being completed in the waste solidification engineering hot cell a t Hanford. T h e level of radioactivity will be increased in succeeding experiments to full-level first-cycle and combined first- and second-cycle Purex waste, each with a burnup of 2 X lo4 M W D per ton of uranium. Demonstration runs with this prototype are expected to provide operating information and process data which will be used in evaluating the process for industrial application. They will also point out areas requiring further development.
Conclusions
Phosphate glasses with good quality and resistance to environmental degradation were made by controlled heating of a mixture of Hap04 and simultated Purex waste to 1200’ C. Continuous production of phosphate glass from several types of simulated Purex waste was demonstrated on a pilot plant scale. Bench-scale studies indicate that high aluminum-containing waste can be similarly processed on a continuous basis. Pilot plant materials of construction are satisfactory, although the useful life of platinum remains to be determined. Bench-scale experiments with high level Purex waste demonstrated that essentially all the fission products including ruthenium remain with the glass. Essentially all of the mercury is collected in the evaporator condensate. Indications are that melter off-gases might be decontaminated for disposal as a low level waste.
Acknowledgment We gratefully acknowledge the assistance of Roy F. Domish for off-gas studies; of Joseph J. Fedelem, Raymond Drager, and G. Champlain in developing the process; and of Me1 Rigterink of Bell Telephone Laboratories for information on phosphate glass properties. References (1) Hatch, L. P., Weth, G. G., Tuthill, E. J., “Waste Fixation in Phosphate Glass with Emphasis on the Continuous Mode of Plant Operation,” International Atomic Energy Agency Pub., Vienna, 1963. (2) Strickland, G., Tuthill, E. J., Drager, R., “Evaporator Studies with Suspended Purex Wastes,” Proceedings of Symposium on Solidification of Long-Term Storage of High Level Wastes, Richland, Wash., February 1966. (3) Tuthill, E. J., Emma, L. C., Weth, G. G., Strickland, G., Hatch, L. P., “BNL Process for Continuous Conversion of High Level Radioactive Wastes to Phosphate Glass,” Proceedings of Symposium on Solidification of Long-Term Storage of High Level Wastes, Richland, Wash., February 1966. ( 4 ) Tuthill, E. J., Weth, G. G., Abriss, A., “Studies on Ion Exchange and Glass Formation as Applied to Ultimate Waste Disposal,” U. S. At. Energy _ . Comm. Rept. TID 7613, 310 (September 1960). (5) Upson, L., Weth, G. G., “Fixation of High Level Waste in PhosDhate Glass.” Battelle Northwest Laboratorv Rebt. BNWL-iOO (November 1965). (6) Weth, G. G., Strickland, G., Tuthill, E. J., “Platinum for High Temperature Crucibles Used in Processing Radioactive Wastes,” Proceedings of Symposium on Solidification of LongTerm Storage of High Level Wastes, Richland, Wash., February 1966. RECEIVED for review July 27, 1966 ACCEPTED February 7, 1967 Work performed under auspices of the U.S. Atomic Energy Commission. BNL 10313. ,
1
SEPARATION OF HYDROCARBONS IN FIXED BEDS OF MOLECULAR SIEVES W. J . SCHUMACHER’AND ROBERT YORK School of Chemical Engineering, Cornell University, Ithaca, N . Y . The liquid-phase adsorption of n-hexane from cyclohexane, benzene, and 2,2,4-trimethylpentane was studied in a fixed bed of Type 5 8 molecular sieves at atmospheric pressure from a 5 to 20% n-hexane feed at flow rates between 4 and 60 feet per hour. The effect of flow rate, adsorbent particle size, temperature, feed concentration, and solvent system on the height of the adsorption zone and over-all mass-transfer coefficient is reported. To a lesser extent, vapor-phase adsorption was studied for the same systems. The effect of velocity on the liquid-phase mass-transfer coefficient is believed to be influenced by convection currents induced by density gradients in the column.
sieves are unique among industrial adsorbents in their ability to separate molecules on the basis of molecular size. Recent applications (2, 3, 5, 7) in separating and purifying gases and liquids have placed molecular sieve adsorption processes among the major separation operations used in some process industries. This study provides basic rate data for the separation of hydrocarbons by Linde Type 5A molecular sieves in a fixedOLECULAR
Present address, Universidad Nacional de Trujillo, Trujillo, Peru.
