Ind. Eng. Chem. Res. 1990,29, 2079-2088
2079
Modeling of Moving-Bed Coal Gasifiers Juan Adinez* and F. Garcia Labiano Znstituto de Carboqulmica,t P.O.Box 589,Zaragoza, S p a i n
A mathematical model of an atmospheric countercurrent moving-bed coal gasifier was developed and used for air + steam gasification. The model considers an adiabatic reactor in a stationary state
with different temperatures for the solid and the gas and uses a shrinking unreacted core model to describe the solid-gas reactions. T h e system of differential equations for the mass and energy balances is solved by using a fourth order Runge-Kutta-Merson method with variable step size. The effect of the reaction scheme in the combustion zone on the longitudinal temperature profiles predicted by the model is studied. A parametric study is performed on the effect of reactivity with steam, C02, and 02,of particle size, and of the emissivity of the ash layer on the longitudinal temperature profiles, the maximum temperature, and the gas composition. The effect of operating conditions on the gas compositions, low heating value, and thermal efficiency of the process is also studied.
Introduction There are numerous gasification processes in which coal reacts with an oxygen-steam or air-steam mixture. Gasification is performed in different types of reactors in which, depending on the type of solid-gas contact, the bed can be moving, fluidized, entrained, or melting salts. Of these, the moving bed is the most widely used due to its high thermal efficiency and high coal conversion (Hebden and Stroud, 1981; Denn and Shinnar, 1987). This type of reactor operates either at medium (Lurgi) or atmospheric pressure (Wellman-Galusha, Woodhall-Duckham, etc.). In these, the coal is fed through the top of the reactor and flows downward under gravity. The gasifying agents, steam and air or 02,are fed in from the bottom and come into contact with the solid as they move upward. Ashes can be removed either fused or, more normally, dry by means of a rotating grate that also acts as a gas distributor. The main disadvantages of this type of gasifier are that they can handle caking coals only by use of stirrers, diluents, or preoxidation and that they cannot be used with coals whose chars have low mechanical resistance. Furthermore, in the most widely used reactors in which ashes are removed dry, the maximum temperature inside the reactor must be lower than that at which the ashes begin to soften. Different approaches have been used to model this type of reactor. There are equilibrium or thermodynamic models such as those by Gumz (1950),Woodmansee (1975), and Kosky and Floess (1980). Yoon et al. (1978) developed a stationary-state model for the Lurgi gasifier, taking the reactor to be an adiabatic core with a boundary layer with heat transfer near the water jacket. Furthermore, they assumed the solid and gas temperatures to be equal. Amundson and Arri (1978) developed a model that took into account the temperature differences between the phases, as well as the structural changes that particles undergo in the combustion zone. This model considers the use of the kinetics proposed by Johnson (1974,1981). Biba et al. (1978) also considered a temperature difference between the solid and the gas, without radial profiles, and used simplified reaction kinetics. Yu (1981) incorporated radial profiles into Yoon et al.’s model, together with the nonstationary state. Data available for gasifiers of this type normally refer to product gas compositions, feeds, and yields. Little differences on off-gas compositons predicted from kinetic and chemical equilibrium models have been found. +
CSIC.
0888-5885/90/2629-2079$02.50/0
Thermodynamic models are easier to make, but the information they generate is restrictive, because they only can predict off-gas compositions. However, both for design and simulation, it is important to know the behavior inside the reactor, and this can only be determined with kinetic models. Longitudinal profiles of solid and gas temperatures, and particularly the peak temperature and its location in the reactor, are basic parameters for operating a moving-bed reactor, since the solid peak temperature in the reactor must not at any time go as high as the ash softening point. But the gasifiers have to operate at the highest possible temperatures, because higher temperatures may result in higher conversion rates, reduced tar formation, lower steam consumption, and lower fines carryover. The predictions of the models for the temperature profiles of the solid and gas and the maximum temperature are affected by the reaction scheme used and by heat transfer, both solid-solid and solid-gas. Table I shows the characteristics of the different models in the literature and the assumptions they make. The different models deal with different combinations of the following reactions:
-
c + 0 2 cop c + y202 co xc + 0 2 2(h - 1)CO + (2 - X)CO2 H2 + 7 2 0 2 H2O co + y*02 c02 C + H2O CO + H2 c + c02 2co C + 2Hp CHI CO + H2O + C02 + H2 +
+
+
(a) (b) (C)
(d) (e) (f)
(8) (h) (i)
This paper develops a model of an atmospheric countercurrent moving-bed coal gasifier and studies the effect of operating conditions on the gas yield and composition, the process efficiency, and the longitudinal temperature profiles. It also analyzes the different factors affecting the temperature profile and their effect.
