Ind. Eng. Chem. Process Des. Dev. 1982, 21, 256-266
258
S = surface area per unit mass of particle = fresh feed temperature Tf = fluid temperature TL = reactor exit temperature To= reador inlet temperature or furnace outlet temperature Tb = heat exchanger outlet or furnace inlet temperature Tp = catalyst temperature t = time u = velocity
&
z = distance
Greek Symbols
6 = mass transfer coefficient y = catalyst porosity y = catalyst porosity (-AH) = heat of reaction
ATad= adiabatic temperature rise c = voidage of bed pf = fluid density pp = particle density u = variance of residence time distribution of thermal wave 7th = thermal wave residence time $F = feed rate to furnace dg = fuel flow rate to furnace $H = feed rate to heat exchanger $R = flow rate of reaction fluid w = frequency w, = resonance frequency of heat exchanger furnace loop wc = frequency at which temperature disturbances in reactor are damped to 0.35 original value Literature Cited
Aris, R., et ai. Chem. Eng. Sci. 1959, 7 7 , 199. Cohen, W. C.; Johnson, E. F. Ind. Eng. Chem. 1956, 4 8 , 1031. Crider, J. E.; Foss, A. S. AIChE J. 1968, 74, 77. Crider, J. E.; Foss, A. S. AIChE J. 1986, 72, 514. Denn, M. M. "Stabllity of Reaction and Transport Processes"; Prentice-Hail. Inc.: Englewood Cllffs, NJ, 1975. Eigenberger, G. Chem. Eng. Scl. 1972a, 2 7 , 1909. Elgenberger, G. Chem. Eng. Sci. 1972b, 2 7 , 1917. GouM. L. A. "Chemical Process Control, Theory and Applications"; AddisonWesley, Reading MA, 1969. Hoiberg, J. A.; Lynch, 8. C.; Foss, A. S. AIChE J . 1971, 1 7 . 1434. Horowltz, I.M. Roc. IEEE1976. 64(1), 123. Llu, S. L.; Amundson, N. R. Ind. Eng. Chem. Fundam. 1062, 1 , 200. Liu, S. L.; Amundson, N. R. Ind. Eng. Chem. Fundam. 1083, 2 , 183. Luss, D.; Amundson, N. R. AIChE J. 1867, 73,279. Mchenry, K. W.; Wllhelm, R. H. AIChE J. 1957, 2 , 83. Parega, G.; Reilly, M. J. Ind. Eng. Chem. Fundam. 1989, 8 , 442. Perlmutter, D. "Stability of Chemlcal Reactors"; Prentice Hall: Engelwood Cliffs, NJ, 1972. Reilly, M. J.; Schmltz, R. A. AIChE J. 1966, 12. 153. Rellly, M. J.; Schmltz, R. A. AIChE J. 1887, 13, 519. Rkter, A. B.; Douglas, J. M. Ind. Eng. Chem. Fundam. 1970, 9 , 21. Roffel, 8.;Rljnsdorp, J. E. Chem. Eng. Sci. 1974, 2 9 , 2083. Rosenbrock, H. H. "Computer Aded Control System Design"; Academic Press: New York. 1974. Root, R. B.; Schmitz, R. A. AIChE J. 1969, 75, 670. Root, R. B.; Schmltz, R. A. AIChE J. 1970, 16, 356. Shinnar, R., et al. Chem. Eng. Sci. 1972, 2 7 , 1627. Shmnar, R. ACS Symp. Ser. No. 72. 1978, 1. Slhrerstein, J. L. Ph..d Thesis, 1978. Slhrersteln, J.; Shinnar, R. Ind. Eng. Chem. Process Des. Dev. 1975, 74. 127. Sinai, J.; Foss, A. S. AIChE. J. 1070, 16, 659. van Doesburg, H.; de Jong, W. A. Chem. Eng. Sci. 1976, 31, 45. van Doesburg, H.; de Jong, W. A. Chem. Eng. Sci. 1976, 3 1 , 53. van Heerden, C. Chem. Eng. Sci. 1958, 8 133. van Heerden, C. Ind. Eng. Chem. 1953, 4 5 , 1242. Yagi, S.; Kunil, D. AIChE J . 1957, 2. 373.
