedings, International Solvent Extraction Conference, The Hague, 1971. Thomas, R. J.; Mumford, C. H. Process Eng. Dec 1072, 54. Treybal, R. E. "Liquid Extractlon", 2nd 4.;McGraw-HIII: New York, 1963. Vljayan. S.; Panter, A. B.; Jeffreys, G. V. Chem. Eng. J . 1975, 70, 145.
Received for review March 16, 1981 Revised manuscript received November 23, 1982 Accepted December 21, 1982
Gasification of High Ash Content Coals with Steam in a Semibatch Fluidized Bed Reactor Martin Schmal, Jor6 Lulr Fonter Montelro, and Humberto Torcanit Univers&de Federal do Rio de Janelro, C O P E N F R J , Program de Engenharla Qdmica, Calx8 Postal 68502, .?GOO Rlo de Janeko, Basil
This work reports a study on gaslficatlon of Brazilian mineral subbltuminous coal with steam in a semibatch fluidized bed reactor. Several tests for the fluidization characteristics of mixtures of coal and ash were performed.
Fluidization velocity was determined from the data of the minimum velocity, calculated at high temperatures and later tested. Experimental results show that flow conditions must be determined experimentally for high temperatures and pressures. The influence of temperature and pressure on product gases during the reaction and on the ratio CO/CO, were determined. The reaction rate is very sensitive to temperature variations between 850 and 1000 O C . For pressures higher than 10 atm the effect of the pressure on reaction rate I s negligible. The experimental results are well described by the unreacted core model above 850 OC where the chemical reaction is the ratecontrolling step. The activation energy was found to be 39 k&l/mol.
Introduction The consumption of gas which already comprises a high percentage on the total energy demand is expected to increase much more in the near future. Natural gas cannot meet this demand on a long term basis. Particularly in Brazil, the sources of coal of high ash content are of the order of 22 million tons and are presently only partially used as metallurgical coal. Since 1976 we have been engaged in the improvement of the conventional methods of gasifying this coal and also developing the techniques to gasify Brazilian coal of high ash content in a fluidized bed under pressure. The Cientec Research Center is also now engaged in developing a demonstration plant for fluidized bed coal gasification at normal pressure, whose main objective is to obtain gases of low BTU. In the meantime, Rio's gas company CEG is engaged in implementing the Lurghi process to obtain medium BTU gases, while the Brazilian Petroleum Company (Petrobrh) is implementing the Koppers-Totzek process to supply a petrochemical complex to be established in the south of Brazil. Both process will use the same high ash content coal (>40% of ash). This work was carried out to determine whether the high ash content coal of Brazil could be gasified with steam +Present address: Fundado de Ci6ncia e Tecnologia CIENTEC, R. Washington Luiz 675, Porto Alegre, Rio Grande do Sul, Brasil. 0196-4305/83/1122-0563$01.50/0
in a dense fluidized bed. The objectives of this work were: (1)to determine if the fluidized bed technique was operable at temperatures and pressures for coal gasification and to determine the effects of process variables on the reaction rate; (2) to compare the reactivity of the high ash content with other coals of different sources, and (3) to provide a design correlation for the reaction rate. Fundamentals of coal gasification were presented by Fredorsdorf and Elliot (1963) discussing the reactions, the thermodynamic equilibrium, reaction rates and mechanisms, and heat and mass transfer in fixed and fluidized bed system. May et al. (1958) studied the gasification of coke with steam, nitrogen, oxygen, and hydrogen in fixed and fluidized bed systems for the determination of the design correlation in the fluid system. Jensen (1975) studied of the kinetics of the residum coal gasification with steam in a fluidized bed for determination of the kinetic parameter. Klei et al. (1975) studied the gasification of active coals with steam in a fixed bed reactor, determining the gas composition and the influence of operational conditions on reaction rates. Also, Chan and Papic (1976) and Shaw and Paterson (1978) studied the cod gasification in a fixed bed and fluidized bed, respectively, in order to determine the kinetic parameters and the reaction rate of coal gasification using coals of very low ash content. No attention was given to the ash content in coal gasification and very little information exist about the product distribution during the reaction as well as the 0 1983 American Chemical Society
564
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 4, 1983
Table I. Segregation Tests ash content samples, % upper middle bottom _. a b Table 11.
