717
Ind. Eng. Chem. Process Des. Dev. 1986, 25, 717-722
Astarb, 0. "Mass Transfer wlth Chemical Reaction": Elsevier: New York,
where
e,
= 1 + q0/2
.
io117 ,--.
(AI())
Rather than fixing the wall temperature, the wall heat flux was used as a parameter in the calculations. The validity of the calculation was established by reproducing the results over a limited range of conditions. The method employed quasi-linearization with function iteration and fourth-order Runga-Kutta integration.
Literature Cited Aibright, L. F. "Processes for Major Addition-Type Plastics and Their Monomers"; McGraw-Hill: New York, 1974; pp 13-34. Albright, L. F.; Crynes, B. L.; Corcoran. W. H. Tyrolysls Theory and Industrial Practice": Academic Press: New York, 1983.
Bird, R. 6.; Stewart, W. E.; Llghtfoot, E. N. "Transport Phenomea": Wiley: New Yo& 1960.
~ ~ ~ ~ ~ ~ i ~ , ~ ~322.. R ~ ~ ~ ~ ~ Brokaw, R. s. J . C / " . phvs. 1061.35, 1569. Danckwerts, P. V. "Gas-Liquid R~ctlons";McGraw-Hill: New York, 1970. Froment. 0. F.; Pijcke, H.; Goethals, 0. Chem. Eng. Sci. 1961, 73, 173. F ~R. R.; Smith, ~ ~ A I C ~ EJ~. 1062,8 ~ , 654. ~ , J. M. Kershenbaum. L. S.; Martin, J. J. AIChE J . 1967, 73, 148. Krieve, W. F.; Mason, D. M. AIChE J . 1961,7 , 281. Schotte, W. Ind. Eng. Chem. 1058,50, 683. Sundaram, K. M.; Froment, F. G. Chem. €ng. Sci. 1979,3 4 , 117.
Received for review April 1, 1985 Revised manuscript received November 8, 1985 Accepted November 18, 1985
Radial Dispersion Characteristics of Two- and Three-phase Fluidized Beds Yong Kang and Sang D. Kim' Department of Chemical Engineering, Korea Advanced Institute of Sclence and Technology,
Seoul 13 1, Korea
Radial dispersion coefficients of a continuous liquid phase in two-and three-phase fluidized beds have been studied in a 10.2 cm i.d. column. The coefficients were determined from the radial tracer (KCI) concentration profiles and the use of an infinite space model. Effects of water (4-12 cm/s) and air (0-10 cm/s) velocities and particle size (1.7-6.0 mm) on the radial dispersion coefficient were determined. The mixing intensity of the liquid phase in the radial direction was examined in terms of energy consumption rate. The Coefficient Increased with the gas velocity and particle size. However, the coefficient exhibited a maximum value at a liquid velocity in two- and three-phase fluidized beds. The coefficient increased with the energy consumption rate in the beds. I t was found that the coefficient in terms of the Peclet number is well-represented by the isotropic turbulence model.
Various hydrodynamic aspects of three-phase fluidized beds have been investigated in recent years, reflecting the increasing application of three-phase fluidized beds to industrial processes. Recently, three-phase fluidization was reviewed by Muroyama and Fan (1985). Among the different modes of three-phase contacting operations, one with solid particles fluidized by cocurrent flow of gas and liquid appears to be the most important mode of operation. Examples of industrial applications of this mode are the H-oil process for hydrodesulfurization of residual oils, the H-coal process for coal liquefaction, and biooxidation of wastewater. In a continuous multiphase processing system, the contacting between phases and the mixing intensity may predominate the other phenomena governing the hydrodynamic characteristics of three-phase fluidized beds. To obtain the information on the mixing intensity in three-phase fluidized beds, several investigators (Schugerl, 1967; Vail et al., 1968; Michelsen and Ostergaard, 1970; Kim et al., 1972; El-Temtamy et al., 1979a, 197913;Kim and Kim, 1983) have studied the axial dispersion of the liquid phase in recent years. The intensity of mixing has been found to depend on the particle size and velocities of fluids. The intensity of liquid axial mixing is found to increase with gas velocity, whereas its variation with liquid velocity may depend on the particle size (El-Temtamy et al., 1979a; Kim and Kim, 1983). For radial dispersion of the liquid phase in three-phase fluidized beds, Vail et al. (1968) reported that the intensity of radial liquid-phase mixing increased with liquid velocity but it decreased with gas velocity. On the other hand, 0196-4305/86/1125-0717$01.50/0
El-Temtamy et al, (1979b) later reported that the liquidphase radial mixing coefficient increased with the gas velocity and the coefficient exhibited a maximum value at a liquid velocity in the beds of 2- and 3-mm glass beads. Also, the radial dispersion coefficients were found to be 1 order of magnitude lower than the axial dispersion coefficients, and its values were higher than the comparable solid-liquid and liquid-gas systems. Most of the previous mixing studies were focused mainly on axial mixing in the beds. However, the intensity of radial mixing is an important parameter in a large diameter column since the radial mixing of the liquid phase can affect the mass and heat transfer in three-phase fluidized beds (Kang et al., 1983,'1984, 1985). In the present study, the effects of liquid and gas velocities and particle size on the radial dispersion coefficient of the liquid phase have been determined. The radial mixing intensity of the liquid phase has been interpreted in terms of the energy consumption rate in two- and three-phase fluidized beds.
