Ind. Eng. Chem. Process Des. Dev. 1981, 20, 675-682
875
Coal Gasification in a Pilot Scale Fluidized Bed Reactor. 1. GasificatOon of a Devolatilized Bituminous Coal Mark J. Purdy, Rlchard M. Felder," and James K. Ferrell Department of Chemical Engineerlng, North Carollna State Unlverslty, Ralebh, North Carolina 27650
A devolatilized bituminous coal was gasified with steam and oxygen in a 15.2-cm i.d. fluidized bed reactor at a
pressure of 790 kPa (100 psig), average bed temperatures varying between 925 and 1025 O C , and molar steam-to-carbon feed ratios between 0.6 and 1.6. Material balances were obtained on total mass and major elements (C, H, 0, N, S), and carbon conversion and product composition data were correlated with a two-stage reactor model. The incorporation of water-gas shift reaction kinetics into the model provides significantly better correlations than are obtained by assuming shift reaction equilibrium.
Introduction Since 1976, the Department of Chemical Engineering at North Carolina State University has been engaged in a research study on coal gasification and gas cleaning sponsored by the U.S.Environmental Protection Agency. The gasification facility used for this research is a 15.2-cm i.d. fluidized bed reactor. The overall objective of the project is to characterize the gaseous and condensed phase emissions from the gasification-gas cleaning process and to determine how emission rates of various pollutants depend on adjustable process parameters. Specific tasks to be performed are the following. (1) Identify and measure the gross and trace species concentrations in the gasifier effluent streams. (2) Correlate measured emission levels with coal composition and gasifier operating variables. (3) Perform material balances around the gasifier, raw gas cleanup system, and acid gas removal system, and determine the extent to which selected species are removed from the synthesis gas in each subsystem. (4) Correlate measured extents of conversion and removal efficiencies for various species with system operating variables. (5) Evaluate and compare the performance characteristics of alternative acid gas removal processes. (6) Use the results to develop models for the gasification and gas cleanup processes. A complete description of the facility and operating procedures is given by Ferrell et al. (1980), and in abbreviated form by Felder et al. (1980). In the initial series of runs on the gasifier, a devolatilized Western Kentucky No. 11coal was gasified with steam and oxygen, and material balances were obtained on total mass and on major elements (C, H, 0, N, SI. In addition, a two-stage model for the gasifier was formulated and used to correlate the results of the char gasification runs. This paper summarizes the modeling and model parameter estimation procedures and compares the experimental results with model predictions. A more detailed description of the experimental data and data analysis procedures is given by Ferrell et al. (1981). Extensions of the research to other coals and elaborations of the model will be presented in subsequent papers. Experimental Section The gasifier is a 15.2-cm (6-in.) id. Schedule 40 pipe (316 SS) enclosed in several layers of insulation and contained in a 61-cm (24-in.) i.d. Schedule 80 carbon steel pipe. The overall height of the unit is roughly 3.7 m (12 ft). The gas feed is introduced into the reactor through three feed nozzles spaced triangularly near the bottom of the reaction 0196-4305/81/1120-0675$01.25/0
chamber. Coal is fed at the top of the reactor and removed at the bottom by nitrogen-purged screw conveyors. The height of the bed above the feed nozzles is normally 1m (3 ft), although occasional runs with a 1.5 m (4.5 ft) bed have been made. The temperature profile in the bed is monitored by means of six thermocouples located at various vertical positions within a central tube in the reactor. Differential pressure taps are set at 38 cm and 89 cm above the feed cones. The level of the fluidized bed is monitored with a nuclear level gauge and is controlled by adjusting the char removal screw rotation rate. In the PCS (particulates, condensables, and solubles removal) system, the make gas is passed through a cyclone separator, where most of the elutriated particles are removed, and then through a venturi scrubber, which removes water-soluble and condensable compounds. The quenched gas stream is fed to a receiving tank in which the condensate is collected, and the emerging gas then passes through a shell-and-tubeheat exchanger, a demister, and a coalescing filter. The gas leaving the filter is either burned in a shielded flare or fed to the acid gas removal system. The condensate in the receiving tank can either be recirculated to the venturi or discarded. Signals from 96 process variable sensors in the plant are sent to a control panel, which in turn sends them to various analog display facilities, a video display terminal, a Honeywell TDC-2000 process control computer, and a PDP11/34-based data acquisition and processing computer. The TDC-2000 regulates 16 different control loops in the plant. Samples of all of the feed and effluent streams are taken at the beginning and end of each gasification run, and gas and liquid samples are also taken periodically during the steady-state operating period. Gas samples to be analyzed for the principal make gas components (CO, H2, C02, N2, CHI) are taken in l-L stainless steel bombs, and samples to be analyzed for sulfur gases are taken in l-L glass bombs. To aid in the analysis of the run data, a computer program has been written which takes as input the operating parameters and chemical analysis results for a run, and calculates and prints out such quantities as the make gas flow rate and composition; the carbon and sulfur conversions achieved in the run; material balances on total mass, carbon, hydrogen, oxygen, nitrogen, sulfur, and selected trace elements, and an energy balance. A representative program output is shown by Felder et al. (19801, and a 0 1981 American Chemical Society