bed adsorber. T h e height of the adsorption zone and the over-all mass-transfer coefficient are taken as measures of the mass-transfer rate, and their values are determined from experimentally measured breakthrough curves a t a fixed set of adsorption conditions. T h e principal variables are feed concentration, flow rate, temperature, adsorbent particle size, and solvent system. T h e adsorption of n-hexane from three solvent systems, 2,2,4-trimethylpentane (TMP), cyclohexane, and benzene, corresponds to the separation of normal paraffins from isoparaffins, cycloparaffins, and aromatics. The emphasis of this study is on liquid-phase adsorption, but some VOL. 6 N O . 3
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vapor-phase runs are made to provide a comparison between liquid- and vapor-phase operation. Adsorber Design Theory
The specification of the diameter and height of a fixed-bed adsorber to accomplish a desired separation requires the prediction of either the complete breakthrough curve or the height of the adsorption zone, and also the adsorbate loading on the adsorbent within the zone. The design method followed in this work is the adsorption zone approach suggested by Michaels ( 7 7). The adsorption zone is defined herein as that portion of the bed over which the fluid-phase concentration of adsorbate changes from 5 to 95y0 of the feed concentration. This approach is based on the important assumption that the height and velocity of the adsorption zone remain constant as it moves through the bed. Glueckauf and Coates (6) point out that the height of the adsorption zone decreases as it moves through the bed if the equilibrium curve is concave to the fluid-phase concentration axis (as is the case with nhexane and Type 5A molecular sieves). However, this “selfsharpening” characteristic is counteracted by a finite rate of mass transfer and axial diffusion, so that, in actual practice, zones of nearly constant height are often observed. The height of the adsorption zone can be calculated from experimental breakthrough data by two independent equations. The first is based on a constant width, constant velocity adsorption zone, and a zone formation time which is proportional to the fractional saturation of the adsorbent within the zone a t the breakthrough point. T h e zone height, Z,, is expressed as:
where Z is the total bed height, wa is the solute-free effluent collected between the breakthrough and the exhaustion points, we is the total solute-free effluent collected u p to the exhaustion point, and f is the fractional ability of adsorbent within the adsorption zone to adsorb additional solute. The second equation for determining the zone height is based on a solute material balance within the adsorption zone. Za = Qa/APsXsf
(2)
where Qa is the quantity of solute adsorbed between the breakthrough and exhaustion points, A is the cross-sectional area of the column, p s is the bulk density of solute-free adsorbent, and X, is the equilibrium loading of the adsorbent. Agreement between Equations 1 and 2 will confirm the validity of the assumptions inherent in Equation 1. The height of the adsorption zone can also be derived from the expression for the rate of mass transfer in the adsorption zone. (3) where G, is the superficial mass flow rate of solvent, K y a is the over-all mass transfer coefficient, N t is the over-all number of transfer units, and Y , Y*, Y,, and Y e are the solvent concentrations of the fluid stream a t any point in the adsorption zone, in equilibrium with the adsorbent, a t the breakthrough point and a t the exhaustion points, respectively. Equation 3 is based upon the assumption that K y a is constant throughout the adsorption zone. The over-all mass-transfer coefficient is determined from experimental data by calculating the zone height from Equation 1 or 2 and solving Equation 3 for Kra. T h e over-all resistance to mass transfer in molecular sieves can be envisioned to be composed of at least three individual 322
I & E C PROCESS D E S I G N A N D DEVELOPMENT