Developing the Model According to Krieb (1973), a moving-bed coal gasifier can conceptually be divided into different zones, with different physical and chemical processes, as shown in Figure 1. In the direction of coal flow, first comes the 1990 American Chemical Society
2080 Ind. Eng. Chem. Res., Vol. 29, No. 10, 1990 Table I. Assumptions Used by Different Authors Yoon et al. shrinking core
particle reaction model kinetics
P, constant to calculate -r, 0.6(-rc) equilibrium at each height c, f, g, h
Biba et al. volumetric model controling -pi as a function of T P, = /(height)
Amundson and Caram and Arri Fuentes shrinking core thermal model IGT
-
P, = f(t,height) P,= /(height)
-rd =
shift reaction reactions on combustion zone different gas and solid temp radial temp profiles particle temp profiles flow pattern for solid and gas radiation heat transfer char reactivity effect on reactions product distribution in devolatilization COAL
reaction rate a, f , g
no no no Plug-Plug
Battacharya et al. this work shrinking core shrinking core fixed bedU -r, = f(Py,T)
equilibrium a t each height b, d , e, i
equilibrium a t each height b, d, e, i
Yes
yes
no
yes Plug-Plug
no no PlUg-PlUg
-rd = 0.6(-rc) reaction rate C
no yes no Pl Ug-PlUg
equilibrium a t each height a, d, e, f, g yes
no no PlUg-Pk
no
no f, g
Loison
proposed
OFF G A S
I
4
Coal Preheating
H Pyrolysis
U H Gasification
Combustion
Gas Preheating
GASIFICATION AGENTS Figure 1. Schematic diagram of the countercurrent gasifier zones. ASH
drying and solid preheating zone, where the coal is heated by the rising gases. Then there is the pyrolysis (devolatilization) zone, where the coal is thermally decomposed and loses its volatile components. The char thus formed goes on to the gasification zone, where it reacts with H20 and COz and where there is no 02. Then there is the residual char combustion zone and finally the ash layer, which acts as a preheater of the reacting gas. In fact, these zones can overlap, and there is no clear distinction between them, particularly in the case of some coals, and the reactions considered by different authors in each zone vary considerably. The model has been developed for adiabatic reactors, although it can be used in those that have a water jacket to generate the steam necessary for gasification. According to Yoon et al. (1978), the thickness of the nonisotermal layer near the wall is small (=lo cm) and can be used without significant error for the adiabatic model in these kinds of reactors. Different solid and gas temperatures are used, because this is a system in which the gas carries part of the heat generated in the combustion zone to the gasification zone in order for the endothermic reactions to take place. Heat transfer from the solid to the gas is by convection, while heat transfer by radiation is only considered between layers
proposed
of the solid and not of the gas. According to Amundson and Arri (1978), in these reactors there are intraparticle temperature profiles, but their consideration excessively complicates the model. Also, there are no important differences between the core and the bulk temperature of the particles working at 24 atm. Consequently, in an atmospheric reactor these temperature differences will be smaller and the coal particles have been considered to be isothermal. For a Reynolds number on the order of 150, which is a value characteristic of atmospheric moving-bed coal gasifiers, it has been found that this reactor can be approximated by 120 stirred-tank reactors in series (Levenspiel, 1972), and then can be considered to be a plug-flow reactor. In numerous industrial gasifiers, on processing coal particles or pellets with considerable ash content, it has been found that they have a sharp interface of reaction and the size variation is negligible. Therefore, to define the reaction rate of the coal particles, the shrinking unreacted core model is used. Although the surface reaction rates with O2 (reactions a and c), HzO (reaction f), and COz (reaction g) are not first order, they can be approximated without introducing significant errors. The shift reaction is a fast reaction and is catalyzed by coal ash according to Kosky and Floess (1980). For this, equilibrium in each differential element of the reactor can be assumed. This assumption was numerically verified by Yoon et al. (1978). Coal pyrolysis is assumed to be instantaneous without thermal effects and to occur during heating. Devolatilization times of 80 s for 1-cm-diameter particles and 170 s for 2-cm-diameter particles (Pillai, 1981) are in the range of coal heating and drying times. Kinetics
To define the reaction rate of the coal particles, the shrinking unreacted core model is used with the three resistances (external and internal diffusion and chemical reaction) in series. According to Szekely et al. (1976),the conversion of the solid is given by the following equation:
L
&A,
J
Ind. Eng. Chem. Res., Vol. 29, No. 10, 1990 2081 its effect on the temperature profile is then studied.
Table 11. Values for the Kinetic Data reaction
c + 02 C + H20 c + coz Hz + l/20* co + '/ZO*
act. energy, kcal/mol 27.0 42.0 59.0 23.9 23.9
" I n kmol/(m2 h atm).
frequency factor 1.25 X IO6" 2.20 X lo6" 1.21 X lo8" 3.09 X l P b 8.83 X loBb
author Sergent and Smith Gibson and Euker Dutta et al. Tesner Haslam
m3/(kmol 8).
