Modeling a pilot-scale fluidized bed coal gasification reactor

Fuel Processing Technology, 19 (1988) 265-290

265

Elsevier Science Publishers B.V., Amsterdam - - Printed in The Netherlands

M o d e l i n g a P i l o t - S c a l e F l u i d i z e d B e d Coal Gasification Reactor ROBERT P. MA, RICHARD M. FELDER* and JAMES K. FERRELL

Department of Chemical Engineering, North Carolina State University, Raleigh, NC 27695 (U.S.A) (Received February 29th, 1988, accepted in revised form June 23rd, 1988)

ABSTRACT A steady-state model has been developed to simulate the North Carolina State University pilotscale fluidized bed coal gasification reactor. The model involves instantaneous devolatilization of coal at the top of the gasifier {freeboard region) and char combustion and gasification in the fluidized bed. A two-phase (emulsion-dilute gas ) representation of the fluidized bed incorporates the phenomena of jetting, bubbling, slugging, and mass and heat transfer between phases, and enables the prediction of individual species flow rates and temperature profiles within the bed. The model has been successfully used to simulate the gasification of a devolatilized Western Kentucky bituminous coal and a New Mexico subbituminous coal and to predict effects of various operating parameters on key gasifier performance variables.

1. INTRODUCTION

Modeling a fluidized bed coal gasification reactor involves simulating devolatilization, combustion, and gasification kinetics and fluidized bed hydrodynamics. Each of these subjects has been extensively studied and reviewed [17] and several models with widely differing levels of complexity have been proposed to describe fluidized bed coal/char gasification reactors [8-21]. However, most of these models either have been tested only against limited coal/char gasification data or have not been tested at all. For the past two decades, the U.S. Environmental Protection Agency (EPA) has carried out a study of the environmental impact of coal conversion processes. As a part of the study, the EPA contracted for the construction of a coal gasification and gas cleaning pilot plant at North Carolina State University. A description of the plant is given by Ferrell et al. [22] and in abbreviated form by Felder et al. [23]. The overall objective of the project is to characterize the effluents from the gasification and gas cleaning processes and to determine *To whom correspondence should be addressed.

0378-3820/88/$03.50

© 1988 Elsevier Science Publishers B.V.

266

how the rates of formation of pollutants depend on coal feedstocks and process operating conditions. One facet of the project has been the development of a mathematical model for the pilot plant fluidized bed gasifier. This paper reports on the final stage of this effort. 2. E X P E R I M E N T A L

The coal gasifier is a 15.2 cm I.D. Schedule 40 316 stainless steel pipe enclosed in several layers of Fiberfrax bulk ceramic insulation and housed in a 61 cm I.D. Schedule 80 carbon steel pipe. Gases are fed to the reactor through three feed nozzles spaced triangularly near the bottom of the reaction chamber. Each nozzle consists of a 0.48 cm orifice inside a diverging cone. Coal is fed at the top of the gasifier and char is removed at the bottom by nitrogen-purged screw conveyors. The temperature profile in the bed is monitored by six thermocouples housed in a central thermowell in the reactor. The thermocouple located 25 cm above the feed nozzles is used for reactor temperature control, with the oxygen feed rate being adjusted to maintain the selected temperature. A reactor pressure tap is located at the top of the reactor. Differential pressure taps on the gas feed line below the feed nozzles, at 38 cm and 89 cm above the feed nozzles, and at the top of the reactor, are used to monitor the pressure drop across the feed nozzles and over two intervals in the bed. The bed level is monitored with a nuclear level gauge and controlled by adjusting the char removal screw rotation rate. Gases are drawn for analysis during the steady-state period of a run. H o t synthesis gas is continuously sampled from a centerline point in the effluent pipeline downstream of the cyclone. The tars and solids in the sampled gas are trapped in a steel wool filter, condensable and water-soluble species are removed in a cold water quench, and the gas is reduced in pressure and either vented or periodically drawn into one-liter coated stainless steel or glass bombs. The solids and tars are separated by a methylene chloride extraction. The synthesis gas composition can be reconstructed from the flow rates and compositions of the samples. Gas chromatography is used to determine concentrations of fixed gases (H2, CH4, CO, C02, and N2), sulfur gases (H2S, COS, CS2, thiophenes, and mercaptans), and light aliphatic and aromatic hydrocarbons. Proximate, ultimate, and sieve analyses are routinely performed on the solid feed and effluent streams. Purdy [13] gives details about process control and gasifier operation and describes gasifier effluent sampling and analysis methodology, and additional procedural details are provided by Rudisill [24], Zand [25], Lee [26], and Rhinehart [27 ]. A detailed description of the pilot plant data aquisition system is given by Willis [28]. Since 1979 a devolatilized Western Kentucky bituminous coal, a New Mex-