47.9 93.5
41.5 93.3
76 93.6
Mass Flow Rates of Steam (g/min) P, atm
T,"C
1
5
10
15
850 900 950 1000
1.99 1.85 1.67 1.55
9.96 9.26 8.35 7.78
19.92 18.52 16.71 15.55
29.9
-
-
-
fluidization characteristics in a fluidized bed reactor. Moreover, no good agreement was found for the reaction models in the different gasification reactions.
Fluidizing Characteristics of Coal To get the basic data for the design of this reactor, several tests of fluidization characteristics of mixed coal and ash were performed in a glass tube of the same diameter used for the reaction system (4 = 45 mm). The objective was to establish the ratio of particle size of coal and ash in mixed form to provide good fluidization conditions and to determine the minimum fluidization velocity for establishing the operational velocities at high temperatures and pressures. The first test showed two problems. One was the slugging flow caused by the use of large particle sizes. The other was the insufficient mean flow rate of steam,to satisfy the stoichiometry of the reactions, caused by the use of very small particle sizes. Therefore, to provide a good fluidization without compromise of the excess of steam for the reactions, intermediate particle sizes were chosen (35-48 mesh Tyler) and so a homogeneous fluidized bed, almost without large bubbles, was achieved. For these experimental conditions also the Kozeny-Carman equation fits the experimental results well and therefore we could also calculate the minimum fluidization velocity using the following equation (Kunni et al., 1969)
where umfis the minimum fluidisation velocity, ps and pf are the density of solid and fluid, dpand 4sare the particle size and esphericity, emf is the bed expansion for umf, and p is the viscosity of fluids. If the minimum fluidization velocity umffor a specific solid and fluid is known from experimental data at normal conditions, we can calculate umfalso for different fluids but for other conditions, if ps of the solid, p and pf of the
fluid are known and assuming that the expansion e& of the bed does not change. Neglecting pf, which is much smaller than ps, we get the ratio umf = -P' (2) U'mf
Then, assuming that this ratio is valid for all experimental conditions, we calculate the operational velocity u LP for different experiments. Thus, as the fluidization velocity was determined in a glass tube using air as the fluidization agent, we calculate, as above, the operational velocity for steam at various temperatures in the reactor. Furthermore, during the fluidization tests we observed segregation when mixtures of coal and ash of the same range of particle sizes were used. After some time of fluidization the gas flux was instantaneously interrupted and samples at different high levels were collected for ash analysis. The results for pure coal shown in Table I, row a, leads us to test several mixtures of coal and ash with different particle sizes. The best results were obtained for a mixture of 20 g of coal of 35/48mesh and 140 g of ash of 42/60 mesh Tyler, as shown in Table I, row b. Therefore, these conditions were used in the gasification tests which would provide an excess of vapor and good fluidization. The height of the bed was 12 cm and the expansion of the order of 3 cm. The pressure drop was determined as a function of the flow rate and hence, the minimum fluidization velocity was 7.0 cm/s and the fluidization velocity was 24 cm/s. We also calculated for comparison the fluidization velocities, using eq 2, and the corresponding mass flow rates for different experimental conditions, as shown in Table 11. The observed discrepances are probably due to the assumption of a constant value of Emf.
The next step was to test these results experimentally for high pressures and temperatures as in a real system, using the established conditions. The axial internal temperature was measured for a constant mass flow rate of steam, and when the profile was isothermal good fluidization was attained. Table I11 shows these results for different mass flow rates and pressures but for a constant temperature control of 900 "C. These results show that the experimental fluidization velocities are approximately twice the calculated values. It makes clear that fluidization conditions cannot quite be established from extrapolated formulas but must be determined experimentally for higher pressures and temperatures. The same procedure was used for other temperatures. No sintering was observed and the entrainement of fines was neglibible. Experimental Procedure The fluid unit (Figure 1) was designed to operate at pressures and temperatures up to 15 atm and 1000 "C,
Table 111. Experimental Mass Flow Rates of Steam and the Corresponding Temperature Profiles steam height of the bed, cm velocity, mass flow, P, atm cmls g/min 0 5 10 15 0.0 10.7 20.2
0.0 1.9 3.6
897 905 903
893 901 902
882 895 900
865 882 895
5
25.8 7.5 13.1
4.6 6.7 11.7
902 902 898
902 899 898
900 878 896
898 858 885
10
20.5 13.1 21.5
18.3 23.3 38.3
903 903 903
904 904 903
901 901 901
900 842 899
1
c1
comments fixed bed fixed bed partial fluidization fluid bed fixed bed partial fluidization fluid bed partial fluid fluid bed
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 4, 1983 585
-
E,
NITROGEN
E2 E, E, E5
-
ROTAMETER
E6
.