resistances to diffusion: (1) external film resistance, (2) pore diffusion resistance, and (3) intracrystalline (zeolite) diffusion resistance. I t seems reasonable to expect that during the initial portion of the breakthrough curve, adsorption will take place in crystals near the surface of the pellet and that the external film resistance will be controlling. O n the other hand, during the final portion of the breakthrough curve, the adsorbent is almost saturated and a molecule may have to diffuse a distance into the pellet before being adsorbed, in which case pore diffusion or intracrystalline diffusion may control the rate of mass transfer. Thus, it is likely that the K y a calculated from Equation 3 is not a constant but represents an average value for the entire adsorption zone. Description and Function of Apparatus
A flow diagram of the experimental equipment is shown in Figure 1. The adsorption column was 2-inch NPS brass pipe packed with molecular sieves to a height of about 63l/z inches and wrapped with Nichrome wire for heating during adsorption runs above ambient temperature and during regeneration. Three sets of iron-constantan thermocouples were located midway up the column and about 6 inches from either end. The thermocouples within each set were spaced 120’ apart and extended 1, I/z, and 0 inch from the column wall. Two additional thermocouples measured the temperature of the column feed and effluent streams. During a run, temperatures were automatically recorded at intervals of 1 minute. The sample taps, located about halfway up the column and just above the top of the bed, were ‘/d-inch copper tubing pinched closed a t one end and perforated with ‘/sz-inch holes. The sample taps extended to the center of the column, so that the samples collected represented an average of any radial concentration gradients. When adsorbing in the vapor phase or a t high liquid temperatures, a coil condenser was attached to the sample tap to condense or cool the sample. The column was repacked with fresh, new sieves a t each change of hydrocarbon feed system and put through a regeneration cycle before the first adsorption run. Flow through the column was upward during adsorption and downward during regeneration. A feed drying column made of borosilicate glass, 4 feet long by 4 inches in diameter, and filled with about 16 pounds of l/s-inch, Type 4A molecular sieves removed any dissolved water from the feed to avoid co-adsorption of water and solute in the adsorber. The feed was heated or vaporized in a 14-foot, double pipe heat exchanger consisting df 3/8-inch copper tubing irkhe ‘/Z-inch brass pipe and the effluent from the adsorption column was cooled in an identical exchanger. A rotameter indicated the flow rate to the column. Nitrogen for regeneration was supplied in cylinders, metered by a rotameter, and heated in a 1-inch stainless steel pipe which first passed through a 29-inch electric tube furnace and was then wrapped with Nichrome heating wire for 47 inches. The desorbate was removed from the effluent purge stream by a water-cooled and a dry ice-acetone-cooled condenser connected in series. The molecular sieves used in the adsorption column were the standard Type 5A, l/16- or l/s-inch cylindrical pellets manufactured by the Linde Co. The cyclohexane, benzene, TMP, and n-hexane were supplied by the Phillips Petroleum Co. with a minimum purity of 99 mole %. The purge nitrogen was the “dry” grade produced by the Air Reduction Co. Experimental Procedure
The experimental procedure consisted of taking periodic samples of the column effluent during an adsorption run at fixed operating conditions. Hydrocarbon feed was circulated through the apparatus, bypassing the adsorption column, until the flow rate and feed temperature were constant a t the desired values, a t which point the flow was switched to the column. A stopwatch indicated the time from this instant,
C.W.
PRODUCT DRUMS
PRODUCT COOLER
1
1
dT I
Nichromc-wrapped line
PURGE GAS FLOWMETER
TUBULAR FUR NAGE
t
ADSORPTION
droin FEED DRYING COLUMN
DRY NITROGEN
COLUMN
n
DE SORBAT E CON DENSE RS F E E D FLOWMETER
F E E D DRUMS
cooling water
hot
dry ice- acetone
1
1
water rteom FEED HEATER C.W.
c FEED PUMP
-.
'I
i
droin
t
DES OR BAT E
t
droin Figure 1.
