314
I&.
Eng. Chem. Process Des. Dev. 1981, 20, 314-318
Heterogeneous Model for Mowing-Bed Coal Gasification Reactors Y. S. Cho and B. Joseph' Chemical Engineering Department, Washington University, St. Louis, Missouri 63 130
A heterogeneous steady-state model for moving bed coal gasification reactors is developed. The resulting set of stiff differential equations is solved using the pseudo-steady-state approximation. Comparisons are made with plant data on the Wesffield gasifier and the Morgantown Energy Technology Center's gasifier. The simulation results show substantial gas-solid energy exchange especially in the combustion zone. The overall gas-solid heat transfer coefficient is smaller than those predicted by correlations for nonreactiie pack&bed systems. Parametric studies using the model indicate that the optimum thermal efficiency is obtained when the conversion of char is just complete.
Introduction A moving-bed gasifier is a vertical countercurrent reactor in which coal reacts with steam and oxygen to produce a gas consisting mainly of C02, CO, H2, CHI, and other hydrocarbons. Coal is fed to the top of the reactor through a lock-hopper and moves downward by gravity. A preheated mixture of steam and oxygen (or steam and air) is fed at the bottom and flows upward reacting with the coal. Ash and unreacted char are removed at the bottom by rotating grates. In the Lurgi gasifier ash is removed as a solid although other moving-bed processes are under development which operate at higher temperatures and remove ash as a slag (slagging reactors). There are several advantages to a moving-bed process over entrained flow and fluidized bed reactors. These include high char conversion, high thermal efficiency, and low pressure drop. As a result this process has received considerable attention in the literature recently. Qualitative models have been reported by Rudolph (1972), Hoogendorn (1973), Elgin et al. (1974),and Hebden (1975). Arri and Amundson (1978) developed a detailed model of the char particle undergoing gasification and applied this model to a countercurrent char gasifier (Amundson and Arri, 1978). Qualitative agreement with experimental data is established although no comparison with plant data was shown. In another study, Yoon et al. (1978) report a comprehensive model for a gasifier including the drying and devolatilization processes. Yoon used a homogeneous temperature model and the results showed reasonable agreement with plant data. In this work, the model developed by Yoon et al. is extended to include unequal gas-solid temperatures. Material balances around the devolatilization zone are used to obtain a more accurate representation of char. Since it is known that the water-gas shift reaction is catalyzed by the mineral matter in coal, we have incorporated kinetic equations for this reaction as well. Although the development of the material and energy balance equations is easy to derive, the resulting system of equations poses a challenging task in obtaining a solution. A pseudosteady-state assumption is invoked to permit a solution within reasonable computing time. Results show close agreement with plant data. Comparisons with some newly reported data on the Morgantown Energy Technology Center gasifier are also made. In the heterogeneous model reported by Amundson and Arri, the reactor is divided into a combustion zone and a gasification zone in which distinct reactions take place. Q796-43Q5/81/112Q-Q314$Q1.25fQ
This results in a multi-point boundary value problem involving the end points of the reactor and a point in between which divides the combustion and gasification zone. By not distinguishing between the combustion and gasification zones (i.e., allowing both types of reactions to take place simultaneously) this problem of determining the interface between the zones is avoided. The results obtained are qualitatively similar although we do not predict as sharp temperature gradients as Amundson and Arri. Development of Equations The material and energy balance equations describing the reactor is developed in this section. Since the development follows that of Yoon et al. (except for the energy