Experimental Section Experiments were carried out in a Plexiglas column 3 m high and 0.102 m in diameter as shown schematically in Figure 1. In order to maintain a relatively constant dynamic liquid level, a concentrically mounted outlet header was attached to the top of the column as a liquid weir. The solid particles were supported on a perforated plate which contains 207 evenly spaced holes of 3-mm diameter which served as a liquid distributor. The distributor was 0 1986 American Chemical Society
c
~
~
718
Ind. Eng. Chem. Process Des. Dev., Vol. 25, No. 3, 1986
.
0 3 3
3
Figure 1. Equipment: (1)main column, (2) tracer tank, (3) pump, (4)tracer inlet, (5) conductivity probe, (6) conductivity bridge, (7) recorder, (8) dc power supplier, (9) pressure tap, (10) rotameters, (11) filter, (12) regulator, (13) liquid tank.
situated between the main column section and a 0.5 m high X 0.102 m diameter stainless steel distributor box into which liquid was introduced through a 2.54-cm pipe from the liquid reservoir. The water flow rate was measured with one of two calibrated rotameters and regulated by means of glove valves on the feed and bypass lines. The oil-free compressor air was fed to the column through a pressure regulator, filter, and a calibrated rotameter. It was admitted to the column through four 0.64 cm i.d. perforated pipes drilled horizontally in the grid. The pipes were evenly spaced across the distributor plate having 88 holes of 1 mm in diameter. Pressure taps were mounted flush with the wall of the column at 0.125 m height intervals. The static pressure at each of these points was measured with a liquid manometer. The solids used were either 1.7-, 4.0-, or 6.0-mm glass beads with a density of 2.5 g/cm3.
Phase Holdup Measurement Gas and liquid were introduced into the bed of solid particles at the desired superficial velocities. When a steady state was reached, the pressure profile along the entire height of the column was measured by water manometers. The bed height was taken as the point at which a change in the slope of the plot was observed. Individual phase holdups of gas, liquid, and solid were determined from the knowledge of pressure drop, expanded bed height, and properties of each phase (Kim et al., 1975; Kat0 et al., 1981; Kim and Kim, 1983). Radial Dispersion Coefficient Measurement Potassium chloride solution (1.0 N), the tracer, was contained in a reservoir which was connected through a micrometering pump to an injection nozzle, a point source, located at the center of the column from 5 cm above the distributor plate. The concentration of the tracer was monitored by four conductivity probes connected to the conductivity bridges. The probe was made of 0.5 mm diameter platinum wires encased in a 5 mm i.d. stainless steel tube which in turn was inserted horizontally into a fluidized bed column up to the column center at 10 cm above the injection nozzle. Four probe tips were installed though the downward holes of the tube. The holes or thus the tips were 1.53 cm apart from each other, starting from the center of the column. The conductivity gain obtained from the response curve was converted into the concentration from the previously calibrated relationship between the conductivity gain and the concentration of KCl. From the obtained concentration profiles, the radial dispersion coefficient of the liquid phase was calculated by using the
X (r/A)
Figure 2. Concentration profiles in radial direction in three-phase fluidized beds. d, = 4.0 mm, U, = 6.0 cm/s, U, = ( 0 )0, (A)2, (w) 4, (v) 8 cm/s.