NITROGEN
W A T E R ' S TANK VAPORIZER PREHEATER
E7
-
PREHEATER
€0
.
GASIFIER
-
E9
-
E,0 Ell
I
I
C O A L FEEDER FILTER CONDENSER
E 1 2 . COLECTOR El,-
FLOWYETER
El,-
CONDENSER
E E17
COLECTOR
TP Y V
- TANK
-
TERMOCOUPLE MANOMETER VALVES
Figure 1. Flowsheet of gasifier unit. Table IV. Ultimate and Elemental Analysis of the Charqueadas Coal orig pyrolyzed coal coal ultimate analysis fixed carbon, wt % volatile matter ash moisture elemental analysis carbon hydrogen sulfur others heat capacity, cal/mol density, g/cm3
26.3 21.3 52.4 7.5
32.3 1.8 65.9 0.7
28.3 1.9 0.4 13.4 3501 1.95
0.2
0
10
PO
x)
40
sa
eo TlUE
a
Data from CIENTEC-Porto A1egreR.G. Sul.
respectively. The reactor was constructed of stainless steel 316 tubing with an inside diameter of 45 mm. The height of the bed was 200 mm and was maintained in the heated section of the furnace. The total height of the reactor was 400 mm. A high ash content subbituminous and pyrolyzed coal was used in the desired particle sizes. This coal came from the Charqueadas mine, Rio Grande do Sul. Inspections on typical samples are presented in Table IV. A charge of 20 g of pyrolyzed coal (35/48#) mixed with 140 g of ash (42/60#) was introduced batchwise in the reactor. Residual volatile matter was eliminated by passing nitrogen at operating temperature before the gasification itself. In the meanwhile, water was vaporized and superheated in the preheaters by passing the bed. When the prescribed temperature was attained, the counter flow flux of nitrogen was interrupted instantaneously and steam passed through fluidising the bed at a specified mass flow rate. Product gases and excess of vapor passes through a filter where fines are retained and then through the condenser and separator. Samples of product gases were collected at different times and analyzed in the chromatograph using molecular sieve and Porapak N columns. The total flow
I.,"
I
Figure 2. Determinationof volume of CO + COz with time; T = 900 O C ; P = 5 atm.
rate of product gases was measured during the reaction. Masses of coal and ash were determined at the end of the reaction. Steam conversion was calculated on the output basis where converted steam was taken as equivalent to the carbon monoxide and carbon dioxide in the product gases. Using the volumetric flow rate of CO + COzwe calculate by integration the volumes of the unonreactedncoal (Vmc) with time. The reason why the corresponding volume of the "nonreacted" coal (Vmc) was taken is that this curve better fits the experimental results at the end of the reaction but not at the beginning. From Figure 2 the "nonreacted" volume at time t6 is As, at time t4is (Ad + A4), and so on. The corresponding volume of reacted coal is obtained by substracting this volume from the total volume of CO + COz, based on coal analysis. From the ultimate analysis of the pyrolysed coal we obtained 32.3% fixed carbon. Multiplying this by the factor 0.92, which relates the fixed carbon to the elemental carbon, we obtain for 20 g of coal 5.94 g of elemental carbon. This factor is an experimental mean value of various analyses of the pyrolyzed coals at different temperatures due the inhomogneiety and incomplete decomposition of volatile matter
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 4, 1983
566
Table V. Conversion vs. Time at 900 "C and 5 atm flow rate value of of nonreacted time, CO + CO,, coal conversion, EA;. L % min Llmin t, =49 0.014 0.082 99.7 t , = 36.1 0.078 0.675 94.4 t , = 25.4 0.180 2.055 83.0
CO
/to*
I
-__I
t , = 15.6 t , = 8.1
0.300 0.410
4.407 7.070
8 0
b -
0
6 -
-
63.6 45.6
4 -
COMPOSITION 8
-
I
t 1
I
40
0
I
20
6
4
I 40
I
I
60
I
I
I
I
e0
0
12
10
TIME
(min )
COMPOSITION 1
Io
I
I
0
(%'
I
1
t
b "
@
n
A
20
40
(10
00 100 CONVERSION (70)
Figure 4. Ratio of CO/C02. (a) P = 5 atm; (b) P = 10 atm.
