gas is higher a t 59.0%. Test 3 was carried out using butane a t 25-atm. pressure and ,a steam ratio of 2.0. T h e low preheat temperature of 440’ (1. gives rise to a high concentration of methane in the product gas. Calculated gas compositions are given in which carbon dioxide has been removed to 1%. T h e calorific values of the resultant gases would lie in the range 807 to 855 B.t.u. per cubic foot. A commercial plant is being built for the production of 25,000 therms per day of methane-rich gas a t 17-atm. pressure from a distillate feedstock. Production of Gas Interchangeable with Natural Gas
Although the process so far described is capable of producing a gas having a calorific value of up to 8.50 B.t.u. per cu. foot, the presence of hydrogen raises the flame speed above the level of certain natural gases. To be interchangeable with natural gas, therefore, the equilibrium needs to be established a t as low a temperature as possible. When liquefied petroleum gases are used as the feedstock, the reaction with steam may be carried out bet\veen 300’ and 400’ C. Gas of the following composition was produced when butane was gasified a t 25atm. pressure using a steam-butane ratio of 1.3 by weight: CO1 17.7.5, CO 0.05, Hz 1.8, CHb 80.45 volume After the removal of carbon dioxide this gas would contain 96,5.5% of methane and have a calorific value of 970 B.t.u. per cu. foot. With high boiling distillate, however, the temperature of the gasification stage cannot be reduced to this low level without causing deterioration of the catalyst, but a subsequent methanation stage can be used. T h e pilot plant was equipped with a n auxiliary methanation tube through which gas from the gasification catalyst was passed a t 360’ C. anld 25-atm. pressure. If carbon dioxide
had been removed, the dry gas produced would have had the following composition : C O S 1.0, CO 0.3, Hz 0.7, CH4 98.0 volume Calorific value 977 B.t.u. per cu. foot
%
When two stages are used it is convenient to condense some of the undecomposed steam as a n intermediate step in order to raise the concentration of carbon oxides in the reaction zone. T h e scale of the experiment was not large enough for the problem of heat removal from the catalyst bed during methanation to be encountered. There are a number of ways in which this may be done, but a discussion of these techniques is outside the scope of this paper. When operating a t 25-atm. pressure, 1 therm of methane can be produced from 0.98 therm of distillate a t a thermal efficiency of 95%. Acknowledgment
T h e authors thank their colleagues who were responsible for the experimental work. literature Cited (1) Atwell, H. V., Schroeder, \V. C., Technical Oil Mission Rept. 5. Office ofTechnical Services Reut. PB 2051 (Mav 15. 19451. (2) ’Cockerham, R. G., Percival, G., Trans. Znst. Gis Engrs. 107, 390-424 . . - 11957-58).
(3) Dent, F. J., MGgnard, L. A,, Eastwood, A. H., Blackburn, W.H.. Hebden, D., Zbid., 95,602-704 (1945-6). (41 Kemball. C., Roonev, J. J., Proc. Roy. . SOC.(London) A 257, 132-45 (1960); A 263,‘567-77 (1961). (5) Pitt. E. E. H., Rupprecht, \V. E., Fuel (London) 43, 417-25 (1964). (6) FVickbold, R., Angew. Chern. 69, 530-3 (1957). RECEIVED for review September 8, 1964 RESUBMITTED March 15, 1966 ACCEPTED March 22, 1966 DiLision of Fuel Chemistry, 147th Meeting, ACS, Philadelphia, Pa., April 1964. Published by permission of the Gas Council. FVork carried out under the direction of F. J. Dent.
DEVELOPMENT OF A CONTINUOUS COUNTERCURRENT FLUID-SOLIDS CO NT A CT 0 R Ion Exchange H. L. SHULMAN, G. R. YOUNGQUIST, AND J .