L'i
cm/s
Figure 3. Effect of liquid velocity on radial dispersion in liquid fluidized beds. d, = ( 0 )1.7, (A)4.0,).( 6.0, (0) (El-Temtamy et al., 1979b) 2.0 mm.
infinite space model (Klinkenberget al., 1953; El-Temtamy et al., 1979b). Since the present experimental conditions satisfy the assumptions of the above model, the radial dispersion coefficient can be derived from the following relations \k
C/C, = - exp(-\II/2{X2) 2{
(1)
where
(3)
X = r/R
(4) in which D, and D, are the dispersion coefficients of the liquid phase in the axial and radial directions, respectively, Z is the height of the measuring point from the tracer inlet, and R is the radius of the column. The radial dispersion coefficient of the liquid phase can be determined from the slope or intercept of the C/C, vs. X plot of the semilogarithmic coordinate. Results and Discussion Liquid Radial Mixing in the Liquid-Solid System. A typical concentration profile in the radial direction in liquid fluidized beds (U, = 0) is shown in Figure 2. The radial dispersion coefficient has been determined from these concentration profiles. As can be seen in Figure 3, the radial dispersion coefficienta go through a maximum point as the liquid flow rate in the beds increases. A same trend has been observed by El-Temtamy et al. (197913).
Ind. Eng. Chem. Process Des. Dev., Vol. 25, No. 3, 1986 719 Table I. Comparison between the Predicted and Experimental Values of Bed Porosity at Which the Maximum Liquid Radial Dispersion Coefficient Is Attained. d,, mm WI,mPa pS, g/cm3 Ut, cm/s n elDr,- (exptl) e ~ (calcd) ~ , clh," ~ (exptl) ~ 1.7 1.o 2.5 21.5 2.65 0.62 0.69 0.68 2.48 0.60 0.61 0.60 4.0 1.0 2.5 33.5 2.43 0.59 0.55 0.57 6.0 1.0 2.5 41.2
It has been known that mixing in multiphase fluidized beds may depend on the degree of turbulence, flow regimes of the liquid-solid system, configuration of the distributor, etc. (Joshi, 1980). Increasing the water flow rate leads to an increase in the turbulence in the bed and consequent increase in liquid-phase mixing in the radial and axial directions. However, at higher liquid flow rates, the hindrance of particle movement would be reduced considerably due to the higher liquid holdup and decrease in solid-phase holdup. Consequently, the intensity of turbulence per unit cross-sectional area would be reduced. The bulk intensity of turbulence can be expressed as a function of the liquid flow rate and phase holdups of the liquid and solid phases (Joshi, 1983): I = 1.5Ult,/€l = 1.5u1(1- €1)/€1 (5) As can be seen in (5), two conflicting effects may exist with variations in the liquid flow rate; Le., the intensity of turbulence may or may not increase with the liquid flow rate since the liquid-phase holdup increases with the liquid flow rate. In addition, the effect of the hindrance of the solid movement on the liquid-phase radial mixing may be a dominant one to that of axial mixing since the presence of solids may promote the radial flow of the liquid element rather than reduce the axial fluid flow. Therefore, the radial dispersion coefficient of the liquid phase may exhibit a maximum value at the optimum solid-phase holdup. The turbulence intensity (eq 5) is similar to the energy assumption rate (eq 6) which was derived from the mechanical energy balance in a liquid fluidized bed (Kang et al., 1985). Ul(1 - 4 ( P , - P l k (6) ED2 = ElPl
The energy consumption rate can be estimated from the liquid-phase holdup, liquid velocity, and properties of the solid and liquid (eq 6) by the Richardson and Zaki's equation (1954)
From the first derivative of eq 7 with respect to the liquid-phase holdup or bed porosity, the liquid-phase holdup at which the maximum energy consumption rate would occur in the bed can be estimated. The resulting optimum liquid-phase holdup can be expressed as (Kang et al., 1984)
in which n is the Richardson and Zaki's index. Since the bulk turbulence intensity can be expressed as eq 7, the bed porosity corresponding to the maximum turbulence intensity may be expressed as eq 8. Therefore, the radial dispersion coefficient of the liquid phase in liquid fluidized beds may attain the maximum value when the energy consumption rate or turbulence intensity exhibits the maximum value. The agreement between the predicted and the experimental values of liquid-phase holdup or bed porosity at which the maximum liquid radial dispersion coefficient
d u
0o
z
6
4
L'g
a
io
cmis
Figure 4. Effect of gas velocity on radial dispersion in three-phase (El-Temtamy et fluidized beds. VI = ( 0 )4,(A)6,).( 8, (v) 10, (0) al., 1979b) 2.0 cm/s. 8,
6 -
/:
/..//"/' /a
720
Ind. Eng. Chem. Process Des. Dev., Vol. 25, No. 3, 1986
21 0
2
4
6
cg
8 cm/5
1
0
I
0
0.3
.