Experimental apparatus and flowsheet
and the time at which the hydrocarbon first reached the sample tap was noted. No attempt was made to maintain isothermal conditions, except that the bed temperature was raised to the feed temperature when necessary before an adsorption run by use of the column heaters. The bed temperature rise during adsorption ranged from 15' to 45' F. for liquid-phase runs and 115' to 180" F. for the vapor phase. Radial temperature differences between the center of the bed and the bed a t the wall were 4' to 17' F. for liquid-phase operation and 75' to 120' F. for the vapor phase. Samples of 10 to 15 ml. were taken approximately every 1/z pound of effluent and sampling was continued until the nhexane concentration in the effluent was the same as that in the feed. After an adsorption run the column was drained, blown down with nitrogen, and regenerated by passing hot nitrogen through the electrically heated bed a t rates between 0.33 and 0.44 SCF per minute. When all thermocouples registered between 600' and 670" F., regeneration continued 1 additional hour while all temperatures were maintained between these limits. Total regeneration time was about 2l/2 hours and only a negligible amount of condensate was collected during the final hour of regeneration. Analysis of the three binary systems was by refractive index with an estimated accuracy of *0.15, *0.30, and A0.60 weight 7 0 n-hexane for the benzene, cyclohexane, and T M P systems, respectively. Experimental Results
T h e experimental breakthrough curves were smooth and S-shaped and differed from an idealized breakthrough curve only in that they were slightly asymmetrical, resulting from a
slower change in the slope of the experimental curve at the trailing edge than at the leading edge. Nonisothermal adsorption conditions produced no significant distortion of the breakthrough curve. The spatial relationship between the breakthrough curve and nonconstant temperature zone in the column is given by Schumacher ( 7 4 ) . Adsorption Zone Height. Experimental zone heights calculated by Equations 1 and 2 agreed within 0.1 foot and did not differ significantly for breakthrough curves determined a t different bed depths (3.19 and 5.30 feet). This agreement indicates that the assumption of an adsorption zone of constant width and speed is justified. The variation of the zone height with superficial velocity is shown in Figures 2 to 4 for liquid-phase adsorption a t three feed concentrations. For a 57* n-hexane feed, the zone height was proportional to approximately the 0.7 power of the velocity. This agrees well with the velocity dependence of zone heights calculated from the breakthrough data of Ziegenhain (78), who studied the adsorption of n-heptane from toluene by '/ls-inch Type 5A molecular sieves. At higher feed concentrations the velocity dependence of the zone height varied from the 0.1 to the 0.8 power at low and high feed rates, respectively. For vapor-phase adsorption (Figure 5), the zone height was proportional to about the 0.64 power of velocity for all solvent systems and particle sizes. This agrees fairly well with the results of Kehat and Rosenkranz (70),whose zone heights for the vapor-phase adsorption of n-hexane from benzene by Type 5A molecular sieves were proportional to about the 0.57 power of velocity. VOL. 6
NO. 3
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323
.u
e
t-
L
2 W
ZIE GENHAIN ,A'
feet bed depth.
DAT+/
0 -
w
e-
0
m-
N
, /
(P-
L
z
CYCLOHEXANE
P -a-
(u-
-
I
I
Figure 2.
I
2
I
I
I l l l I I
4
I
1
I
l
l
8 IO 20 3040 60 80 SU PERF1CI A L V E L 0 C I T Y ( f t./hr.) 6
Effect of velocity on adsorption zone height
4.35 to 5.1 4% n-hexane in solvent shown, 80' F. feed, '/ls-inch sieves Symbol Solvent Bed Depth, Ft. Feed Concn., % 0 Cyclohexane 5.27 4.35-4.5 Cyclohexane 3.19 5.14 Q Benzene 5.30 5.0-5.1 Ziegenhain data ( 7 8); 5.25% n-heptane in.toluene for 4-ft. bed depth
-I
I 2
I I I I IIII 3 4 5 6 8 1 0
I 20
I
I I I 406000
J
SUPERFICIAL V E L O C I T Y (ff./hr.)
Figure 4.
Effect of velocity on adsorption zone height
17.1 5 to 19.1 5% n-hexane in solvent shown, 80" F. feed. '/ls-inch sieves Symbol Solvent Bed Depth, Ft. Feed Concn., % 0 Cyclohexane 5.30 17.25-1 7 . 7 6 5.30 17.15-1 7 . 8 4 0 Benzene A 2,2,4-TMP 5.30 17.59-1 9 . 1 5
CYCLOHEXANE
/
FZIEGENHAIN DATA - m P ? A
_.
IC
i e e t bed deoth
2 p o i n t s , 3.19 and 5.30 f t . bed depth
W
z
0
N
I
I
I
I
i i I I l
I
I
/
I
I
L
9-
6
Figure 3.
