Ind. Eng. Chem. Res. 1991,30, 1582-1588
1582
75-15-0;17, 109-99-9;18, 110-86-1; 19,98-95-3;20,91-22-5;21, 872-50-4.
Literature Cited Angelovich, J. M.; Pastor, G. R.; Silver, H. F. Solvents Used in the Conversion of Coal. Ind. Eng. Chem. Process Des. Dev. 1970,9 (l),m. Bade, K. D. The Structure and Composition of Coal Tar and Pitch. Rev. Pure Appl. Chem. 1972,22,79. Barton, A. F. M. Handbook of Solubility Parameters and other Cohesion Parameters; CRC Press, Inc.: Boca Raton, FL, 1983. Blanks, R. F.; Prauenitz, J. M. Thermodynamics of Polymer Solubility in Polar and Nonpolar Systems. Ind. Eng. Chem. Fundam. 1964,3, 1. Borwitzky, H.;Schomburg, G. Separation and Identification of Polynuclear Aromatic Compounds in Coal Tar by Using Glass Capillary Chromatography Including Combined Gas Chromatography-Mass Spectrometry. J. Chromatogr. 1979,170,99. Chawla, B.; Davis, B. H. Effect of Temperature and Solvent on Coal Extraction under Mild Conditions. Fuel Process Technol. 1989, 23, 133. Dryden, I. G. C. Action of Solvents on Coals at Lower Temperatures II-Mechanism of Extraction of Coals by Specific Solvents and the Significance of Quantitative Measurements. Fuel 1951,30,39, 145.
Fiacher, P.; Stadelhofer, J. W.; Zander, M. Structural Investigation of Coal Tar Pitches and Coal Extracts by '8c n.m.r. Spectroscopy. Fuel 1978,57,345. Guillh, M. D.;Blanco, J.; Canga, J. S.; Blanco, C. G. Study of the Effectivenew of 27 Organic Solventa in the Extraction of Coal Tar Pitches. Energy Fuels 1991,5,188. Hansen, C. M. The Three Dimensional Solubility Parameter-Key to Paint Component Affinities. 11. Dyes, Emulsifiers, Mutual Solubility and Compatibility, and Pigments. J. Paint Technol. 1967,39,505. Hildebrand, J. H.; Wachter, J. M.; Scott, R. L. Regular and Related Solutions; Van Nostrand Reinhold: New York, 1970; Chapters 1 and 2. Riggs, D. M. Carbon Fiber from Solvent Extracted Pitch. Preprints-American Chemical Society, Division of Petroleum Chemistry (Symposia), St. Louis, MO; American Chemical Society: Washington, DC, 1984;Vol. 29, p 480. Van Krevelen, D. W. Chemical Structure and Properties of Coal. XXVIII-Coal Constitution and Solvent Extraction. Fuel 1966,44, 229. Weishauptova, Z.;Medek, J. Pitch Coke Structure and its 'Pramition into Graphite. Fuel 1985,64,999. Received for review July 3, 1990 Revised manuscript received December 5, 1990 Accepted December 24, 1990
Extraction in a Single-Stage Mixer-Settler Abu Baker S. Salem* Department of Chemical Engineering, Qatar University, Doha, P.O.Box 2713, Qatar
A new design of a singlestage mixer-settler is presented. A horizontal mixing compartment, 5.5-cm diameter and 12 cm long, leading to a similar vertical coalescence compartment reduced the settler volume to a minimum. Almost complete solute transfer occurred in the mixer for the system used, while the vertical settler, 3.5-cm diameter and 40 cm long, served merely as a flash separator. This provision may prove that only a single properly designed single stage may be enough for many extraction duties with the result of low overall investment compared to other contactors. The contactor performance was investigated by the test system toluene-acetic acid-water at 25 OC. Data analysis using a flash vaporization technique gave very good agreement with the experimental results.
Introduction Different types of equipment have been developed for liquid-liquid extraction, and the choice among them can involve many parameters. These include the stages required, flow rates, floor space available, residence times, type of solvent used, scaleup reliability, capital, and economic factors. Logsdail and Lowes (1971) have presented an excellent review on this subject. Robbins (1979) proposed a decision network based on the criterion that the least complicated contactor that will perform the extraction with low maintenance is preferred for industrial use. Mixer-settlers are usually preferred over tower extractors,with or without mechanical agitation in cases of long residence times and limited head-room space. Shaw and Long (1957) reported that mixer-settlers typically exert a high agitation intensity, Le., tip speed of 170 m/min, and employ a settler throughput of up to about 4 m3/h.m2of interface area. However, mixer-settlers are usually arranged in a horizontal pattern, therefore requiring large floor space, high inventory of liquids, independent agitation for each stage, and interstage pumping (Laddha and Degaleesan, 1976). Nevertheless, they are considered one of the most efficient *Present address: P.O. Box 5933,Helliopolis, West Cairo, Egypt.