3 a5
a4
1
0.6
Et
Figure 6. Effect of gas velocity on radial dispersion in three-phase fluidized beds. d, = 6.0 mm, VI = ( 0 )6, (A)8, (m) 10, (V)12 cm/s.
Figure 7. Relationship between liquid and solid holdups in threephase fluidized beds. d, (mm), U, (cm/s);).( 1.7, 4.0; (A)1.7, 8.0; ).(
4.0, 8.0; (V)6.0, 8.0.
Table 11. Volume Fraction of Fluidized Region and the Ratio of the Liquid Phase in the Wake-to-LiquidPhase in the Fluidized Redon in Three-phase Fluidized Beds d,, mm pl, mPa s U,, cm/s Cf m 1.7 1.7 4.0 4.0 6.0 6.0
1.0 1.0 1.0 1.0 1.0 1.0
0.910 0.880 0.890 0.800 0.860 0.745
4.0 8.0 4.0 8.0 4.0 8.0
0.087 0.124 0.105 0.163 0.111 0.176
phase fluidized beds (El-Temtamy et al., 1979a; Kim and Kim, 1983). This motion of bubbles and wakes may also promote radial mixing since nonuniform velocity profiles and gas holdup would occur locally between the bubble zone and the continuous liquid phase due to the rising bubbles. The three-phase fluidized bed may be composed of three different functional regions, namely, fluidized, wake, and gas bubble regions (Bhatia and Epstein, 1974; Henriksen and Ostergaard, 1974; Darton and Harrison, 1975; Khang et al., 1983) as written in the form tf + E , + Eg = 1.0 (9) in which subscripts f, w, and g denote fluidized, wake, and gas bubble regions, respectively. The fluidized region contains solid particles and liquid, while the wake region may consist of liquid alone, since the solid content in the wake is found to be far smaller than that in the fluidized region (Bhatia and Epstein, 1974). When the ratio of the liquid phase in the wake to the liquid phase in fluidized region is m, the following relations hold: Ef = lf' + Ef, = tfl + t, (10) t, = mefl (11) The overall fraction of the liquid phase can be written as tl = tfl + t, = (1 + m)tfl (12) Combining eq 10 and 12, we have =
€1
- Efl = Ef - -
(13) l + m Thus, from the plot of solid-phase holdup vs. liquid-phase holdup, the ratio of the liquid phase in the wake to the liquid phase in fluidized region, m,and the volume fraction of the fluidized region can be determined (Figure 7). The obtained values of m and Ef are summarized in Table 11. The amount of bubble-driven liquid in the wake region increased with the gas velocity in all the cases studied (Table 11). This may suggest that the increase in the gas velocity may intensify the liquid-phase mixing both in the radial and axial directions due to the increase in the ratio of the liquid in the wake to the fluidized regions. Therefore, the radial dispersion coefficient of the liquid phase, D,, may t,
Ef
6
10
!
Ut, c m / s
Figure 8. Effect of liquid velocity on radial dispersion in threephase fluidized beds. d, (mm), U, (cm/s);).( 1.7, 4.0; (A)4.0, 4.0; (+) 6.0, 4.0; (0) 0.96, 4.0; (A)2.0, 5.0; (0) 3.0, 6.0. 0 , A, = present study; 0, A, 0 = El-Temtamy et al. (1979b).