I
1
I I I IIII
I
I
Effect of velocity on adsorption zone height
8.62 to 10.79% n-hexane in solvent shown, 80' F. feed, '/'&-inch sieves Symbol Solvent Bed Depth, Ft. Feed Concn., % 0 Cyclohexane 5.27 8.62-9.85 Cyclohexane 3.19 9.60-9.80 Q 0 Benzene 5.30 9.30-9.77 Benzene 3.18 9.33-9.64 A 2,2,4-TMP 5.30 10.05-1 0 . 7 8 A 2,2,4-TMP 3.19 10.71 -1 0.79 Zieghenhain data (78); 1 1 % n-heptane in toluene for 4-ft. bed depth
T h e effect of molecular sieve diameter on the zone height is shown in Figures 5 and 6 for vapor and liquid-phase adsorption, respectively. O n doubling the particle size, the zone height increased roughly two- to threefold. The effect of feed temperature is illustrated in Figure 7 . The zone height was unchanged or slightly decreased by doubling the feed temperature from 80' to 160' F. The effect of feed concentration on the zone height was determined by making a crossplot of the data in Figures 2 to 4 a t a constant velocity. For all solvent systems, the relationship was linear, with the zone height increasing at an average rate of approximately 0.2 foot per weight per cent increase in feed concentration. The effect of the solvent system is shown in Figures 2, 3, and 4 for liquid-phase adsorption and in Figure 5 for the vapor phase. The liquid-phase zone heights decreased in going from benzene to cyclohexane to T M P solvent systems, with the 324
l&EC PROCESS D E S I G N A N D D E V E L O P M E N 1
benzene zone heights 30 to 100% greater than those for the T M P system. For vapor-phase adsorption the decrease was benzene to T M P to cyclohexane, with the benzene system zone heights about double those for the cyclohexane system. Over-all Mass-Transfer Coefficient. Both experimental and theoretical values for K y a are presented in Figures 8 to 11. The coefficient is based on a weight ratio driving force in the fluid phase and is expressed in units of pounds of hexane/ (hr.) (cu. ft.) (lb. hexane/lb. solvent), abbreviated to pounds of solvent/(hr.) (cu. ft.), The experimental values were calculated from Equation 3. The number of transfer units, N t , was calculated assuming irreversible adsorption as indicated by the n-hexane adsorption
f
INCH SIEVES
+l/6
2
1 3 4
6 8 IO
20
40 6 0 8 0
SUPERFICIAL VELOCITY (ft./hr.)
Figure 8. Increase of over-all mass transfer coefficient with velocity for n-hexane in cyclohexane 80° F. feed Bed Depth, Sieve Size, Ft. Inch
Symbol
5.27 3.19 5.30 3.18 5.30 5.30
0
%
's A
Feed Concn., n-Hexane 4.35- 4 . 5 0 5.14 8.62- 9 . 8 5 9.60- 9 . 8 0 8.62- 9 . 8 0 17.25-17,76
70
'/l6
'/I6 '116
l/16
'I0
lI16
7.2.4- T M P
CYCLOHEXANE I
2,2,4- T M P
A
60
80
100
120
140
160
F E E D TEMPERATURE, O F
Figure height
0
I 2
SI-
7. Effect of feed temperature on adsorption zone
1
5.30-ft. bed depth, '/winch sieves Symbol
0 0
A A
Solvent Cyclohexane Benzene 2,2,4-TMP 2,2,4-TMP
Feed Concn., % n-Hexane 8.90-9.80 9.27-9.47 10.05-10.93 17.66-18.70
I
IIIII
I
I
I 1
6 810 20 40 6 0 8 0 SUPER F IC IA L VELOCITY (f 1.) hr3
Feed Concn.,
Symbol
Bed Depth, Ft. 5.30 5.30 3.19 5.30
% n-Hexane
5.00- 5 . 1 0 9.30- 9 . 7 7 9.33- 9 . 6 4 17.15-17.84
8-
z
Q)
5
s
-
0-
K y a (theoretical)
f $
s
\
(4)
The area, a, on which the theoretical coefficient was based is the external surface area of the pellets, which was experimentally estimated by measuring the average length of a sample of pellets and calculating the surface area assuming smooth, perfectly cylindrical particles. The external film coefficient (Equation 5) was determined from j factor correlations presented by Treybal (75) for the flow of liquids and gases through packed solids.