contactors still in use for industrial applications. In the past decade, considerable effort has been expended on improving the design of mixer-settlers, and many new units have been reported in an attempt to reduce the space requirement and to increase throughput per unit volume without introducing interstage pumping. Reference can be made to the works of Treybal(1964) and others (Cheng, 1979). Horvath and Hartland (1985) introduced a mixel-settler extraction column that has the advantages of both mixer-settlers and column contactors. The column can contain a number of stages placed vertically one above the other. A typical stage contains one mixer and a settler. The stirrers of the mixers were attached to one common shaft along the axis of the column. Stage efficiencies of up to 170% have been obtained at 375 rpm agitation speed and 2.1-min residence time per stage. They reported that this was due to repeated complete coalescence and redispersion in each stage and also to the large interfacial area produced by the intensive mixing combined with low back-mixing in the column. However, Smith (1963) indicated that efficiencies greater than 100% are possible in cases where each stream leaving the stage is not of the same composition due to incomplete contact between the phases. The object of the present work was to develop a contactor that may combine the advantages of columns with
08SS-6SS6/91/2630-1582~~2.50/0 0 1991 American Chemical Society
Ind. Eng. Chem. Res., Vol. 30,No. 7,1991 1683 u
Figure 1. Flow diagram for the extraction system. Figure 2. Photograph of the system. Table I. Te&nical Data of the Developed Single-Stage Mixer-Settler no. item technical data 1 mlventtank d G , capacity 30 L 2 feedtank same as 1 3 feedpump centrifugal, 1/6 hp 4 solvent pump same as 3 5 solvent flow meter scale reading (OAW), float material St. St 316 max flow 304 mL/min, min flow 13 mL/min 6 feedflowmeter scale reading (0150), float material St. St 316 max flow 304 mL/min, min flow 13 mL/min 7 feedvalve St. St 316 8 drain valve sameas7 9 solvent valve same as 7 10 mixer two sections St. St 316 (a) vertical section, 5.5-cm diameter, 12-cm high, 285 cm3 capacity; (b) horizontal section, 5.5-cm diameter, 12 cm long, 285 cm3 capacity 11 agitator speed range: speed 1,45-300 rpm; speed 2,300-2000 rpm; regulation of speed, control knob with a scale-hollow shaft inside diameter 10.2 mm with two turbine stirrer blades, 1-cm width, and 1.5 cm long 12 settler cylindrical glasa tube, 3.5-cm diameter, and 40 cm high, 385 cm3 capacity 13 discharge valve glass, packed with Teflon 14 vent same as 13 15 valve same as 13, used for interface control 16 extractdrum glass cylinder, 3.4-cm diameter, 10-cm length 17 raffinatedrum same as 16 18 extract flow meter same as 5 19 raffinate flow meter same as 6 ~
those of mixemttlers. The proposed contactor is simply
a singlestage mixer-wttler in which the mixer has to perform essentially two functions, mixing of the phases and then d e s c i n g of the dispersed phase droplets. The settler in this case may serve as a separating vessel that can be a vertical tower with relatively small size compared to the mixer. Conventionally the settler size is usually larger than the mixer (Treybal, 1963).
Experimental Section Developed Contactor. A simplified flow diagram of the extraction system developed is shown in Figure 1. Technical data of the unit are given in Table I. Light and heavy saturated liquid phases are fed from two ground-level tanks by centrifugal pumps through flow meters to the mixer, where they are agitated. The dis-
Table 11. Physical Data for the Liauids Usedo acetic item water toluene acid 18 92.13 60.05 mol w t 1.049 density, g/mL a t 25 O C 1.00 0.866 viscosity, CPa t 25 "C 0.91 0.59 1.221 118.1 boiling point, O C 100 110.8 1.367 refractive index a t 25 O C 1.334 1.4962
toluene15% acid
0.890 0.064 1.4777
oInt.erfacial tension of the toluene-acetic acid-water system in the range of concentration studied and a t 25 O C = 8-28 dyn/cm.