Aris, R. Chem. Eng. Sci. 1957, 7 , 8. Ark, R.; Amundson. N. R. Chem. Eng. Sci. 1958a. 7 , 121. Ark, R.; Amundson. N. R. Chem. Eng. Sci. 1958b, 7 , 132.
Received for review April 2 , 1980 Accepted October 10, 1981
Kinetics of Coal Gasification Martin Schmal, Jos6 Lulz Fontes Montelro, and Jorge Lulr Castellan' COPPENFRJ, Program de Engenharia Quimica, Centro de Tecnologia, Bloco Gs/G1 15, Ilha do Fundo, Rio de Janeiro. Brazil
This work reports a kinetic study on the gasiflcation of Brazilian mineral coal with steam using a thermobalance. The coal is a high ash content (>50 wt %) subbituminous, run of mine coal (Charqueadas). Isothermal runs were made at temperatures between 800 and 1000 OC and at atmospheric pressure, using -14 +20 mesh Tyler size particles. The coal was devolatilized before the gasification. Initially, the influence of the particle size of coal and the flow of steam were studied in order to determine if diffusive effects are important. For certain conditions these effects are negligible and chemical reaction is the rate controlling step. The experimental results are well descrlbed by the unreacted core model above 850 OC and by the continuous model below this temperature. The gaseous products were determined at different times and temperatures and contained essentially H, CO, and CO,. Comparison of the present results with other data published shows that this coal is highly reactive.
Introduction Coal reserves in Brazil are of the order of 12 billion tons, which represents approximately 90% of the domestic fossil potential energy. It therefore is a major source of energy. So far, all Brazilian coals, with the exception of metallurgical coal, have been used exclusively for generation of electrical power. However, because of the energy crisis, there now is interest in using it as a source for energy to substitute for petroleum. Gasification is one of the important alternatives. CIENTEC, R. Washington Luiz 675, Port0 Alegre R. G. Sul, Brasil. 0196-4305/82/1121-0256$01.25/0
It is well known that gasification is the reaction between the carbonaceous matter of coal and the gasification agents (steam, hydrogen, carbon dioxide, oxygen, air, or mixtures of them). These reactions mainly produce hydrogen, carbon monoxide, carbon dioxide, and methane, depending on the gasification agents and the operational conditions. These products are used mainly for fuel gases for domestic and industrial consumption, reduction gases in siderurgical processes, and production of ammonia, methanol, and liquid fuels. Gasification involves two consecutive steps, namely devolatilization (including pyrolysis) and gasification. The products of pyrolysis will depend on the temperature, heat transfer, pressure, agents of gasification, and thermal 0 1982
American Chemical Society
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982 257
history of devolatilization. During the gasification of coal with steam at high temperatures and normal pressure, the two main reactions are C
+ H20(g)
CO
+ H,O(g)
-
CO
+ H2 (AH= 31.4 kcal/mol)
s C02
+ H2
(AH= -9.8 kcal/mol)
Chan and Papic (1976) studied the kinetics of subbituminous coal with steam in a differential fixed bed reactor. At normal pressure and temperature in the range 800 to 1000 "C, the hydrogen, carbon monoxide, and carbon dioxide compositions were found. The reaction rate was also found to be sensitive to temperature variations. These experimental data were well represented by the shrinking core model of chemical reaction control. On this basis, the appropriate value of activation energy for the reaction was found to be 32.7 kcal/mol. Pilcher et al. (1955) studied the influence of the temperature and the partial pressure of steam on the reaction rate using a thermobalance. The investigation covered the temperature range 800-1450 "C and partial pressures from 30 to 360 mmHg. An activation energy of 40.7 kcal/mol was found between 1000 and 1100 "C. For higher temperatures the diffusion rate was found to be significant. The apparent reaction order for steam was 0.66, and the CO/CO2 ratio increases with temperature. Riede and Havesian (1975) used the same experimental procedure described above but for temperatures between 500 and 900 "C and partial pressures between 70 and 120 mmHg. They varied the gas velocity from 5 to 15 cm/s and found the reaction rate to be the controlling step, with diffusion representing less than 2% of the total resistance. This resulted in an activation energy of 16.4 kcal/mol, but this seems to be a rather low value and suggests that diffusional effects may have been present. Hunt et al. (1953) used an integral fixed bed reactor to measure rates of gasification of char with steam for temperatures between 980 and 1370 "C at atmospheric pressure. A t lower