267 TABLE 1 Analyses of coal feedstocks (wt.%) West Kentucky bitum, char

New Mexico subbit, coal

Texas lignite

86.0 2.4 0.9 10.7

35.2 31.7 10.5 22.6

23.1 29.1 24.3 23.5

83.8 0.6 2.2 0.1 2.6 10.7

52.5 4.8 18.1 1.2 0.8 22.6

39.2 4.2 32.1 0.5 0.5 23.5

proximate fixed carbon vol. m a t t e r moisture ash

ultimate carbon hydrogen oxygen nitrogen sulfur ash

ico subbituminous coal, and a Texas lignite have been gasified over a wide range of operating conditions. The average bed temperature varied from 1017 K for the Texas lignite to 1281 K for the Western Kentucky char; the span for steam-to-carbon feed ratio was 0.68 for the Kentucky char to 2.27 for the Texas lignite; and the percentage carbon conversion varied from 16.2 for the Kentucky char to 94.0 for the New Mexico subbituminous coal. In most of the runs the system pressure was maintained at 765 kPa (100 psig), and in a few runs the pressure was as low as 550 kPa. The average feed particle diameter and density were 0.2-0.6 m m and 1.4-1.8 g/cm 3, respectively, and the comparable figures for particles in the fluidized bed were 0.05-0.4 m m and 1.8-2.6 g/cm 3. Table 1 lists typical proximate and ultimate analyses of the feedstocks. Complete experimental results are given by Purdy [13] and Rhinehart [27]. 3. S I M U L A T I O N M O D E L

The gasifier model consists of two stages: instantaneous devolatilization of coal at the top of the reactor (freeboard region) and char combustion and gasification in the fluidized bed. A two-phase representation of the fluidized bed incorporates the p h e n o m e n a of jetting, bubbling, slugging, and mass and heat transfer between phases. The following sections outline the principal features of the model; complete details are given by Ma [29]. 3.1 Freeboard zone

Coal drying and devolatilization occur in the freeboard region above the fluidized bed. Devolatilization yields of various species were determined from the

268 TABLE 2 Devolatilization yields f r o m N e w Mexico s u b b i t u m i n o u s coal

species

Mi/Mdaf~Ai-'b B i T ai

CO H2 CH4 C02 H2S COS Tar

-0.625 - 4.93 X 1 0 - 2 - 0.325 - 9 . 9 5 x 10 -2 - 2 . 2 8 X 10 -2 -2.79X10 3 0.1175

Bi 6.52 X 10 -4 5.02 X 10- s 3.43 × 1 0 - 4 1.43X 10 -4 2.29X 10 -~ 3.16X10 6 0.0

gasifier data by subtracting species flow rates at the top of the bed (determined using an in-bed sampler with water as a tracer) from flow rates in the reactor synthesis gas. The devolatilization yields were then correlated with average bed temperature by linear regression [8,27]: Mi

Maaf

-Ai+B~T

(1)

where M~( kg/h ) is the devolatilization yield of species i and Mdaf (kg/h ) is the dry ash-free coal feed rate. Table 2 lists values of Ai and Bi for the New Mexico subbituminous coal feedstock.

3.2 Fluidized bed Assumptions regarding the bed hydrodynamics are listed below: 1. The fluidizing gases enter the bed in jets, which degenerate into bubbles that rise through the bed and coalesce with other bubbles. Slugging occurs when the average bubble diameter becomes larger than half of the reactor diameter. 2. The fluidized bed consists of a dilute phase (jets, bubbles, and slugs) and an emulsion phase that can be further divided into an interstitial gas phase and a solid phase. Mass and heat exchange take place between the dilute phase and the interstitial gas and between the interstitial gas and the solids. 3. The dilute and interstitial gases are ideal and in plug flow, and the solids are perfectly mixed. The fluidized bed is usually operated at a bed height four to six times the bed diameter. Using a simple two-phase model with plug flow in the dilute phase and dispersed flow in the emulsion phase, Mireur and Bischoff [30] derived a correlation for the emulsion phase dispersion coefficient in terms of the bed aspect ratio (bed height divided by bed diameter). Estimates of the emulsion phase P~clet number for the NCSU coal gasification reactor are greater than 3, which indicates that the interstitial

269

gas is neither in plug flow nor well-mixed but that plug flow should be a better assumption than perfect mixing. 4. The emulsion phase is at minimum fluidization conditions with all excess flow and bed expansion being accounted for by the dilute phase. 5. Conduction of heat through the particles is fast enough so that the temperature is uniform within the particles, and the solid temperature is uniform throughout the bed due to the large specific heat of the solids and the assumption of perfect mixing. Figure 1 is a schematic diagram of the model. Merry's correlation for jet penetration depth [31 ] and mass and heat transfer correlations of Behie et al. [32,33] were used for jetting zone calculations. A modified bubbling bed model [34] incorporating a correlation for bubble coalescence [35 ] was used to describe the bubbling bed region, and the slugging region was described with a modified two-phase model [ 36 ] with mass and heat transfer coefficients given by Hovmand and Davidson [37] and Raghuraman and Potter [36]. The reactions considered in the fluidized bed are:

Dilute phase 2C0+O2=2C02

(R1)

2H2 + 02= 2H20

(R2)

CO -4--H20 = CO2 A-H2

(R3)

COS + H20 = H2S + C02

(R4)

Interstitial gas phase 2C0 + O2 = 2C02

(R5)

2H2 + 02 = 2H20

(R6)

CO + H20 = C02 + H2

(R7)

COS + H20 = H2S + CO2

(R8)

Solid phase C + a O 2 = 2 ( 1 - a ) C O + ( 2 a - 1)CO2

(R9)

C+H20=CO+H2

(R10)

C + 2H2 = CH4

(Rll)

270

C + 0.5H2 + 0.5H20 = 0.5CO + 0.5CH4

(R12)

fuel-H

(R13)

'H2

fuel-O + 2C

,2CO

(R14)

fuel-S + H2

,HzS

(R15)