0
I
0
I
I
I
12
6
I
I
10
I
I
I
H
3a T I M E I miml
I
Figure 3. Composition with time, (a) T = 950 OC; P = 10 atm. (b) T = 850 OC; P = 15 atm.
in the coal a t operational temperature. Therefore, for a total volume of CO COPat 25 "C equal to 12.11 L (VT), the conversion is calculated by
+
(3) Typical data of the total rate and curves of the conversion with the reaction time are shown in Table V and Figure 2, respectively. Data of product gases were used to calculate the ratio of the partial pressures of the water-gas shift reaction to compare with the theoretical equilibrium constant and to verify how far this reaction is from the equilibrium. The partial pressure of each component was calculated from the total flow rate and the gas composition data. Typical data from the operational system are shown in Table VI. Complete data for all other conditions were obtained by Toscani (1979). Results and Discussion Influence of Temperature and Pressure. The influences of temperature and pressure on product gases during the reaction are shown in Figure 3a and b. The
percent composition of hydrogen is of the order of 62% and is higher than the other components. During the reaction it initially decreases and then increases slowly with increasing time of reaction. The opposite occurs with carbon monixide, whose composition is of the order of 30%. I t is higher initially but decreases slowly with increasing time of reaction. The carbon dioxide is present only in smaller quantities. The last column of Table VI show composition data of May et al. (1958) for a similar condition and reaction time. Results are close to those obtained here. These data can also be analyzed in view of the carbonsteam and water-gas shift reactions which are considered as the most significant gasification reactions of coal with steam. The carbon-steam reaction generates H2 and CO whereas the shift reaction uses a part of CO to produce C 0 2 and Ha. Thus,the increase in CO and the diminishing in H2 contents would indicate that the shift reaction rate decreases in relation to the steam reaction rate and therefore the C02content decreases. Otherwise if Hzincreases and CO decreases the opposite will occur, increasing the C02 content as shown in Figures 3a and b. Table VI1 presents their mean values and some equilibrium data for comparison and better visualization of the temperature and pressure influence on the composition of product gases. Observe that C 0 2 and H2 increase for higher temperatures whereas CO decreases. The CHI content does not change much with temperature but increases for higher pressures. The H2 content is mainly constant with pressure variations. Figure 4a and b show changes of the CO/C02 ratios with conversion for different pressures and temperatures. Observe that this ratio increases up to 50% conversion but
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 4, 1983 587 Table VI. Operating Conditions and Typical Data 900 "C reactor temperature 5 atm pressure 20 g (35/48 mesh Tyler) mass of coal 140 g (42/60 mesh Tyler) mass of ash 20.5cm/s superficial velocity of steam 15.6 25.4 8.1 time, min product gass composition, vol %
co
COZ CH.
outlet gas rates, L/min total
% COZ CH, outlet gas rates above the bed, L/min HZO
co
COl HZ partial pressures, atm H*O
co
COZ HZ ratios H,/CO CO/CO, conversion, %
reactor temperature pressure mass of coal mass of ash time, min product gas composition CO COZ CH,
24.8 4.9 1.2 46.7 6.3 15.7
24.0 5.6 1.2 48.1 5.7 15.5
911" 7.5
22.4 3.9 1.0 40.6 5.7 26.5