temperatures it was suggested that the reaction rate was the controlling step, with an activation energy of 45 kcal/mol. However, at higher temperatures activation energy decreases and diffusion becomes the controlling step and a mechanism of simultaneous reaction and diffusion control is proposed. Jensen (1975) studied the kinetics of the gasification of residual coal with steam in a fluidized bed at temperatures from 1000 to 1300 "C. The reaction rate was found to be controlling and the activation energy was given as 19.8 kcal/mol. The results fit the shrinking core model of reaction and are independent of the particle diameter in the range of 0.07-0.59 mm. They found that the best temperature for this reaction is between 1100 and 1200 "C,because the reactivity is higher and the C02 content of the product gas is very low. However, the activation energy appears to be rather low and no reference was made to alternative interpretations of the data which would made a positive discrimination in favor of chemical reaction control. Johnson (1974) studied the kinetics of the gasification of a bituminous coal with hydrogen, steam, and mixtures using a thermobalance and a fluidized bed. The studies covered temperatures from 815 to 1100 "C and pressures from 1 to 70 atm. The results have been correlated in terms of a specific model which includes three consecutive steps devolatilization, fast reaction of methane, and a slow rate of gasification. The second step arose from the influence of the pressure of hydrogen. The gasification of the Brazilian coals has been studied by the Gas Development Corporation (1976) using a thermobalance employing hydrogen and steam as well as mixtures as gasifying agents. The pressure extended up 35 atm over the temperature
Table I. Proximate and Elemental Analysis of Charqueadas Coala proximate anal. moisture volatiles ash carbon
a
5.8 19.3 56.0 24.7
elemental anal, C H, S N, + 0, ash
28.3 1.9 0.4 13.4 56.0
ash anal.
SiO, AlzO, Fe,O, MgO CaO K,O Na,O other
61.1 25.3 4.1 3.3 2.7 1.7 0.6 1.2
Heat capacity, 2960 cal/g; specific weight, 1.95 g/cm3.
Figure 1. Thermobalance system: 1,gas (N2);2, tank; 3, pump; 4, vaporizer; 5, reactor; 6, condenser; 7, separator; 8, collector; 9, chromatograph; 10, spring; 11, rotameter; 12, thermometer; 13, basket.
range 870-980 "C but only gas compositions were determined. There is clearly difficulty in general in establishing the rate-controlling mechanism. Most authors have not presented information on a scheme for positive identification of the reaction rate as the rate-controlling steps. Rather, they have only tested for sufficiency in terms of the compatibility with the functional form of the equations describing chemical reaction control. Some tests for positive discrimination in favor of a particular mechanism need to be made. In particular, attention needs to be directed at the values of the activation energy particularly since this is a good indicator of possible diffusional effects. However, in comparing results between various authors it is necessary to bear in mind the fact that they are working with different types of coal. In this paper the kinetics of the gasification of coal having a very high ash content have been studied at 1atm and temperatures between 800 and 1000 "C in a thermobalance. Particular attention has been paid to establishing experimental conditions such that the influence of mass and diffusion have been eliminated. The results are shown to be correlated well by heterogeneous reaction models. Information is also obtained on the gas composition.
Experimental Section Coals from the IIF layer of the Charqueadas mine location of Rio Grande do Sul have been used in these experiments. This is subbituminous coal having a very high ash content and is typical of most Brazilian coals. Table I shows the elemental and proximate analysis of the coal. All experiments were carried out in a thermobalance (Figure 1) at isothermal conditions and atmospheric pressure. The thermobalance comprises an electrical furnace with an internal tube of Mullite of 40 mm diameter and 1000 mm height. The upper part of 500 mm length is of Pyrex and connected to the tube. Inside this tube the displacement of a quartz spring is allowed during the reaction. Distilled water is introduced through a water pump at a constant mass flow rate of 500 g/h and is vaporized in the pre-heater. The coal sample of particle size -14 +20 mesh Tyler is previously dried and placed in a
258
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982 1.0 X
0.0 0.8
0.7 G p
0.6
0
(O/h)
0.I
H/2Ro
1994
170
5
a
2
1423
170
4
0
3
1002
170
3
o
7
0673
170
2
0
8
0437
170
I
A
12
0 513
170
I
0.3
0.2
,
I
0.1 0.4
Mo (0)
6.0 -. -
1
0
IO
I5
20
t (min) Figure 2. Influence of mass on reaction rate at 1003 "C.