kya = jU,apb/(N~C)a~58
L f
8OoF. feed, '/ls-inch sieves
isotherms of Barry (7) and Roberts (73). By this assumption, Y* in Equation 3 is zero, and the value of N t between the limits of 0.05 and 0.95 is 2.94 for all experimental runs. A theoretical K y a (Equation 4) was calculated assuming that the over-all resistance to mass transfer may be expressed as the sum of an external film resistance and a constant, velocityindependent, pore diffusion resistance.
+ l/kPa)-l
I 3 4
Figure 9. Increase of over-all mass transfer coefficient with velocity for n-hexane in benzene
Superflcial Velocity, Ft./Hr. 9.5-9.9 10.1-1 0 . 6 11.6-12.3 11.7-12.4
0
Kya = ( l / k y a
I
0
0 0(u
0
>
Y
0
2-
(5)
T h e solid-phase coefficient (Equation 6) was assumed to be equal to an over-all diffusion coefficient divided by an average distance for diffusion. VOL. 6
NO. 3
JULY 1 9 6 7
325
~
Ky a (theoretic al, cyclohexane, 1/16” sieves)
_________----------
-E
BENZENE
Y
0 0
a1
02
0.4 0.6
LO
2.0
4.0
SUPERFICIAL VELOCITY (ft./sec.)
Figure 1 1. Increase of over-all mass transfer coefficient with velocity for vapor phase adsorption of n-hexane Av. Bed Temp., Symbol
8 $ A 0
Solvent Cyclohexane Cyclohexane Cyclohexane Cyclohexane Benzene 2,2,4-TMP
F. 230 230 240 240 230 250
Bed Depth,
F t. 5.27 3.19 5.27 3.19 5.30 5.30
Sieve Size, Inch ‘116
’18
l/g ‘/le
‘/is
Feed Concn., % n-Hexane 7.60-9.59 9.40-9.59 9.10-9.40 9.10-9.40 9.46-9.60 9.76-10.43
Wheeler (77) gives a n equation for the over-all diffusion coefficient in a porous material and the average distance for diffusion is taken as one sixth of the pellet diameter. Liquid-phase diffusivities for the hexane-benzene system have been measured by Johnson and Babb (9). T h e diffusivities for the cyclohexane and TMP systems were calculated from the Scheibel equation as suggested by Reid and Sherwood (72). O n the basis of this simplified model, the internal resistance, with calculated values 10 to 20 times greater than the external film resistance, controls the rate of mass transfer, and the over-all theoretical Kya is nearly velocity-independent. Comparison of the theoretical and experimental coefficients should be made with caution, however, as the actual internal resistance in a molecular sieve pellet may be a resistance to pore diffusion, a resistance to transfer through the apertures of the sieve crystal, a resistance to intracrystalline diffusion, or a combination of all three. Also, this internal resistance is not likely to be constant, but probably increases as the pellet becomes more and more saturated and a solute molecule has to diffuse farther into the pellet before finding a n available adsorption site. Thus, the assumption of a constant, “film-type” internal pore diffusion resistance to make possible the calculation of a theoretical K y a represents a great simplification of the actual internal resistance. The experimental values of the liquid-phase K y a agree in order of magnitude with the theoretical values but show a strong dependence on velocity which is not theoretically predicted. At low liquid velocities especially, the velocity dependence is significantly greater than that predicted even if only the external resistance is assumed to control the rate of mass transfer. Recent work by Roberts (73) shows that the intracrystalline resistance is significant a t higher flow rates and that existing correlations underestimate the liquid diffusivity of n-hexane in cyclohexane by 40%. Taking these two factors into account brings the experimental and theoretical coefficients into closer agreement a t the higher flow rates. 326
I&EC PROCESS D E S I G N A N D DEVELOPMENT