persion is transferred to the column settler, where it is allowed to separate into a heavy and a light phase. The outlet phases are then collected in glass drums before transferring to receivers through two outlet flow meters. The valve leading to the heavy phase drum served as a means for controlling the interface. Samples were taken from the sample points of the drums. The mixer was made of two sections. A horizontal section that receives the feed and solvent phases and an integral vertical section of similar size that provides a suitable space for the dispersed phase drops to coalesce during their passage upward via the outlet tube to the settler. The lower section was intentionally made horizontal in order to provide it with a coil if some systems needed to be heated or cooled during extraction. The agitator was mainly a shaft provided at its end with two turbine blades. The blades were positioned in the central part of the horizontal section. A Teflon ring was fitted to the shift through the top cover of the mixer to prevent liquid leakage. Dimensions of the mixer and other items are given in Table I. The settler in most mixer-settler designs is usually a horizontal cylinder. Its main function is to separate the dispersion mixture into two hopefully pure phases. Almost as much as 30% mass transfer can take place in the settler [Rochaet al., 19861. However, in this work, the settler was designed for separating the phases only since coalescence and solute transfer was expected to be completed in the vertical compartment of the mixer. Therefore, the settler size was largely reduced in order to minimize the holdup of liquids. Figure 2 shows a photograph for the whole setup. Materials Used. The chemicals used in this study were toluene, acetic acid, and distilled water. Toluene was obtained from BDH Chemicals Ltd.,England, and acetic acid from Koch-Light Laboratories Ltd., Colnbrook, England. The distilled water was produced in our department. Some properties of the liquids used are given in Table 11.
1584 Ind. Eng. Chem. Res., Vol. 30,No. 7, 1991 Table 111. Solubility Data of the System Water-Acetic Acid-Toluene at 25 "C water-rich phase
toluene-rich phase
wt%
water 100 90.477 82.611 76.013 70.391 65.543 61.314 57.603 54.315 48.282 43.308
acetic acid 0 9.49 17.332 23.921 29.536 34.377 38.591 42.297 45.581 50.648 54.517
wt%
toluene 0 0.031 0.057 0.066 0.073 0.079 0.096 0.099 0.103 1.070 2.175
RI 1.3340 1.3410 1.3463 1.3502 1.3540 1.3572 1.3604 1.362 1.3640 1.3670 1.3690
toluene 100 89.14 80.41 73.18 67.05 61.89 53.55 44.48 39.82
acetic acid 0
10.79 19.48 26.60 32.49 37.48 45.41 53.88 57.88
water 0 0.06 0.11 0.22 0.46 0.63 1.04 1.64 2.30
RI 1.4962 1.4832 1.4730 1.4638 1.4562 1.4505 1.4387 1.4280 1.4225
Table IV. Equilibrium Data for the System at 25 "C wt, B
no. 1 2 3 4 5 6 7 8 9
water 50 50 50 50 50 50 50 50 50
toluene 50 50 50 50 50 50 50 50 50
acetic acid 10 15 20 30 40 50 60 70 80
water phase w t E, g RI 58.77 1.3450 63.29 1.3500 68.25 1.3524 77.67 1.3570 85.96 1.3615 96.36 1.3644 106.24 1.3670 115.98 1.3692 126.00 1.3780
solute total balance balance E + R, g error, % error, % 108.92 0.982 0.135 113.62 1.200 5.191 -2.162 118.92 0.900 -5.494 0.900 128.83 137.65 -4.250 1.678 -3.647 0.593 149.11 -3.308 0.700 158.87 -1.791 0.521 169.13 -1.350 0.560 179.00
toluene phase y, %
w t R, g
16.10 23.50 27.00 34.00 41.30 46.00 50.40 54.60 57.80
50.15 50.31 50.67 51.16 51.96 52.75 52.63 53.15 53.00
Equilibrium Data. The solubility data for water and toluene phases containing acetic acid were determined at 25 O C by the titration method reported by Othmer and co-workers (1948).The water-phase data were determined by titrating different mixtures of known compositions of water and the solute with toluene until the appearance of the first permanent turbidity. The toluene-phase data were determined by a similar procedure where different mixtures of known composition of toluene and the acid were titrated with water until the appearance of the first permanent turbidity. The refractive indexes of these phases were measured by an Abe refractometer. The solubility data obtained are given in Table 111. Equilibrium data of the system were determined at 25 OC. Q u a l amounts of water and toluene were stirred with a certain weight of solute for 2 h. The mixture was then allowed to settle in a separatory funnel for 24 h. The phasea were separated and weighed, and the solute content in each phase was determined by refractometry. Data obtained are given in Table IV. Material balance calculation showed that errors in these data did not exceed 6%, and good agreement with Woodman's data (1926)was found. Typical Run. The saturated feed solutions of either the water or the toluene phase were made up beforehand. The feed was toluene phase containing 15% by weight acetic acid. Pure water was the extracting solvent. All adjustments to the operating conditions were made during the early stages of the run. Feed-flow meter calibrations were checked by a measuring cylinder and a stopwatch. The agitation speed was checked by Flash-Tac stroboscope. The p h a m were fed to the mixer at the specified flow rates until it was approximately half full. At that point the agitator was started at the specified speed. When the dispersion was transferred to the settler, the interface was controlled by the valve leading to the heavy phase drum. The acetic acid concentrations in both streams leaving the contactor were monitored from time to time during the run by refractometry. Steady states have been confirmed after about 12 min from the startup point in most runs. Samples of inlet and outlet streams of both phases were