R. C O V E R T
Clarkson College of Technology, Potsdam, N . Y . The development and testing of new equipment for continuous countercurrent ion exchange are discussed. Mass transfer coefficients characterizing the equipment performance were obtained and correlated. Results for copper exchange with the magnesium form of Amberlite IR-120 are presented. The equipment appears to be suitable for a wide range of applications. HE recovery, concentration, or separation of liquids and Tgases is of considerable industrial importance. Methods frequently used a t present are distillation, extraction, and ion exchange for liquids, and absorption and adsorption for gases. Three modes of operation are commonly employed for ion exchange or adsorption where the separation involves contact of a fluid and solid. Where small quantities of fluid are to be
treated, the fluid may be brought into contact with the solid in a batch or stagewise system where equilibrium is achieved. For treatment of larger quantities of fluid, semicontinuous fixed-bed or continuous countercurrent devices may be used. In fixed-bed operation, fluid containing the adsorbate is passed through the adsorbent bed until “breakthrough” is achieved, “breakthrough” being defined as the appearance of the maxiVOL. 5
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JULY 1 9 6 6
257
mum allowable adsorbate concentration in the effluent stream. The adsorbent must then be replaced or regenerated before further use of the equipment. Continuous processes using fixed-bed equipment are designed around the cyclic operation of two or more beds. Continuous countercurrent devices operate a t steady state with continuous movement of both fluid and solid through the equipment. Relatively few continuous countercurrent ion exchange devices are in use a t present. The Dorr Hydrosoftener ( 2 ) is one such unit that has been used commercially. Several other units have been tested on a laboratory scale. A pulsedbed column for use with ion exchange resins was developed by Higgins and Roberts ( 2 ) . Use of a continuous resin-impregnated belt has been investigated by McCormack and Howard ( 3 ) , Selke and Muendel (9), and Mihara and Terasaki (6). Selke and Bliss (8) developed a continuous system to recover copper, in which the ion exchange resin was fluidized in the exchange columns. McNeil and coworkers ( 5 ) have evaluated the use of ore-dressing jigs for continuous ion exchange. The present paper concerns the development of equipment for continuous ion exchange. Results obtained for the exchange of copper ion with the magnesium form of Amberlite IR-120 demonstrate the effectiveness of the equipment. Performance is characterized and correlated by liquid phase mass transfer coefficients.
Experimental Equipment and Procedure
The contacting device, detailed in Figure 1, consisted of a rotating frame having two 6-inch steel disks held 18 inches apart by six steel ribs, which could be adjusted to allow variation of the frame diameter. T h e frame was supported by two sleeve bearings, each welded to a supporting plate. The frame was mounted on a tilting table and rotational speeds between 0.196 and 1.02 r.p.m. were possible. T h e contacting zone consisted of plastic tubing, 0.475-inch Ld., wrapped around the drum. T h e ends of the tubing were
Figure 1 . 258
connected to rotary joints at each end of the rotating frame to permit passage of the resin and solution to and from the contacting zone. These rotary joints were constructed from 28/12 ball and socket joints, which were placed through the support bearing a t each end and were well lubricated by stopcock grease to allow easy rotation. The resin, the magnesium form of Amberlite IR-120, was stored under water in a hopper suspended over the feeding device. The feeding device was manually operated, and consisted of two glass tees connected by a short section of glass tubing. The side arms of the tees were fixed to pass water but not resin by inserting a porous plug in the side arms. T h e top side arm was connected to the storage tank for flush water. The bottom side arm was connected to a drain. The outlet from the feeding device was connected to the joint leading to the contacting zone. All four openings to the feeding device were controlled by pinchclamps on short sections of plastic tubing. Solution was fed continuously from a constant-head system and metered with rotameters. The feed solution containing 635 p.p.m. (20 meq. per liter) of copper ion was brought to the contacting zone through the bottom rotary joint and the effluent was removed through the top rotary joint. Resin was fed to the contacting zone by manual operation of the