The ratio of CO to C02 in the products of char combustion is uncertain [3946]. It is therefore assumed that combustion proceeds according to reaction (R9), where a is an adjustable parameter: a=0.5 indicates formation of CO only and a = 1.0 signifies that only CO2 is formed. Rate laws for the homogeneous combustion of CO (R1, R5 ) and H2 (R2, R6 ) are given by Haslam [47]. Karim and Mohindra's rate equation [48] was used for the dilute phase water-gas shift reaction (R3) and a modified rate equation of Singh and Saraf [49] was used for the emulsion phase catalytic water-gas shift reaction (R7). Hydrolysis of carbonyl sulfide (R4,R8) is assumed to reach equilibrium based on experimental results [29 ]. A rate expression proposed by Field [50] was used for the char combustion reaction (R9). Rates of char gasification reactions (R10-R12) are described by correlations given by Johnson [4]. The fractional conversions of H, O, and S in the fuel are assumed to equal the carbon conversion. 3.3 Conservation equations The steady-state mass and energy balances for each phase but the solids are coupled first-order ordinary differential equations, with axial position z being the independent variable. The dependent variables are the species molar flow rates (f~ denotes the molar flow rate of species i in phase p) and phase temperatures. In the model notation the subscripts i = 1,2,'- .,9 denote the species 02, CO, H2, CH4, COe, N2, H2S, COS, and H20; subscriptsj = 1,2,---,15 denote reactions (R1) through (R15); and superscripts p = J , B, SL, E, and S represent jet gas, bubble gas, slugs, emulsion gas, and solids, respectively. The complete set of gas-phase equations and boundary conditions is given by Ma [29]. For illustrative purposes, the equations for the jet phase and the emulsion phase parallel to it are given below. In these equations, ~ is the volume fraction of the bed occupied by the dilute phase; ~rnfis the bed voidage at minimum fluidization conditions; A is the bed cross-sectional area; and Aij represents the stoichiometric coefficient (positive for products, negative for reactants) of component i in reaction j, which proceeds at a rate rj. The other variables and physical properties in the equations are defined in the Notation section.

271

~-z'=3A[ KJE(Cf-C~)+~AJj ]j=l dfl~=A dz

i = 1,- • -,9

3KjE ( C ~ - CE) "~- (1-- 3)~mfj

4

]

Ajar1

i=1,---,9

(2)

(3)

9

3A hjE(TJ--TE)+ ~rjAI2I~ + ~IYt~(df~/dz) dT j dz

j=l

_l i=l

(4)

m

~=lf~[c~g.~+(TJ-T~)(d%t,,~/dT) 1

dT s

TERM 1+ TERM2 + TERM3 - TERM4 - TERM5

dz

TERM6

TERM1

=3AhjE ( T J- T s )

TERM2 = ( 1 - 8 )

(1-emi)Aas2~gp(TS-T K)

TERM3 = ( 1 - 8 ) ( 1 -

~mi)Aa~a[es (T s )4--eE (T E )4]

(5) (6) (7 (8

8 T E R M 4 = (1 -g)Aem[ ~ r~3/~°i

j=5

9

(9

dfE 10

TERM5 -- i=~lt~E ~

TERM6=~lfI~[cpei+ = ,

(T E - T r ) _~_[dcve,~ 7d

11

The overall mass and energy balances for the fluidized bed are

Fcoal /eed"]'- [ 2fi ]gas /eed=Fspent char "[- [ 2fi ]gas products

[~f iI2Ii]in~t= [~,f iIYti ]out,¢t+ Q

12 (13)

where F and fi denote mass flow rates and/~i denotes specific enthalpy of species i. The adjustable model parameter Q is the net heat loss from the fluidized bed to its surroundings, and includes heat loss to the reactor environment, radiant heat loss to the freeboard, energy required to completely dry and pyrolyze the devolatilized coal, and energy needed to heat the devolatilized char

272

to the temperature of the solids in the bed. The value of Q depends on the fluidized bed and ambient temperatures, the thermal conductivity of the reactor insulation, the coefficient for heat transfer from the fluidized bed to the inside reactor wall, and various properties of the feed coal.

3.4 Implementation of the model The model contains two nested iteration loops. The convergence variables are temperature of the bed solids ( T s) in the outer loop and fractional conversion of nondevolatilized carbon (Xnc) in the loop. W h e n the model is implemented, T s is assumed and the extent of devolatilization of the feed coal in the freeboard region is determined using eqn. (1), assuming that devolatilization occurs at T s. The residual char from this step is taken to be the feed to the fluidized bed, where gasification and combustion take place. The bed calculation is carried out from the bottom (z--0), where flow rates and temperatures of the fluidizing gases are known. The emulsion phase is assumed to be at minimum fluidization conditions, so that division of the feed gas between the jet phase and the emulsion phase can be determined after the emulsion gas superficial velocity is estimated. A value of nondevolatilized carbon conversion, Xnc (a variable in the kinetic correlations included in the model) is assumed and taken to be constant throughout the bed in keeping with the assumption that the bed solids are perfectly mixed. Dilute gas phase and emulsion gas phase temperatures and gas species flow rates are determined at the boundaries of each increment until the top of the bed is reached. The species flow rates and temperatures at each bed increment are determined iteratively. The flow rates and temperatures calculated in the previous step (or entering the reactor for the first step) are taken as inlet values to the current increment. Outlet values of these variables are assumed, and conditions in the increment are taken to be uniform at a point midway between the inlet and outlet values. Quantities including the minimum fluidization velocity, bubble diameter or slug length, fraction of the bed occupied by the dilute phase, dilute gas velocity, and heat and mass interchange coefficients are calculated from the appropriate correlations (i.e. for the jetting, bubbling, or slugging regime). Mass balances are then solved to obtain new outlet flow rates for the increment, and energy balances are solved to obtain a new pair of dilutephase and emulsion gas-phase temperatures. Convergence for that increment is declared if the differences between successively determined values are each less than 1%. If convergence has not been attained, new estimates of the convergence variables are generated by modified successive substitution combined with the dominant eigenvalue method [51]. Following a complete pass through the bed, the nondevolatilized carbon conversion (Xnc) is calculated from the carbon flow rates in the gases entering and leaving the reactor and that in the coal feed, and the result is compared with