So, in a stationary state, with a constant solid rate (us), eq 2 is obtained
PSFP
In (1 - X ) 2XAP2De
)
+psvp XA&G
(2)
The mass-transfer coefficient in the film around the solid is found from the correlation given by Gupta and Thodos (1963) for calculating the j factor. For Reynolds numbers of about 200 (a typical value for fixed beds), Yoon et al. (1978) reorder this correlation as follows:
Since the temperature and composition of the gas phase are constantly changing, an average value for the diffusivity of the gaseous species is used. The effective diffusion coefficient for the ash layer is calculated by using the equation proposed by Walker et al. (19591, as a function of the ash porosity De = DitC2 (4) To calculate the surface reaction rate, assumed to be first order, the values of the preexponential factors and activation energies proposed by Sergent and Smith (1973)) Gibson and Euker (1975), and Dutta et al. (1977), shown in Table 11, are used. For the kinetics of the gas-phase homogeneous reactions (d) and (e), the Tesner (1960) and Haslam (1923) proposals are used, in which the reaction rate is first order with respect to each reactant 1
a i
- kCACB Vj' = --V dt -
(5)
Heat Transfer This model considers the possibility that the solid and gas are at different temperatures. To calculate the coefficient of convection between the solid and the gas, the equation given by Gupta and Thodos (1963) is used
Furthermore, heat is assumed to transferred between solid layers by radiation, and gases did not absorb heat by this way. To calculate the heat transmission per unit area and time between two parallel differential elements, each of which behaves as a gray surface, eq 7 is used (7) where TIand T2are the temperatures of the solid in two successive solid elements. Values close to 0.9 are taken for the emissivity of coal (Amundson and Arri, 1978), and
Mass and Energy Balances For the purposes of the model, the gasifier is divided into three zones, in ascending order: ash zone or gas preheating zone; combustion-gasification zone; and pyrolysis-solid preheating-drying zone. The ash zone is modeled simply as a countercurrent heat exchanger. Combustion-Gasification Zone. In this zone, three different reaction schemes are considered: assumption 1, reactions a, d, e, f, g, and i; assumption 2, reactions c, d, e, f, g, and i; assumption 3, reactions c, f, g, and i and instantaneous reactions d and e. CO and H2 react in the gas phase with O2according to reactions d and e, with first-order kinetics in each component. Table I1 shows the preexponential factors and activation energies used for each reaction. The shift reaction (i) is considered to be in equilibrium in each differential element of the reactor. Likewise, the amount of methane generated by reaction h is assumed to be negligible, since we are dealing with a gasifier working at 1atm. The mass balances, on a per mole of carbon basis, in a differential element of the reactor of height dz are given by the following equations, for assumption 2: dF,/dz = (u, + uf + up) (8)
-M -C O -dz
UCO,
-= dz
- 1)
x
u,
+ 2ug + U f - u,'
(y).+, - ug
dFH,o/dz
-Uf
+ Ud
dFH,/dz = U f - Ud'
u,'
(10)
(11) (12) (13)
For the heterogeneous reactions, the apparent rates vi are defined as follows:
a)a)1
Vi=-=--
dz
dt us
(14)
where u, is the rate at which the solid falls, which is constant since the particles are assumed not to vary in size and not to vary with the porosity of the bed. Since the lower end of the gasifier is taken as the origin and the upward direction is positive, the velocity at which the solid falls is negative. In turn, d@/dt is defined by the kinetics of the unreacted core model for each reaction and is given by eq 2. For the homogeneous reactions (d) and (e), the moles reacted in each differential element are taken into account. To simplify the calculation, the shift equilibrium (i) is introduced at the end of the calculation of each differential element. Since different temperatures are used for the solid and gas, the heat balances for each phase must be performed. The corresponding differential equations are solid
2082 Ind. Eng. Chem. Res., Vol. 29, No. 10, 1990
where
Table 111. Devoiatilization Data
(16)
Loison fract distrib of volatiles, % tar chem water coal gas composition of coal gas CH4 HZ
co
COZ other (CzH4,CzHB,SHz)
where
AC,"' = CPpraluct. - C P m &
(18)
Drying, Pyrolysis, and Solid Preheating Zones. In a moving-bed gasifier, coal heating and drying are fast and can be considered to be instantaneous (MacIntosh, 1976). Devolatilization times depend upon the coal type, temperature, and particle size. By using parameters typical of the conditions in atmospheric moving-bed coal gasifiers, a devolatilization time of the order of 120 s was estimated (Pillai, 1981). This time is small in comparison to a solid residence time. To model this zone, the same equations as those used for the gasification-combustion zone are used, although in this case, since the gasification rates are very low, the result is similar to that of a countercurrent heat exchanger. In pyrolysis, the coal is assumed to lose the weight of its moisture and volatiles, measured by proximate analysis. The distribution of the pyrolysis products affects the composition of the gases produced by the gasifier. Table I11 shows the distribution of volatile products proposed by Loison and Chauvin (1964) and another determined experimentally using a subbituminous coal from the Teruel basin (Spain). Also, Table IV shows the proximate analysis of this coal and the distribution of the pyrolysis products obtained in similar conditions to those existing in a moving-bed coal gasifier. As can be seen, the char yield is different to that in proximate analysis. All gases generated in the pyrolysis zone are assumed to be part of the gas product at the end of the gasification zone, i.e., when X, = 0. Application to Gasification at Atmospheric Pressure The model above was used to design and simulate a moving-bed air + steam gasifier. The system of differential equations was solved for use in design, as a problem of initial conditions, and in simulation, as a boundary value problem. The Runge-Kutta-Merson fourth-order method with variable step size was used. By using the initial conditions, Le., using the design model, the equations are solved for the lower part of the reactor, given the air and steam flow rates, the coal conversion, and the output temperature of the ashes. Thus, the temperature profiles, for both solid and gas, are determined, as well as the variation in gas composition in the reactor and the relationship between coal conversion and height. To apply the model, a subbituminous coal from the Teruel basin was used (proximate analysis shown in Table IV). Design Longitudinal Temperature Profiles. A critical parameter in the operation of moving-bed gasifiers with dry
