of the hard-base substances such as an alkanolamine to remove all the desired species either as a mixture or in successive contacting. An alternative might be to choose oxidation reactants from the list of borderline and soft acids. Unfortunately, there will always be some absorption of aromatics in any solvent which removes mercaptans because they are both in the same (soft-base) category. Although the hard and soft-acid base generalizations are overly simplistic and without further development cannot be quantified to the point of predicting effective activity coefficients, they do classify the processes in a consistent pattern and could be helpful for screening new absorbents and oxidants for potential use in removing acid gas impurities of fluid hydrocarbons.
Literature Cited Frazier. J.. Hydrocarbon Process., 49 (4), 101 (1970). Hochgesand, G., I d . Eng. Chem.. 62 (7), 37 (1960). Hoffmann, H., Hydrocarbon Process., 53 (9), 103 (1974). Jockel. H., Triebskorn, B. E., Hydrocarbon Process., 52 (l),93 (1973). Kasai. K., Hydrocarbon Process., 54 (2). 93 (1975). Pearson, R. G., Chem. Br., 3 (3), 103 (1967). Pearson, R. G.. J. Amer. Chem. SOC.,85,3533 (1963). Pearson, R. G., Chem. Eng. News, 43 (26), 90 (1965). Swaim. C. C., Jr.. Hydrocarbon Process., 49 (3), 127 (1970). Wall, J., Hydrocarbon Process., 52 (4), 87 (1973). Wall, J., Hydrocarbon Process., 54 (4), 79 (1975). Williams, W. W., Pet. Refiner, 43 (7), 121 (1964).
Received for review September 12,1974 Accepted May 26,1975
Kinetics of the Activated Carbon-Steam Reaction Herbert E. Klei,' James Sahagian, and Donald W. Sundstrom Department of Chemical Engineering, University of Connecticut, Storrs, Connecticut 06268
A bed of activated carbon was reacted with steam at temperatures between 1400 and 150OoF and steam concentrations between 10.7 and 53.4 mol YO to determine the kinetic rate equations which predict the reaction rate. In addition, since activated carbon is a porous solid, the effect of gas flow rate through the bed was studied to see if mass transfer was important. The reaction rate was found to vary with the 0.58 power on the steam concentration with an activation energy of 63.6 kcallg-mol. The gas rate did have some effect on the reaction rate but it appeared not to be sufficient to shift the controlling step to mass transfer.
Introduction Activated carbon with its high surface area and affinity for water contaminants is becoming widely used in the purification of wastewater. In order to lower costs, the carbon is regenerated in a furnace near 1500OF in an atmosphere of steam and inert gases. During this regeneration step, the adsorbed organic materials are burned off the carbon restoring its adsorptive capacity. Unfortunately about 10% of the carbon is consumed per pass by reacting with the steam to form a variety of gaseous products. Several pilot plant studies have been done noting the effect of temperature and gas composition on activated carbon losses and on the absorptivity of the regenerated activated carbon (Loven, 1973; Battelle, 1970). Generally these studies were empirical and the kinetics of the reaction were not determined. For the steam-carbon system, however, a considerable amount of kinetic data are available. In contrast with activated carbon, the carbon or graphite used was nonporous and had a well-defined surface area for reaction. For the steam-carbon system the principal reactions are C
CO
+ Ha0 = CO + Hz
+ HzO = COz + Hz C
C
+ COZ= 2CO
+ 2Hz = CH4
AH = 31.14 kcal
AH = -9.65 kcal
AH = 40.79 kcal AH = -17.87 kcal
(1)
(2) (3) (4)
Reactions 1 and 2 are the significant ones a t the regeneration temperatures, particularly where the COz partial pres470
Ind. Eng. Chem., Process Des. Dev., Vol. 14, No. 4, 1975
sure is small compared to the steam (Walker, 1959; Long, 1948; Gadsby et al., 1946). Since the overall steam-carbon reaction is endothermic, difficulties in maintaining the isothermal conditions throughout the reaction zone were encountered (Derman et al., 1960; Malinauskas, 1970; Abel and Holden, 1962). Activation energies for this reaction range from 19.8 to 88 kcal/g-mol (Walker, 1959; Derman et al., 1960; Malinauskas, 1970; Long, 1948; Pilcher, 1955; Shchebrya et al., 1965). It is generally accepted that the hydrogen directly inhibits the rate of the steam-carbon reaction by occupying sites that would otherwise enter into the reaction. The effect of carbon monoxide is uncertain with some data showing inhibition while other data show no effect. Since the activated carbon is a porms solid, the steam reaction may be controlled by one or more of the transfer steps involving mass transfer across the stagnant gas film around the particle, pore diffusion within the particle, and surface reaction within the pore. At low temperatures the reaction rate is controlled by the chemical reactivity of the solid with an effectiveness factor equal to 1. At medium and high temperatures, steam concentration gradients exist within the activated carbon and the effectiveness factor is much less than 1. This paper presents a study to determine whether mass transfer is important in the reaction of steam with the activated carbon and the kinetic rate equations which predict the reaction rate.
