Application of a Kinetic Model for Comparison of Catalytic Cracking in a Fixed Bed Microreactor and a Fluidized Dense Bed Benjamin Gross," Donald M. Nace, and Sterling E. Voltz Research Department, Mobil Research and Development Gorp., Paulsboro, New Jersey 08066
A kinetic model has been applied to t h e catalytic cracking of gas oils of widely different molecular composition in both fixed and fluid bed reactors. T h e model describes t h e conversion and selectivity behavior of all stocks. Charge stocks with t h e highest content of paraffins and/or naphthenes have t h e highest rate constants for cracking and gasoline formation and the lowest rate of catalyst deactivation. Conversion in t h e fixed bed is higher than t h e fluidized bed for all stocks, while catalyst deactivation
and gasoline selectivity were independent of reactor configuration. B y comparing the same set of gas oils and catalyst on two different reactors, a single scale factor to account for inefficiency of fluid bed contacting could be determined. This allows one to m a p the fixed bed kinetics over to a specific fluid bed a n d t h u s increase t h e utility of the fluidized bed for scale-up to commercial FCC operation.
Introduction Numerous experimental studies have shown that a fluidized reactor generally requires more catalyst than a fixed bed to achieve a given conversion level under similar reaction conditions. The differences between the two reactors are normally attributed to the flow pattern of the gas in the fluidized reactor which usually involves some bypassing. Kunii and Levenspiel (1969) have recently reviewed much of this work, which has been mostly concerned with catalytic reactions of pure compounds. This paper compares the catalytic cracking of several gas oils in a fixed bed microreactor and a fluidized dense bed reactor. Comparisons of conversion, gasoline selectivity, and catalyst decay have been made. Direct comparisons of the cracking of cumene over silica-alumina catalysts in fixed beds and fluidized reactors have been reported by Mathis and Watson (1956), Gomezplata and Schuster (1960), and Iwasaki, et al. (1965). These workers showed that their fixed beds gave considerably higher conversions over wide ranges of reaction conditions than the various types of fluidized reactors used. The latter included reactors of different bed heights, length/diameter ratios, and fluidization methods. The conversion levels in the fluidized reactors usually decrease at v e r y . 1 0 ~and high gas velocities as would be expected. Reactor models which treated the fluidized reactors as two parallel reactors were developed to explain the results. Ishii and Osberg (1965) found that the isomerization of cyclopropane to propylene over a silica-alumina catalyst was first order in a fixed bed. The react,ion rate in a fluidized reactor was strongly dependent on linear gas velocity and bed height. The rate constants for the fixed bed were significantly greater than those for the fluidized reactors. Lewis, et al. (1959), reported higher efficiencies for fixed beds than fluidized reactors for the hydrogenation of ethylene over supported nickel catalysts. Baffles increased conversion in the fluidized reactor. They derived a mathematical model to correlate the results. Erofeev and Skringan (1969) established that the equilibrium constants for the dehydrogenation of cyclohexanol determined in fluidized reactors and fixed beds were similar. Johnstone, et al. (19551, studied the catalytic oxidation of ammonia and found that the ratio of the reaction rate constants in the fluidized reactor to those in the fixed bed was related exponentially to the superficial gas velocity. Massimilla and Johnstone (1961) clearly established the
high conversions in fixed beds and showed the effectiveness of baffles in a fluidized reactor. Shen and Johnstone (1955) found the reaction rate for the catalytic decomposition of nitrous oxide greater in a fixed bed than in a fluidized reactor. Transfer coefficients between phases in the fluidized reactor were calculated from the fixed bed results. Weekman and Nace (1970), using time-dependent decay terms, studied the selectivity and conversion behavior in fixed, moving, and fluid bed reactors for the catalytic cracking of gas oils. They showed that when both the gas oil and gasoline cracking functions of the catalyst decay at the same rate, the selectivity behavior of fixed, fluid, and moving bed reactors are identical. Nace, et al. (1971), showed that the fixed bed model could be used to determine rate constants for the contained fluidized bed. However, they did not directly compare the same gas oil and catalyst on two different reactors so that differences due to inefficient contacting not accounted for in the fixed bed model would be hidden in the calculated rate constants. Wojchiechowski, et al. (1971), and Campbell, e t al. (1969), have also used some variations of Weekman's model to describe gas oil cracking. In spite of the lower efficiencies and other problems associated with fluidized bed reactors, they are often the best reactors available for kinetic, process, and catalyst studies of catalytic cracking. Their advantages include ease of handling large amounts of fluidized catalysts, excellent heat transfer, nearly isothermal conditions throughout the reactor, uniform coking and catalyst decay, and suitability for large-scale operations. Fixed bed reactors are generally recommended for kinetic studies, but their utility in catalytic cracking is somewhat limited by large temperature gradients due to the highly endothermic nature of cracking and pressure drops across extremely fine fluid catalyst beds. Both of these effects can be minimized to some degree by using small amounts of catalyst in shallow beds ( i . e . , a microreactor). In this study a microreactor was used to obtain kinetic data for three different gas oils and the results compared with those obtained in a fixed (contained) fluidized bed. The kinetic model of Weekman (1968) was used to compare the gasoline selectivity, conversion, and catalyst decay in the two reactors. By using the relatively simple kinetic model of Weekman, the catalytic cracking of gas oils of widely different Ind. Eng. Chem.,Process Des. Develop., Vol. 13, No. 3,1974
199
properties has been described in both types of reactors. It has been established that catalyst decay and gasoline selectivity are similar in both fluidized dense bed and fixed bed reactors, whereas large differences exist in conversion. These results have been used to describe the behavior of laboratory risers and commercial FCC units more precisely.