+
increase with the gas velocity since the liquid-phase mixing may be caused mainly by the movement of wakes. The radial dispersion coefficient exhibits the maximum value with an increase in the liquid velocity as can be seen in Figure 8. It can be anticipated that the overall mixing intensity in the bed may increase with fluid velocities. However, it can be seen from Figure 8 that there is an optimum liquid velocity which provides a maximum radial mixing of liquid phase in three-phase fluidized beds. For a given fluidizing condition, the flow pattern and velocity of particle movement will be governed by the bed porosity which will be affected by fluid velocities and the geometry of the distributor, etc. The bed porosity increased with the liquid velocity due to the higher bed expansion. Thus, the space between the particles will be widened with the liquid velocity. The solid particles have an inertia force which was rendered by fluid velocities. This force may contribute to the uniformity of three-phase fluidized bed by generating the form turbulence (Schlichting, 1968) and it may disintegrate the rising bubbles (Kim et al., 1977). These two conflicting effects may cause the maximum radial mixing of the liquid phase with liquid velocity. Kang et al. (1983) reported that the flow pattern of the solid phase would change from circulation to random motion at the optimum bed porosity when the maximum heattransfer coefficient is attained in the beds. The optimum bed porosities at the maximum radial dispersion coefficient of the liquid phase were in the range of 0.57-0.7 (Figure 9). The magnitude is similar to those of the maximum heat-transfer coefficient (h") in three-phase fluidized beds (Kang et al., 1985) as shown in Table I. This maximum liquid radial dispersion with variations in the liquid velocity (Figure 8) has indeed been observed by El-Temtamy et al. (1979b). The radial dispersion coefficient of the liquid increased with the particle size (Figures 8 and 9) except for the lower liquid velocity near the minimum fluidizing velocity. This
Ind. Eng. Chem. Process Des. Dev., Vol. 25, No. 3, 1986 721
I
'1 ,./
a~
,
$4
/a
I
1
El* €0
Figure 9. Effect of bed porosity on radial dispersion in three-phase fluidized beds. d, (mm), U,(cm/s); ( 0 )1.7,4.0; (A)4.0,4.0; (A)4.0,
Pe.1 cal
6.0, 4.0. 8.0; (0)
Table 111. Energy Consumption Rate per Unit Mass of Continuous Liquid Phase in Three-Phase Fluidized Beds
Edg, d,, mm
VI, cm/s
U,,cm/s
e,
1.7 1.7 1.7 4.0 4.0 4.0 4.0 6.0 6.0 6.0
8.0 8.0 8.0 8.0 8.0 10.0 10.0 8.0 8.0 10.0
4.0 8.0 12.0 4.0 8.0 8.0 12.0 4.0 8.0 12.0
0.075 0.096 0.135 0.067 0.082 0.127 0.189 0.111 0.186 0.223
el 0.589 0.597 0.563 0.525 0.530 0.535 0.478 0.428 0.370 0.388
e,
cm/s
0.336 0.307 0.265 0.408 0.388 0.338 0.333 0.461 0.444 0.389
29.1 36.5 43.5 35.3 45.3 46.3 60.3 44.3 64.0 77.0
may be due to the larger particles which have the effective hindrance potential to generate the form turbulence and break the bubbles in three-phase fluidized beds. The hindered fluid element and the broken small bubbles can easily redistribute in the radial direction which may lead to an increase in mixing intensity in the radial and axial directions. The intensity of mixing in the bed can be predicted from the knowledge of the eddy length and its velocity scale, since the energy associated with small-scale eddies and high velocity can promote the turbulent intensity in the flow field (Clift et al., 1978) and to enhance the phase contact between the eddy and the continuous liquid phase. The size of the eddy, Le, and the velocity of the eddy, U,, can be written as (Hinze, 1958) Le=($) 114 E3PI
in which E3 is the energy consumption rate per unit mass of the continuous liquid phase in three-phase fluidized beds. This energy consumption rate can be derived from the overall energy balance (Kang et al., 1985) as (VI + U g h g P g+ ElPl + %P,) E3 = (16) CIP1
From the knowledge of gas and liquid velocities and individual-phase holdups, the energy consumption rate, E3, can be determined. The resulting values of E3 are summarized in Table 111. As can be seen, the energy consumption rate increases with the particle size and gas velocity. From eq 14 and 15, the increase in the energy consumption rate may indicate an increase in the eddy velocity and a decrease in the eddy length in the bed. Since the microeddy with high velocity is more effective to produce turbulence and to reduce contact time with
Figure 10. Comparison between the experimental and calculated values of the Peclet numbers. d, = (0)1.7, (A)4.0, (0) 6.0, ( 0 )0.96, (A) 2.0 mm. 0, A, 0 = present study; 0 , A = El-Temtamy et al. (1979b).