balance), not all the assumptions are stated and justified here. The interested reader is referred to the article by Yoon et al. The reactor is conceptually divided into various zones (Figure 1)where distinct chemical and physical events take place. Both drying and devolatilization are treated as instantaneous steps. The time for these two steps is of the order of a few seconds, whereas the total residence time for the coal is of the order of hours. Hence the distance over which these two steps are quite small in the reactor and for modeling purposes we can treat these steps as going to completion at the top of the gasifier. In this work a correlation proposed by Gregory and Littlejohn (1965) is used to predict the distribution of volatiles as a function of temperature. The term char is used to represent coal after the drying and devolatilization process. We approximate char by CH,O, where a,P are calculated from the ultimate analysis of coal and material balances around the devolatilization zone. The sulfur in the coal is assumed to be converted to H2S and the nitrogen to NH3 in the devolatilization process. The nitrogen and sulfur content in a typical Illinois coal is of the order of 3-4% by weight on a dry basis. More than 50% of N2 and S is released in the devolatilization process. The rest of the N2 and S which may be chemically bonded with carbon would be released slowly during the gasification and combustion. However, the impact of this on the reactor performance is small due to the low percentage of the compounds present. I t would, of course, be easy to include these as well if needed. The reactions taking place in the combustion and gasification zones are shown in Table I. The reaction of char with hydrogen and carbon dioxide is assumed to be kinetically controlled. The char-stream and char-oxygen reactions are fast and are affected by diffusion resistance. 0 1981 American Chemical Society
Id.Eng. Chem. Process Des. Dev., Vol. 20, No. 2, 1981 315 Coal
P r o d u c t Gas Reactions
steam: F'
bl a = (1 - P)r3 - 8(r4 + r5) - - r6 + r8
HaO dz
Fl '-1
Wet c o a l
+
Dry c o a l + H 2 0
Dry c o a l
+
char
+
volatile products
CO
+
H2
2
(M2)
Drying
Devolatilization
Gasification
+
H20
Char
+
C02
+
CO
+
Char + H 2 * CH4 H 2 0 + CO 2 H 2
+
Char
Combustion
Ash
Char
H20
+
O2
CO
+
CO
+
C02
+
C02
H 2 t C02
co2: ~
methane:
Table I. Major Reactions in the Gasifier
oxygen:
events
notation
reactions
drying devolatilization gasification
wet coal -+ dry coal + H,O dry coal + CH,O + YP CH,Op t (1- P)&,O 2 CO + ( 1 - P + ( 4 2 ) ) H , CH,O@ t CO, 8 2CO PHZO + ( ( 4 2 ) - P)H, CHaOp ( ( 4 - (Y 2P)/ 2)H, 2 CH, + PH,O CH,Op + 7 0 , -+ ( 2 P + (a/2))CO + (27 + P - ( 4 2 ) -1)CO, t (4 2)HZO H, + 1 / 2 0 , + H,O H,O + CO 2 H, + CO,
+
+
combustion
+
R1 R2' R3
+
R6
R7 R8
CO/CO2 = K exp(-E/RT)
proposed by Rossberg (1956) is used. There is some doubt as to whether the hydrogen-oxygen reaction does take place in the gasifier. Yoon et al. neglected this reaction while Amundson and Arri allowed this reaction to take place if the oxygen concentration exceeded a certain minimum (denoted by fon,,,in). With these considerations, the local mass balance equations can be written as follows (z = 0 at the bottom of the reactor). (A) When fo, < foz,min(i.e. with no water formation reaction; r7 = 0)
+ r4 + r5 + rg)
8)
R5
For these two reactions the shrinking core model ( h i s and Amundson, 1973) is used. Yoon et al. have indicated that better predictions are obtained by this model than an ash-segregated model. In the char-oxygen reaction, the main difficulty is in predicting the molar ratio of CO to C 0 2produced. Yoon assumed a 1:l ratio of CO to COPwhereas in this work a correlation of the form
dxc FO, - = -(r3 dz
where x , = (F', - F J / P , ; XI = ( F o z o i FH$~/J'"H,;x2 = FH /FoH2; X 3 = FCO/FoH20; x4 COz/ HzO; xs = FC / b H Z O ; x6 = ( P O ,- F O , ) / P S . When f% 1 foz& (r5 = 0 and r7 goes to completion) eq M2, M3, M6 are replaced by the following equations.
R4
CH,Op denotes char. VP denotes volatile products.
char:
-r4
Steam & Oxygen (or A i r )
Figure 1. Schematic of the physical and chemical events taking place inside the reactor.