basket of platinum of 48 mesh Tyler. This basket is suspended by a spring of quartz from a platinum string. The quartz spring is fixed on the upper tube and cooled with a secondary flux of nitrogen. The weight loss of the sample is proportional to the displacement of the spring of quartz measured with the aid of a cathetometer. Initially the sample is placed in the upper part of the reactor (cold region) and nitrogen flows upward through the sample to eliminate air. When the temperature is attained in the reaction zone (isothermal region), the reactor is moved upward until the sample is located in the reaction zone and then closed. At this time, in the nitrogen atmosphere, pyrolysis begins. This step is very fast at the reaction temperature. When this step is completed, steam is introduced and the residual carbon is gasified. The weight loss is followed during the reaction and at the end the content of ash in the sample is determined by complete combustion of the final residue. Also, during the gasification several samples of gases are collected and analyzed chromatographically using two columns (molecular sieve of 5 A and the Porapak N). The excess of steam is condensed in a solution of sodium chloride (40%) to avoid dissolution of gases.
Results and Discussion The first step in the analysis of these experimental results involves establishing a relationship between the measured fraction of weight loss and conversion X. The conversion is defined as the ratio of the gasified carbon to the initial carbon in the coal, i.e.
X =
(Mo- M)/Mo - v 1-v-2
(1)
The influence of the mass of coal on the reaction rate was first studied in order to observe the limits to be expected on diffusion and thermal effects. These experiments have been done for a constant flow rate of steam and at the highest temperature level (lo00 "C), where these effects would tend to be more pronounced. For an initial mass of coal of 1.9 g it was observed, through the weight loss of sample, that the degree of conversion changed during the reaction. The results are presented in Figure 2. It can be seen that diffusion and the thermal effects on the reaction rate are not important when the ratio of the height to particle diameter is less than or equal to 2,
in other words, if the weight of the sample should be less than 0.7 g. To eliminate this effect all further experiments have been carried out using as a criterion that H / 2 R < 2. In addition, it is necessary to investigate the influence of the steam flow on the reaction rate and the importance of film diffusion at the external surface of the particle. Initially the Reynolds numbers were estimated and the corresponding mass transfer coefficientsj, were calculated. The values show that the mass transfer coefficient increases approximately 3 times by chainging the mass flow rate from 85 to 340 g/h. Therefore for coal samples dispersed on a layer of 0.20 cm in a basket of 2.5 cm diameter the mass effects could be negligible. The experimentswere also conducted at 1000 "C by varying the mass flow rate of steam (85,170, and 340 g/h). These results are shown in Figure 3, which shows that the gas film diffusion effect is negligible in this range of flow. Therefore, all subsequent tests were carried out within these limits, so that the reaction rate and /or the diffusion of gases through the ash layer are the only controlling steps during gasification. Results for the influence of the temperature on the reaction rate of gasification are shown in Figure 4. These runs were at a constant steam flow rate of 170 g/h and for temperatures between 800 and 1000 "C. They show that the reaction rate is very sensitive to temperature variations, which suggests that the chemical reaction is probably the rate-controlling step. The products of gasification were analyzed during the reactions but only for temperatures of 900 and 1000 OC. Hydrogen, carbon monoxide, and carbon dioxide were produced in significant amounts. The results are shown in Figures 5 and 6. These results suggest that the most important reactions are C + HzO CO + Hz (1) CO + HzO s COZ + HZ (11) Reactions I and I1 are considered to occur at the solidgas interface, but reaction I1 may also occur in the gas phase (Von Fredersdorf and Elliot, 1963). Figure 5 shows that the molar ratio CO/C02 increases as the conversion increases and varies between 2 and 3 for temperatures between 900 and 1000 "C. This can be interpreted that reaction I is more significant than reaction 11. It should be noted that this ratio is not influenced very much by the temperature. +
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982
259
1.0 XO.9
0.8 0.7 Exp
Mo
9
0.516
340
II
0.521
340
7
0.673
I70 170
12
0.513
170
IO
0.516
85
0.6 0
0.5 0.4
0
0.3 0.2
A
-a
tml-(l-x)8’’
0. I
5
0
I5
IO
t (min)
2o
Figure 3. Influence of mass flow on reaction rate at 1003 “C. 1.0 X
09
0.8 0.7 0.6
ua
T (OC)
0.4
0.3
0
801
0
851
b
9 50
I
I003
900 0.2
--t *
0. I
I-(l-X)
In
0
IO
20
30
40
50
t (min)
Figure 4. Influence of the temperature on reaction rate for G = 170 g/h and Mo= 0.437-0.725 g.