676
Ind. Eng. Chem. Process Des. Dev., Vol. 20, No. 4, 1981
program documentation, summary of calculation procedures, and complete source code listing are given by Ferrell et al. (1981). The reactor is started up by blowing preheated nitrogen through a shallow bed and then gradually adding more coal and oxygen and cutting down the nitrogen flow until the bed ignites, next adding steam, and finally commencing the removal of spent char when the bed height reaches the desired level. About 3-4 h is required to achieve steadystate operation. The steady state is maintained for 1-2 h, during which time several sets of gas and liquid samples are taken, and the reactor is then shut down. The superficial velocity in the bed at steady state is on the order of 0.3 m/s (1.1ft/s), which is three to five times the minimum fluidization velocity at the reactor conditions. There is a good likelihood that the bed is in a slugging mode of operation at steady state, but any consequent oscillations are too small to be detected by the system instrumentation. Energy balances on the reactor indicate that operation is essentially adiabatic. The feed to the gasifier in the series of runs discussed in this paper was a Western Kentucky 11bituminous coal that had been devolatilized at 1100 OC and sized to 10 X 80 mesh (mean particle diameter = 0.5 mm). On a moisture-free basis the char contained 85% fixed carbon, 2% volatiles, and 13% ash. The ultimate analysis showed 81% C, 2% 0, 0.4% H, 1% N, 2.6% S, and 13% ash. Gasifier Model The primary function of the North Carolina State University gasification reactor is to provide a reproducible and realistic feed gas for studies of the potential environmental impact of coal gasification processes and to determine the performance characteristics of various gas cleaning processes. The development of a correlative and predictive model of the gasifier was felt to be an indispensable adjunct to planning and implementing the overall experimental program. The first step toward the achievement of this goal was to formulate the simplest possible model which incorporated the principal gasification reactions and the gross physical characteristics of the reactor, and to determine the degree to which the gasifier performance could be correlated by this model. The model adopted is basically that developed by the Institute of Gas Technology (1978), with several modifications to be described. It treats the gasifier as two perfectly mixed stages, with combustion taking place in the first stage and gasification in the second. It does not include effects of coal pyrolysis, so to get an accurate prediction of make gas composition for the gasification of a raw coal a third stage would have to be added. This was not required in the present study, which involved a devolatilized coal as the feedstock. Gasification Kinetics. The following six reactions are presumed to take place C + H2O = CO + H2 (1) C + 2H2 = CHI (2) 2C + H2 + H2O = CO + CHI (3) (4) CO + H2O = COZ + Hz
c + 1 / 2 0 2 = co c + 02 = coz
(5) (6)
Reactions 5 and 6 are the oxidation steps required to supply heat for the remaining reactions. These two reactions are assumed to occur instantaneously in a zone of negligible volume separate from the gasification zone. All oxygen in the feed gas is assumed to be consumed to form
CO and COz, with the combustion product distribution being governed by the relation c + a02 = (2 - 2a)CO + (2a - 1)COZ (7) where a, the combustion product distribution coefficient, is an adjustable model parameter. A value of a = 0.5 indicates that all CO is formed, while a = 1.0 indicates that only COz is formed. Reactions 1, 2, and 3 are the reactions with which Johnson (1974) at the Institute of Gas Technology correlated gasification kinetics data. Reaction 1 is the conventional steam-carbon reaction. Reaction 3 is assumed to be an independent reaction, although it is attainable as a linear combination of 1 and 2. The correlation used by Johnson to describe the carbon conversion is given by
r = fLkT(l- f c ) 2 / 3 exp(- 6f:) (8) where r is the rate at which the carbon is gasified, kT is the sum of the rate constants for reactions 1 , 2 , and 3, f, is the fractional carbon conversion, and 6 is the kinetic parameter which depends on gas compoeition and preasure. Expressions for the rate constants and 6 are given as follows. kI
= exp(9.0201- 31705/T)
(
1- -
/ )E: : :
1+ 1
exp(-22.2160
+ 44787/T)
X
=
kII
PH?exp(2.6741 - 33076/T)
(
1- /E n);:p
[ I + PH,exp(-10.4520 +19976/T)] (10) kIII
= P H 2 1 / 2 P H , ~exp(12.4463 - 44544/T) (1 -
PcH4pco
)/[
PHJ'H~OKIII~
X
1 + exp(-6.6696
+
where Pi = partial pressure of component i, atm, T = bed temperature, OR,and KI,Kn, Km = equilibrium constants. The relative reactivity factor, f L , is determined from f~ = fo exp(8467/Td (13) where Tois the maximum temperature to which the coal has been exposed prior to gasification. The reactivity coefficient, fo, which is an adjustable parameter whose values depend on the particular char used, has values ranging from 0.3 for low-volatile bituminous coal char to about 10 for North Dakota lignite (Johnson, 1974). Reaction 4 is the water gas shift reaction, often assumed to be at equilibrium in gasification processes. Results to be described indicate this may be a bad assumption, leading to the necessity of incorporating shift kinetics into
Ind. Eng. Chem. Process Des. Dev., Vol. 20, No. 4, 1981 877
Eigenvalue stability is considered to have been achieved when both of the following conditions are satisfied
Table I. Empirical Constants Obtained from Least-Squares Fit of Equilibrium Dataa react ion C t H,O= CO + H, C + 2H, = CH, 2C + H, + H,O= CO t CH, CO t H,O= CO, + H, a
0 0
a1
-29010.6 17744.4 -11266.2 8395.2
17.09 -12.23 4.86 -4.29
(a)
In K E = a,(l/T) t a , , where T i s in degrees Rankine.