RI 1.4948 1.4937 1.4932 1.4912 1.4890 1.4872 1.4855 1.4838 1.4817
x, %
1.00 1.80 2.30 3.80 5.60 7.30 8.50 10.20 11.50
Table V. Extraction Experimental Runs 1-5 on the Contractor at Different Speeds of Agitation (a) Operating Conditions item data feed (toluene phase) 51 mL/min = 45.34 g/min flow rate F feed concn 0.15 mas8 fraction 105 mL/min = 105 g/min, pure solvent solvent (water) flow rate S phase ratio F / S 1:2.3mas8 ratio stage operating capacity 850 cms residence time T 5.45 min
(b) Results water phase RI (at speed, rpm 25 "C) y, % 0 1.3386 6.70 500 1.3388 7.00 800 1.3385 6.50 1300 1.3385 6.50 2000 1.3386 6.70
toluene phase
RI (at 25 "C) 1.4930 1.4945 1.4949 1.4952 1.4954
%
x*, %
Em, %
2.30 1.25 0.98 0.70 0.50
0.31 0.33 0.30 0.30 0.31
86.45 93.73 95.37 97.28 98.71
X,
Table VI. Extraction Experimental Run 6 on the Contractor at 2000 rpm (a) ODeratina Conditions item data feed (toluene phase) 103 mL/min = 92 g/min flow rate F feed concn 0.15maw fraction solvent (water) flow 208 mL/min = 208 g/min, pure solvent rate S phase ratio F / S 1:2.26 mass ratio stage operating capacity 850 cms 2.73 min residence time T
speed, rpm 2000
(b) Results water phase toluene phase RI (at RI (at 25 "C) y, % 25 "C) x , W x * , % Em, W 1.3387 6.80 1.4950 0.98 0.32 95.50
taken in all runs. In some runs, additional samples were withdrawn from the mixer soon after shutdown of the
Ind. Eng. Chem. Res., Vol. 30, No.7, 1991 1586 Table VIL Extraction Experimental Run 7 on the Contactor at 2000 rpm
(a) Operating Conditions item data feed (toluene p h ) flow rate F 120 mL/min = 108.68g/min feed concn 0.16 maw fraction Bolvent (water) flow rate S 40 mL/min 40 g/min, pure Bolvent p h ratio FIS k0.376 "a ratio stage operating capacity 860 cma rwidence time T 5.31 min
-
(b) Reaulta water phase
toluene phase
speed, RI (at Em, RI (at W 26OC) x , W rpm 26OC) y, 7% y*, % 2000 1.3520 26.50 29.30 90.44 1.4927 2.70
XI,
W
Em, 96
2.23
96.32
90.0
I0
1
0.5
1
1.5
2
25
I
3
Phase Ratio
system or during its operation.
Figure 3. Effect of solvent-to-feedratio on the efficiency of the unit.
Results and Discussions The results of some typical operation runs on the contactor are given in Tables V-VII. The conditions of operation are shown in these tables for the specific runs. These conditions include feed and solvent flow rates, feed concentration, phase ratio, operating capacity, and the residence time in the stage. The lower part of each table shows the results of the run. Table V shows the results of five runs. The unit was operated at zero speed of agitation and then at different speeds up to 2000 rpm. The table shows also the concentration indexes of the phases at 25 OC. The equilibrium value ( x * ) of the toluene phase corresponding to the (y) value of the extract phase was also included to facilitate calculating the Murphree stage efficiency for each run, which has been also included. The Murphree stage efficiency of the raffinate, Em, was calculated as follows:
Table VIII. Dispersed Phase Holdup in the Unita 6 = h/(H+ h) speed, rpm H,g h?g (maas fraction) 0.195 662.70 160.19 500 156.84 0.187 800 680.38 0.186 682.74 156.48 1300 0.189 677.20 157.93 2000
This table shows that the efficiency of the stage has increased by an increase of the speed of agitation of up to about 2000 rpm. The table also shows that agitation was responsible for about 13% more solute extraction efficiency for the system under consideration. However, 2000 rpm was the maximum speed of the motor used, and therefore it was not possible to explore the effect of increasing the speed above this limit. In the subsequent runs the unit was operated only at that speed. In Table VI the results of run 6 are given, where the phase ratio was approximately constant around 1:2.3 mass ratio, but the rate of flow of each phase was doubled to reduce the residence time in the stage. The results show a decrease in the efficiency from 98.7% to 95.5%. The results of run 7 are given in Table VII. The residence time was kept at about 5.3 min, but the phase ratio was different (1:0.375) mass ratio. In this case the amount of solvent used was one-sixth that of the first five runs. This was done to find out the range of phase ratios that can be handled by the unit. The raffhate efficiency of this run was 96.32% as shown in this table. The extract efficiency, Em, was calculated for this run by using the concentration of the extract phase and the result is included also in Table VII. The table shows that this efficiency was lower than the raffinate efficiency. This may be due to the fact that the liquid in each of the two phases was not of the same composition and the contact between them was not usually the same at each point of the contactor as was indicated by Smith (1963). The efficiency is plotted against the phase ratio (SIF) in Figure 3 for both runs 5 and 7. The figure shows that a phase ratio of (1:2) or more can be used without too much
"RS= 1:2:3 mass ratio.