Table I1 shows the comparison of the data of CO/COz for different conversions and temperatures with other results in the literature. For higher temperatures the conversion increases, which may favor the reaction C + COz = 2CO too. This has not been taken into account and may explain the increase of the ratio CO/COz. It can also be observed that the temperature and the nature of coal have only a small influence on the ratio CO/COz as does the conversion of steam and solid and the pressure. The results of Pilcher et al. (19551, May et al. (1958), Goring et al. (1953), and Klei et al. (1975) all indicate that this CO/COz ratio is inversely proportional to the partial pressure of steam for their types of coal, while the results of Pilcher et al. (1955) and Hunt et al. (1953) show that the CO/C02 ratio is directly proportional to the conversion of steam. The data here show that this ratio increases as the conversion increases, which is in agreement with the
results of Goring et al. (1953) and May et al. (1958) and in disagreement with the results of Chan and Papic (1976). Johnson’s (1974) resulta also agree with our data at higher temperatures. It is clear that at this time there is no coherent pattern which can be identified as far as product gas composition is considered for the range of coals so far studied. There seems to be some evidence that the resulting balance of CO and C02 is a specific function of the type of coal and perhaps even processing conditions. Gasification is usually classified as an irreversible gassolid reaction. It is, therefore reasonable to use, as a first attempt of interpretation, the shrinking core model described by Levenspiel (1967) and Wen (1968) for the experimental results reported here. At the beginning, the reaction will occur at the external surface of the particle and the front gradually moves inside leaving an ash layer behind. At intermediate conversions of the solid, there is
260
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982
Table 11. Influence of Coal, Temperature, Pressure and Conversion upon the Ratio (CO/CO,)
6
0
present work (subbituminous)
'03
0
.0 _ +
0 2
(r
Goring et al. (coke)
---
i
I003
Chanand Papic (subbituminous)
Klei et al. Pilcher et al. (graphite)
Hunt et a1
May et al. 00
10
0 5
Conversion, X Figure 6. Gas composition as function of conversion.
essentially a shrinking core of nonreacted solid which diminishes as the reaction proceeds. There are three resistances for the reactant gas to pass through to reach the reacting surface: (1)diffusion through a stagnant gas film around the particle; (2) diffusion through the ash layer;
x,
temp,
%
"C
30 50 70 50 10 30 50 10 30 50 10 30 50 70 90
900 900 900 1003 870 870 870 870 870 870 900 900 900 900 900 760 843 1000 1200 1400 1100 1100 1100 1100 1204 1204 1204 1204 874 874 933 933 933
pressure, atm total
H,O
1.0
1.0
1.0 1.0 1.0 1.0 1.0 1.0 6.0 6.0 6.0
1.0 1.0 1.0 0.9 0.9 0.9
1.0
1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0
1.0 1.0 1.0 1.0 1.0
9.6
5.4 5.4 5.4 1.0 1.0
2
mol/ mol
1.9 2.7 3.4 2.0 2.3 3.8 6.7 1.0 2.5 4.0 4.0 2.3 3.6 1.0 0.6 6.0 1.9 4.0 3.4 11.8 12.0 8.7 4.0 2.4
1.0 1.0 1.0 0.3 0.9 0.19 0.19 0.19 0.04 0.12 0.29 0.47 0.01 0.06 0.21 0.36
47.8 18.9 5.6 3.3
0.30 0.70 0.26 0.60 5.90
1.0 0.7 3.7 1.5 1.2
(3) chemical reaction at the surface. One or all of these steps might be rate controlling in the gasification process. All three mechanism have been tested separately (Castellan, 1978). However, the results show that only the reaction at the surface is the controlling step. This was also found by Jensen (1975) and Chan and Papic (1976)
06
0.2
0 .I
0.0 t (min)
Figure 7. The shrinking core model.
eo/ co
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982 261
Figure 8. The continuous model. T(OC) 1000
950
850
900
800
I
I
1
Y
c A E -8
-
:
K O =
39,5 K c o l / m o l 1350 molr/cm2. mtn.