Part of the observed velocity dependence is probably the result of convection currents set up in the column because of the different densities of the heavier, solute-free effluent a t the top of the column and the lighter feed a t the bottom of the column. Such behavior would tend to broaden the adsorption zone and reduce the experimental value of the mass-transfer coefficient. The convection effect would be more marked a t low velocities, since the relatively mild convection currents would be overcome a t higher bulk flow rates. Dryden, Strang, and Withrow (4)show a distinct difference in the j factor for upflow and downflow at low Reynolds numbers in a packed bed where 2-naphthol or benzoic acid is dissolved by water. They attribute the variation to back-diffusion and free convection resulting from concentration gradients in the void channels. For vapor-phase adsorption, the experimental coefficients are less than predicted by theory and also show a greater than expected velocity dependence. The theoretical values assume bulk diffusion in the pores of the sieve pellets, which is expected, since the distribution of pore diameters in a pellet peaks sharply a t 10,000 A. (76). However, Henry et al. (8) report that Knudsen diffusion controlled the rate of mass transfer in porous catalyst pellets despite a large fraction of void volume in macropores. If Knudsen diffusion actually did control the masstransfer rate in the pellet, the theoretical coefficients would be approximately halved, bringing them into closer agreement with the experimental values. The theoretical values would also be lowered by the existence of an intracrystalline resistance to mass transfer. Summary
Adsorption zone heights and over-all mass-transfer coefficients are calculated for the liquid- and vapor-phase adsorption of n-hexane by Type 5A molecular sieves in fixed beds. The zone height is dependent on flow rate, particle size, feed concentration, and solvent system, but nearly independent of temperature. The over-all mass-transfer coefficients agree in order of magnitude with the theoretical values but have a greater than expected velocity dependence, possibly because of convection currents and axial diffusion in the column a t low flow rates. Acknowledgment
The authors thank the Linde Co. and the Phillips Petroleum Co. for donating the molecular sieves and hydrocarbons used in this work. Nomenclature
effective mass-transfer area between fluid and solid, sq. ft./cu. ft. of tower volume A = cross-sectional area of bed, sq. ft. d p = diameter of sieve pellet, ft. D L = diffusion coefficient in liquid phase, sq. ft./hr. = fractional ability of adsorbent within the adsorption f zone to adsorb additional solute G, = mass flow rate of solute-free effluent, lb. solute/(hr.) (sq. ft.) = j-factor for mass transfer, dimensionless j k p = mass-transfer coefficient for internal diffusion, lb. solute/(hr.) (sq. ft.) (lb. solutejlb. solvent) k y = mass-transfer coefficient for external diffusion, lb. solute/(hr.) (sq. ft) (lb. solute/lb. solvent) K Y = over-all mass-transfer coefficient, lb. solute/(hr.) (sq. ft.) (lb. solutejlb. solvent) Nso = Schmidt number, dimensionless a
=
Nt = over-all number of transfer units based on weight ratio
driving force in fluid phase = quantity of solute adsorbed between breakthrough point and exhaustion point, lb. = superficial velocity of feed, ft./hr. = W e - W b , lb. = quantity of solute-free effluent a t breakthrough point, lb. W e = quantity of solute-free effluent a t exhaustion point, lb. = fraction of voids within a sieve pellet x = concentration of solute on adsorbent in equilibrium X8 with feed of concentration Yo,lb. solute/lb. adsorbent Y = concentration of fluid stream a t any point in column, lb. solute/lb. solvent = concentration of fluid stream a t breakthrough point, y b lb. solute/lb. solvent Ye = concentration of fluid stream a t exhaustion point, lb. solute/lb. solvent Y* = concentration of fluid stream in equilibrium with adsorbent a t any point in column, lb. solute/lb. solvent Z = height of packed bed, ft. za = height of adsorption zone, ft. = density of solvent, lb./cu. ft. Pb = bulk density of activated sieves, lb./cu. ft. Ps