feeding device as follows. Pinchclamps between the resin storage hopper and the feeder and the clamp on the lower side tube were removed, permitting a resin slurry from the storage hopper to fill the feeder. These clamps were then replaced and the other pinchclamps removed, permitting flush water from the reservoir to flush the batch of resin into the contacting zone. This cycle of operation was repeated so as to maintain one batch of resin per loop in the contacting zone. Rotation of the frame, which was tilted at 25' from the horizontal, caused the resin to move through the zone by screw action. The resin feed rate was determined from the capacity of the feeding device. Concentrations of the influent and effluent streams were analyzed for copper colorimetrically by standard techniques (70) using a Klett-Summerson colorimeter. Regeneration of the resin was carried out batchwise in a 2-inch by 4-fOOt column using 10% magnesium sulfate solution followed by distilled water wash. Since the data of Selke and Bliss ( 8 ) indicated the equilibrium relationship to be linear for the Cu+Z-Mg+2resin system, it was necessary to determine only one equilibrium value for analysis
Schematic diagram of equipment
I & E C PROCESS D E S I G N A N D D E V E L O P M E N T
of the d a t a obtained. T h e maximum capacity of the resin was determined by saturating a small bed of resin with solution containing 20 meq. per liter of CUSOI. T h e saturated bed was washed with distilled water to remove residual copper sulfate solution, then eluted with 19% HzSOa and washed with distilled water; the total volume was collected and analyzed for copper, from which the equilibrium copper content of the resin was calculated. Maximum capacity of the resin was observed to 4.83 meq. per gram. Selke and Bliss (8) reported a value of 4.9 meq. per gram. Results and Discussion
Performance data (obtained for a wide range of operating conditions are summarized in Table I. Removals of copper ion ranging from 19.7 to 93.2y0 were obtained from feed solution containing 635 p.p.m. (20 meq. per liter). Solution feed rates were varied from 0.222 to 11.0 liters per hour. Rotational speed of the co:ntactor, which in part controls the fluidsolid contact time, wais varied from 0.196 to 1.02 r.p.m. T h e amount of resin per loop of tubing was varied from 1.14 to 3.65 grams. T h e loop diameter used was either 0.647 or 0.917 foot and either 8.0 or 5.7 loops were employed. The angle of tilt for the contacting device was kept constant a t 25", a value giving smooth flow of the solids. Other angles might be more favorable for other ranges of operating conditions. The batches of resin moving through the contacting zone did not fill the cross section of the column; hence, channeling of the liquid stream past the resin was observed a t all flow rates. T h e maximum liquid rate possible was limited by fluidization and carryover of resin particles. Nevertheless, satisfactory mass transfer rates were observed, resulting in high removals of the exchanging ion. T h e rate of ion exchange in the equipment used here was characterized by the use of a mass transfer coefficient, kD5, obtained using the equations commonly employed for a conventional mass transfer column. For a differential section, the exchange rate is given by
-VdC
=:
kDS(C
- C*)dZ
=
Rdq
(3) AC = C
where
- C*
and
Values for kDs were calculated using Equation 3 from experimental values of concentrations, solution flow rates, and bed size. The mass transfer coefficient, kDs,exhibited dependence only on the relative velocity between the solution and the resin. Resin slug size, rotation speed, loop diameter, and column length had no apparent effect. Figure 2 shows the observed dependence on relative velocity: kDs = 0.126 Uo.87. For fixed beds, the dependence of kDS on liquid velocity is somewhat lower-for example, from 0.493 to 0.673 power as reported by McCune and Wilhelm (4)and 0.7 power as reported by Selke and Bliss (7). Gilliland and Baddour ( I ) , however, report 0.84 power dependence on velocity. T h e observed values of kDs'are about an order of magnitude lower
I
1.0
c
I
1
I
1 1 1 1
I
I
1
1
I I
I
I
IIII
.2
.I
I-
A
>. O ? l
(1)
Values for kDS may be obtained by integration of Equation 1, which gives
zs, V
kDs =
-dC (C
- C*)
For the case of linear equilibrium, Equation 2 becomes
Run No. 1 2
8 9 10 11 12 13 14 15 16 17 18 19 20
V, LiterslHr. 1.87 0.960 3.870.484 0.222 0.220 0.960 3.87 5.88 1.40 5.88 11 .o 5.88 11 . o 1.40 5.47 3.20 1.33 9.64 1.92
Loop Diam., Ft. 0.647 0.647 0,647 0.647 0.647 0,647 0,647 0.647 0.647 0.647 0.647 0.647 0.647 0.647 0,647 0.647 0.647 0.647 0.917 0.917
,021
.I
I
I
.2
I
I
.3 .4 .5
U
Figure 2.