(
atm
)n
-
-9-
-
-10-
-
-11-
-12
>
795
I
I
I
s,o
895
990
Figure 9. The Arrhenius equation for the shrinking core model.
although their basic reaction scheme appears to be different. In this case, when the reaction is the controlling step the following equations may be applied
where (3)
The corresponding experimental data are plotted in Figure 7. From this figure we obtain values of .j- at different temperatures and can find the value of K using eq 3. A t high conversions and lower temperatures there is
a tendency for the experimental data to drift from the theoretical line, but this is not unexpected in view of the quasi-steady state assumed in deriving the model. As the conversion increases, the depth of the ash layer increases, so that the diffusion resistance of the steam through the thin layer also increases. These increases imply that the controlling step may change during the reaction. A t first sight this might be seen as the reason for the drift of experimentaldata noted above. However, if this were true, the deviation would also increase because the diffusional resistance is more significant at higher temperatures. This fact was not observed here. Therefore, it seems unlikely that this would be an alternative explanation of the tendency for the results to deviate from the theoretical curve at high conversions.
262
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982
T(T) Y
I
1000
950
900
8 50
800
1
I
1
I
I
Figure 10. The Arrhenius equation for the continuous model. eon
0
.
0
n
0
T = 1003.C EIP
0
A
9
0,516
340
0,521
340
7
0,673
170
12
0,513
170
IO
0,516
85
----
1
I
1
I
1
5
1
I
1
I
1
I
I
1
1
D
I
1
1
0
(a/h)
II
I
Me
(a)
1
MODEL
I
MODEL
E
I
20
I5
I
I
t (min)
Figure 11. Comparison of models and experimental data at 1003 "C.
An alternative model proposed by Wen (1968) assumes that all of the subparticles in the system react uniformly. This is similar to the dusty gas model for catalyst effectiveness factor and represents the other extreme. This model is represented by the expression
_ dX - kP1" (1- X) dt
(4)
Since the partial pressure of steam is constant, this equation can be integrated, and the conversion vs. time expression is -In (1 - X) = (kP,")t
(5)
The predicted behavior is shown in Figure 8 together with the corresponding experimental data a t different temperatures. It can be seen that this model is also a good fit to' the experimental data and is more sensitive at lower temperatures, exactly the opposite of the shrinking core
model. The calculated value of k for different temperatures using the results of Figure 8 are included in Table I11 for comparison with the other model. From the results presented in this table the Arrhenius equation can be plotted for K and k in terms of reciprocal temperature, as shown in Figures 9 and 10. It can be seen that the experimental data fit well by both equations and have an activation energy of 39.5 kcal/mol. This leads to the following expressions for K and k . a. shrinking core model
K = 1350 exp(-39500/RT)
(mol/cm2 min atmn)
(6)
b. continuous model
k = 2.4
X
lo6 exp(-39500/RT)
(l/min atm") (7)
It is interesting to note that the activation energies are the same for both models and have magnitudes which
Ind. Eng. Chem. Process Des. Dev., Voi. 21, No. 2, 1982
_----
263
- .-
T = 950.C EXP
Mo
G
(a)
(g/h)
13
0,519
I70
8
14
0,567
170
A
IS
0.718
170
- ---
MODEL
I
MODEL
E
Figure 12. Comparison of models and experimental data a 950 "C. 1P
X
0.9.