36.1
49.0
15.2 3.4 0.6 31.0 5.7 44.0
4.0 1.6 0.1 13.7 5.7 74.7
21 40 5.4 5.5 0.6 16.1 72.4
1.36 0.65 0.33 0.08 0.016
1.02 0.48 0.25 0.05 0.012
0.68 0.28 0.15 0.03 0.007
0.4 2 0.13 0.064 0.014 0.003
0.26 0.04 0.010 0.004 0.001
19.076 0.26 0.063 0.511
19.209 0.197 0.040 0.378
19.363 0.118 0.023 0.224
19.484 0.051 0.011 0.103
19.555 0.008 0.003 0.032
4.794 0.064 0.016 0.128
4.847 0.049 0.010 0.094
4.910 0.029 0.006 0.055
4.959 0.013 0.003 0.026
4.989 0.002 0.001 0.008
2.0 4.3 41.9
1.9 5.0 63.9
2.0 4.5 94.8
1.8 5.7 83.4
3.4 2.5 99.7
850 "C 1atm 20 g (35/48 mesh Tyler) 140 g (42/60 mesh Tyler) 10.7 26.8 47.6
77.5
115.6
168.5
18.8 9.4 0.8 56.1 5.5 9.5
27.5 3.9 0.9 47.2 6.6 13.9
24.7
17.1 2.7 0.5 33.5 7.3 38.8
24.3 7.3 1.0 52.9 6.0 8.6
28.0 5.3 1.0 52.9 6.4 6.4
3.1 0.8 42.9 6.8 21.7
b 47.8 1.25 0.62 50.2
outlet gas rates, L/min, above condenser
3 outlet gas rates above the bed, L/min HZO
co
COZ HZ partial pressures, atm
Et? CO, HZ ratio HJCO
CO/CO,
0.191 0.064 0.032 0.0027
0.171 0.079 0.024 0.0032
0.154 0.082 0.015 0.0029
0.086 0.050 0.007 0.0016
0.050 0.029 0.004 0.0009
0.021 0.011 0.00 2 0.0003
25.35 0.241 0.121 0.720
25.426 0.298 0.090 0.645
25.490 0.309 0.057 0.581
25.146 0.188 0.026 0.324
0.959 0.009 0.005 0.027
0.961 0.011 0.003 0.024
0.964 0.012 0.002 0.022
0.979 0.007 0.001 0.012
0.988 0.004 0.0006 0.007
0.995 0.00 2 0.0003 0.003
3.0 2.0
2.2 3.4
1.9 5.3
1.7 7.1
1.7 8.0
2.0 6.3
24.5
37.8
55.3
74 .O
88.2
98.1
25.882 0.109 0.015 0.188
25.991 0.042 0.007 0.079
conversion, % '
X a
Data of May et al. (1958).b Equilibrium composition at 850 "C, 1atm; Fredersdorff (1963).
then decreases. It means that the steam reaction rate gets slower than the shift reaction up to 80%. Then as the CO/C02 ratio decreases for temperature and pressure, increasing the difference between their reaction rates is
smaller. These results are in agreement to the data of Klei (1975) and Marcilio et al. (1981). The excess of vapor improves the shift reaction and, therefore, the ratio CO/C02 is much smaller for higher conversions.
588
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 4, 1983
Table VII. Composition of Product Gases at Different Temperatures and Pressures. Equilibrium Data in the Brackets (Fredersdorf, 1963) P = 1 atm co, H2
co
T
32 (47.8) 32 (49.4) 27 26
850 900 950 1000
6 (1.25) 7 (0.31) 9.5 9.5
61 (50.2) 60 (50.3) 63 64
P = 5 atm CO, H,
P = 15 atm CO CO, H, CH, 34
CH,
CO
1.1 (0.62) 1.1
35
5.5
58
1.7
35
6
58
2.0
32
6.5
59
1.5
30
7
60
1.9
1.0 0.8
27 22
62 64
1.8 1.9
26
10
62
2.0
-
9 12
CH,
P = 1 0 atm CO CO, H, CH,
Table VIII. Partial Pressures for Shift Reaction partial pressure, atm
--
ratio PcOpH,I
t , min
pHqO
pco
pco,
PH,
8.1 15.6 25.4 36.1 49.0 4.0 8.6 14.7 20.0
4.79 4.84 4.91 4.95 4.98 0.708 0.750 0.818 0.874
0.064 0.04 9 0.0 29 0.013 0.002 0.068 0.06 5 0.049 0.03 1
0.128 0.094 0.055 0.026 0.008 0.193 0.161 0.116 0.085
6.5 13.5 20.8 32.8 43.9
9.86 9.89 9.91 9.94 9.96
0.042 0.038 0.031 0.020 0.012
0.016 0.010 0.006 0.003 0.001 0.03 2 0.024 0.018 0.011 0.012 0.007 0.005 0.003 0.002
57
1.8
operating conditions
pH,dCO
T = 900 "C P = 5 atm
0.0066 0.0040 0.0023 0.0012 0.0008 0.128 0.079 0.052 0.035 0.0023 0.0012 0.0008 0.0005 0.0003
0.080 0.063 0.050 0.033 0.020
7
T = 1000 "C P = 1 atm
T = 850 "C P = 10 atm
HZ/CO
A
-
-
0
05O.C 000%
-
0.12
-
0.00
-
-
0.04
I
10
I mtm
-
2.0
-
0.10
3.0
1OOO.c
lcO0.C
0 -
-
4.0
-
1
1
I 40
20
0
l
I
I
I 0
60
-
I 100
CONVERSION 1%)
n2/co
f 0
0 4.0
3.0
-
-
8% * C
e
Y)
I
A
' 06
*
I 0 8
I
10
Figure 6. Variation of Pco,)'H,/PH~O~CO.