the model. The rate expression used is that given by Wen and Tseng (1979)
r4 = 1.6652 X 104V(l - t)fwgexp(-25147/T)PG
(14)
where V = bed volume (cm3),G = yco - y ~ g c q / & & r ~ ) , t = bed void fraction, f,, = adjustable shift reactivity parameter (varies from char to char), Kw = equilibrium constant, and P = reactor pressure. The equilibrium constants for the water-gas shift reaction and for reactions 1,2, and 3 were taken from Lowry (1963), and were fit to the equation In K = ao/T + al (15) by linear regression (Alexander, 1978). The results are given in Table I. Reactor Simulation. The simulation program treats the fluidized bed as two perfectly mixed stages, with combustion proceeding to 100% oxygen consumption in the first stage and gasification taking place in the second. In the combustion stage, the molar flow rates of carbon and oxygen are reduced by the appropriate amounts from their feed values, and the flow rates of CO and C02 are calculated as
Qco' =
PO-
a)/aIqo2
(16)
qco; = [(2a - l)/alqo2
(17)
The effluent from the second stage is next estimated by assuming a carbon conversion of 20% in this stage (based on the carbon flow out of the first stage), with all conversion taking place by the steam-carbon reaction, eq 1. The methane effluent rate is estimated to be 1% of the combined molar flow rates of CO, C02, H2, and H20. The calculation of the effluent composition is performed iteratively. The rates of the four reactions determined in a previous iteration or assumed for the initial iteration are substituted into the perfect mixer balance equation, which with the reaction stoichiometry yields the product gas component flow rates. The flow rates are in turn used to calculate the composition of the gas in the reactor, and the reaction rates are recalculated. Convergence is said to be achieved if the percentage differences between successively determined rates are all less than 0.1%. In initial attempts to implement this model, simple successive substitution was used to proceed from one iteration to the next. Convergence was found to be exceptionally slow, typically requiring hundreds of iterations, and in some instances instability was observed. In the procedure developed to overcome these difficulties, new rates are calculated by modified successive substitution, and convergence is accelerated by the Dominant Eigenvalue Method (Orbach and Crowe, 1971). In each iteration, the dominant eigenvalue E of the vector of reaction rates is calculated
where the superscripts indicate the iteration number.
[(E&- Ek-')/Ekl
I0.01
(19)
(b) n 1 5 (20) where n is the number of iterations since eigenvalue stability was last achieved. When this condition is met, new rates are calculated as
r t f l = rt-l
+ p [ ( r t - rF-l)/(l - Ek)]
(21)
where p is a damping factor. When eigenvalue stability is not achieved (i.e., when eq 19 and 20 are not satisfied), new rates are calculated by modified successive substitution
r t + l = qri,:
+ (1 - q)ri,:
(22)
where r i , t is a rate calculated in the kth iteration from the kinetic rate equations, ri,: is an estimated rate for the kth iteration, and q is a second damping factor. Preliminary studies showed that damping factor values p = 0.7 and q = 0.3 provide rapid and stable convergence; these values were therefore incorporated into the model program. The new rates are used to calculate new product flow rates for the principal gas components qco = QCO'
+ M C k l+ 0.5r3)- r4
(23)
+ r4
(24)
qcoz = qco;
q H z = Mc(rl - 2r2 - 0.5r3) + r4
+
~ H = ~ ~OH ~ o ' Mc(-rl -
QCH, =
(25)
0 . 5 ~-) r4
(26)
K(r2 + O W
(27)