273

the previously assumed value. If the estimates fail to agree within 1% a new estimate of Xnc is generated using Wegstein's method [52] and the bed calculation is repeated starting from the bottom. W h e n this loop converges the solid temperature, T s, is determined eqn. (13) and the calculated value is compared with the previously guessed value. Wegstein's method with a 1% convergence criterion is also used for this loop. In principle the devolatilization calculation should be repeated for each new bed temperature; however, it was found expedient to recalculate the devolatilization yields only after the bed calculation converged. A combustion zone exists in the lower portion of the bed if the feed gas contains oxygen. Rate laws for homogeneous combustion (R1, R2, R5, R6) and heterogeneous combustion (R9) are employed to calculate consumption rates of 02, CO, H2, and char within an incremental volume in the combustion zone. Oxygen concentrations in both dilute and emulsion phases are checked at the end of an incremental bed calculation, and once a point is reached at which oxygen is essentially gone from both phases the combustion calculations are omitted.

3.5 Model code options and computation times Several options were included in the model to make it more flexible. T h e y are the jetting zone option, the water-gas shift equilibrium option, and the isothermal bed option. In the NCSU coal gasification reactor the fluidizing gases are introduced through nozzles, leading to a distinct jetting zone. However, bubbling hydrodynamics would be obtained at the gas inlet if a porous plate or perforated plate were used as the gas distributor. As an option of the model, either jetting or bubbling hydrodynamics can be chosen for the bed entrance region. It is general practice in gasifier modeling to assume shift equilibrium when calculating synthesis gas compositions. However, this assumption has not always been found to be warranted [ 12 ]. The model offers the option of assuming reaction equilibrium or integrating a rate equation for the shift reaction. A homogeneous rate law is used for the dilute phase and a catalytic rate law coupled with an adjustable parameter (water-gas shift reactivity) is used for the emulsion phase. In the third option, the average bed temperature is specified and assumed constant throughout the bed and the corresponding mass balance equations are solved. This mode of operation reduces the required computation time significantly but provides no information about the effects of operating conditions on the average temperature or about local temperature distributions within the bed. The model code was written in F O R T R A N and implemented on a VAX 11/ 750 computer. The CPU time for a run varied from four to nine minutes for

274 the nonisothermal bed calculations and was less than one minute for the isothermal bed option. The isothermal bed algorithm always converged; however, in the nonisothermal bed calculations if the initial guess of the solid temperature was too low, the calculated solid temperature could be higher than the valid temperature range of the kinetic equations, leading to arithmetic overflows and underflows. The problem could easily be solved by shifting the initial guess of the solid temperature to a higher value.

3.6 Parameterestimation The model has four adjustable parameters: the carbon combustion product distribution parameter (a); the char reactivity; the reactivity parameter for the emulsion phase catalytic water-gas shift reaction; and Q, the loss to the surroundings from the fluidized bed (expressed as a fraction of the total feed enthalphy). The parameters were estimated by nonlinear regression of reactor data, using a Pattern Search technique [53 ] to minimize a function of the sums of squared deviations between predicted and measured values of gasifier performance variables. The objective function was

y=l_nt=l ~ { Wl(xc~-xc"~2+ x~\ ]

(~/am) 2 G ~l

(14)

W3[~[YP, I--Y~,,I)2

t-- L k

2 +

\

y- m

+ 2(~Y~z--Y: ~ 2]

\ Y~mt / J

/

(T:vg,,--Tamvg,,~ \ Tamg,l ])

where n is the number of runs in the data base; Xc is the fractional carbon conversion; G is the synthesis gas volumetric flow rate; Yi is the fraction of species i in the synthesis gas; Tavg(K) is the volume-average bed temperature; and W,. is a weighting factor. Superscripts p and m indicate predicted and measured values, respectively. The runs with the best mass balance closures were chosen to provide the data base for the parameter estimation. The weight for the average bed temperature term, W4, was difficult to assign as this term was typically two orders of magnitude smaller than the other terms in eqn. (14). A two-step implementation of the regression analysis was accordingly carried out. In the first step values of a, the char reactivity, and the reactivity of the catalytic water-gas shift reaction were varied with WI~-W2=W3:W4~I, and in the second step Q was varied with W4 ~-~1. Pilot plant data for the gasification of devolatilized Western Kentucky bituminous coal and New Mexico subbituminous coal were correlated with the

Wl=W2=W3=Oand

275

model after small subsets of the data were used to estimate the adjustable model parameter values. The results are summarized in the next section. 4. R E S U L T S