Experimental Section Equipment. A thin layer of activated carbon (Nuchar
WV-W, 8-30 mesh) was placed in a 2 in. i.d. ceramic tube shown in Figure 1. Argon was used as an inert carrier gas. By controlling the rate of steam generation and argon flow, gas mixtures of various steam concentrations and flow rates were obtained. Prior to mixing in a heated tee, both of these gases were preheated in separate furnaces. The gas mixture entered the lower end of the reaction tube and passed through a bed of inert alumina sand. Here the gas mixture reached the desired temperature prior to passage through the bed of carbon. Carbon bed temperatures were monitored by four chromel-alumel thermocouples placed along the radius and length of the bed. After passing through the reaction bed, the gas (now composed of a mixture of Ar, H20, CO, CO2, H2, and CHI) left the reactor tube and was cooled in a water condenser where essentially all of the unreacted steam was condensed and measured. The gas then went to a tee where a desired amount was channeled through a Dry Ice condenser and analyzer train while the balance was vented. The Dry Ice condenser was used to remove any trace water thus preventing it from affecting the packing in the gas chromatograph. After passing this condenser, the C02 in the gas was monitored by a Lira Model 300 infrared analyzer. Once through this analyzer the C02 was selectively adsorbed out of the stream, again to prevent damage to the molecular sieve packing in a gas chromatograph. The Ar, CO, H2, and CH4 were then measured with a Carle Model 800 gas chromatograph. Since hydrogen was one of the gases measured, an 8.5% hydrogen-91.5% helium carrier gas was used. Procedure. Three sets of experiments were carried out. In the first set, the bed temperature was varied from 1400 to 1550°F. During these runs inlet steam concentration, inlet gas flow rate, bed height, and residence time were kept the same. Data obtained from these runs were used to calculate the apparent activation energy and frequency factor. In the second set, the flow rate was varied from 4.5 to 13.9 l./min a t 1400"F, 4.5 to 16.9 l./min a t 1418°F and 4.4 to 14.0 l./min at 1550°F. During this set, temperature and inlet steam concentration were kept constant. Bed height was varied in order to keep the residence time constant. The effect of mass transfer around the carbon particle was studied using these data. In the third set, the steam concentration was varied from 29.1 to 53.4% (mole) a t 1400"F, and 10.7 to 44.0% a t 1550°F. The apparent reaction order was determined from these runs. Prior to steam generation, a hot argon purge was used to desorb C02 from the carbon bed since activated carbon adsorbs C02 from the atmosphere. An Inconel wire screen separated the carbon bed from the alumina sand. In this way the carbon could be aspirated from the reaction tube after each run. Between runs, a hot argon purge was used to evaporate any water remaining in the reaction tube. Results The output concentrations during a typical run are given in Figure 2, showing that the concentration of CO and H2 in the exit gas were approximately equal and that steadystate conditions were obtained after 15-20 min from the initial feeding of the water. The rate of carbon gasification was taken to be the sum of the steady-state formation rates of CO, C02, and CH4. The exit steam concentration was obtained from the oxygen balance: ( H Z O ) ~=" ~(H20)in - CO - 2C02. Values of the effluent hydrogen concentration were lower than those predicted by the balance: H2 = CO 2C02 - 2CH4. If steady-state conditions are assumed, a material balance on water around an element of the reactor bed gives
+
+
F - (F + dF) rS d W = 0
(5)
T TEMPERATURE CONTROLLER
J';
4RGON
Figure 1. Apparatus.
TIME AFTER H 0 INJECTION (minl 2
Figure 2. Concentration in reactor effluent during a run.