loss of efficiency. The resulta of measuring the dispersed phase holdup for runs 2-5 are given in Table VII. The table indicates that agitation speed has little effect on the value of the dispersed phase holdup, which oscillated around a value of 0.19 mass fraction. Although the organic phase in these rune was flowing at about 45% of the aqueous-phase rate, it was observed that the contents of this organic phase did not exceed 2% of the mixer contents. In the whole stage the toluene phase was less than 20% of the stage contents as shown in Table VIII. This result shows that estimation of the dispersed-phase holdup in mixer-settlers cannot be usually considered as the phase ratio used since the value of this holdup largely depends on the unit design and system used. It may also be understood that the large amount of the aqueous phase in the mixer relative to the amount of the organic phase may be responsible for the obtained high efficiency of the unit. The presence of the large amount of the heavy phase in the mixer may be attributed to the inability of the agitator to lift this amount of liquid and transfer it to the settler. In an attempt to disperse the heavy phase, some suitable changes in the design of the mixer may have to be done. These changes may include feeding this heavy phase in the middle of the vertical compartment and providing the impeller with a lifting mechanism. Samples from the outlet of the mixer were taken, instantaneously separated, and analyzed for the acid concentration in both phases. It was realized that the effluent concentrations from the mixer were exactly the same in most cases as those of the settler. Therefore, it could be concluded that complete solute transfer between the phases of the system studied actually occurred in the mixer. The settler served only as a separating drum simii to separators in flash vaporization methods. Hence, there was no need to enlarge the volume of the settler and merely a vertical tower of suitable size can be used for the purpose of separating the phases of similar systems. The measured high rate of solute transfer in the mixer may be due to the damping effect induced on the flow from the mixing compartment by the integrated vertical coalescing compartment. This sort of provision seemed to
1586 Ind. Eng. Chem. Res., Vol. 30, No. 7, 1991 Table IX. Comparison of Some Newly Developed Contactors contractor type reference
co-current column mixer Karr (1984)
mixer
column
diameter, cm height, cm capacity, L settler capacity, L stage capacity, L no. of stages total capacity of the system, L flow rate, L/min throughput, ms/(mz/h) speed of agitation, rpm residence time in mixer, min in settler, min total extractor, min system used” direction of mass transfer phase ratio, S I F feed concn, wt % Murphree stage efficiency
2.5 91.0 0.446 44.0 19 52.446 1.8-16.4 220-2010 0-400
mixer-aettler column sievetray horizontal mixer-vertical (MIXET) extractor mixersettler settler Horvath Rocha (1986) Salem (1990) present work (1985) column column cylindrical vessel two cylindrical sections 4.57 15.2 10 5.5 12.0 11.0 75.0 300 0.55 0.175 0.36 0.3 0.85 0.475 2.72 4 1 5 10 1.9 13.6 0.85 0.6-1.8 0.05 0.156-0.311 2-6 3.05 4-8 1800-2850 200-400 0-2000
4.13-37.76 2.00
2.1
-
7.55-22.67 W-A*-T
MIBK-A-W
+
-+
X-A-W
0.0625
0.667 1 6.0 63% at zero speed; 99% at 400 speed 50% at 200 30% speed; 170% at 375 speed
3.05 6.0 36 T-A-W
-
3 15 59% at 1800 speed; 96.7% at 2550 speed
1.77-3.52 1.16-2.31 3-6 T-A-W 4
0.375-2.3 15 83% at zero speed; 99% at 2000 speed
OA = acetic acid, A* = acetone, MIBK = methyl isobutyl ketone, W = water, X = o-xylene.