4.1 New Mexico subbituminous coal Six gasifier runs with New Mexico subbituminous coal as the feedstock were chosen as the data base for parameter estimation. The method described in the previous section yielded the estimated values a = 1.0, char reactivity = 0.85, shift reaction at equilibrium, and Q = 0.027. The value a-- 1.0 indicates that CO2 is the only combustion product. A similar result was obtained by Ross et al. [44], LaNauze and Jung [45 ], and Smith [46 ]. J o h n s o n proposed that char reactivity can be estimated as 6 . 2 Z ( 1 - Z ) , where Z is the dry-ash-free mass fraction of carbon in the original coal. This formula predicts a value of 1.05 for the New Mexico subbituminous coal, reasonably close to the estimated value of 0.85. The New Mexico subbituminous coal char catalyzes the water-gas shift reaction to an extent that equilibrium is established in the emulsion phase. The heat loss (Q) amounts to roughly 3% of the feed enthalpy. With these parameter values incorporated, the model was used to simulate all acceptable subbituminous coal runs, where "acceptable" signifies that the overall and elemental carbon mass balance closures were between 92% and 108%, elemental hydrogen and oxygen mass balance closures were between 90% and 110%, and there were no major upsets in the reactor during steady state operation. Plots of predicted versus measured values of average bed temperature, carbon conversion, dry synthesis gas flow rate, and sweet gas (dry synthesis gas with CO2 and H2S removed) heating value are shown in Figs. 2 through 5. The close proximity of the data to the 45 degree lines indicates the model's ability to correlate the performance of the gasifier, particularly for runs with good mass balance closures. The model also did a good job of predicting evolution rates of individual species. Figs. 6 through 8 show plots of predicted versus measured evolution rates of carbon monoxide, carbon dioxide, and hydrogen. The points are close to and evenly distributed about the 45 degree line. 4.2 Devolatilized Western Kentucky bituminous coal Six gasification runs provided the parameter estimation data base for devolatilized Kentucky bituminous coal, with the results being a--0.53, char reactivity = 0.33, shift reactivity = 5.7 × 10-~, and Q = 0.055. The estimated value of a indicates that 94% of the carbon oxidized forms CO, unlike the case of subbituminous coal which yielded CO2 only. This result also has precedents in the literature [39-42], and an equation by Arthur [43]

276

MAKE-GAS

COAL

t GASEOUS PRODUCTS OF PYROLYSIS

FREEBOARD

I

f

DEVOLCHAR

BED MAKE-GAS

t

t

DILUTE GASES

EMULSION GASES • MASS

EMULSION SOLIDS

•

4

HEAT

MASS

,

= HEAT :

FLUIDIZING GASES

SPENT CHAR

Fig. 1. Schematic diagram of the fluidized bed coal gasification reactor. 1050 •

C,H,O balance closure better than 4-6%

o

C,H,O balance closure between " 6 % and "t-8% C,H,O balance closure worse

x 1000

/

J /

/ /

than - I - 8 %

E N

x

m

950

/

•

~

~

o

o

o

o

o °

900

850 850

' 900

' 950

1 000

1 050

E X P T ' L AVG. BED TEMP., ° C

Fig. 2. Predicted versus measured average bed temperature (NM subbituminous coal).

277 1.0

0.8

,,.I • C,H,O balance c l o s u r e b e t t e r than ± 6% / I o C,H,O balance c l o s u r e b e t w e e n / 4- 6% and 4- 8% x/° x C,H,O balance c l o s u r e w o r s e / ~ than4-8% ..~,~:. v

d

xo

z o o z o m r,-

0.6

0 ,4 uJ

0.4

•

I I I ]

o

x•

0.2

0.0 0.0

i

i

i

i

0.2

0.4

0.6

0.8

1.0

EXPT'L CARBON CONV., FRAC.

Fig. 3. Predictedversus measured carbon conversion (NM subbituminous coal). 0.7

O o3 u~

* C,H,O balance c l o s u r e b e t t e r than .4- 6% o C,H,O baIance c l o s u r e b e t w e e n 4-6% and -4-8% x x C,H,O balance c l o s u r e w o r s e than 4- 8% ¢ , J

/ ~ / /

0.6

x • ,,',

x

o to*

0.5

rr a 0.4

0.3 0.3

i

i

i

0.4

0.5

0.6

0.7

EXPT'L DRY MAKE-GASRATE,SCMM

Fig. 4. Predicted versus measured dry synthesis gas production rate (NM subbituminous coal).

predicts values of a between 0.52 and 0.53 for the estimated combustion zone temperatures, remarkably consistent with the estimated value of 0.528. However, in view of the complexity of the combustion process and the variations in CO/CO2 distribution from one study to the next, it is probably not advisable to attach much mechanistic significance to the results.

278 17.0

tO

•

C,H,O balance closure better than 4- 6 %

o

C,H,O balance closure between 6 % and 4- 8 %

x

C,H,O balance closure worse than -t.- 8 %

••o **

15.0 x

°e

•

•

a,.

/

16.0

x

~ z

/

o

14.0

i

13.0 13.0

i

14.0

r

15.0

16.0

17.0

EXPT'L HEATING VALUE, MJ/SCM

Fig. 5. Predicted versus measured sweet gas heating value (NM subbituminous coal). 12.0

•

C,H,O balance closure better than 4- 6 %

+

C,H,O balance crosure between 4. 6 % and -4- 8 %

x

C,H,O balance closure worse

J x

10.0 than 4- 8 %

x x

,,.¢

8.0

O

I-

e•

rr O tO 6.0 ILl ¢

•

•

4.0

2.0 2.0

,

i

4.0

,

i

6.0

,

i

,

8.0

i

10.0

,

12.0

EXPT'L CO RATE, KG/HR

Fig. 6. Predicted versus measured CO production rate (NM subbituminous coal).