water-gas shift
h 4
H , Odz =
(MU
In deriving the energy balance equations certain assumptions regarding the contribution of the heat of reaction terms are necessary. Reactions R3 through R7 all take place on the surface of the solid and it seems reasonable to assume that the heats of reaction contribute to the solid phase energy balance equation. Further, we assume that the reactanta enter the solid phase at the bulk gas temperature and leave at the temperature of the solid. Reaction R8,the water-gas shift reaction, is assumed to take place solely in the gas phase. The exchange of energy between the gas and solid phase through conduction, convection,and radiation is represented by a gas-aolid heat transfer coefficient. Unfortunately, no correlations are available for estimating the coefficient when reactions are taking place. Hence it was taken to be proportional to the gas-solid heat transfer coefficient for nonreactive systems estimated from the correlation of Gupta and Thodos (1963). The constant of proportionality (E) is established from comparisons with plant data. Based on the above assumptions the local energy balance equations are dT, 7 solid: FsCp, = C riAHi(T,) + i=R
Ind. Eng. Chem. Process Des. Dev., Vol. 20, No. 2, 1981
316
I300
'g
,
'Im/ 900
I W +
Table 111. Comparison of Product Gas Distribution with Westfield Plant Data for Illinois coal
1,:
component ,
Toon's work; This work
-0-
50114 Temp
Gor Temp
0
02
04
Yoon (1978)
44.4 19.5 27.5 6.5 2.1
39.1 17.3 31.2 9.4 3.0
45.0 14.4 33.7 5.8 1.1
co CO,
CH,
P
700
plant data
H2
, -0-0-
this work
others
08
06
IO
Z/L
Table IV. Typical Operating Parameters for t h e METC Gasifier (Desai, 1978)
Figure 2. Typical temperature profile for Wyoming coal. Operating conditions of Table I1 with .$= 0.3, and no water formation reaction. Table 11. Operating Conditions of the Gasifier a t Westfield (from Yoon et al., 1978) parameter data reactor length reactor diameter particle size
3.048 m 3.658 m 0.01 m 13040 kg/h feed rate Of 18300 kg/h feed flowrate of oxygen 183 k m z / h . 1757 kmol/h feed rate Of steam 1244 kmol/h temperature of feed gas 644 K operating pressure 25 atm
for Illinois coal for Wyoming coal for bdth coals for Illinois coal for Wvomina- coal
parametera
run no. 70.6
run no. 71.5
reactor length reactor diameter particle diameter coal feed rate air feed rate steam feed rate operating pressure
1.981 m 1.067 m 0.005 m 552.5 kglh 1733 kg/h 14.5 kg-mol/h 6.13 atm
1.981 m 1.067 m 0.005 m 497.6 kg/h 1863 kg/h 12.6 kg-mol/h 5.17 atm
580 K was used for the temperature of feed gas. Table V. Simultion Results (Operating Conditions of Table) plant data
product gas, % V
The boundary conditions associated with these equations are x , = 0.0 at z = L
I-1,
at z = 0 xi = 0 (i = 1, 6) T g =TBf a t z = O
N2
T , = T,, at z = L Since boundary conditions are split, it is necessary to use iterative techniques to solve the equations. Attempts to solve the above split boundary problems by the shooting method (as used by Yoon, 1978) proved unsuccessful. This was due to the stiffness of the problem which prevented a solution even for extremely small step lengths. This stiffness was traced to the small heat capacity of the solid stream near the bottom of the reactor. In the combustion zone, a large amount of heat is generated on the solid surface which is transferred to the gas phase by energy exchange. Even small changes in the gas-solid heat transfer coefficient can cause drastic changes in the temperature of the solid with this set of equations. Further calculations showed that the increase in the heat energy content of the solid stream was small compared to the energy released by reaction. Thus the contribution of the derivative of the solid temperature in the energy balance equation is negligible compared to the other terms. This led to the conclusion that the solid temperature can be calculated by setting the temperature derivative equal to zero. This pseudo-steady-state assumption for the solid phase eliminated the stiffness problem and a plution could be obtained by the shooting method. Details of the iterative algorithms are given by Cho (1980). Results The simulation model was validated by comparison with data reported on two experimental reactors. The first reactor is a gasifier operated at Westfield, Scotland, described by Elgin et al. (1973). The operating conditions of this gasifier are shown in Table 11. Figure 2 compares the temperature profiles predicted by this model and by Yoon. The solid temperature profile is close to the ho-
______
this work
run 70.6 run 71.5 run 70.6
co CO,
CH, others total product gas flow (kg-mol/h) char conversion
12.8 22.3 6.9 2.4 54.5 1.1 90.1 0.862
15.7 16.2 12.2 3.6 52.5 1.0 90.4
10.0 21.5 6.8 2.2 58.5 1.0 89.5
0.927
0.864
-
run 71.5 12.9 20.3 7.7 2.8 55.2 1.1 92.8 0.955
Solid Temp
_ - - G m Temp 0
d ,'
0
02
0
04
06
I l l Coal w y o cool
08
IO
Z/L
Figure 3. Typical temperature profile for Illinois and Wyoming coals. Operating conditions of Table I1 with [ = 0.3 and no water formation reaction.
mogeneous temperature predicted by Yoon. The gas temperature is substantially smaller in the combustion zone. The gas-solid heat transfer parameter, f , was set equal to 0.3 after some experimentation. The effect of this coefficient is reported later in this section. There is not much difference in the predicted composition of the product gas as seen in Table 111, which also reports the experimental data. Both models predict lower methane in the product gas. It is possible that the methane formation reaction is catalyzed by some of the mineral constituents in the coal. Comparison was also made with the plant data reported for the gasifier operated by the Morgantown Energy Technology Center (Desai and Wen, 1978). The operating conditions for this reactor are shown in Table IV. This reactor operates at a much lower pressure (6 atm vs. 25 atm) than the Westfield plant. In addition, a stirrer was provided inside the reactor to prevent coals from agglom-
Ind. Eng. Chem. Process Des. Dev., Vol. 20, No. 2, 1981 317 - Solid Temp
___ Gar Temp
098
-
096-
2 r
094-
8 LT
0920
0
02
04
06
08
10
Z/L
090
Figure 4. Temperature profile with 5 = 5.0. Operating conditions of Table I1 (with Illinois coal).