this work
20 23 57
25.6
50.3 13.1 20.6 6.1 9.9
15.34 6.67 40.43 29.80 7.76
74.4
Table IV. Parameters Used in the Design 2 molar ratio C/Oz molar ratio H,O/Oz 1.5 373 K ash off temp particle diameter 2 cm 2m reactor diameter 1000 kg/h coal mass feed flux coal emissivity 0.9 prox anal. of coal moisture 16.8% 21.8% ash 35.7% fixed carbon 25.7% volatile matter pyrolysis prod distrib 67.09% char 12.01% gas tar + HzO 20.90%
ash extraction is the longitudinal temperature profile, since the temperature inside the reactor must not exceed the ash softening point at any time. The design of this type of reactor depends heavily on the reactions assumed to occur. The different possibilities are analyzed below. An extreme case in the temperature profile is that in which the only reaction occurring in the combustion zone is that with 02, leading to COz. The gasification reactions begin when the oxygen is exhausted. However, these assumptions, mentioned above, do not explain the effect of coal reactivity on the longitudinal temperature profile, as found experimentally by Hebden (1975). In this case, since the total reaction rate in the particle is controlled mainly by the internal and external mass transfer, the reactivity would not affect practically the temperature profile nor the maximum temperature reached in the combustion zone, although it would cause an effect in the gasification zone. To explain the effect of the coal reactivity on the longitudinal temperature profile found by Hebden, it is necessary to assume that the combustion and gasification reactions occur simultaneously. Furthermore, the temperature profile shape will depend on the distribution of products produced by combustion (CO, COz) and on whether CO is burned to give COPand where this occurs. There are different expressions for calculating the distribution of the combustion products (Bhagat, Arthur, etc.). According to Bhagat (1980), for atmospheric beds this is given by CO/C02 = 218 exp(-7250/TS)
(19)
Different situations for CO oxidation can be considered, depending on whether oxidation is considered to occur inside the solid, in the gas phase with a finite reaction rate, or instantaneously. Thus, Arri and Amundson (1978), using the double-film theory, believe that the oxidation of CO occurs inside the solid. A different situation arises if gas-phase reactions are taken to be instantaneous, as assumed by Amundson and Arri (1978), or to have specific reaction rates, as assumed by Weimer and Clough (1981)
Ind. Eng. Chem. Res., Vol. 29, No. 10, 1990 2083 1500
3l
I
Hypothesis 1
2W
1300
-
- 2 E
Hypothesis 2
1
n
5c
c
3 z 2
e
=
e
I? W 1 1100
-
1000
I
1.5
I
I
I
I
1.75
2.0
2.25
2.5
MOLAR RATIO
U
800
900
TEMPERATURE
C/O2
Figure 2. Maximum temperature as a function of C/Oz molar feed ratio with HzO/Oz = 1.5.
to model gasification in a fluidized bed. These three different assumptions predict different temperature profiles in the reactor and, therefore, different peak temperatures. Figure 2 shows the solid peak temperature as a function of the C/Oz molar ratio, Le., the amount of saturated air at 337 K introduced into the reactor. An increase in the C/O2 ratio, corresponding to a drop in the gasifying agent flux with the same ratio H20/02,is accompanied by a drop in the maximum temperature. The highest temperatures correspond to case 1,in which CO is burned to C 0 2inside the solid, since the heat generated in the reaction increases the temperature of the solid. The intermediate temperature corresponds to case 2, in which the oxidation of CO to C02is assumed to take place in the gas phase at a finite reaction rate. In this case, the peak temperatures are lower since, under normal conditions using the Bhagat (1980) equation, CO predominates over C02,and the amount of heat generated in the solid drops. In this situation, it is not possible to operate with C/O2 ratios below 2.0, since, because of the distribution of final products, more CO is produced instead of COz and therefore the specific O2 consumption to gasify a given amount of coal is lower. But as long as the heat balance can be maintained, increased CO and decreased C02 per unit of O2consumed is desirable. The third case, in which CO is burned instantaneously in the gas phase, predicts lower peak solid temperatures, since the heat generated in the reaction directly increases the gas temperature. A t first, there are no clear reasons for distinguishing between the three cases. However, there are differences with regard to the distribution of CO and C 0 2 in combustion. According to Wen and Dutta (1979), the distribution depends on particle size, and for particles greater than 0.1 cm (such as are used normally in moving-bed gasifiers), only C02 is produced. Furthermore, since the objective of modeling the longitudinal temperature profile of solids is to prevent ash softening, it would be more reasonable to use a slightly more conservative criterion. Therefore, assumption 1 is adopted, which will be used from now on in this paper. Figure 3 shows the longitudinal temperature profile of the solid. The temperature is seen to reach a peak in the combustion zone. This maximum is situated in a band of the gasifier where the combustion reaction rate is highest. Figure 4 shows the temperature difference between the gas and the solid in this case, which has two peaks. The first corresponds to the ash cooling zone, in which the downward-moving solid gives up part of its heat to the
1100
1000
1200
1300
1400
(K)
Figure 3. Solid temperature profile for conditions given in Table IV.