Experimental Section The fixed bed microreactor consisted of a Vycor glass tube divided into a preheater section packed with glass beads and a catalyst bed with approximately 1 g of catalyst. The two sections were separated by a glass frit. The preheater-reactor assembly fitted inside an electrically heated glass core furnace with separate temperature controls for both preheater and reactor. The reactor was kept a t 900°F and the preheater at 950°F. Temperature fluctuation was quite small during the reaction period. Thermal cracking was negligible (as determined from experiments with inert quartz chips). The charge stock was pumped downward into the preheater section by means of a heated syringe pump. Due to the small size of the unit and the shallow catalyst bed, pressure drops were very small. Total product collection was effected in a liquid nitrogen trap. The liquid and gaseous products were then separated and analyzed on a Perkin-Elmer 900 temperature programmed gas chromatograph. Regeneration of catalyst was accomplished in situ by flowing air through the preheater and reactor while elevating the reactor temperature to >1000"F. Catalyst coke was determined volumetrically from the gases produced during regeneration. The fluidized dense bed reactor (contained-no replacement of catalyst) consisted of a stainless steel reactor with a diameter of 1.6 in. and 14 in. long with a separation zone of 7 in. In an experiment, the gas oil was vaporized in a separate preheater and by-passed around the reactor while the catalyst was fluidized with nitrogen. At the appropriate time, the switch from nitrogen fluidization to gas oil cracking in the reactor was practically instantaneous. The gas oil was normally diluted with 10 mol % of nitrogen during the run; the same nitrogen dilution was also employed in the fixed bed microreactor. Temperatures in the bed were measured by thermocouples a t three positions in the reactor and recorded periodically during the run. At the end of the cracking experiment, the catalyst was immediately stripped with steam for at least 10 min. The products were collected in a series of cold traps. The gaseous products were analyzed by mass spectrometry and the liquid products were distilled in a microstill with 10 in. of Cannon packing operated at a constant reflux ratio of 3:l to determine the Csf-430"F gasoline. Catalyst coke was determined by regenerating part of the catalyst in oxygen and determining C02 by Orsat gas analysis. The liquid products were also analyzed by the same chromatograph used with the fixed bed microreactor to compare analytical procedures. For both units, a commercial FCC zeolite catalyst was used. The on-stream times (catalyst residence time) of oil through the reactors were 1.25 and 5 min. The weight ratio of catalyst to total oil charged was varied from 0.5 to 7 by changing catalyst volume in the bed and/or the feed rate of gas oil. Since catalyst was not added during a fluidized or fixed bed run, the activity and thus cracking rate decreased with on-stream time for both systems. For both units products were collected over the entire experiment and are thus time-averaged products. The properties and molecular compositions of the three gas oils which were used in the study are given in Table I. P3 was an Amal gas oil, relatively high in paraffins. 200
Ind. Eng. Chern., Process Des. Develop., Vol. 13, No. 3,1974
Table I. Properties of Gas Oils Gas oils ~~~
P3
PA331
s39
467 901 276 0.10 0.03 13.48 0.02 0.6
447 886 332 0.42 0.09 11.44 0.22 3.0
559 928 312 0.56 0.07 12.48 0.10 1.8
46.2 35.1 18.6
17.7 26.2 56.2
28.5 36.2 35.3
66.7 25.0 8.2
47.9 29.5 22.6
60.7 25.0 14.3
~
ASTM distillation 5% 95%
(OF)
( O F j
Average molecular weight Sulfur (wt %) Nitrogen (wt %) Hydrogen (wt %) Conradson carbon (wt %) Bromine number Molecular structure by mass spectrometry Paraffins (wt % ) Naphthenes (wt % ) Aromatics (wt %) n-d-M Analysis CP (wt %) C N (wt %) C A (wt %)