where M , is the molar holdup of carbon in the fluidized bed and q{ is the flow rate of component i leaving the first stage. The flow rate of H2S is estimated by assuming that the sulfur conversion equals the carbon conversion, and all converted sulfur forms H2S. (Provisions for predicting the evolution of several sulfur gas species, including COS and CS2,are currently being developed for incorporation into the model.) The newly determined product rates are then used in the kinetic rate expressions to calculate reaction rates, and convergence is again checked. This procedure is continued until the convergence criteria are satisfied. After the make gas composition and component flow rates have been determined, the carbon conversion in the gasifier is calculated as the flow rate of atomic carbon leaving as CO C02 + CHI divided by the rate at which carbon enters in the feed coal. The make gas flow rate and the fractional steam consumption are also calculated. Results The bituminous coal char used as the feedstock in these studies was gasified at a pressure of roughly 790 kPa (100 psig), average bed temperatures varying between 925 and 1025 "C, and various steam, oxygen, and coal feed rates. Material balances were obtained on total mass and major elements (C, H, 0, N, S). Fifty-six runs were completed, of which the first 13 were used primarily for the development of operating and sampling procedures and refinement of analytical methods. The data from the remaining 43 runs are presented and discussed by Ferrell et al. (1981). All subsequent references will be to the latter runs. A summary of the principal experimental results is shown in Table 11. In processing the data from these runs, the purge nitrogen flow rate was adjusted to close the nitrogen balance on the gasifier. This was done to correct
+
678
Ind. Eng. Chem. Process Des. Dev., Vol. 20, No. 4, 1981
Table 11. Summary of Char Gasification Run Dataa mol of H.V. of H,O/ solids dry make sweet p? mol of holdup, %C gas rate, gas, run T,"F pug C lb conv SCFM Btu/SCF 37.5 13.44 267.0 14.6 0.63 GO-14 1790.3 103.0 12.23 283.0 14.9 44.9 0.79 GO-15 1786.8 100.8 44.2 19.8' 10.70 304.6 98.8 0.70 GO-16 1778.8 13.00 332.2 15.0 30.5 1.15 GO-17 1778.6 102.0 42.4 18.67 339.4 13.7 1.15 102.4 GO-18 1872.0 12.90 326.0 23.3' 49.6 0.97 GO-19 1761.3 101.9 13.38 323.3 22.0' 59.6 102.0 1.15 GO-19A 1777.9 21.36 321.7 15.8' 104.3 1.26 GO-20 1881.5 106.5 17.58 313.3 17.1' 99.5 103.7 1.05 GO-22 1868.6 42.1 16.61 307.9 30.1 1.01 GO-24 1878.1 105.0 40.2 15.83 306.1 26.8 1.13 GO-25 1862.2 101.7 40.4 10.35 319.9 1.00 24.0' GO-26 1714.3 101.0 17.08 335.9 17.3 56.9 1.28 GO-28 1820.5 104.0 335.9 16.97 24.3' 47.1 1.13 GO-29 1816.5 104.0 16.19 336.0 14.6 50.9 GO-30d 1830.2 103.5 1.27 19.11 340.1 1.03 15.5 39.9 GO-31d 1832.7 106.2 10.74 196.6 22.2 16.4 103.9 0.73 GO-34 1693.7 21.1 22.2 10.54 273.5 103.4 0.94 GO-35 1683.0 15.18 243.1 19.1 26.2 104.7 0.68 GO-36 1782.8 14.71 314.4 17.3 32.9 1.31 GO-37 1786.9 105.4 16.72 323.7 15.9 56.0 104.9 1.33 GO-38A 1829.7 15.9 54.3 16.22 324.7 1.33 1829.8 104.6 GO-38B 17.25 319.0 16.8 65.1 GO-43 1819.6 104.5 1.31 11.73 296.0 18.4 31.6 101.6 1.31 GO-44B 1699.8 14.4 51.2 18.17 315.9 106.1 1.32 GO-27 1806.7 15.97 121.7 20.4 20.3 103.1 0.81 GO-45 1696.6 42.8 17.51 212.7 0.88 16.6 103.5 GO-46 1822.7 18.41 216.9 0.82 17.3 42.9 GO-47 1814.6 104.2 54.2 16.38 315.5 16.3 1.33 GO-48 1818.3 103.8 13.70 294.1 0.83 15.1 46.6 78.4 GO-49 1818.3 40.0' 13.07 305.1' 0.83 16.7 GO-51 1815.3 78.2 18.88 336.0 16.7 40.7 GO-53 1818.7 104.2 0.98 11.67 315.3 16.5 37.0 GO-55 1738.8 77.5 1.05 15.3 55.3 16.61 329.3 GO-56 1825.1 104.0 1.55
-
%mass rec 103.7 100.3 98.8 103.3 105.5 101.3 100.5 101.4 100.5 103.9 103.3 99.7 103.5 94.2 99.3 102.4 97.6 100.0 98.8 97.6 95.5 95.6 103.7 98.9 103.7 99.1 96.3 97.0 102.2 103.2 104.0' 101.2 101.8 98.2
A T accept worst % rec 116.8 (H) 105.3(H) lO9.l(H) 116.8(H) 120.1(H) 111.1(H) 106.8(H) 107.3(C) 106.6(H) 107,6(H) 106.1 (H) 104.8(H) 107.7(0) 87.9(C) 93.3 (C) 106.7(H) 107.4 (H) 106.8(H) 118.0(H) 109.1(H) 94.3 (0) 95.6(0) 92.2(H) 101.8(H) 107.2(C) 103.1(H) 9O.O(C) 95.6(H,O) 104.6(C,H) 88.6(H) 108.7'(0) 119.4(H) 106.0 (C) 97.9(H)
?