Y m cn u)
c
E 0
-25 0,o
091
032 HEIGHT
0,3
Os4
0,s
(m)
Figure 4. Temperature difference between solid and gas for the conditions given in Table IV.
rising gasifying agents. The other peak occurs in the combustion zone, since the large amount of heat generated in the reaction is not transmitted to the gas. It should be recalled that this difference is lower than in the ash zone, since the heat transfer from the solid to the gas by convection increases with temperature, so this transfer is greater in the combustion zone. Later, in the gasification zone, due to the fact that the reactions that take place are endothermal, the gas temperature is slightly higher than that of the solid. Both the solid and gas temperatures fall further up the gasification zone, and they tend to equalize as the gasifying agents are exhausted. Effect of Emissivity. The solid and gas temperature profiles predicted by the model depend on the parameters used. Heat transfer is one of the most important factors to be considered in modeling a moving bed. Thus, due to the high temperatures achieved in a moving bed, radiation is one of the most important mechanisms of heat transfer in the solid and, therefore, cannot be neglected. Coal has a high emissivity, between 0.75 and 0.9, according to Amundson and Arri (1978). The value of this parameter is decisive in modeling the gasifier, so much so that if low emissivity values are used, or if this parameter is not taken into account at all (e = 0), then excessively high temperatures will be predicted inside the reactor. Figure 5 shows the effect of this parameter on the solid peak temperature in the model proposed here as a function of the C/O2 ratio for particles of 1-cm diameter. Thus, a drop in the emissivity gives a lower axial heat flow by radiation in the solid so that higher temperature
2084
Ind. Eng. Chem. Res., Vol. 29, No. 10, 1990 1600
I 2.5
1
1200
1.5
1.75
2.0
2.25
2.5
0,o 850
950
1050
1150
MOLAR RATIO
C/O2
Figure 5. Maximum temperature as a function of C/O, molar feed ratio for different coal emissivities. (H,O/O, = 1.5).
1250
1350
(K)
TEMPERATURE
Figure 7. Solid temperature profile as a function of coal reactivity and conditions given in Table IV.
250
g
200 150
m
a I8
100
0
t
I
..
0,oo
1100 0,05
0,lO HEIGHT
0,15
0,20
(m)
Figure 6. Temperature difference between solid and gas for different coal emissivities and conditions given in Table IV.
differences between the solid and gas are found, as shown in Figure 6. Furthermore, the height at which the peak temperature is achieved is affected by this parameter, since an increase in the solid temperature produces higher combustion reaction rates. Effect of Coal Reactivity. The reactivity of coals is one of its most important characteristics, since it influences the rate a t which it will react under different conditions and atmospheres and therefore determines the reactor size. Highly reactive coals, such as lignites and subbituminous coals, will therefore need lower residence times than low reactive coals under the same operating conditions. The gasification reactions, particularly the one between carbon and steam, are affected more by a particular coal reactivity than the combustion reaction, which is controlled basically by film diffusion. To introduce reactivity into the model, the coal reactivity is defined in terms of the reaction rate with H20, 02,or C02, as follows: reactivity = reaction rate of coal used/ reaction rate of a standard coal* (20) (* given in Table 11) where in both cases the reaction rate is controlled by the chemical reaction. Figure 7 shows the longitudinal temperature profiles of the solid, found when gasifying coals with reactivities ranging from 1 to 10. The figure shows that the higher reactivity produces a faster drop in the temperature and then the reactor may need a shorter length to achieve complete conversion. This is because the gasifying agents
1150
1200
1250
TEMPERATURE
1300
1350
(K)
Figure 8. Solid temperature profile as a function of particle diameter and conditions given in Table IV.
(C02,H20J are consumed faster, since the reaction rates are higher. In highly reactive coals, the endothermal gasification reactions are responsible for the lower peak temperature, which agrees with the data given by Hebden (1975). Effect of Particle Size. The particle size used affects the total reaction rate of the solid. Figure 8 shows the effect of the particle size on the longitudinal temperature profile of the solid. As the particle diameter increases, the width of the combustion zone is greater, and a larger combustion zone and gasifier are required to achieve complete coal conversion. This is because of a drop in the overall reaction rate due to slower mass transfer in the ash layer. Furthermore, the size hardly affects the temperature difference between the solid and gas in the ash cooling zone. In contrast, Figure 9 shows that it does influence both the solid-gas temperature difference in the combustion zone and the height of the combustion zone. An increase in particle size implies an increase in the height of the combustion zone and a broadening of the temperature difference peak. A large amount of heat is generated in this zone so that the bigger the particle, the greater the difference in T, and Tgdue to the higher resistance to heat transfer in the ashes. Effect of the H 2 0 / 0 2Ratio. One fundamental variable in operating the gasifier is the H 2 0 / 0 2molar ratio fed into the reactor. This ratio is directly related to the saturation temperature of the blast air entering the gasifier. For a given flux of coal, the peak temperature in the re-
Ind. Eng. Chem. Res., Vol. 29, No. 10, 1990 2085 , -"
100 I
u 80 c VI
60
-20
o,o
0,2
0,4
1,O
0,s
0,6 HEIGHT
0,o 0
1,2
2
1
(m)
HEIGHT
Figure 9. Temperature difference between solid and gas in combustion zone as a function of particle diameter and conditions given in Table IV.