3 ,
-
1.0..
0.4
CONVERSION (%I
- sw-C - 0 - 0 5 0 *C
2 0 -
0.1
I 50
I
I
TO
I
I
W CONVERSION 1%)
Figure 5. Ratio H2/C0. (a) P = 5 atm; (b) P = 10 atm.
The H2/C0 ratios as function of conversion are shown in Figures 5a and b for different temperatures and pressures. It decreases by increasing the conversion up to 80% and is more pronounced for higher temperatures and pressures. These data are also consistent with the inter-
pretation of the C O / C 0 2 data above. There is considerable kinetic information of the gas shift reaction but its rates cannot be predicted in the gasification reactions. Most of them indicate that this reaction is fast and that equilibrium is reached in a reacting steam-carbon system. Product gas composition was used to calculate the which were than compared to ratios (PCO*PH,/Pco~P~,o) the equilibrium constant of the water-gas shift reaction. These data were plotted in Figure 6 for two extreme cases at 850 "C, 10 atm and at lo00 "C and 1atm. All other data are comprehended in these limits as shown in Table VI11 for intermediate pressures and temperatures. None of these data reach the theoretical equilibrium constants which are at 1atm 0.95 and 0.55 for 850 "C and lo00 "C, respectively. These seem to be reasonable because here a large excess of steam has been used. May et al. (1958) attained equilibrium conditions for lower conversions.
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 4, 1983 569 Table IX. Values of
7
pressure, atm
/
T,"C
1
5
10
850 900 950 1000
250 103 72 40
141 65 32 16
75 48 24
15 75
-
-
-
1.0
0 8
-
.
0 6
E
-
u
0 4
0 2
/
0 6
,
0 0 5
0
1
min
Figure 8. Flow rate. CONVERSION 0.9
I
-
0 01 0
25
I5
I
:! 0.2
10
I
I
m
90
I 50
I 40
I
50
TIME I min )
Figure 7. Determination of
- SHRINKING
7.
CORE
(I50.C
Reaction Model for the Carbon-Steam Reaction. The rate of the gasaolid reactions should be represented by the shrinking core model or the continuous model as discussed by Wen (1968). Both models have been tested previously for the coal gasification by Schmal et al. (1982), the conclusion being that for higher temperatures the shrinking core model fits the experimental results better but for temperatures up to 850 "C the continuous model also fits the experimental data, where diffusion and mass effects were eliminated in the thermobalance. As the fluidization conditions were established to provide effective mixing in the bed, the diffusion and mass effects could be disregarded and the shrinking core model could be applied. The chemical reaction was found to fit the experimental results better and therefore it should be the controlling step for the carbon-steam reaction in a fluidized bed. Hence, the following expression was used (Kunni et al., 1969) t/T
= 1 - (1 - X)1/3
(4)
where r = Cs&,/r is the reaction rate in mol/(cm2 min), Cso = apcf/12 is the initial concentration of carbon in mol/cm3, Rois the mean radius of the particle (0.0178 cm), a is the ratio of carbon and fixed carbon (0.921, p is the particle density (1.84 g/cm3), cf is the fixed carbon (0.323), X is the conversion [ l - (r/Ro)3], T is the time of total consumption, and t is the reaction time. Ploting [ l - (1 - X)1/3]vs. time as shown in Figure 7, we obtain the values of 7 from the experimental conversion data (see Table IX) based on the final time of total consumption according to the procedure used before shown in Figure 2, and together in Figure 8. Figure 9 shows the resulta of the experimental
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data for different temperatures and pressures. Results show that the reaction rate is very sensitive for temperature variations. It doubles for an increase of 50 OC independently of the pressure, which suggests that the chemical reaction is the controlling step. The reaction rate is also very sensitive to pressure variations as shown in Figure 10. It also double as the pressure is varied from 1 to 5 atm and from 5 to 10 atm. For higher pressures the effect on the reaction rate is negligible, suggesting that the raction order is zero. At 1000 "C the reaction rate is very fast for pressures of 1 and 5 atm even for higher pressures as shown in Figure 11. As has been seen, the shrinking core model with chemical reaction controlling step fits all experimental data very well. Then the reaction constants were determined for different temperatures and using the Arrhenius equation the k values were plotted in terms of reciprocal temperatures as shown in Figure 12. The activation energy was found to be 39 kcal/mol which, is compatible with most values given in the literature and very close to the results of Schmal et al. (1982) and Marcilio (1981) for bruitfan coals obtained previously in a thermobalance. These results imply that the experimental data of the carbon-steam reaction in a fluidized bed can be discrim-
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Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 4, 1983
I
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0
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loo
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120
140
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100
100 TIME
( min.)