I
I
1.0 2 (CMISECJ
3
I l l 1
7
4 5
I I l l
7
IO
Dependence of kos on relative velocity
Table 1. Summary of Experimental Results Contactor No. of Slug Size, R, G./Loop G./Hr. Z , G. Loops R.P.M. 0.449 30.7 9.12 8.0 1.14 30.7 9.12 0.449 8.0 1.14 30.7 9.12 8.0 1.14 0.449 30.7 9.12 8.0 1.14 0,449 30.7 9.12 8.0 1.14 0.449 13.4 9.12 8.0 1.14 0.196 13.4 9.12 0.196 8.0 1.14 13.4 9.12 0.196 8 .O 1.14 0.196 13.4 9.12 8.0 1.14 0.196 13.4 9.12 8.0 1.14 0.449 30.7 9.12 8.0 1.14 0.449 30.7 9.12 8.0 1.14 1.02 8.0 1.14 69.7 9.12 1.02 8.0 1.14 69.7 9.12 1.02 69.7 9.12 8.0 1.14 0.449 8.0 2.67 71.9 21.4 0.449 8.0 2.67 71.9 21.4 0.449 8.0 2.61 71.9 21.4 0.449 5.7 3.65 98.4 20.8 0.449 5.7 3.65 98.4 20.8
%
Removal 42.5 55 . O 32.0 74.6 93.2 82.7 43.2 20.8 19.7 33.9 28.2 24.7 38.6 30.2 66.9 51.4 59.4 81.4 48.9 76.3
VOL. 5
U,
Cm.lSec. 0.896 0.674 1.38 0.558 0.496 0.246 0.427 1.13 1.62 0.534 1.87 3.12 2.43 3.68 1.34 1.71 1.21 0.765 2.95 1.07
NO. 3
kDs,
LiterlG. Hr. 0.118 0.0877 0.180 0.0748 0.0668 0.443 0.0644 0.115 0.176 0.0692 0.247 0.432 0.341 0.485 0.175 0.204 0.144 0.109 0.351 0.134
JULY 1 9 6 6
259
than results obtained using fixed beds (7). T h e smaller value of k D s and the greater effect of velocity observed here are probably due to the unusual geometry of the contactor and the fact that the resin does not fill the entire cross section of the column.