3
(m)
Figure 12. Coal char and HzO conversion profiles for conditions given in Table IV. 35
1400
1
30
25 20 15
cn
10
a
0
1275 1 ,oo
1,25
1,50
1,75
Figure 10. Maximum solid temperature as a function of HzO/02 molar feed ratio. (C/Oz = 2.25).
9
1
1
25
co
v)
a O
5 0
0
0.5
1
2
1.5 HEIGHT
0 0.5
1 .o MOLAR RATIO
MOLAR RATIO H2 0 I 02
30
5
2.5
(m)
Figure 11. Gas composition profiles for conditions given in Table IV.
actor can be controlled by varying the amount of steam introduced into the reactor. Figure 10 shows the effect of the H 2 0 / 0 2 ratio on the peak temperature, for a C/O2 molar ratio of 2.25. Composition of the Gas Product. Figure 11shows the variation in gas composition with the height of the gasifier in gas preheating and combustion-gasification zones. A zone of 15-30 cm, depending on the coal size and operating conditions, at the bottom of the gasifier, which corresponds to the ash cooling zone, can be seen, where no reactions
1.5
2.0
H 2 0 / 02
Figure 13. Gas composition (vol %) as a function of HzO/O, molar feed ratio (C/Oz = 2).
occur, and the composition does not change. Then there is a zone with a maximum slope in the Oz and COz concentration curves, which corresponds to the highest rate for the combustion reaction. The steam concentration is seen to vary very little in the combustion zone, since coal reactivity has been taken to be 1. When the oxygen is almost totally exhausted, the gasification zone begins. There is a progressive drop in the concentration of gasifying agents H20 and CO, (faster in the case of HzO) and a progressive increase in the amounts of products CO and Hz. Likewise, Figure 1 2 shows the variation of the solid and the HzO conversion with the position in the reactor. The variation in coal conversion is greater in a small band in the combustion zone, where the reaction rate is higher as a result of the high temperatures. This leads to the consumption of all the oxygen entering this band. Figure 13 shows the effect of the H20/Oz molar ratio on the composition of the gas product. There is an increase in the percentage of CO, and H, as the ratio increases and a drop in the percentage of CO; this is due mainly to the effect of the increase in the water partial pressure on the shift equilibrium. Figure 14 shows the variation, with the C/Oz molar ratio, of the composition of the gas product. To keep the ratio constant, an increase in the C/O2 ratio is accompanied by a drop in the total amount of gasifying agents, both air and steam, entering the reactor, since the coal feed is constant. An increase in the ratio produces an increase in the percentage of CO and H,,thus giving a gas with higher heating
2086 Ind. Eng. Chem. Res., Vol. 29, No. 10, 1990
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.- za -
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PARTICLE DIAMETER
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2.75
C / O2
Figure 15. LHV and cold thermal efficiency of gas produced, as a function of fixed carbon to oxygen molar feed ratio (HzO/Oz= 1.5, d, = 2 cm, reactivity = 1).
value and lower temperatures in the reactor. Simulation The model proposed here can be used to simulate an atmospheric moving-bed gasifier using air + steam as gasifying agents. A 2-m-diameter, 3-m-high reactor is studied. In this case, the solution of the model becomes a boundary value problem. At the bottom of the reactor, the solid conversion is known, its temperature is the ash output temperature, and the gas temperature is that of the gasifying agent feed. In the upper part of the reactor, the condition is that the solid conversion is zero. To fit the limiting conditions, the Fibonacci search method is used. Effect of Operating Conditions on Gas Characteristics. The thermal efficiency and the gas heating value depend on the operating conditions, i.e., C / O 2and H 2 0 / 0 2 molar ratios, coal reactivity, and particle size. By defining cold thermal efficiency without tar (F,) as F1 =
LHVdry
gas without WQG dry gas without tar
ii
i 1300 1275
z 55
(21)
LHVcoaI~cod
for a fixed coal flux of 0.057 kg/(m2 s), a typical value for atmospheric gasifiers using saturated air a t 337 K with spherical particles of 2-cm diameter and a reactivity of 1, both the thermal efficiency (Fl)and the low heating value go through a peak, as a function of the amount of gasifying agent feed (C/O, molar ratio), as shown in Figure 15. This is because an increase in the C / O , ratio means less carbon
?i
1300
o
1 1.25
-
1350
l i "
f
5 J
4
0,046
"E
z
(cm)
1500
I-
E
P
Figure 16. LHV (- -) and cold thermal efficiency of gas produced (- - -1 and mass feed flux of fixed carbon (-), as a function of particle diameter (H,O/O, = 1.5).
2.50
Figure 14. Gas composition (vol % ) as a function of C / O z molar feed ratio ( H 2 0 / 0 2 = 1.5).