other phases (Clift et al., 1978), these microeddies can increase the contacting frequencies between the fluid elements and consequently increase the continuous liquid phase mixing in the bed. Therefore, the radial dispersion coefficient of the liquid phase increases with the particle size and gas velocity in the bed. Since the liquid-phase axial dispersion coefficients were well-represented by the isotropic turbulence model (Baird and Rice, 1975; Kim and Kim, 1983) in two- and threephase fluidized beds, the present data of radial dispersion coefficients in terms of the Peclet number are also correlated by the isotropic turbulence model (Kim and Kim, 1983) as
Per =
(F)
= B.3(
;)( L)"' Ul+
ug
(17)
where the correlation coefficient is 0.93. The goodness of fit between the experimental data of the present work and the previous study of El-Temtamy et al. (197913) and calculated values of D,can be seen in Figure 10. Equation 17 covers the range of variables, 0.019 C d ID C 0.06 and 0.33 < (Ul/(Ul+ U,)) C 0.86 as shown in figure IO. Conclusions The radial dispersion coefficient of the liquid phase in liquid fluidized beds increases with the particle size. But the coefficient exhibits the maximum value at a liquid velocity. This maximum value occurred at the bed porosity at which the energy consumption rate or turbulent intensity in the bed is maximum. Introducing the gas into the liquid fluidized beds enhanced the liquid radial mixing, and the coefficient in the three-phase fluidized beds increased with the gas velocity and particle size. However, it attains its maximum value at a liquid velocity. The coefficient increase with an increase in energy consumption rate per unit mass of the continuous liquid phase. The coefficients are well-represented by the ratio of the fluid velocities and of particleto-column diameters based on the concept of the isotropic turbulence model. Acknowledgment We acknowledge a Grant-in-Aid of research from the Korea Science and Engineering Foundation. Nomenclature A : cross-sectional area of the column, cm2 C : concentration of the tracer, mol/L d : particle diameter, mm $: column diameter, cm
fnd. Eng. chem. Process Des. Dev. lQ86,25, 722-728
722
4
D, : axial dispersion coefficient, cm2 s D,: radial dispersion coefficient, cm s
ED : energy consumption rate, cm2/s g : gravity acceleration, cm/s2 h : heat-transfer coefficient, W/(m2 K) H : bed height, cm Z : turbulent intensity, cm/s Le : eddy length, cm m : ratio of the liquid phase in the wake-to-liquid phase in the fluidized region M : weight of solid, g n : Richardson and Zaki’s index AP : pressure drop in the bed, N/mz r : radial coordinate, cm R : radius of the column, cm U : fluid superficial velocity, cm/s U, : eddy velocity, cm/s Ut : terminal velocity, cm/s X : dimensionless radial coordinate 2 : axial coordinate
Greek Letters : viscosity, mPa s p : density, $cm3 : phase hol up : modified Peclet number { : dimensionless axial coordinate Subscripts 2 : two phase 3 : three phase a : axial direction f : fluidized region g : gas bubble region 1 : liquid phase max : maximum o : average value r : radial direction s : solid phase jt
w : wake region