0.e
T =900°C ElP
6
8
A
I6 17
- ---
0.3
I 10
I
I
I
I
20
I
30
I
I 40
Mo (0)
G (O/h)
0,716 0,725 0,548
170 170 85
MODEL MODEL
I II
I
I
50
1
t (min)
Figure 13. Comparison of models and experimental data at 400 O C . Table 111. Comparison of Specific Velocities from the Shrinking Core Model and the Continuous Model
T,"C
7,
min
L/min
K , mol/cmz min atm"
h , L/min atm"
801 851 900 950 1003
158 74 35.1 18.0 8.8
0.022 0.055 0.106 0.206 0.414
1 . 3 X lo-* 2.8 X 5.9 X 11.5 X 23.5 X lo-'
0.022 0.050 0.106 0.206 0.414
kP,"
suggest that diffusion effects are not present. Both models and respective experimental results are shown in Figures 11-15 at different temperatures. The experimental data for lower conversions are fitted well by both. For higher temperatures, above 900 "C, the
shrinking core model gives a better representation than at lower temperatures. A t temperatures less than 850 "C, the continuous model is better. At lower temperatures the reaction is slower and steam diffuses faster through the pores of the particle. Otherwise, at higher temperatures the reaction is so fast that steam is consumed immediately on penetrating into the particle so the reaction is concentrated a t the surface. At intermediate temperatures (850-900 "C), there is a concentration gradient of gas through the particle and therefore the behavior is intermediate between the shrinking core and continuous models. These results imply that the experimental data cannot provide a definitive discrimination between the two mechanisms and emphasizes the need for a critical appraisal of data in order to reach useful conclusions.
--&
X
039
-
098
-
07-
0,6 -
T = 851 O EKP.
0
10
20
30
40
60
TO
I 80
Mo
G
( 0 )
(glhl
19
0,502
170
18
0,877
170
- --0
C
MOOEL
I
MODEL
II
1
1
1
90
too
I10
t
i
[min)
Figure 15. Comparison of models and experimental data at 801 "C.
Figure 16 shows the comparison of reactivity of Charqueadas coal with other well-known results in the literature. The comparisonhas been carried out a 900 "C, using the shrinking core model with reaction control. I t can be seen that for the same conversion, the time required for the C h a r q u e a h coal is less than for other coals, and this is an indication of greater reactivity although the others have a minor ash content. Figure 17 indicates the relative reactivity of the Charqueadas coal as function of the temperature in the range of 700 to lo00 "C. The relative reactivity is defiied as the ratio of the reaction rate of gasification of the Charqueadas coal compared with others from various sources in the literuatre but for the same conditions and at fixed con-
version, i.e., X = 0. On this basis the Charqueadas coal is half the reactivity of the other coals at 950 "C (Johnson, 1974)) but is 1to 5 times more reactive than most of the other coals in the range of 700 to 1040 "C, at atmospheric pressure.
Conclusions From these results it is possible to draw the following conclusions. (1)The devolatilization of coal is very fast and complete. (2) The reaction rate is very sensitive to temperature variations. (3) The results suggest that the carbon-steam reaction and the shift reaction are the most important reactions. (4) The resulting balance of CO and
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982 265
lo
0
40
t (min)
Figure 16. Comparison of coal reactivities at 900 O C . b(6).
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References
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Figure 17. Relative reactivity of this coal as function of temperature.
COz is a specific function of the type of coal and perhaps even processing conditions. (5) Both models of shrinking core and continuous model are well represented by the experimental data. These results imply that the experimental data cannot provide a definitive discrimination between the two mechanisms and emphasizes the need for a critical appraisal of data in order to reach useful conclusions. (6) The Charqueadas coal is more reactive than most of the other coals in the range of 700 to 1040 "C at atmospheric pressure although the others have a minor ash
content. Nomenclature A = frequency factor of the Arrhenius equation, cm/min Cf = fraction of fixed carbon (weight) C, = concentration of water vapor, mol/cm3 = initial concentration of carbon, mol/cm3 activation energy, kcal/g-mol G = mass flow rate of steam, g/h H = height of the bed, cm K = specific velocity per reaction surface, mol/cm2 atm" min
266
Ind. Eng. Chem. Process Des. Dev.