kDs = mass transfer coefficient, liter/g. hr. Q = concentration of exchanging ion in resin, meq./g. dry resin R = resin rate, g. dry resin/hr. u = relative velocity between resin and liquid, cm./sec. liquid rate, liters/hr. z = bed size, g. of dry resin
v =
Conclusions
Equipment for continuous countercurrent ion exchange has been developed and tested. I t is simple in design and operation and scale-up may be possible. Mass transfer coefficients characterizing the performance of the equipment have been obtained for a wide range of operating variables. Mass transfer in the liquid phase is indicated as the rate-controlling factor for Cu+2-Mg+2 exchange with Amberlite IR-120 and mass transfer coefficients have been correlated as a function of the relative velocity between the liquid and resin. Nomenclature = = C* = C, =
a
C
ultimate exchange capacity, meq./g. dry resin concentration of exchanging ion in liquid, meq./liter equilibrium concentration, meq./liter concentration of exchanging ion in feed solution, meq./ liter
Literafure Cited
(1) Gilliland, E. R., Baddour, R. F., 2nd. Eng. Chem. 45, 330 (1953). (2) Hiester, N. K., Phillips, R. C., Chem. Eng. 61,161 (1954). (3) McCormack, R. H.,Howard, J. F., Chem. Eng. Progr. 49, 404 (1953). (4) McCune, L. K., Wilhelm, R. H., 2nd. Eng. Chem. 41, 1124 (1949). (5) McNeil, R., Swinton, E. A., Weiss, D. E., AIME Congress, New York. 1954. (6) Mihara,’S., Terasaki, Y . , Japan. Patent 2223 (1951). ( 7 ) Selke, W. A,, Bliss, H., Chem. Eng. Progr. 46, 509 (1950). (8) Ibid., 47, 529 (1951). A., Muendel, C. A., U. S. At. Enerm Energy (9) . . Selke, W. A,, .,, Comm., AEC Document NYO-962 (Feb. 15,1953). (lO).Snell, F. D., Snell, C. T., “Colorimetric Methods of Analysis,” Van Nostrand, New York, 1949. RECEIVED for review June 16, 1965 ACCEPTED March 16, 1966
DISPERSED PHASE DISTRIBUTORS IN A PACKED COUNTERCURRENT, LIQUID=LIQUID EXTRACTION COLUMN CHI-HA1 SHIH1 AND RICHARD R. KRAYBILL
Department of Chemical Engineering, University of Rochester, Rochester, N . Y. Rates of extraction were experimentally determined for distributor nozzles of l/8-, 1/i6-, l/32-, and ‘/ts-inch i.d. in a 3-inch i.d. countercurrent, liquid-liquid extraction column packed with 58 inches of ‘/s-inch Raschig rings. Similar tests were made for the unpacked column. Diethylamine was transferred from toluene phase drops to the continuous water phase. A statistical analysis of the results correlated by Colburn’s HTUequation showed that the initial dispersion had negligible effect on rate of extraction in the packed column. The rate of extraction in the unpacked column increased greatly as the size of nozzle decreased. A dependence of HTU upon flow rate of dispersed phase is explained on the basis of drop size-jet velocity relationships.
of packed liquid-liquid extraction columns Most investigators either have used only one dispersed phase distributor or have neglected to report its design. T h e generally accepted belief is that the size of the droplets of the dispersed phase is determined by the size of the packing material and not by the design of the dispersed phase distributor (2, 75, 20). However, in a spray column a marked influence of liquid distributor design on rate of mass transfer has been confirmed by many investigators (7, 9, 70, 76, 77, 20). T h e objective of this investigation was to verify the supposedly minor effect of the design of the dispersed phase distributor in a packed column and to seek an optimum nozzle diameter, if one existed, wherein highest extraction rate is FFICIENCIES
E depend upon many variables.
Present address, Elastomer Chemicals Department, Experimental Station, E. I. du Pont de Nemours & Co., Wilmington, Del. 260
l & E C PROCESS DESIGN AND DEVELOPMENT
attained. A secondary objective was to compare the mass transfer rates in the empty spray column for the distributors. T h e system diethylamine-toluene-water was chosen, inasmuch as previous investigators (72, 73,74) had shown the applicability of Colburn’s H T U correlation (4,5) for low solute concentrations. In addition, extensive equilibrium data were available (74,23). Materials
Diethylamine of 55’-56’ C. boiling point and 0.704 specific gravity was purchased from the Matheson Coleman & Bell Division, Matheson Co., Cincinnati, Ohio. Toluene of technical grade, 0.862 specific gravity, was purchased from Mallinckrodt Chemical Works, St. Louis, Mo. T a p water was used from the City of Rochester system. T h e packing material was standard nonporous and unglazed porcelain Raschig rings, 3/&ch 0.d. X 3/*-inch length X ‘/l,-inch wall thickness, manufactured by the United States Stoneware Co., Akron, Ohio.