8 62 31 26 37 53 26 19 37 28 56 43 31 58 21 22 46 51 49 39 36 33 35 36 13 44 39 50 43 25 50 54 29 26
no yes no no no no yes yes yes yes yes yes yes no yes yes yes yes no no yes yes yes yes yes yes no yes yes no no no yes yes
Runs 14-27 A T is for 10-
a Note: runs 21, 23, 32, 33, 39-42, 50, 52, and 54 not included due to equipment failures. 6-cone distributor. 40 in.; GO-28 A T is for 10-35 in. ' 52-in. bed.
for belatedly discovered calibration errors in the nitrogen flowmeten and for suspected purge nitrogen leakage in the system. The other quantities used in the run data analysis reflect actual operating and analytical data. Mass Balances. The raw data for all runs were reduced using the previously described data processing program. Criteria for acceptance of a run were arbitrarily chosen following inspection of the mass balance results. A run was considered acceptable if the total mass recovery was between 95% and 105%, and if the worst of the recoveries of elements C, H, and 0 was between 92% and 108% (see Table 11). Based on these criteria, 22 of the 34 runs reviewed were judged acceptable and are designated by crossed circles in the model correlation plots to be described (Figures 1-8). Points with filled circles on these plots correspond to runs with totalmass recoveries between 95% and 105% and worst element recoveries between 94% and 106% . Open circles are used for all other runs. Temperature Effects. The effect of the average bed temperature on the dry, nitrogen-free make gas flow rate is shown in Figure 1. For the points shown, the molar steam to carbon ratio varied from 0.92 to 1.15. The plot indicates that the make gas flow rate is highly sensitive to the average bed temperature, with scatter due mainly to the small steam-to-carbon ratio differences and differing feed rates. The high sensitivity makes determination of the average bed temperature crucial for good model predictions. The effect of temperature on carbon conversion cannot be clearly shown due to the coupled effects of other variables, such as gas velocity in the bed and coal feed rate,
OF^
1
I8 Molar Steam t o Carbon R a t i o o f 0.92 t o 1.15
Q ;
I
O 1
worse than 8% w l t h i n 8%
1680
1120
1160
Ql
0
@ A l l e l n n e n t mass balances
1800
Average Bed Tenperature.
I840
I ea0
OF
Figure 1. Effect of average bed temperature on make gas flow rate.
which were difficult to hold constant from one run to another. However, the model for coal gasification developed in this study does a reasonably good job of correlating the data, including temperature effects, as will be shown subsequently. Steam/Carbon Feed Ratio Effects. The effect of the steam-to-carbon feed ratio on the make gas flow rate is
Ind. Eng. Chem. Process Des. Dev., Vol. 20,
8o
18
N
-g
17
t
O E l e n e n t mass balance worse than 8% @All e l n e n t mass balances u l t h i n 8% @All e l m e n t mass balances w i t h i n 61
e @
t3 @@
'Ot
'
No. 4, 1981 679
e
a/
0
Molar Steam t o Carbon R a t i o
Figure 2. Effect of steam-to-carbon feed ratio on make gas flow rate.
shown in Figure 2. At a given temperature the effect of increasing the steam feed rate is to increase the make gas flow rate. An added benefit to operating with relatively high steam-to-carbon ratios in the fluidized bed gasifier is a reduced tendency for the char to clinker. During the first three runs,in which the steam-to-carbonratio was less than 0.8, it was not uncommon for the removal screw to jam because of a clinker lodging in the screw housing. During the remainder of the runs, with the steam-tocarbon ratio greater than 1.0, few problems arose due $0 clinker formation. The effect of steam-to-carbon ratio on carbon conversion cannot be shown graphically for the same reasons given in the section on temperature effects. Sulfur Conversion. Measured sulfur conversion, assumed to equal the carbon conversion by the model, is plotted against carbon conversion in Figure 3. As this figure shows, in most cases the sulfur conversion was greater than the carbon conversion. Studies are currently underway to put the prediction of sulfur gas evolution on a firmer theoretical basis. Model Parameter Estimation. In ita present form, the model has three adjustable parameters: (1)the coal reactivity coefficient, fo; (2) the combustion product distribution coefficient, a, which specifies the split between CO and C02 in the products of the combustion stage of the gasification; and (3) the water gas shift reactivity parameter, fw These parameters were evaluated by using a Pattern Search routine (Beightler et al., 1979) to minimize a function of the sum of squared deviations between predicted and measured values of gasifier performance variables. The function is 2
Y=
+
( ) (Yco~c#-aw)]] Yco,~ - YCO,m
+
YCO,m
+
where wl,w2, and w 3are arbitrarily set weighting factors, 4 is the dry make gas flow rate, f is the total carbon conversion, yi is the mole fraction of gas i in the product gas, and the subscripts p and m indicate predicted and mea-
Percent Carbon Conversion
Figure 3. Percent sulfur conversion vs. percent carbon conversion.