PA331 was prepared by blending P3 and an aromatic extract. S39 was a Mid-Continent gas oil. The three charge stocks consisted of both light and heavy fuel oil fractions as shown by the ASTM distillations and molecular weights. The sulfur content varies from 0.1 to 0.6 wt 70; past experience has established that these concentrations of sulfur have no significant effects on the cracking kinetics. Basic nitrogen compounds are kn,own to reduce the cracking activities of acidic cracking catalysts; major catalyst deactivations would not be expected for the nitrogen concentrations listed in Table I. In the structural analysis by mass spectrometry, all of the compounds which contain aromatic rings are classified as aromatics. Thus, even large alkyl groups and naphthenic rings attached to aromatic rings are reported as aromatics. The n-d-M method is a correlation that estimates the concentrations of carbon atoms in various structures. For example, carbon atoms in alkyl side chains attached to aromatic rings would be reported as Cp (paraffinic carbons). The molecular structures of the three gas oils are significantly different and variations in cracking rates have been reported by Nace, et al. (1971).
Kinetic Model In the catalytic cracking of gas oils, comparisons of results are difficult since the catalyst deactivates rapidly and the implications of time-averaged data must be factored into the kinetics. Rate constants were calculated for each charge stock in both reactors with a kinetic model for catalytic cracking proposed by Weekman (1968). This model can be represented schematically as K
A, 4 A2 ,/&
K L
A3 A1 represents the gas oil charged, A2 is the gasoline produced, and As is the butanes dry gas coke. The gas oil cracking is assumed to be second order and gasoline cracking is first order; the justifications for these orders have been presented previously by Weekman. The differential reaction rate equations can be written as
+
-2 dA = -K1A12e-~& K,J12e-nk d to
-
+
I
-(Ki
+
(3)
where to = oil residence time, tc = catalyst residence time, and a = catalyst decay constant. The term ,-at, is a first-order decay function which accounts for the catalyst deactivation. These equations were written under the assumption of an isothermal, fixed bed reactor with plug flow of the vapor phase and negligible interparticle diffusion. In addition, the vapor transit time through the fixed bed reactor (-1-2 sec) is usually much less than the catalyst residence time. Individual oil molecules traverse the length of the reactor so fast relative to catalyst decay that they see catalyst of essentially uniform age, and the catalyst bed thus decays uniformly. Weekman previously derived the integrated rate equations which account for catalyst decay and time averaging effects. The calculated rate constants from this kinetic model are: a = catalyst decay constant; K O = (K1 K3) = rate constant for gas oil cracking; K1 = rate constant for gasoline formation; Kz = rate constant for gasoline cracking; K3 = rate constant for gas oil cracking to A,; K1/Ko = selectivity ratio for gasoline formation; and K2/Ko = selectivity ratio for gasoline cracking. The rate equations derived for fixed beds are also applicable to the contained fluid bed reactor with plug flow in the gas phase. The catalyst activity in the fluid bed also decays uniformly through the entire bed during a run due to the high rate of catalyst mixing. In addition, the fluid bed is generally less efficient than a fixed bed for various reasons, among them incomplete mixing and bubbling effects. Hopefully there is a region of operating conditions where optimum fluidization exists. Within this regime, efficiency is still lower than a fixed bed; however, if hydrodynamic conditions do not greatly change the efficiency over this range, it should thus be possible to adjust the kinetics by a scale factor to map from one reactor to another. Equations 1, 2, and 3 for the fixed bed would thus be modified by a scalar efficiency factor, i.e.