The estimated char reactivity of 0.33 compares well with values reported by Johnson [54 ]--0.3 for a low volatile bituminous coal char and 0.5 for a Disco char. Its value relative to the value of 0.85 obtained for the New Mexico subbituminous coal conforms to Johnson's observation that reactivity generally

279 20.0

+ C,H,O balance closure better than 4- 6%

/

o C,H,O balance closure between 4- 6% and 4- 8% x C,H,O balance cSosureworse than 4- 8% + x

18.0

/

/ / ./

16.o!

14.0

n,-. r~

¢

~

¢

¢

12.o

10.0

8.0 8.0

,

I

,

10.0

I

,

I

12.0

.

14.0

I

,

16.0

I

,

18.0

20.0

EXPT'L CO 2 RATE, KG/HR

Fig. 7. Predicted versus measured C02 production rate (NM subbituminous coal). 1m3

a:

1.1

• C,H,O balance closure better than 4- 6% ¢ C,H,O balance closure between x 4- 6% and .4,-8% x C,H,O balance closure worse than 4- 8% f x

0.9

%

/

]

/ / /

(5

::~cc

/J° o

0.'/

0.5

,

0.5

I

0.7

.

%

x

o

•

.

I

0.9

,

I

1.1

.

1.3

EXPT'L H 2 RATE, KG/HR

Fig. 8. Predicted versus measured H2 production rate (NM subbituminous coal).

decreases with increasing coal rank. The estimated value of Q indicates that the heat lost to the surroundings and freeboard from the fluidized bed plus the heat required to completely devolatilize and dry the partially pyrolyzed coal in the bed plus the energy needed to heat the char entering the fluidized bed to the bed solid temperature amounts to 5.5% of the feed enthalpy.

280

1450 ~ .

COMBUSTION ZONE ~ ' ~

1350

1150

1050

~ -

-

-

750 ' 0.0

'

i 0.2

JETI-ING Z Q N E

~ 0.4

-

,

0.6

0.8

1.0

z/~ 1400 - o

calculated measured

1300 l , o o

TOTAL FLUIDIZED BED 1200 0.0

E 0.2

0.4

0.6

0.8

1.0

z/L t

Fig. 9. Temperature profiles in the fluidized bed (GO-38A, Char).

With these parameter values incorporated, the model was used to simulate all char runs with acceptable results. Generally speaking, predicted values of average bed temperature, carbon conversion, dry make-gas flow rate, sweet gas heating value, and evolution rates of carbon monoxide, carbon dioxide, and hydrogen are better than those for the New Mexico subbituminous coal simulation, probably due to the absence of the difficult-to-simulate devolatilization stage; however, the model slightly underpredicts the hydrogen evolution rate. Complete results are given by Ma [29].

4.3 Temperature profiles A typical plot of predicted temperature profiles during the gasification of Western Kentucky char is shown in Fig. 9. The emulsion gas temperature (broken line) rises rapidly to the solids temperatures due to homogeneous combustion of CO and He and heat transfer from the solids. The combustion raises

281

the emulsion gas temperature slightly higher than the solids temperature but the heat is rapidly dissipated to the dilute and solid phases. The temperature of the dilute gas (solid line) rises more slowly than the emulsion gas temperature. Initially the temperature rise is due primarily to heat transfer from the emulsion gas, and then as the concentrations of CO and H2 in the dilute phase increase, homogeneous combustion becomes the primary heat source. Homogeneous dilute phase combustion is also responsible for the high dilute gas temperature relative to the emulsion gas and solids temperatures in part of the combustion zone. The peak temperature appearing in the dilute phase implies that local hot spots exist in the bed that may lead to clinker formation if the temperature exceeds the fusion temperature of the coal. (This phenomenon was in fact occasionally observed in the gasifier). Oxygen is depleted well within the jetting zone, and heat transfer then leads to the attainment of a common constant temperature in the dilute gas, emulsion gas, and solid phases. The prediction of isothermality above the combustion zone is the result of the model assumption that the solids are well-mixed and that they are an efficient heat sink/source, absorbing heat released from heterogeneous char combustion and supplying heat required by the heterogeneous char-steam gasification reaction. Experimental bed temperatures shown in Fig. 9 decrease along the bed and bracket the predicted bed temperature. Each of the measured temperatures is a time average of dilute gas and emulsion gas temperatures [34]. The model is seen to determine the average bed temperature accurately, but predicts a smaller temperature gradient than was actually found above the combustion region. A variety of explanations of this result could be suggested, not least Of which is that the solids are in fact not perfectly mixed. No data were obtained in the combustion region that could provide a basis of comparison with the model prediction. In simulations of the gasification of New Mexico subbituminous coal, the emulsion gas temperature profile is similar to that shown in Fig. 9; however, the dilute gas temperature rises monotonically to the solids temperature. This behavior results from the relatively slight occurrence of CO and H2 combustion in the dilute phase.