-
,
I
I
- Solid Temp
---
0
02
04
06
Gar Temp
08
10
Z/L
Figure 5. Temperature profile with 5 = 0.2. Operating conditions of Table I1 (with Illinois coal).
erating inside the reactor. Table V compares the product gas distribution. Run 71.5 shows good agreement whereas the predicted value of COz in run 70.6 is only half the experimental value. Since the operating conditions of both reactors are similar, it is hard to explain this difference. The solid temperature profile is strongly influenced by the reactivity of the coal as shown in Figure 3 which compared the temperature profiles of a highly reactive coal (Wyoming) to the moderately reactive Illinois coal. As mentioned earlier, the gas-solid heat transfer coefficient was computed as fh where h is given by the Gupta and Thodos correlation. Figures 4 and 5 compare the temperature profile for two extreme values of f , 5 = 5 and f = 0.2. When f = 5, there is almost no difference in temperature of the solid and the gas. This case approximates the homogeneous temperature case. For f = 0.2, the temperature difference is substantial. A value of f = 0.3 was found to give results that agreed most closely with the Westfield plant data. One of the differences between this model and Yoon's model is in the treatment of the water-gas shift reaction. Yoon assumed that since this reaction was fast, equilibrium is attained at all points in the reactor. Wen (1972) has compared the kinetics of this reaction at various conditions. The reaction rate in the presence of a commercial catalyst is 1000 times faster than in the presence of charcoal. Coal contains many minerals including iron compounds which can catalyze this reaction. Figure 6 compares the effect of this reaction rate on the conversion of char. At high rates of reaction there is no effect but the char conversion drops dramatically when this reaction rate is reduced below kmol/m3 kPa2 h. Since a high conversion implies higher efficiency, it would be advantageous to promote this reaction in the gasifier. As mentioned earlier, there is some doubt as to whether the water formation reaction does take place in the gasifier. As indicated by Amundson and Arri (1978), a certain minimum oxygen concentration may be required before this reaction can proceed. Figures 7 and 8 compare the composition and temperature profiles obtained for two different values of this minimum oxygen concentration.
02
0
04
08 0 0 ZIL
04
08
Figure 7. Concentration profiles with f o l d = 0.001 and 0.1. Operating conditions of Table I1 for Illinois coal. I400
1
_--
0
--- fOzmin=OOO1 fOzmin=Ol ,'I
02
04
06
08
IO
Z/L
Figure 8. Comparison of solid temperature profiles for the two cases = 0.1 and fold = 0.001. Operating conditions of Table I1 for % inois + l coal.
For the case foz,min = 0.001, the char conversion is incomplete and this lowers the concentration of gasification products significantly. The temperature profile drops sharply when the oxygen concentration is extinguished. The results with fo . = 0.1 (almost no water formetion) are closer to plant indicating that this reaction does not occur to any significant extent in the gasifier. The thermal efficiency, defined as the heating value of the product gas per unit mass of coal feed, is a strong function of the coal to oxygen feed ratio and the coal to steam feed ratio. Results using the heterogeneous model follow trends similar to those reported by Yoon et al. (see Figure 9). One observation was that in almost all cases the maximum efficiency was obtained when the conversion of char is just complete with a very small ash layer in the reactor. Appearance of char in the ash indicates insufficient gasification. The maximum solid temperature is another important variable in the operability of the gasifier. In the Lurgi
E,
318
Ind. Eng. Chem. Process Des. Dev., Vol. 20, No. 2, 1981
, 2:
1
30
35
FIXED CARBON TO OXYGEN RATIO
Figure 9, Effect of fixed carbon to oxygen feed ratio on char conversion, thermal throughout and heat recovery. Operating conditions of Table I1 for Illinois coal. ?UCC
c
i;
c
i
L J
zsr 7 -
3 1
1 -
r IEI “ECv T I
OXYSE’I FE‘‘
VT
Figure 10. Effect of fixed carbon to oxygen feed ratio on maximum temperatures and ash layer thickness. Operating conditions of Table I1 for Illinois coal.