prevent the impeller from inducing high turbulence which usually results in small rigid dispersed-phase droplets. The continuation of the liquids in both compartments may have also allowed for circulation of the droplets, their repeated breakage, and coalescence,with the result of high mass-transfer rates. Therefore the phases left the mixer almost at equilibrium and needed merely to be separated from one another. The high rate of mass transfer taking place in this type of mixer may be also attributed to the possibility of consecutive phase inversion from time to time or from part of the mixer to the other. The presence of the very small amount of organic phase in the mixer compared to the water phase was observed in all runs in spite of the fact that the flow rates of the toluene phase in the seventh run was 3 times larger. This cannot be confirmed unless some facilities for determining the occurrence of phase inversion in a contactor have been provided. However, Robbins (1979)recommended the use of the Sellker and Sleicher correlation for predicting which phase is dispersed and the possibility of phase inversion in a contactor. This correlation is given as
The light phase is probably dispersed if the value of 8 = 0 . 3 4 . 6 , whereas the heavy phase is probably dispersed at B = 2-3.3. Phase inversion is most probable if B = 0.5-2.0. The light phase is always dispersed when B is less than 0.3 and the heavy phase is always dispersed if it is larger than 3.3.
With this criterion for the available system, B was found to equal to 0.42 for the first five runs and 3.22 for the seventh run. This means that in the first five runs the toluene phase is probably the dispersed phase. In the last run it is not possible to confirm which phase is dispersed and whether or not phase inversion had occurred. It was generally observed that the dispersed phase droplet sizes coming out of the mixer were relatively large,
in the range 5-10 mm. As soon as these droplets travel to the settler, they merge in the interface, which was kept at the entrance in its middle section. Scaleup of the Contactor. As for the scaleup of a liquid-liquid mixing vessel, Nishikawa et al. (1987)recommended the use of the variable (4Np/4(n3d2D).This variable coincides with the average energy dissipation rate per unit mass of liquid in the mixing vessel when the vessel scale is fixed, and the dispersed phase fraction is kept constant. Eckert et al. (1985)reported that the impeller tip speed (nd)can have the controlling effect in drop formation. Nishikawa et al. (1987)added that in such a case the impeller-to-vessel diameter ratio should also be taken into account. Horvath and Hartland (1985)indicated that it is essential for scaleup of liquid contactors that the range of solute concentration be kept similar. However, since in this single-stage contactor no back mixing or entrainment can take place, the stage efficiency should be independent of the mixer diameter or the feed concentration providing that the power input per unit volume is kept constant. The vertical coalescence compartment and the settler can be scaled up by throughput and residence times. Comparison with Other Contactors. A comparison between the developed contactor with some other newly proposed types is given in Table IX. The comparison items include contactor geometries, capacity, liquids flow rates, throughput, residence times, agitation speed, system used, direction of mass transfer, phase ratio, feed concentration, and stage efficiencies. Karr (1984) used a 19-vertical-stage column for co-current mixing of the phases. The dispersion was then introduced into a settling vessel of a capacity 5.2times that of the mixer. They indicated that agitation was responsible for 3.6 transfer units relative to one transfer unit produced at zero speed. Horvath and Hartland (1985)used a five-stage vertical column. Each stage contained one mixer and a settler.
Ind. Eng. Chem. Res., Vol. 30, No. 7, 1991 1587 Table X. Results of Extraction Data Analysis by Flash Calculation distribution coeff run a=SIF xu KA IO'KT KW 1 2.3 0.023 11.68993 6.459 226 101.1175 89.513 44 2 2.3 0.0125 17.40274 5.670 106 3 2.3 0.0098 20.797 65 5.416 675 86.905 24 4 2.3 0.007 27.1335 5.209 506 84.281 51 5 2.3 0.005 35.770 76 5.056 511 82.615 42 6 2.26 0.0098 20.797 65 5.416 675 86.965 24 106.3047 7 0.375 0.027 10.639 2 6.838 44 Solvent, Ys
- I
6 = EjF exptl 2.5 2.49 2.35 2.6 2.58 2.2 0.45
calcd 2.441 902 2.442 123 2.442 273 2.442 51 2.442 737 2.402 097 0.497 858
70 error -2.5 -2.0 +3.8
-6.0 -5.3 +9.1 +10.6
Extract Phase I E, Y i I Equation 9 has to be solved to get the value of B. This can 1 be done by Newton-Raphson iteration method. The it-
Raffinate Phase (R,Xi)
Figure 4. Single-stage mixer-settler.