+
d A, - - -EfKA,nie-afC dt
(4)
For the fixed bed the equations for conversion are
t'
(instantaneous)
[ ']
1 + -KO e-"Lc
(5)
S
1+-
=
LYtc
In
(time-averaged)
(6)
1 + KyO e -atc
where c and i: are, respectively, instantaneous and timeaveraged conversions and S = space velocity. Instantaneous gasoline yields are calculated by the following equation which can be derived from eq 1 and 2
Ein
(2)+ I)&( Ein
(7)
where
Equation 5 can be used with eq 7 to obtain the instanta-
neous gasoline yields since
t=l-Al (9) Similarly the time-averaged gasoline yields are calculated with eq 5 and 7 as
For the fixed fluid bed the equation for conversion is D
(11)
The equation for instantaneous gasoline is unchanged from eq 7 since only the ratios of rate constants appear. The gasoline is thus only dependent upon conversion for fixed ratios of rate constants.
Results and Discussion Kinetic Rate Constants. Rate constants for the three charge stocks were calculated from the fixed bed cracking data. Equations 6 and 7 were used to obtain a , KO,K1, and K2. For the fluid bed eq 12 and 7 were used to obtain a , Ko', KI', and Kz', where Ko', Kl', and K2' are "apparent" rate constants; Le., EPK, = K,'. The apparent rate constant term is not an actual rate constant but a number that combines the actual rate constant with the efficiency of the given fluid bed relative to the fixed bed. The best fit of rate constants to the data was obtained using a nonlinear least-squares technique of Marquardt (1963). The values for the rate constants for the three charge stocks are given in Table I1 and the ratios of the rate constants in the fixed bed to the "apparent" ones in the fluidized dense bed reactor are shown in Table 111. The ratios of the catalyst decay constants are approximately 1 for the three charge stocks, which signifies that catalyst decay is independent of the type of reactor. The ratios of these rate constants for the fixed bed to the apparent ones for the fluidized dense bed reactor are slightly greater than 2 . This result is in agreement with the numerous literature references cited earlier in this paper. The selectivity ratios, K1/Ko and K2/Ko, are approximately 1 for each charge stock. A limited amount of experimental data was also obtained with a very highly aromatic stock. Accurate rate constants could not be calculated due to the limited data, but the general results were similar to those shown in Tables I1 and 111. Charge stock P3 has the lowest catalyst deactivation rate in both reactors which is consistent with its low concentration of aromatics. The values of KO,K1, and Kz are in the order P3 > S39 > PA331 in both reactors. Other detailed cracking studies by Nace, et al. (1971), and Voltz, et al. (1971), have shown that highly paraffinic and naphthenic charge stocks have lower values of a and higher values of KO, K1, and Kz than highly aromatic charge stocks. A direct comparison of the conversion (time-averaged) us. WHSV (weight hourly space velocity) for charge stock P3 in the fixed bed and fluidized dense bed reactors is shown in Figure 1. The points are the experimental data and the solid lines were computed from the kinetic model with the rate constants listed in Table 11. The conversion Ind. Eng. Chem., Process Des. Develop., Vol. 13,No. 3,1974 201
IO0
A
ac +
3
80 60
z
0
2
w
>
u
40
z
z
0
D E N S E BED
20
u
I
20 01
0 0
10
20
0
30
I
I
I
I
IO
20
30
40
W H S V ( W T / W T . - HR.)
W H S V (wT./wT.-HR.)
Figure 1. Plot of conversion time = 5.0 min).
us.
WHSV for P3 (catalyst residence
Figure 2. Plot of conversion dence time = 5.0rnin).
Table 11. Rate Constants for Cracking of Gas Oils.
Quantity CY
KO K O’ Ki Ki ’ K2 K2
’
KdKo (Ku’Ko)’ KdKo WdKo) ’
US.
WHSV for PA 331 (catalyst resi-
80
0 FLUIDIZED DENSE BED
Reactor
P3
PA331
539
Fixed Fluidized Fixed Fluidized Fixed Fluidized Fixed Fluidized Fixed Fluidized Fixed Fluidized
24.6 30.5 73.9 34.0 62.9 28.0 6.10 1.86 0.85 0:82 0.08 0.05
38.1 36.4 31.1 15.6 24.5 12.6 3.51 2.66 0.79 0.81 0.11 0.17
40.0 36.1 64.1 27.4 47.4 22.5
4 F I X E D BED
h
60
5
v
W
z 0
3
*..
...
a For fluid bed the calculated number is “apparent” rate constant; Le., it is the actual rate constant multiplied by a scale factor to account for contacting inefficiency of a fluidized bed.