1200 1100
5r s
1100
I-
0 1.50
' LII
-
LL
5
1250
1200 1
2
3
4
5
6
7
8
9
10
REACTIVITY
Figure 17. LHV (-), mass feed flux of fixed carbon (---), and maximum temperature (-.-), as a function of coal reactivity (H,O/O, = 1.5).
is used in combustion and more in gasification, with the same above-optimum amount of oxygen. However, there comes a point in which the oxygen feed does not generate enough heat for the endothermal gasification reactions, so the coal conversion is incomplete. This point coincides with the peak in thermal efficiency and in the low heating value. The size of the coal particles is another parameter that affects the reactor yield. Obviously, for a fixed flow of gasifying agent, the smaller the particle is, the greater the amount of coal is that can be gasified. Figure 16 shows its effect on the cold thermal efficiency (FJ, the low heating values, and the fixed coal mass flow that can be completely gasified using 509.3 (N m3)/(m2h)saturated air at 337 K as the gasifying agent. We can see that as the particle size increases, the amount of coal that can be gasified drops. Furthermore, the amount of feed oxygen is moving away from the optimum, using less and less carbon in the gasification reaction. This leads to a drop in the cold thermal efficiency and the low heating value of the gas product of the reactor. The type of coal used in gasification is another of the parameters that most influences the solid behavior of the gasifier. Figure 17 shows the low heating value and peak temperature obtained for a gasifying agent flux of 509.3 (N m3)/(m2h) saturated air at 337 K, as a function of coal reactivity. Under certain operating conditions, highly reactive coals, such as lignites and subbituminous coals, need shorter residence times than do coals with lower reactivities. For a fixed gasifying agent flux, the more reactive the coal, the more coal that can be processed. Therefore, the C / O z molar ratio increases, and more fixed carbon is used in gasification. This has the double advantage of improving the reactor efficiency while reducing the peak temperature.
Ind. Eng. Chem. Res., Vol. 29, No. 10, 1990 2087 Table V. Comparison between Predicted and Experimental Data Rosebud River King
(subbitu-
(bitumi-
minous)
nous)
Benton (lignite)
Uv
4200
72
4000
70
3800
>
23.0 29.1 12.6 35.3
10.3 35.3 9.3 45.1
32.8 34.9 6.4 25.9
E "
3600
Y
k
Y
66
45.4 27.4 18.3 7.1 1.4 0.4
49.5 23.7 16.5 7.2 1.6 0.4
49.2 25.9 15.7 7.5 1.3 0.3
42.5 23.0 21.4 10.6 1.7 0.5
40.3 24.6 23.0 9.3 2.1 0.5
5.1
13.7
11.7
11.2
14.6
kg/ 100 kg
coal
Comparison with Operating Gasifiers. The data available for moving-bed atmospheric gasifiers generally refer to the product gas composition, yields, and consumptions, and there is little data available on the longitudinal temperature profiles and the data also are not comparable. To partially check the validity of the model, the data so predicted were compared with real data on the product gas composition for various coals, and good agreement was found in all cases. Table V shows the comparison with plant-operating data from Thimsen et al. (1985) using coals of different rank, subbituminous and bituminous coals and lignites. The slight deviations of the composition are due to the fact that a general distribution of pyrolysis products is used with a different coal. The main differences between experimental results and those predicted by the model are in the cold thermal efficiency without tar. This can be due to the difference between the distribution of products obtained in the pyrolysis and the proximate analysis. So, for the subbituminous coal from the Teruel basin (Table IV), there is a great difference between the fixed carbon determined by proximate analysis (35.7%) and that determined by pyrolysis (45.3%), because pyrolysis conditions strongly affect the distribution of the products obtained. By this, the utilization of fixed carbon determined by proximate analysis implies a lower carbon flux to gasificate and a lower flux of gas obtained and lower cold thermal efficiencies without tar ( F J . To clarify this effect, in Figure 18 are shown the differences obtained by converting volatile matter into fixed carbon, corresponding to the proximate analysis. In this figure, a coal feed flux of 1500 kg/h has been used with a molar ratios C/O2 = 2.75 and H20/02= 1.5. It can be observed that an increase in the percentage of the volatile matter converted into fixed carbon produces an increase in the cold thermal efficiency without tar and in the dry gas flux obtained. Conclusions The proposed model predicts the yields and operation characteristics of an atmospheric countercurrent movingbed coal gasifier with realistic suppositions. The reactivity
u)
t: c r
w
E
2 CI
3400
1100 2070 2600 1306 3499 2157 steam, knlh 317 768 785 real computed real computed real computed
E
X -I 3
0
coal compn moisture volatile matter ash fixed carbon operation condition8 coal throughput, kg/h air, N m3/h
dry gas comp, % v/v N2 44.0 co 28.4 H2 18.0 COZ 6.0 CH, 1.6
74
3200
62
3000 0
10
20
30
40
50
VOLATILE MATTER CONVERTED INTO FIXED CARBON (%)
Figure 18. Effect of volatile matter converted into fixed carbon over thermal efficiency and off-gas flux.
of the coals and the emissivity of the ash layer must be known accurately, since they have a strong influence on the temperature profiies and maximum temperature in the reactor and on its capacity for processing coal, although they have little effect on the composition of the gas product. Also, to get a suitable prediction of the efficiencies and off-gas flux with the model, it is necessary to use a distribution of products of pyrolysis obtained under similar conditions for the existing moving bed, and not those only corresponding to the fixed carbon of the coal, determined by the proximate analysis.