They reported remarkable difference in stage efficiency with the direction of mass transfer. Rocha et al. (1986)developed a sieve tray column containing 10 stages: 30% stage efficiency was obtained with this conventional column contactor. They reported little effect of the direction of mass transfer on the contactor for the four systems studied. Salem and Shierah (1990)introduced a four-stage horizontal mixer-settler. The capacity ratio of settler to mixer was 1.7. Interface control was a problem due to the existence of a narrow dispersion band in such horizontal settlers. This was the reason for intermittent back mixing and entrainment problems that reduced the contactor efficiency in many runs. In the present work most of the above-mentioned problems have been minimized or eliminated. However, other systems have to be used to test ita general applicability.
Data Analysis Referring to Figure 4, the feed F containing a solute and a diluent is contacted with a solvent S. After mixing, the phases are allowed to separate in the settler into an extract E and a raftinate R. Considering the settler to be a flash separation drum, the calculation procedure may be similar to that for an isothermal flash process in vapor-liquid systems (Ried et al., 1987). An overall mass balance on the stage gives F+S=R+E (3) A component balance gives (4) h , F + SYi,E = RXi,R + Then the raffinate and extract compositions can be estimated as follows: Xi,F + aYi,S (5) = 1 + a + B ( K i - 1)
eration procedure is continued until two consecutive values for B agree within a certain specified tolerance. Prausnitz et al. (1980) used the UNIQUAC equation for estimating the activity Coefficients of the components in both phases ya and ym These coefficients were then used in estimating the equilibrium distribution for each component. They observed that convergence by this method is very slow, particularly near the plait point. Raman (1985) recommended using experimental equilibrium values if they are available. In the present work, since the experimental equilibrium data were available, they have been used to estimate the equilibrium values Ki. The results of this calculation are given in Table X. This table shows a very good agreement between the experimental and the calculated B values. Conclusions On the basis of the present investigation, the following may be concluded: A single-stage mixer-settler may be sufficient for extraction duties of simple systems if the mixer is provided with a coalescence compartment. A vertical tower settler of a relatively smaller size than the mixer can be used with the consequence of requiring only simple means for interface control. Scaleup of such contactors may be simple since the mixer is the only equipment that needs to be scaled up. The coalescence compartment and the settler are scaled up by throughput and residence times values. Drop size in such a mixer will be in the range of oscillating and circulating drops, which allows for high masstransfer rates. The settler in this case is merely a device for separation of the phases. This leads to minimum inventory of liquids. Residence times in this single-stage contactor is relatively small compared with other types of contactors. Stage efficiency was higher than 95% in a wide range of operation with the system of liquids used. This assumes the stability of contactor operation and ita flexibility to similar extraction applications. Heating and cooling facilities can be easily fitted to this contactor, which allows for investigations of the effect of temperature and temperature profile on extraction. A requirement that is usually needed in petroleum refining applications. A laboratory size contactor of this type may be useful for evaluating the solvents used in extracting aromatics from naphtha. Such solvents are usually heavier than water.
Acknowledgment During his final year project, Ishaq Rahma has been involved in the experimental part of this work. His contribution is acknowledged.
Ind. Eng. Chem. Res. 1991,30, 1588-1593
1688
Nomenclature d = impeller diameter, mm D = vessel diameter, mm Em = Murphree stage efficiency of the extract phase Em = Murphree stage efficiency of the raffinate phase F = feed flow rate, g/min h = heavy phase dispersed phase holdup, g H = continuous phase holdup, g i = component K = distribution coefficient, y / x 1 = light phase n = impeller speed, s-l Np = power number Qh = flow rate of heavy phase, mL/min Q1= flow rate of light phase, mL/min rpm = revolutions/min R = raffinate, g (raffinate flow rate, g/min) RI = refractive index S = solvent flow rate, g/min T = residence time, min x = raffinate phase concentration, mass fraction x* = raffinate phase concentration in equilibrium with a certain extract phase concentration, mass fraction xo = feed concentration, mass