F" d X = -rS d W
(6)
Assuming that the rate may be expressed by
eq 6 may be integrated to give
kSW
d n )= F" The only terms in eq 9 not known from the data are k and n. If a value of n is assumed, then the reaction rate constant, k, may be calculated for each point. The best n is the one which gives the smallest ratio of variance to mean. Such an analysis of the data of Table I a t 1550°F and 9.09 l./min, gave the best value of n to be 0.54. This method of integral analysis was not sensitive enough to determine a value of n a t 1400'F where the conversions were very low. To find the reaction order at this temperature a differential reactor analysis was made. Again assuming steady-state operation, a balance on the water gave C"XU" = kSW(C")" (1 - X)"
(10)
If C"Xuo is plotted versus C o ( l - X) on log-log coordinates, the slope of the line is n. Figure 3 shows the data plotted in this manner at 1400°F and gives n equal to 0.58. This result is in close agreement with the reaction order found from the 1550°F data. Taking n = 0.58 for the whole temperature range, values of the rate constant were calculated for all temperatures Ind. Eng. Chem.. Process Des. Dev.. Vol. 14. No. 4. 1975
471
Table I. Run Conditions S,
T,
OF
1400 1400 1400 1400 1400 1400 1400 1424 1455 1400 1550 1550 1550 1520 1493 1550 1550
oo, l./min
w,g
H,O consumed/ C consumed
9.85 9.85 9.85 9.85 4.88 15.03 9.85 9.85 9.85 16.88 9.09 9.09 9.09 9.09 9.09 4.40 14.01
88 .O 82.1 85.0 84.0 41.4 124.2 83.8 83.2 81.4 151.8 80.8 79.1 80.8 82.4 82.4 40.3 126.6
1.25 1.20 1.19 1.14 1.34 1.30 1.34 1.34 1.30 1.23 1.12 1.16 1.14 1.20 1.09 1.15 1.14
____
3r m0
1
co x
F- x 102,
103,
X'
mol/l.
mol/min
0.0403 0.0440 0.0329 0.0381 0.0216 0.0457 0.0712 0.0938 0.136 0.0633 0.325 0.432 0.392 0.301 0.283 0.325 0.537
4.44 2.87 5.82 3.84 5 -82 5.82 2.22 2.19 2.15 6.30 4 -82 2.34 3.12 2.37 2.40 3.12 3.12
4.33 2.83 5.73 3.88 2.84 8.75 2.19 2.16 2.12 10.62 4.38 2.13 2.84 2.16 2.18 1.37 4.37
-3 o
c
2 1
X 0
15r
I 2
I
4
3 CO(I--X)
5
6
7 8 9 1 0
103
.
.
Figure 3. Differential analysis plot.
.-r
o TOF
~
~
'
'5 ~
'
'
'
T=141E0F
T=14OOOF IO'
"
'
"
1"5
'
"
'
V = FLOW RATE( I/rnin)
Figure 5. Dependence of carbon loss on gas flow rate.
0001
00050
00052 I/T
00054
OR-'
Figure 4. Arrhenius plot for carbon loss.
using eq 8. An Arrhenius plot was made in Figure 4. The energy of activation was found to equal 63.6 kcal/g-mol. The final kinetic rate expression became 472
Ind. Eng. Chem., Process Des. Dev., Vol. 14, No. 4, 1975
The rate of activated carbon gasification was found to increase with increasing gas flow rate as shown in Figure 5. This dependency indicates that the external mass transport of reactants has an effect on the reaction rate. However, since the activation energy is around 63 kcal/g-mol, it does not appear that the external mass transfer is the dominant resistance in this temperature range. An estimation of the mass transfer rate was made using the mass transfer correlations for packed beds (McCabe and Smith, 1967) and showed that the observed rate of reaction was l,&-3%of the mass transfer rate, again indicating that the external diffusional resistance is not dominant. Riede and Hanesian (1975) also found this effect and they concluded that their reaction of graphite with steam was also surface reaction controlling also by calculating the mass transfer resistance for their graphite sample. When the external mass transport resistance is greater than that of pore diffusion or surface reaction, the rate of steam-activated carbon reaction may be expressed as r = k, (C - C,)
If external mass transfer was controlling, the reaction would have been first order with respect to steam concen-
tration. Since the reaction was found to have an apparent order of reaction near 0.6 and to be flow rate dependent, the reaction occurs in the intermediate region where both chemical kinetic and mass transfer effects are present. As mentioned above, however, chemical kinetics appears to be the more important resistance. Conclusion The reaction rate of activated carbon with steam can be represented by the steam concentration to the 0.58 power. The activation energy was found to be 63.6 kcal/g-mol. Although it appears that the dominant resistance is the reaction rate, external mass transfer does have some effect. Nomenclature C = concentration of water in vapor phase, g-mol/cm3 F = water feed rate, g-mol/min k = reaction rate constant n = orderofreaction r = mol of carbon consumed/min-g of carbon R = gas constant, 1.98 cal/g-mol OR S = mol of water consumed/mol of carbon consumed T = temperature of reaction, OR u = gas flow rate, cm3/min at reactor temperature and pressure W = weight of carbon bed, g X = fraction of water converted