Nomenclature AR = cross-sectional area of the reactor, m2 A, = external particle surface, m2 C, = specific heat, kcal kmol-' K-' D = diameter of the reactor, m De = effective diffusion coefficient in the ash, kmol m-' atm-' h-' d, = particle diameter, m D, = effective diffusivity of gas i, kmol m-' atm-' h-' e = coal emissivity, dimensionless F, = form factor, dimensionless F = molar feed flux, kmol h-' m-2 F , = cold thermal efficiency without tar (eq 21) H , = heat-transfer coefficient, kcal h-' m-2 K-' H,= reaction enthalpy, kcal kmol-' K = thermal gas conductivity, kcal m-' h-' K-' k = intrinsic reaction rate coefficient, kmol h-' m-2 atm-' KG = bulk film mass-transfer coefficient, kmol h-l m-2 atm-' N = moles n = orfier of the reaction, dimensionless P = pressure, atm Q = heat flux, kcal h-' m-2 QG = dry off gas flux, N m3 h-' R e = Reynolds number, dimensionless Sc = Schmidt number, dimensionless t = time, h 7' = temperature, K u, = solid down rate, m h-l u = apparent carbon reaction rate, kmol h-' m-3 u' = homogeneous reaction rate, kmol h-' m-3 V , = particle volume, m3 X = solid conversion, dimensionless 2 = height, m Greek L e t t e r s X = stoichiometric coefficient, dimensionless c
= bed porosity, dimensionless = ash porosity, dimensionless
tC
u
= Stefan-Boltzmann constant, 4.88 loW8Kcal h-' m-2 K4 = gas viscosity, Kg m-l s-'
2088 Ind. Eng. Chem. Res., Vol. 29, No. 10,1990 ps
= molar density of carbon, Kmol m-3
Superscripts
o = initial conditions (1) = initial conditions in the differential element Subscripts
C = carbon s = solid G = gas
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1981;Second Supplementary Volume, Chapter 24. Johnson, J. L. Kinetics of Bituminous Coal Char Gasification with gases containing Steam and Hydrogen. Adu. Chem. Ser. 1974, 131.
Johnson, J. L. Fundamentals of Coal Gasification. In Chemistry of Coal Utilization; Elliot, M. A., Ed.; Wiley: New York, 1981; Second Supplementary Volume, Chapter 23. Kosky, P. G.; Floess, J. K. Global Model of Countercurrent Coal Gasifiers. Ind. Eng. Chem. Process Des. Deu. 1980,19,586-592. Krieb, K. H. Combined Gas and Steam-turbine Process with LURGI Coal Pressure Gasification. I.G.T. Symp. Pap. 1973, 127. Levenspiel, 0.Not Ideal Flow. In Chemical Reaction Engineering, 2nd ed.; Wiley: New York, 1972;Chapter 9. Loison, R.; Chauvin, F. Pyrolyse Rapide du Charbon. Chem. Ind. (Paris) 1964,91, 269. McIntosh, M. J. Mathematical Model of Drying in a Brown Coal Mill System. 1. Formulation of Model. Fuel 1976,55,47. Pillai, K. K. The Influence of Coal Type on Devolatilization and Combustion in Fluidized Beds. J. Inst. Energy 1981,54,142-150. Sergent, G.; Smith, J. Combustion Rate of Bituminous Coal Char in the Temperature Range 800 to 1700 K. Fuel 1973,52,52. Szekely,J.; Evans, J. W.; John, M. Y. Reactions of Nonporous Solids. Gas-solid Reactions; Academic Press: London, 1976;Chapter 3. Tesner, P. A. 8th Symp. Combust. 1960,807. Thimsen, D.; Mouser, R. E.; Lin, B. Y. H.; Pui, D.; Kittelson, D. Fixed-Bed Gasification Research using U.S. Coals; DOE: Washington, DC, 1985;Vol 19,Executive Summary, DOE/ET/ 10205-T1. Walker, P. L.; Rusinko, F.; Austin, L. G. Gas Reactions of Carbon. Adu. Catal. 1959,2, 134. Weimer, A. W.; Clough, D. E. Modelling a Low Pressure SteamOxygen Fluidized Bed Coal Gasifying Reactor. Chem. Eng. Sci. 1981,36, 549. Wen, C. Y.;Dutta, S. Rate of Coal Pyrolysis and Gasification Reaction. In Coal conversion Technology; Wen, C. Y., Lee, T., Eds.; Addison- Wesley: Reading, MA, 1979. Woodmansee, D. E. Modeling of Fixed Bed Gas Producer Performance. Abstracts of Papers, 5th Conference on Synthetic Fuels from Coal, University of Oklahoma, Stillwater, May 1975. Yoon, H.; Wei, J.; Denn, M. M. A Model for Moving-Bed Coal Gasification Reactors. AIChE J. 1978,24,885-903. Yu, W. C. Ph.D. Theeis, University of Delaware, Newark, 1981. Received for review July 10,1989 Revised manuscript received May 17, 1990 Accepted May 21, 1990