sured values. The nine runs with the best mass balance closures (GO 26-28, 44B-49) were chosen to provide the data base for the parameter estimation. The values of fo and a were first determined by setting w1 = 1,w2 = 1, and w 3 = 0. This weighting scheme was chosen since the predicted make gas flow rate and carbon conversion are highly sensitive to the values of foand a, and relatively insensitive to the value of fw The results were fo = 0.50 (29) a = 0.95
(30) Holding fo and a constant at the above values, the function of eq 28 was again minimized, with w1 = w 2 = 0 and w 3 = 1,in order to determine the optimal value of fW This weighting scheme was used as fv has a large influence on the predicted product gas composition, but only a small effect on the predicted carbon conversion and make gas flow rate. This gave the result of fwg
= 9.9 x 104
(31)
Substitution of this value of a into eq 7 indicates that 90% of the carbon oxidized in the combustion stage forms C02 and 10% forms CO. An equation by Arthur (1951) predicts a = 0.57 at 1400 O F and a = 0.52 at 2000 O F , while several gasification studies have assumed a = 1.0 (Caram et al., 1979; Institute of Gas Technology, 1978; Pukanic et al., 1978; Sundaresan and Ammundson, 1979). Johnson (1975) developed a correlation for char reactivity f o = 6.2yC(l- YJ (32) where yc is the dry, ash-free carbon fraction in the original raw coal. Equation 32 predicts a value of fo = 1.1, which is larger than that determined in this study. The difference may be due to the differences between the microbalance used by Johnson and the fluidized bed of this study. This value of fw = 9.9 X 10" indicates that the shift reaction rate is approximately five orders of magnitude leas than the rate typically obtained in cataytic shift reactors. Wen and Tseng (1979) used a shift reactivity value of 1.7 X lo4 in modeling the gasification of a bituminous coal by the Synthane process.
680
Ind. Eng. Chem. Process Des. Dev., Vol. 20, No. 4, 1981 10
/
1. Carbon Conversion
0
Heating valve of Sweet Gar (Btu/SCFl
360
0 Element mass
balance worse than 8I $ a l l e l m o t mass balance5 withln 8 1 O b 1 1 element mass balances within 6%
320
5c
-
280
,zu L
T
a. 240 c
200
O E l e m n t ma55 balance worse than 8%
$ A l l element mass balances w i t h i n 8%
160
@ A l l e l m e n t mass balancer w i t h i n 6% 20
50
40
30
60 120, 2w
160
240
280
320
360
Experimental Experimental
Figure 4. Predicted vs. experimental carbon conversion.
Figure 6. Predicted vs. experimental heating value of sweet gas. Dry Make Gas Flaw Rate (SCFM)
@
0.8
0
Y
Y
H2
OElement mass balance @ A l l element mass balancer w i t h i n 8%
0.7
-
0.6
-
co2
K'wq O E l e m e n t ma56 b a l a n c e worse than 8%
/ /
$All
element mass balances u l t h i n 8%
.All
e l e n e n t mass balances w i t h i n 6%
0
0
; c.5
ia
P
Zc
O
0.4
a
I
1
I
t
12
14
16
18
@ % / " & o@ @
8
*/,2
/
0.3
IO
0
@ @
0
63%
0
@
0
0
e
@/ 0.2
0.3
0.4
0.5
0.6
0.7
Expermental
Figure 5. Predicted vs. experimental dry make gas flow rate.
Experimental
Figure 7. Predicted vs. experimental K value,
Model Predictions. Using the optimal parameter values, the model was run for all gasifier runs listed in Table 11. Plots of predicted vs. measured values of carbon conversion, dry make gas flow rate, and sweet gas heating value are shown in Figures 4-6. (The sweet gas is defined as the dry make gas with the C02and H&3removed.) The reasonably close proximity of most points to the 45' line is gratifying in view of the simplicity of the model. The proximity of the points corresponding to the "best" runs (from the standpoint of satisfying mass balances) is even more satisfying. For each run, the ratio K = ([CO,l /([COl (H201) (33) was calculated, where [I is the mole fraction of the indicated species in the product gas. This quantity would equal the water-gas shift equilibrium constant at the reactor temperature if the reaction proceeded to equilibrium. A plot of the predided w experimental values of this ratio is given in Figure 7. The substantial degree of scatter seen in Figure 7 may be attributed to a combination of factors, including the
simplicity of the model and the interdependence of the constituents of K, whereby an experimental error in one of them affecta the values of the others. Perhaps the major factor, however, is the sometimes considerable departure from isothermality in the bed, with temperature differences from top to bottom as high as 50 O C occasionally being observed. If this is in fact responsible for deviations between measured and predicted compositions, use of a multi-stage model which allows for temperature gradients should lead to marked improvements. Development of such a model is currently under way. The significance of the plot of Figure 7 emerges when it is compared with that of Figure 8, which shows the values of K predicted assuming shift equilibrium at the bed exit temperature. This assumption leads to the overprediction of K by as much as a factor of 2. It might initially be thought that the product gases are shifting to equilibria at lower temperatures as the gases cool prior to analysis; however, lowering the equilibrium temperature would make the value of K even higher for the exothermic shift reaction, which would increase the discrepancy.
Ind. Eng. Chem. Process Des. Dev., Vol. 20, No. 4, 1981 681 0.8 L
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Figure 8. Predicted vs. experimental K value assuming shift equilibrium.