Superscripts O = inlet f = outlet L i t e r a t u r e Cited Abel, W. T., Holden, J. H., U S . Bur. Mines Rept. Invest., No. 8000, 1-22 (1962). Battelle Memorial Institute, Columbus Laboratories, Columbus, Ohio, March 1970. Derman, 6. M., Rogailin, Farberov, I. L., Akad. Nauk SSSR, Tr. lnst. Goryuchikh iskopaemyhk, (13), (1960). Gadsby. J., Hinshelwood. C. N., Sykes. K. W., Proc. Roy. SOC.London. Ser.. A 187, 129-151 (1946). Long, F. J., Sykes, K. W., Roc. Roy. SOC.London, Ser. A, 193, 377-399 (1948). Loven, A. W., Chem. Eng. Progr., 69 (1I), 56-62 (1973). Malinauskas, A. P., Nucl. Eng., Chem. Eng. Progr. Symp. Ser., Pari XXl, 86 (104). 81-93 119701. McCabe, W. L.: Smih, J. C.. “Unit Qperations of Chemical Engineering, McGraw-Hill. New York, N.Y., 1967. Pilcher, J. M., Walker, P. L., Jr.. Wright, C. C., lnd. Eng. Chem., 47 (9). 1742-1749 (1955). Ride. 6. E., Hanesian, D., lnd. Eng. Chem., Process Des. Dev.. 14, 70-74 ( 1975). Schchibrya, G. C., Morozov, N. M., Temkin, M. I., Kinet. Katal.. 8 , 1057 (1965). Walker, P. L.. Rusinko, F. Jr.. Austin, L. G., Catalysis, 11, 133-221 (1959).
Received f o r review February 10,1975 Accepted June 30,1975 This work was supported in part by funds provided by the U S . Department of the Interior Office of Water Resources Research through the Institute of Water Resources, University of Connecticut, as authorized under the Water Resources Research Act of 1964, PL-83-379, as amended.
Mass Transfer in Packed Beds with Two-Phase Flow S. Goto,’ J. Levec,2and J. M. Smith’ university of California, Davis, California 956 76
Mass transfer coefficients were measured at 25’C and 1 atm for cocurrent liquid (water) and gas flow in beds of small particles (0.054-0.29 cm) of napthalene and CuO-ZnO. The coefficients for transfer between particle and liquid are not greatly different for the three arrangements: upflow, downflow (trickle bed), or liquid full. At high gas rates and low liquid rates upflow gives somewhat higher transport rates, while trickle beds are favored over liquid-full operation at high liquid rates. Desorption of oxygen was measured to obtain liquid-to-gas coefficients kLa when mass transfer in the liquid phase controlled interphase transport. For downflow such coefficients were independent of gas flow rate, but in upflow kLa increased with gas rate. Upflow gave higher coefficients at high gas and liquid rates. Desorption of napthalene from water to air was employed to evaluate the gas-side coefficient in liquid-to-gas transfer for trickle-bed operation. The results showed a high sensitivity to gas flow rate and a modest sensitivity to liquid rate.
In three-phase reactors involving gas and liquid flow as well as the solid catalyst particles, interphase mass transfer resistances are likely to be more important than for the simpler fluid-solid catalytic system. For example, Goto and Smith (1975b) and Baldi et al. (1974), in studies of the oxidation of aqueous solutions of formic acid, found both gasliquid and liquid-solid mass transfer resistances to be significant in trickle-bed operation, while the mass transfer resistance was negligible for liquid-full (no gas phase) operOn leave from University of Nagoya, Japan. On leave from University of Ljubljana, Yugoslavia.
ation. Therefore, reliable information for mass transfer coefficients is often needed for design of packed-bed reactors with gas and liquid flow. Further, such information may help in deciding whether to use a cocurrent upflow or downflow (trickle-bed) flow arrangement. The data reported here were obtained to help fulfill these needs by supplementing the available information, which is particularly meager for beds of small particles. In particular, measurements were made for upflow and downflow in the same apparatus in order to obtain a direct comparison for these two methods of operation. Particle--liquid mass transfer coefficients were deterInd. Eng. Chem., Pracess Des. Bev., Val. 14, No. 4, 1975
473