fraction y = extract phase concentration, mass fraction y* = extract phase concentration in equilibrium with a certain raffinate phase concentration, mass fraction yo = solvent concentration, mass fraction a = solvent-to-feed ratio 8 = extract-to-feed ratio y = activity coefficient 8 = coefficient defined by eq 2 p = viscosity, CP p = density, g/mL 4 = dispersed phase holdup, mass fraction Literature Cited Cheng, L. T. Commercial Liquid-Liquid Extraction Equipment. In Handbook Of Separation Techniques For Chemical Engineers; Schweitzer, P. A., Ed.; McGraw-Hill: New York, 1979;Section 1.10,pp 1-293. Eckert, R. E.; Laughlin, C. M.; Roushton, J. H. Liquid-Liquid Interfacial Areas Formed By Turbine Impliers in Baffled Cylindrical Mixing Tanks. AZChE J. 1985,31,1811-1820. Hooper, W. B.; Jacobs, L. J. Decantation. In Handbook Of Sepa-
ration Techniques For Chemical Engineers: Schweitzer, P. A., Ed.; McGraw-Hill: New York, 1979; Section 1.11, pp 1-345. Horvath, M.; Hartland, S. Mass-Transfer Efficiency and Entrainment. Znd. Eng. Chem. Proc. Des. Deo. 1985,24, 1220-1225. Karr, A. E. Reciprocating Plate Column As a Cocurrent Mixer. AZChE J. 1984,30,697-699. Laddha, G. S.; Degaleesan, T. E. Extraction Equipment Classification and Selection. In Transport Phenomena in Liquid Extraction; McGraw-Hill: New York, 1976;Chapter 8; p 215. Logsdail, D. H.; Lowes, L. Industrial Contacting Equipment. In Recent Advances in Liquid-Liquid Extraction; Hanson, C., Ed.; Pergamon: Oxford, 1971;Chapter 5; p 142. Nishikawa, M.; Mori, F.; Fujieda, S.; Kayama, T. Scale-up Of Liquid-Liquid Phase Mixing Vessel. J. Chem. Eng. Jpn. 1987,20, 454-459. O t h e r , D. F.; White, R. E.; Treuger, E. Liquid-Liquid Extraction Data. Ind. Eng. Chem. 1948,33,10-33. Prausnitz, J. M.; Anderson, T. F.; Grens, E. A.; Eckert, C. A.; Hsieh, R.; O'Connell, J. P. Calculation Of Equilibrium Separations in Multicomponent Systems. In Computer Calculations For Multicomponent Vapor-Liquid and Liquid-Liquid Equilibria; Prentice-Hall: New Jersey, 1980;pp 116-129. Raman, R. Mass Transfer Operations. In Chemical Process Computation; Elsevier Applied Science Publishers: London, 1985;pp 165-167. Reid, R. C.; Prausnitz, J. M.; Poling, B. E. Fluid Phase Equilibria in Multi Component Systems. In The Properties Of Gases and Liquids, 4th ed.; McGraw-Hill, New York, 1987;pp 365-368. Robbins, L. A. Liquid-Liquid Extraction. In Handbook Of Separation Techniques For Chemical Engineers; Schweitzer, P. A,, Ed.; McGraw-Hill: New York, 1979;Section 1.9,pp 1-268. Rocha, J. A,; Humphrey, J. L.; Fair, J. R. Mass Transfer Efficiency of Sieve-Tray Extractors. Ind. Eng. Chem. Process Des. Deu. 1986,25,862-871. Salem, A. S.; Sheirah, M. A. Dynamic Behavior Of Mixer-Settlers. Can. J. Chem. Eng. 1990,68,867-875. Shaw, K. G.; Long, R. S.Solvent Extraction of Uranium. Chem. Eng. 1957,64,251-256. Smith, B.D.Binary Distillation. In Design Of Equilibrium Stage Processes; McGraw-Hill; New York, 1963;pp 152-154. Treybal, R. E. Equipment For Stage-Wise Contact. In Liquid Extraction, 2nd ed.; McGraw-Hill New York, 1963;pp 285-289. Treybal, R. E. Versatile New Liquid Extractor. Chem. Eng. Prog. 1964,60,77-82. Woodman, R. M. The System Water-Acetic acid-Toluene: Triangular Diagram at 25 "C, with Densities and Viscosities of the Layers. J. Phys. Chem. 1926,30, 1283-1286. Received for review May 8, 1990 Revised manuscript received October 23, 1990 Accepted March 5,1991
Controlled Cooling Crystallization of (NH4)#04 in the Ternary System (NH4)2SO4-NH4NOs-H20 Cheong-Song Choi* and Ik-SooKim Department of Chemical Engineering, Sogang University,
C.P.O.Box 1142, Seoul, Korea
The maximum allowable undercoolings of ammonium sulfate in the ternary system ammonium sulfateammonium nitratewater were measured. The nucleation parameters, which may be obtained from maximum allowable undercooling measurements, are shown to be dependent on the saturation temperature of the solution in the ternary system. These data were employed along with a supersaturation balance to calculate the optimum cooling curves in a batch cooling crystallizer. The results show that controlled cooling that takes into account the effect of the cooling rate on the maximum allowable undercooling improves the product crystal size distribution. Introduction Bakh-owrabd cooling crysuizers are widely used in the industry, because they me and flexible,
* Author to whom correspondence should be addressed.
require less investment, and generally involve less process development. Such systems are useful in small-scale operations, especially when working with chemical systems that are difficult to handle due perhaps to their toxic or highly viscous properties. However, batch crystallizers generally yield a poor quality nonuniform product. This
0888-5885/91/2630-1588$02.50/0Q 1991 American Chemical Society