These results indicate the necessity of incorporating shift kinetics into the model. Summary and Conclusions A devolatilized bituminous coal was gasified with steam and oxygen in a fluidized bed reactor at varying temperatures and steam-to-carbon feed ratios, and the conversion and product composition data were correlated with a two-stage reactor model. The average bed temperature and the molar steam-tocarbon ratio are the two most important factors in determining the make gas flow rate. Small changes in the bed temperature can cause significant changes in the make gas flow rate. Increasing the steam-to-carbon ratio, in addition to increasing the make gas flow rate, reduces the tendency to produce clinkers. The reactor model describes the reactor performance reasonably well. The optimal model parameter valueschar reactivity fo = 0.50, CO/C02 distribution coefficient a = 0.95, and shift reactivity fw = 9.9 X lO+-yield predicted values of the carbon conversion and make gas flow rate which compare well with measured values. The incorporation of shift reaction kinetics into the model gives improved results compared to the assumption of shift reaction equilibrium. A rough approximation of the sulfur gas production rate may be made by assuming that the sulfur conversion equals the carbon conversion. However, this approximation generally underpredicts the rate of sulfur gas production. Future Plans As part of the gasification experiments described in this report, measurements were made of emissions of various sulfur gases and trace elements from the gasifier. The results of these measurements and of measurements on acid gas removal will be included in subsequent publications. Beginning in the Fall of 1980, gasification runs were commenced using a New Mexico subbituminous coal as the feedstock. Measurements of the type described in this paper will be performed in this phase of the project. In addition, extensive studies will be made of condensed phase (wastewater)pollutant emissions,which are expected to be far more significant than those observed for the
relatively clean devolatilized coal used in the first series of runs. A large and rapidly growing body of literature deals with the modeling of the gasification of coal in various types of reactors. Studies that specificiallyconcern fluidized bed gasification include the works of Caram and Amundson (1979) and Sundaresan and Amundson (19791, who compared several alternative kinetic schemes and hydrodynamic models; Weimer (1978), who considered jet effects near the gas inlet and particle size distribution effects; and Wen and Tseng (19791, who developed a two-phase three reaction-zone model and included the effects of devolatilization. Although the models developed in these studies are well-conceived and rigorous in their application of sound engineering principles, there is some question as to whether their level of sophistication is justified by the precision of available gasification data. Much of the existing data base was generated in pilot units as the Institute of Gas Technology (Hygas Process) and the Pittsburgh Energy Research Center (Synthane Process) under conditions that could be only approximately defined, with levels of reproducibility that do not lend themselves t~ discrimination between models. Future research in the present program will be devoted to comparing the abilities of the models cited above to correlate the conversion data generated in the pilot plant reactor and to determine the degree to which they improve on the correlations obtained with the simple model presented in this paper. In addition, rate laws for pyrolysis and evolution of sulfur gases will be incorporated into the model, which will then be used to correlate gaseous pollutant levels in the reactor make gas. Acknowledgment Among the many people who contributed significantly to this work are Professor Ronald Rousseau, Bill Willis (computer operations), Robert Kelly and S. Ganesan (plant operation and data analysis), Gary Folsom (plant operation), Larry Hamel, Kwin Black, Kathy Steinsberger, and Mary Minogue (laboratory analysis), and Terrie Cavanaugh (report preparation). Nomenclature a = combustion product distribution coefficient ao, al = constants in eq 15 E = dominant eigenvalue f, = fractional carbon conversion f L = relative reactivity factor fo = constant in eq 13 f = shift reactivity parameter F=variable in eq 14 k , = rate constant for reaction i kT = combined gasification rate constant K,= equilibrium constant for reaction i M, = molar holdup of carbon in bed n = number of iterations since last acceleration step p = damping factor in eq 25 P,= partial pressure of species i q = damping factor in eq 26 q, = molar flow rate of species i qr' = q1 at outlet of combustion stage r, = rate of reaction i T = average bed temperature To= coal pretreatment temperature V = bed volume yc = dry, ash-free carbon fraction in feed coal y, = mole fraction of species i Greek Letters 6 = variable in eq 8 c = bed void fraction 4 = dry make gas flow rate
602
Ind. Eng. Chem. Process Des. Dev. 1901, 20, 682-685
Literature Cited Alexander, D. W. Ph.D. Thesis, North Carolina State University, Raleigh, NC, 1978. Arthur, J. R. Trans. Fam&ySoc. 1951, 47, 184. Bel@bf, C. S.; pMl%ls, D. T.; WBde, D. R. “Foundations of Optimization”, 2nd ed.; RsntlceHel: Englewood CWs, NJ, 1979. Caram, H. S.; Amwrdson, N. R. Ind. Eng. Chem. Process Des. Dev. 1979, 13, 80. Felder, R. M.; Kelly, R. M.; Ferrell, J. K.; Rowrseau, R. W. €nnvkon. Sci. Techno/. 1980, 14, 858. Ferreti, J. K.; Felder, R. M.; Rowseau, R. W.; McCue, J. C.; Keliy, R. M.; WWS, W. E. ”CoalGadtlcatknlGas Cleanup Test FadRty: V d I. Desalption and Qperation”; EPA-800/7-80-046a, 1980. Ferrdl, J. K.; Rrdy, M. J.; Fdder, R. M.; KeWy, R. M. “Coal Gasification/Gas Cleanup Test Fadlfty: Vol. 11. Environmental Assessment of Operatkn wlth DevdatRized Bihrmhous Coal and ChHled Methanol”; EPA report In preparatlon, 1981. Instltute of Gas Techmbgy ”Coal Conversion Systems Technical Data Book”; prepared for U.S. oeparhmtnt of Energy, Contract No. EX-76C01-2286, Report NO. HCP/T2288-01,W-90, 1978.