20
0
1.94 0.74 0.82 0.07
40.
40 60 CONVERSION (WT.
20
0
80
100
$1
Figure 3. Comparison of gasoline selectivities from charge stock P3 in fixed bed and fluidized dense bed (catalyst residence time = 5.0 min).
0
F L U I D I Z E D D E N S E BED
A FIXED BED
Table 111. Ratios of Rate Constants for Cracking of Gas Oils in Fixed Bed to Apparent Constant from Fixed Fluidized Bed
Ratio (fixed bed) bed) Ko(fixed bed) Ko’(fluidizedbed) K l (fixed bed) K1’(fluidized bed) K2(fixedbed) K?’(fluidized bed) CY
cx (fluidized
K1/Ko(fixed K, ‘/Ko’(fluidizedbed) K2’Ko(fixed K2’/K0‘(fluidized bed)
3
v
P3
PA331
S39
Average
0.81
1.05
1.11
0.99
2.17
1.99
2.34
2.17
2.25
1.94
2.11
2.10
3.28
1.32
...
2.30
1.04
0.98
0.90
0.97
1.60
0.65
, ,
.
1.12
in the fixed bed is significantly higher than that in the fluidized dense bed over the range of space velocities which were studied. The conversions in Figure 1 are timeaveraged values for a catalyst residence time of 5.0 min; similar data for other catalyst residence times showed the same effect. The corresponding data for charge stock PA331 are shown in Figure 2. Again the conversion in the fixed bed is higher than that in the fluidized dense bed. The conversions in both reactors are lower for this charge stock than the corresponding ones for P3. These differences in conversion are also reflected in the rate constants in Table I1 and are associated with the respective molecular compositions of these two charge stocks. The gasoline selectivities for charge stock P3 are shown 202
/ I 0
Ind. Eng. Chern., Process Des. Develop., Vol. 13, No. 3, 1974
I
I
I
40 60 80 CONVERSION (WT. %)
20
V 100
Figure 4. Comparison of gasoline selectivities from charge stock PA 331 in fixed bed and fluidized dense bed (catalyst residence time = 5.0 rnin). in Figure 3. The points are the experimental data from the fixed bed and fluidized dense bed reactors. The solid line was computed from the kinetic model with the rate constants for the fluidized dense bed listed in Table 11; the computed selectivity curve for the fixed bed is quite similar. As expected, the gasoline selectivities from both types of reactors are essentially the same. It should be noted that both the gasoline yields and conversions in Figure 3 are time-averaged data for catalyst residence time of 5.0 min. The gasoline selectivities in the two reactors a t other catalyst residence time ( i e . , 1.25 min) are also quite similar; the time-averaged gasoline selectivities increase with decreasing catalyst residence times. The corresponding gasoline selectivities for charge stock PA331 are plotted in Figure 4. Again the selectivities in the fixed bed and fluidized dense bed reactors are about the same. They are significantly lower than the gasoline selectivities for charge stock P3. These differences are related to the molecular compositions of the two charge stocks.
Summary and Conclusion A kinetic model has been applied to the catalytic cracking of gas oils of extremely different molecular compositions and properties on both a bench-scale fluidized dense bed and a micro-sized fixed bed. We have been able to describe the conversion, selectivity, and catalyst decay behavior and also to successfully adjust kinetics determined in the fixed bed by one scale factor on the rate constants to map over to the fluid bed behavior. While this mapping is valid only over the fluid bed operating conditions for which the hydrodynamic conditions do not change efficiency, the data can be helpful in describing FCC commercial behavior.
Acknowledgment The authors wish to thank Vern W. Weekman, Jr., and Solomon M. Jacob for helpful comments and discussions.