gasifier this temperature must be kept below the softening point of the ash. Sensitivity studies using the simulation model show that the maximum solid temperature decreases as the conversion of char is increased reaching a minimum when the conversion is just complete. Beyond this point the maximum solid temperature increases. This trend is illustrated in Figure 10. The initial decrease in maximum solid temperature is caused by the small decrease in solid heat capacity flux with increasing conversion. The other parameters influencing the maximum solid temperature, such as oxygen partial pressure and char concentration in the combustion zone, remain nearly constant. After the unreacted char begins to appear in the ash, the maximum solid temperature increases because the conversion of char is less in the combustion zone. Conclusions The heterogeneous model of moving-bed gasifier indicates that substantial energy exchange occurs between the solid and gas stream, especially in the combustion zone. The temperature profile is strongly influenced by the heat transfer coefficient. A value smaller than that predicted by the Gupta and Thodos correlation for nonreactive systems is indicated for the gasifier. The split boundary value problem resulting from the material and energy balance equations poses severe nu-
merical difficulties due to the small heat capacity of the solid stream. This difficulty can be overcome by assuming that the sensible heat carried away by the solid stream is negligible. This was confirmed by numerical calculations at selected points in the reactor. The effect of the minimum oxygen concentration required for the combustion of hydrogen to form water was investigated. Better agreement with plant data is obtained if this value is set high &e., no water formation reaction). Maximum thermal efficiency (defined as the heating value of the product gas) is attained in the neighborhood where char conversion is nearly complete and a small ash layer is present. This optimum operation also favors a lower maximum solid temperature in the gasifier. Nomenclature aij = stoichiometric coefficient of reactant i in reaction j bij = stoichiometric coefficient of product i in reaction j C, = heat capacity fo, = oxygen concentration Fi = flux of component i P i = flux of component i in the feed h = gas-solid heat transfer coefficient H . ( T ) = enthalpy of species j at temperature T Ahi(T) = heat of reaction at temperature T N = number of components in the gas phase ri = rate of reaction S , = wall surface area/unit volume S, = solid surface arealunit volume T = temperature of gas stream = temperature of solid stream T, = wall temperature U = gas-wall heat transfer coefficient x , = conversion of char z = distance from bottom of the reactor
6
Creek Letters a = number of hydrogen atoms per carbon atom in char
0 = number of oxygen atoms per carbon atom in char
y = stoichiometric coefficient of oxygen in char-oxygen re-
action
$. = multiplicative factor for gas-solid heat transfer coefficient
Literature Cited Amundson, N. R.; Ami, L. E. A I C M J . 1978, 24, 87. Ani, L. E.; Amundson, N. R. A I C M J . 1978, 24, 72. Ark, R.; Amundson, N. R. “Mathematical Methods in Chemical Englneering”, Voi. 2, Prentlce-Hall; Englewood Cliffs, N.J., 1973; pp 318-328. Cho, Y. S. M.S. Thesis, Washington University, St. Louis, 1980. Desai, P. R.; Wen, C. Y . “Simulation of a METC Fixed Bed Gaslfier”, presented at AIChE Meeting, Mlaml, 1978. Eigin, D. C.; Galland, J. M.; Dinsmoor, B.; Tsang, T. “American Coals in Lurgi Pressure-Gastfication Plant at Westfieid, Scotland’, Sixth Synth. Pipeline Gas. Symp., 1974, p 247. Gregory, D. R.; Llttiejohn, R. F. The BWRA Monthly Bull. 1085. 29, 173. Gupta, A. S.; Thodos, G. AIChE J . 108S, 9 , 751. Hebden, D. “High Pressure Gasfflcation Under Slagging Conditions", presented at Seventh Synth. Pipeline Qas Symp., Chicago, Oct 1975. Hoogendorn, J. C. ”Gas from Coal with Lurgi Gasification at Sasoi”, IGT Symp. Papers, “Clean Fuels from Coal”, 1973, p 11 1. Rossberg. M. 2. 2. flektrochem. 1958, 80. 952. Rudolph, P. ”The Lurgi Process, The Route to S.N.G. from Coal”, Proc. Fourth Symp. Synthetic Pipeline GRS,Chlcago, 1972. Wen, C. Y. “Optlmization of Coal GasiRcation Processes”, R and D Report No. 66, Interim Report No. 1; Offife of Coal Research, 1972. Yoon, H.; Wel, J.; Denn, M. M. AIChEJ. 1978, 24, 885. Yoon. H. Ph.D.Thesis, University of Delaware, 1978.
Received for review April 21, 1980 Accepted October 16, 1980