M
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Figure 10. Nonreacted core model.
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Figure 12. Arrhenius equation.
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Figure 11. Nonreacted core model.
inated by a definitive mechanism of reaction in order to reach useful conclusions. Conclusions From these results it is possible to draw the following conclusions. Fluidization tests show that to provide an excess of vapor, good fluidization, and to avoid segregation it is necessary to use mistures of coal and ash of different mass ratios and of different particle sizes. Furthermore, the fluidization parameters of mass flow rate cannot be established from the common extrapolated formulas but must be determined experimentally for higher pressures and temperatures. Hydrogen and carbon dioxide formation increases by increasing the temperature whereas the CO content decreases. Hydrogen does not change much with pressure, but methane increases for higher pressures. The carbon-steam and the shift reactions are the main reactions of this system, being more significant since the C O / C 0 2 ratio was in all cases greater than one. The C O / C 0 2ratio increases up to 80% conversion and then decreases, which implies that the steam reaction rate decreases much leas than the shift reaction rate. This ratio decreases also by increasing the temperature and pressure. The gas shift reaction does not reach equilibrium because of the large excess of steam used. The reaction rate is very sensitive to temperature and doubles for an increase of 50 OC independently of the
pressure. It is also sensitive to pressure up to 10 atm. For higher pressures the effect on reaction rate is neglegible, which suggest that then the reaction order is zero. The shrinking core model for chemical reaction control fits very well the experimental data of gasification in a fluidized bed for different temperatures and pressures. Their activation energy was 39 kcal/mol which is consistent with most values for low ash content coal gasification and very close to those obtained for high ash Brazilian coal gasification.
Literature Cited Chan, H. E.; Paplc. M. M. Can. J . Chem. Eng. 1978, 54, 645. Frederadorff, C. G.; Elbt. M. A. “Chemlstry of Coal Utlllsatlon” Suppl. Vol. Wiey: New York, 1963; 882-1022. Jensen, G. A. Ind. Chem. procesS D e s . Dev. 1975, 14, 470. Klel, H. E.; Sahaglan, J.; Sundstrom, D. W. Ind. €47. Chem. Process Des. Dev. 1975. 14, 308. Kunni. D.: Levensplel, 0. “Fiuldbtlon Englneering”. Wiley: New York, 1969. MarcUb, N. R.; Castellan, J. L.; Montelro, J. L. F.; Schmal. M. Proceedings, Internatlonel Conference of Coal Science, Diisseldorf, 1981; p 203. May, W. 0.; Mueler, R. H.; Sweetser, S. B. Ind. €ng. Chem. 1958, 50, 1289. Schmal. M.; Monteko, J. L. F.; Castellan. J. L. Ind. Eng. Chem. Process Des. D e v . 1982. 21, 258. Shew, J. T.; Paterson, N. P. ”Studies of the Gaslficatbn of Solid Fuels in a FMdlzed Bed at Atmospheric Pressure Fluidisation”, Cambrldge University Press, Cambrldge, 1978, p 229. Toscanl. H. M.Sc.The&, C o ” & o doe Programas de Pbs-GraduacHo de Engenharta de Universkkde Federal do R b de Janeiro COPPE/UFRJ, R b de Janelro. Bra& 1979. Wen, C. Y. Ind. €ng. Chsm. 1988, 60, 34.
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Receiued for reuiew February 12, 1981 Reuised manuscript received November 23, 1982 Accepted March 16, 1983