Johnson, J. L. Adv. Chem. Ser. 1974. No. 131. Johnson, J. L. ”Relatlonshlp Between the Gasfflcatlon ReactMUes of Coal Char and the Physlcal and Chemlcal Ropertles of Coal and Coal Char”; presented at American Chemical Society, Mvislon of Fuel Chemistry Coal Gaskatbn Symposium, Chlcago, IL, 1975. Lowry, H. H., Ed. “Chemistry of Coal Utllizatlon”; Wiley: New York, 1983. Orbach, 0.; Crowe, C. M. Can. J. Chem. Eng. 1971, 49, 509. Pukank, 0. W.; Cobb, J. T.;McMlchael, W. J.; Haynes, W. P.; Strekey, J. P. “Mathematlcal Modeling of the SYNTHANE Gasifier”; presented at 71st Annual AIChE Meetlng, Miami, FL, 1978. Sundaresan, S.; Amundson, N. R. Chem. €ng. Scl. 1979, 34, 345. Wekner, A. W. M. S. Thesis. University of Cobrado, Boulder, CO, 1978. Wen, C. P.; Tseng, H. P. “A Model for F l W d Bed Coal Gasification Simulation”; presented at 72nd Annual AICM Meeting, San Francisco, CA, 1979.
Receiued for reuieu December 1, 1980 Reuised manuscript received May 11, 1981 Accepted May 11, 1981
Kinetics of Coal Devolatllization and Hydropyrolysis Bharat L. Bhatt and Edward N. Zlegler’ Depgrtment of &”I
Engilneerlng. Polytechnic Insmute of New York, Brodclyn, New York 1120 1
A kinetic study of coat devolatilization and hyropyrolysis is conducted in a batch reactor at various heating rates by passing electric current through stainless steel screens containing dry pulverized (53-149 pm) North Dakota liiite particles. Y i M on a carbon converted bask of I 11.5 % CO,5 % C02, 40% CH, 4% C&, and 3% C& were obtained from hydropyrolysis experiments in the range of 538-910 OC, 3.5-10.4 MPa, 20-625 OC/s, and 0.1-25 s at reaction temperature. Yields of 1 4 % C02, 4% CH, and 0.5% C& were obtained from devdapiliratkn experiments in the range of 720-845 “C,0.7-10.4 Ma, wtth 20 “CIS heating rate and 25 s at reaction temperature. For order of reaction n = 6, a reasonably good correlation was obtained. The activation energy was estlmated to be about 205 kJ/mol, which is consistent with the proposed reaction controlled model.
Background and Objectives Interest in producing hydrocarbons from coal has undergone a resurgence in recent years. Although much practical experience was garnered over 50 years ago (see Lowry, 1963), the primary commercialization has been in gasification. Liquefaction is a difficult process in that subtle rearrangements of coal’s chemical structure are required to incorporate additional hydrogen. Liquefaction of coal in the true sense of that term is in a developmental stage. Existing commercial processes gasify the coal first and then synthesize a liquid product. The current work deals with the kinetics of direct hydrogenation of coal. The procedure described works satisfactorily up to temperatures of 900 “Cand pressures of 10.4 MPa. The procedures described could logically be extended to bituminous, subbituminous coals, and other solid-gas reaction systems. A general model for the coal hydropyrolysis, based on a is developed. single coal particle surrounded by H2, a. Experimental Methods. Anthony and Howard (1976) and Belt and Bissett (1978) discussed some of the techniques used in the past for coal devolatilization and hydrogasification studies. A stationary method was chosen in the current study to avoid the difficult estimation of residence time for a particle undergoing significant density changes due to swelling and gasification. Most of the work done previously, for kinetics study, has been at lower pressures, apropos pipeline gas production. Higher pressures are required for coal liquefaction. Some of the experimenters relied on weight loss as a basis for their models
because of lack of facilities to analyze products. Some workers had to rely just on the product gas analysis because their experimental setup did not allow weighing of the coal. The present investigation attempts to ameliorate these difficulties. b. Kinetics and Mechanisms. The exact description of the complex decomposition and transport phenomena involved in coal devolatilization is not yet available. A simple model proposed by many authors is a first-oder decompositiiton occurring uniformly throughout the particle. The rate constant is typically correlated with temperature by an Arrhenius expression. Anthony and Howard (1976) in their assessment of previous work, find little agreement on the observed rate constant, with values of apparent activation energies varying from 8.4 to over 209 kJ/mol. Differences in equipment and experimental procedures seem to play an important role. Steinberg et al. (1978) developed a multistep chemical reaction model. The current study intends to generalize that model. Experimental Apparatus and Procedure The experimental technique used by Anthony et al. (1976) was employed, with various modifications, for coal devolatilization and hydrogasification experiments. A small amount of pulverized coal was contained in an electrically heated metal screen envelope. Figure 1 is a schematic diagram of the experimental equipment. About 15 to 20 mg of dry North Dakota lignite (cut size, 53 to 149 pm) was loaded in the screen envelope. The coal was stored under vacuum in presence of silica gel 0 1981 American Chemical Society