4.4 Jetting region effects It is well known that the region near the gas distributor may play a critical role in determining reactor performance [55]. The importance of the jetting zone in the pilot plant gasifier can be seen from the model predictions. For the gasification of Kentucky char, the jetting region occupies about nine percent of the fluidized bed but accounts for thirty to forty percent of the carbon gasified (not counting the conversion due to combustion). Similarly, for the gasification of New Mexico subbituminous coal the jetting region occupies nine

282 TABLE 3 D e v i a t i o n s of model p r e d i c t i o n s f r o m data ( W e s t K e n t u c k y c h a r ) m e a n % deviation

carbon conversion dry m a k e - g a s rate CO p r o d u c t i o n rate CO2 p r o d u c t i o n rate H2 p r o d u c t i o n rate

jetting

no j e t t i n g

- 0.14 - 3.90 6.84 - 1.41 - 8.11

- 0.89 - 5.19 7.81 -3.51 - 10.5

to twelve percent of the bed and accounts for forty to fifty percent of the carbon gasification. The disproportionately high carbon conversion in the jetting zone is due to the high mass transfer rate between the jets and the emulsion phase; the dilute phase acts as a source of reactants and a sink for products (some of which have a retarding effect on the steam-char gasification reaction). On the other hand, the low mass transfer rate between the phases in the region above the jetting zone makes most of the dilute gas bypass the bed without contacting the solids, thereby decreasing the overall gasification rate. The effect of the jetting region on estimated model parameter values was studied using the model in isothermal mode. Table 3 lists the mean percentage differences between predicted and measured values of key performance variables with and without a jetting region incorporated into the model. The model with a jetting region does a better job of correlating data. In addition, model parameter estimation was carried out on both versions of the model: for the Kentucky char gasification the estimated char reactivities are 0.349 for the model with the jetting region and 1.03 for the model without the jetting region. As was noted earlier, the value of 0.349 is quite close to values estimated by other investigators for similar coal chars, indicating that ignoring the jetting region leads to inaccurate estimates of at least one key model parameter. 5. P A R A M E T R I C S T U D I E S

Parametric studies of the effects of various operating variables on reactor performance have been performed using the model. The operating variables considered were steam-to-carbon feed ratio; oxygen-to-carbon feed ratio; bed height (a measure of the solids mean residence time if the coal feed rate is constant) and reactor pressure, and the performance variables were average bed temperature, carbon conversion, dry synthesis gas production rate, equivalent methane production rate, He/CO ratio in the synthesis gas, and extent of the water-gas shift reaction. The New Mexico subbituminous coal was taken

283 1400 P = 8 . 1 5 e 5 Pa L f ~ 0.97 m H20/C 1300

1.O 1.5 2.0

1200

i,,Q 1100

1000

90C 0.0

i 0.1

i 0.2

, 0.3

0.4

FEED MOLAR O2/C

Fig. 10. Effect of molar oxygen/carbon feed ratio on average bed temperature with s t e a m / c a r b o n feed ratio as a parameter. 1.0 P ---- 8.15e5 Pa

z/~

/~/ ;,

Lf ----- 0 . 9 7 rn ..... 0.8

/ J///j'

gases + t a r

--

gases

~f/__

///

s~t/

,,=

,;',

0.6

:,,,

.2o:c 8

,o •

=~

o4

,.5 20

,',:////

,,'7// ,/// ii I

2.0 0.2

r

0.0

0.0

0.1

i

0,2

I

0.3

0.4

FEED MOLAR O2/C

Fig. 11. Effect of molar oxygen/carbon feed ratio on carbon conversion with s t e a m / c a r b o n feed ratio as a parameter.

284

to be the feedstock. In studying the effect of steam and oxygen feed rates, the nitrogen feed rate was adjusted so that the total gas feed rate remained the same, thus minimizing variations in fluidized bed hydrodynamics. Figure 10 is a plot of average bed temperature versus molar oxygen/carbon feed ratio with molar s t e a m / c a r b o n feed ratio as a parameter. The bed temperature varies almost linearly with the oxygen/carbon ratio for a given s t e a m / carbon ratio, thus enabling relatively precise control of the bed temperature through simple adjustment of the oxygen/carbon ratio (as has been done in the N C S U coal gasifier operation). The s t e a m / c a r b o n ratio has a moderate inverse effect on the average bed temperature. The effect of the oxygen/carbon feed ratio on carbon conversion with steam/ carbon feed ratio as a parameter is shown in Fig. 11. The carbon conversion increases almost linearly with an increasing oxygen/carbon ratio. At a high oxygen/carbon ratio, carbon combustion replaces the char-steam gasification reaction as the main carbon-consuming reaction and the carbon conversion approaches 100%. A 97% conversion of carbon to tar and gas products is achieved at an oxygen/carbon ratio of 0.225 and a steam/carbon ratio of 1.0. The s t e a m / c a r b o n feed ratio has a moderate negative effect on the carbon conversion. It would be ideal to have a molar H J C O ratio in the-synthesis gas of exactly three to one if methane is the intended final product or two to one if methanol is the intended product. Figure 12, a plot of molar H2/CO ratio versus oxygen/ carbon ratio with s t e a m / c a r b o n ratio as a parameter, shows that a high s t e a m / carbon ratio and low oxygen/carbon ratio favor a higher molar H2/CO ratio. The plot also shows that it is possible to obtain a ratio of either two ot three to one with an appropriate combination of oxygen/carbon ratio and steam/carbon ratio, suggesting that suitable reactor control could enable the elimination of a downstream shift reactor in a synthetic natural gas process plant. Full details of the parametric studies are given by Ma [ 29 ], and the conclusions derived from the studies are summarized in the concluding section of this paper. 6. S U M M A R Y AND C O N C L U S I O N S