k = specific velocity per volume, l/min atm" Mo = initial mass of coal, g M = mass of coal at some instance, g P = pressure, atm R = constant of gases, cal/mol K Ro = mean radius of the particles, cm r = reaction rate, mol/min cm2 t = time of reaction, min T = temperature, "C or K U = fraction of moisture of coal (weight) V = fraction of volatile matter (weight) X = conversion of solid 2 = fraction of ash in coal (weight) bs = dry basis T = time for total conversion, min Literature Cited
1982,21, 266-272
Castellan. L. M.S. Thesis, Federal University of Rio de Janeiro, COPE, 1978. Chan, E. M.; Papic, M. M. Can. J . Chem. Eng. 1976, 5 4 , 645. Gas Development Corporation Report, Chicago, Sd.15 p, 1976. Goring, G. E.; Curran, G. P.; Zielke, C. W.; Gorin, E. Ind. Eng. Chem. 1953, 4 5 , 2586. Hunt, B. E.; Mori, S.;Katz, S.; Peck, R. E. Ind. Eng. Chem. 1953, 45, 677. Jensen, 0. A. Ind. Eng. Chem. Frocess Des. Dev. 1975, 14, 308. Johnson, J. L. Adv. Chem. Ser. 1974, 131, 145. Kiei, H. E.; Sahaglan, J.; Sundestrom, D. W. Ind. Eng. Chem. Process Des. Dev. 1975, 14. 470. Levenspiel, 0. "Chemical Reaction Engineering"; Wiley: New York, 1967. May, W. G.; Mueller, R. H.; Sweetser, S. B. Ind. Eng. Chem. 1958, 5 0 , 1289. Piicher, J. M.; Warker, P. L., Jr.; Wright, C. C. Ind. Eng. Chem. 1955, 47, 1742. Ride. B. E.; Havesian, D. Ind. Eng. Chem. Process Des. Dev. 1975, 14. 70. Von Fredersdorf, C. G.; Elliot, M. A. "Chemistry of Coal Utilization"; Wiiey: New York, 1963; p 892. Wen C. Y. Ind. Eng. Chem. 1968, 60.34.
Received for review October 31, 1980 Accepted August 31, 1981
Batchelder, H. R.; Busche, R. M.; Armstrong, W. P. Ind. Eng. Chem. 1953, 4 5 , 1856-78.
Evaluation of Inferential and Parallel Cascade Schemes for Distillation Control Nandklshor 0. Patke' and Pradeep B. Deshpande' Chemical and Environmental Engineering Department, University of Louisviiie, Louisville, Kentucky 40292
Adam C. Chou Mobil Research and Development Corporation, Princeton, New Jersey 08540
This paper presents a comparative assessment of inferential and parallel cascade schemes for controlling the top-product composition of a simulated multicomponent distillation tower. The steady-state data employed in the simulation study pertain to an industrial type column. The paper compares the closed-loop response of the single-temperature feedback scheme with the inferential and parallel cascade schemes. The findings of this work support the implementation of the inferential control scheme for this column.
Introduction Distillation columns are widely used in petroleum and chemical process industries. Improved composition control of these units is important not only from quality control considerations but also because distillation units consume vast amounts of energy. The columns are generally subjected to disturbances in the feed, and the control of product quality is achieved by maintaining a suitable tray temperature near its set point. This type of single-temperature feedback control is not always effective since maintaining a constant tray temperature does not necessarily result in constant product composition (Jones, 1971). Cascade control is used to improve the dynamic response of the column. In this type of cascade scheme, the top product composition is measured and then fed to a composition controller whose output serves as the set point of the tray-temperature controller. This scheme is referred to as the parallel cascade control system (Luyben, 1973) 'Department of Chemical Engineering,Ohio University, Athens, Ohio 45701.
because the manipulated variable, reflux flow, affects the top product composition and the tray temperature through two parallel transfer functions. The parallel cascade control system requires an on-line composition analyzer. The on-line composition analyzers introduce undersirable time delays in the control loops (Meyer et al., 1979). In addition, they are expensive and can be difficult to maintain (Scott, 1968; Hadley, 1969; Painter et al., 1978). The inferential control scheme circumvents many of the problems associated with the composition analyzers. The control scheme, which has been proposed by Brosilow and co-workers (1978), uses the easily available tray temperatures to estimate the product composition and uses the estimated composition to determine the required control effort. Many industrial multicomponent columns similar to the one on which this study is based employ single-temperature feedback control schemes. For the reasons cited earlier, this conventional feedback control scheme is sometimes inadequate. The purpose of this study is to assess through simulation, the benefits of the inferential control scheme and compare it with the parallel cascade
0196-4305/82/1121-0266$01.25/00 1982 American Chemical Society