Literature Cited Balrd, M. H. 1.; Rice, R. 0. Chem. Eng. J . 1975, 9 , 171-174. Bhatle, V. K.; Epstein, N. Roc. Int. Conf. FIuM. Its Appllcat. 1974, 380-392. Clift, R.: Grace, J. R.; Weber, M. E. “Bubbles, Drops, and Particles”; Academic Press: New York, 1978; Chapter 7. Darton, R. C.; Harrison, D. Chem. €ne. Scl. 1975, 3 0 , 581-586. El-Temtamy, S. A.; ECSharnoubi, Y. D.; ECHalwagl. M. M. Chem. Eng. J . 19798, 78. 151-159. ECTemtamy. S. A.; ECSharnoubi, Y. D.;Ei-Halwagl, M. M. Chem. Eng. J . 1979b, 78, 161-168. Henrlksen, H. K.; Ostegaard, K. Chem. Eng. J . 1974, 7 , 141-146. Hlnze, J. 0. “Turbulence”; McGraw HIii: New York, 1958. Joshi. J. 8. Trans. Inst. Chem. Eng. 1980, 58, 155-165. Joshi, J. B. Chem. Eng. Res. Des. 1983, 67, 143-161. Kang, Y.; Suh, I. S.; Klm, S. D. Proc. P A C K IIZ 1983, 2 , 1-8. Kang, Y.; Suh, I . S.; Kim, S. D. chem. Eng. Commun. 1985, 3 4 , 1-13. Kang, Y.; Lee, H. K.; Kim, S. D. Proc. APCCHE 111 1984, 7 , 159-164. Kato, Y.; Uchlda, K.; Kago, T.; Morooka. S. Powder Technol. 1981, 28, 173-179. Khang. S. J.; Schwartz, J. 0.; Buttke, R. D. AIChE Symp. Ser. 1983, 79, 47-54. Kim, S. D.; Baker, C. G. J.; Bergougnou, M. A. Can. J . Chem. Eng. 1972, 50. 895-701. Kim, S. D.; Baker, C. 0. J.; Beraouanou, M. A. Can. J . Chem. Ens. 1975, - 53, 134-139. Kim, S. D.; Baker, C. 0. J.; BergoUgnou, M. A. Chem. Eng. Scl. 1977, 32, 1299- 1306. Kim, S. D.; Kim, C. H. J . Chem. Eng. Jpn. 1983, 76,172-178. Kllnkenberg, A. A.; Krabnbrlnk, H. J.; Lauwerler, H. A. Ind. Eng. Chem. 1953, 4 5 , 1202-1208. Michelson, M. L.; Ostergaard, K. Chem. Eng. J . 1970, 1 , 37-46. Muroyama, K.; Fan, L. S. A I C M J . 1985, 37, 1-34. Richardson, J. F.; Zaki, W. N. Trans. Inst. Chem. Eng. 1954, 32, 35-53. Riquarts, H. P. Ger. Chem. Eng. 1980, 3 , 286-295. Riquarts. H. P. Ger. Chem. Eng. 1981, 4 , 18-23. Schllchtlng, H. ”Boundary Layer Theory”; McGraw-Hill: New York, 1968. Schugerl, K. Proc. Int. Symp. NuM. 1987, 782-796. Veil, Yu. K.: Manokov, N. Kh.; Manshllin, V. V. Int. Chem. Eng. 1968, 8 , 293-296.
Received for review April 22, 1985 Revised manuscript received October 28, 1985 Accepted November 18, 1985
Kinetics of Methyl Oleate Catalytic Hydrogenation with Quantitative Evaluation of Cis-Trans Isomerization Equilibrium Rlcardo J. Grau,’ A l M o E. C8ssano,t and Mlguel A. Baltanls’’ f N E C , 3000-ante
Fe, Argentha
A precise quantitative determinetion of the relative values of hydrogenation and isomerization rates of methyl oleate during catalytic hydrogenation at 398-443 K and 370-647 kPa is given. The proposed reaction model and mathematical methodology of solution also yleld accurate estimates for the cis-trans isomerization equilibrium, in good agreement with the experimental resub generally found by the practitioners in the refinlng and oil processing industry.
Partial hydrogenation of vegetable oils and fatty acids are typical heterogeneous gas-liquid-solid catalytic processes developed mainly to enhance the chemical stability
* To whom correspondence should be addressed.
t Research Assistant from CONICET. *Memberof CONICETs Scientificand TechnologicalResearch Staff and Professor at U.N.L. 11 Instituto de Desarrollo Tecnoldgicopara la Industria Quimica. Universidad Nacional del-Litoral (U.N.L.) and Consejo Nacional de Investigaciones Cientificas y TBcnicas (CONICET).
O796-4305/86/ 1125-0722$OI.50/0
and, in many cases, the melting point and texture of the reaction products. Thus, edible oil, margarines, and shortenings are obtained, as well as stable fatty acids suitable for paint formulations or other industrial applications that take advantage of their interfacial activity. Geometric and positional isomerizations occur simultaneously with a reduction in the degree of unsaturation during hydrogenation reactions. Most often, a limitation in the extent of cis-trans isomerization is as important as the control of the hydrogenation itself, given the deleterious effect of trans isomers on the physical properties of fatty compounds (Duncan, 1984; Beckman, 1983; Stingley 0 1986 American Chemical Soclety