A steady-state two-phase model has been developed to simulate a small pilotscale fluidized bed coal gasification reactor. The model has been used to simulate the gasification of a devolatilized Western Kentucky bituminous coal and a New Mexico subbituminous coal. The model provides good correlations with operating parameters of overall carbon conversion, total synthesis gas production, rates of formation of individual species, and average bed temperature. Parameter estimates indicate that the New Mexico subbituminous coal char is more reactive that the Kentucky bituminous char, a result consistent with Johnson's [54] observation that reactivity increases with decreasing coal rank,

285 4.0 P =

8 . 1 5 e 5 Pa

L f == 0 . 9 7 m

3.0

0 ¢D "lrr

F[20/C 2.0

2.0

C)

1.5 1.0 1.0

0.0

i

0.0

0.1

i

0.2

i

0.3

0.4

FEED M O L A R O2/C

Fig. 12. Effect of molar oxygen/carbonfeedratio on synthesisgas molar H~/COratio with steam/ carbon feed ratio as a parameter. and t h a t New Mexico subbituminous coal char is also more effective in catalyzing the water-gas shift reaction t h a n is the K e n t u c k y char. Carbon monoxide appears to be the dominant product in the K e n t u c k y char combustion while carbon dioxide predominates in combustion of the New Mexico subbituminous coal. However, these results could be artifacts of the modeling; no mechanistic inferences should be drawn from them. The jetting region in the reactor occupies nine to twelve percent of the fluidized bed but can account for thirty to fifty percent of the carbon gasification. Ignoring this region in a simulation can lead to inaccurate estimates of system parameters and a decrease in the predictive capability of the model. On the other hand, unless the intent of the model is specifically to determine the existence and extent of hot spots in the reactor, the assumptions of isothermality in the bed and instantaneous char combustion in a zone of negligible volume improve the model efficiency considerably without an a t t e n d a n t loss in predictive capability. Effects of the molar steam/carbon feed ratio, molar oxygen/carbon feed ratio, bed height, and reactor pressure on gasifier performance have been studied parametrically. The molar oxygen/carbon feed ratio has significant positive effects on the average bed temperature, carbon conversion, dry synthesis gas

286

production rate, and equivalent methane production rate. The average bed temperature varies almost linearly with the oxygen/carbon feed ratio, thus enabling relatively precise reactor temperature control. A high oxygen/carbon ratio also drives the water-gas shift reaction closer to equilibrium. The molar steam/carbon feed ratio has moderate negative effects on the average bed temperature, carbon conversion, and equivalent methane production rate. The approach to water-gas shift equilibrium is favored by lowering the steam/carbon ratio. The dry synthesis gas production rate varies inversely with the steam/carbon ratio at low temperatures and the variation reverses itself at high temperatures. It is possible to obtain a molar H2/CO ratio of either two or three to one in the synthesis gas with an approporiate combination of oxygen/carbon ratio and steam/carbon ratio, thereby eliminating the need for a downstream shift reactor in a synthetic natural gas production plant. The effects of moderate changes in bed height and reactor pressure on the steady state reactor performance are small, implying that the reactor can be operated at any bed height and pressure near the targeted operating conditions without significantly affecting its output. NOTATION

q

A Ai aS

Bi Ci

Cpg,i e

F G hjE

kg

Combustion product distribution parameter: a=0.5 and a = 1.0 respectively signify CO and CO2 as the sole product bed cross-sectional area, m 2 coefficient used in eqn. (1) to correlate coal devolatilization yield of component i particle specific surface area, m2/m 3 coefficient used in eqn. (1) to correlate coal devolatilization yield of component i concentration of component i, mol/m 3 heat capacity of gaseous component i, J / m o b K average particle diameter, m emissivity mass flow rate of solids, kg/s molar flow rate of gas species i in phase p, mol/s dry synthesis gas flow rate, m 3 ( S T P ) / s specific enthalphy, J/mol (gases) or J / k g (solids) heat transfer coefficient between jet and emulsion gases, J/m'~.s •K gas thermal conductivity, J / m . s-K volume interchange coefficient between jet and emulsion gases, m3/ m3.s

Lf Mdaf Mi

height of fluidized bed, m dry ash-free coal feed rate, kg/h devolatilization yield of species i, kg/h

287 n

P

Q r1 T T~vg

Tr W~ Xc Xnc Y Yi Z

Z 5 ~0 AHri {~mf

Aij G

number of runs used as data base for parameter estimation reactor pressure, Pa heat loss from fluidized bed to surroundings, J / s rate of reaction j, mol/m 3. s temperature, K volume-average bed temperature, K reference temperature for enthalpy calculation, 298.15 K objective function weighting factor overall fractional carbon conversion fractional conversion of nondevolatilized carbon objective function mole fraction of gas component i distance above gas distributor,m mass fraction of carbon in original coal (dry ash-free basis) volume fraction of the fiuidized bed occupied by dilute phase heat of reaction j at 298.15 K, J/mol bed voidage at minimum fluidization condition stoichiometric coefficient of component i in reaction j Stefan-Boltzman constant-- 5.669 × 10- s j / m 2. s- K 4

Superscripts B E J m p S SL

bubble phase interstitial gas phase jet phase measured value predicted value solid phase slug phase

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