Nomenclature A1 = weight fraction of unreacted gas oil A2 = weight fraction of gasoline A 3 = weight fraction of C1-C4 and coke A, = time-averaged yield of c = coke-on-catalyst, wt 7'0 Ef = efficiency of fluidized bed contacting K1 = gasoline formation rate constant, (hr-l)(wt fraction of gas oil in charge)K2 = gasoline cracking rate constant, hrK3 = A3 formation from A1 rate constant, (hr-I)(wt fraction of gas oil in charge)KO = K1 + K S = overall gas oil cracking rate constant, (hr- l)(wt fraction of gas oil in charge)n, = order of reaction for ith component
S = weight hourly space velocity, (wt of gas oil)/(& of cat.) (hr) tc = catalyst residence time, hr t o = oil residence time, hr tr = catalyst residence time at end of fixed bed run, hr Greek Letters a = decay velocity constant, hr- 1 c = instantaneous weight fraction converted Z = time-averaged weight fraction converted 0 = normalized time-on-stream, tc/tr
Literature Cited Campbell. D. R . , Wojchiechowski, B. W., Can. J. Chem. Eng. 47, 413 (1969). Erofeev, E. V., Skringan, E. A , , Russ. J. Phys. Chem. 43,71 (1969). Gomezplata, A . , Schuster, W. W., A.l.Ch.E. J., 6, 454 (1960). Ishii,T., Osberg,G. L.,A.l.Ch.E. J., 11, 279 (1965). Iwasaki, M., Kooya, I., Sueyoshi, H., Shirasaki, T., Echigoya, E., Kagaku Kogaku (Chem. Eng. Jap.), 29,892 (1965). Johnstone, H. F., Batchlor, J. D., Shen, C. Y., A./.Ch.E. J., 1, 318 (1955). Kunii, D.,Levenspiel, O., "Fluidization Engineering," pp 227-251, Wiley, New York. N. Y., 1969. Lewis W. K.,Gilliland, E. R., Glass, W.,A.l.Ch.E. J., 5, 419 (1959). Marquardt, D.W., J. SOC.Appl. Math., 11, 431 (1963). Massimilla, L., Johnstone, H. F., Chem. Eng. Scl., 16, 105 (1961) Mathis, J. F., Watson, C. C.,A./.Ch.E. J., 2, 518 (1956). Nace, D. M., Voltz, S. E., Weekman, V. W., Jr., lnd. Eng. Chem.. Pfocess Des. Develop., 10, 530 (1971). Shen, C.Y., Johnstone, H. F., A./.Ch.E. J., 1, 349 (1955). Voltz, S. E., Nace, D. M., Weekman, V . W., Jr., Ind. Eng. Chem., Process Des. Develop., 10, 538 (1971). Weekman, V. W., Jr., Ind. Eng. Chem., Process Des. Develop., 7, 90 ( 1968). Weekman,V. W., Jr., Nace, D.M.,A./.Ch.E. J., 16, 397 (1970). Wojchiechowgki, 8.W., Pachovsky, R. A , , Can. J. Chem. Eng., 49, 365 (1971).
Receiued for reuiew January 13,1972 Accepted April 9,1974
The Effects of a Surfactant on a Mechanically Agitated Extraction Column Richard P. Ruskan' Chemical Engineering Department, Ohio University, Athens, Ohio 45707
A dispersed benzene phase was used to extract acetone from a continuous aqueous phase in a YorkScheibel mechanically agitated extraction column. Introduction of 0.000, 0.001, 0.005, 0.025, and 0.625 wt YO Aerosol OT surface active agent in the solvent stream, entering in the column bottom, was found to increase the interfacial area available for mass transfer and decrease the overall mass transfer coefficient. I t was observed that Aerosol OT reduced overall stage efficiency with additions up to 0.025 wt YO,indicating that the reduction in the overall mass transfer coefficient, rather than the increase in the interfacial area, was the controlling mechanism in extraction. With surfactant additions greater than 0.025 wt YO in the solvent, overall extractor stage efficiency increased with increasing surfactant concentrations, indicating that increases in interfacial area dominated decreases in the overall mass transfer coefficient.
Introduction A surface active agent, sometimes termed surfactant, emulsifier, wetting agent, or soap, is any substance that has both hydrophile and hydrophobe molecular sections and, thus, is attracted to the interface between two unlike liquid phases. Collection and alignment of these molecules a t an interface reduce interfacial tension, a phenomI Present address, Arthur G. McKee and Company, Cleveland, Ohio 44131.
enon widely understood. If a surface active agent were present in an extraction process which involved dispersion of phases, one would then expect, due to lower interfacial tension than without surfactant, greater interfacial area across which extraction could occur and thus greater extraction rates. However, it is known that surface active agents also have a significant effect on the overall mass transfer coefficient in the extraction process. Johnstone (1939) noted that addition of a small amount of surface active Terigitol decreased extraction of acetic acid from isopropyl ether by Ind. Eng. Chem., Process Des. Develop., Vol. 13, No. 3, 1974
203