(2) Carson, D. B., Broughton, D. B., Petrol. ReJiner 38 (4), 130-4 (1959). \ . _ _ _
(3) Cheml Eng. 70 (19), 69 (1963). (4) Dryden, C. E., Strang, D. A., Withrow, A. E., Chem. Eng. Proer. 49 (4). 191-6 (1953). (5) Furanz, Wl‘R., Christens‘en, E. R., May, J. E., Hess, H. V., Petrol. Rejner 38 (4), 125-9 (1959). (6) Glueckauf, E., Coates, J. J., J . Chem. Sac. 149, 1315 (1947). (7) Griesmer, G. J., Rhodes, H. B., Kiyonaga, K., Petrol. ReJiner 39 (6), 125-9 (1960). (8) Henry, J. P., Chermakesavan, B., Smith, J. M., A.Z.CI1.E. J . 7 ( I ) , 10 (1961). (9) Johnson, P. A., Babb, A. L., Chem. Reus. 56, 387 (1956). (10) Kehat, E., Rosenkranz, Z., IND.ENG.CHEM.PROCESS DESIGN DEVELOP. 4, 217-20 (1965). (11) Michaels, A. S., Znd. Eng. Chem. 44, 1922-30 (1952). (12) Reid, R. C., Shenvood, T. K., “Properties of Gases and Liquids,” pp. 283-99, McGraw-Hill, New York, 1958. (13) Roberts, P. V., Ph.D. thesis in chemical engineering, Cornel1 University, Ithaca, N. Y . ,1966. (14) Schumacher, W. J., Ph.D. thesis in chemical engineering, Cornell University, Ithaca, N. Y . , 1964. (15) Treybal, R. E., “Mass Transfer Operations,” pp. 54-5, 504-8, McGraw-Hill, New York, 1955. (16) Tsuruzumi, Akie, Bull. Chem. Soc. Japan 34, 1457 (1961). (17) Wheeler, A., Advan. Catalysis 3, 249-327 (1951). (18) Ziegenhain, W. C., Refining Engr. 29, C-6-C-12 (August 1957).
literature Cited (1) Barry, H. M., Chem. Eng. 67 (3), 105-20 (1960).
RECEIVED for review July 25, 1966 ACCEPTEDFebruary 2, 1967
OLEFIN CHLORINATION IN HOMOGENEOUS AQUEOUS COPPER CHLORIDE SOLUTIONS M. L. SPECTOR, HEINZ HEINEMANN, AND K. D. MILLER
M . W. Kellogg Co., Piscataway, N . J . The application of liquid phase homogeneous catalysis to the oxyhydrochlorination of ethylene is characterized by virtually perfect selectivity. This is attributed to the ease with which a uniform temperature is maintained and to the mild reaction conditions necessary to achieve economic space time yield.
react with aqueous copper halide solutions to multihalogenated paraffins. T h e reaction of ethylene with copper chloride in aqueous solution to produce 1,2-dichloroethane has been examined in detail. Simultaneous introduction of oxygen and hydrogen chloride regenerates the system. Thus, the over-all system is a homogeneous catalyst for the oxyhydrochlorination of ethylene as described below : LEFINS
0 produce
+ 2CuClz 2CuC1 + 2HC1 + ‘ / z C2H4
+
CzH4C12
0 2
2CuC1
(1)
HzO
(2)
+ HzO
(3)
+ 2 C ~ C l zf
OVER-ALL. Aq. CuCI, CuCli
CzH4
+ 2HC1 + ‘/z
0 2 + CzH4Clz
I n view of the established process for thermally decomposing 1,2-dichloroethane to produce vinyl chloride and hydrogen chloride, the oxyhydrochlorination of ethylene as described in Reaction 3 is a vital link in the production of vinyl chloride from ethylene. This investigation has formed the basis for
further development, which in turn has led to a commercial route. Experimental
Materials. ACS reagent grade salts. C.P. ethylene and oxygen. Apparatus. T h e reactor system (Figure 1) consists of a glass reaction vessel, A , suspended from the top portion of a metal flange, B, connected to an exit line, C, by means of a metal ball socket welded to the head of the flange. T h e glass reactor is encased in an electrically heated metal shell, D. The heat input, and consequently the temperature, are controlled by an automatic temperature regulator. Gaseous hydrocarbon is introduced into the metal shell through an appropriate opening, E, and enters the glass reactor by passage through a glass tube, F, which is fitted with a fritted-glass foot. This arrangement provides for pressure equilization on both sides of the glass reactor. The hydrocarbon gas comes in contact with the chlorinating solution a t the bottom of the reactor and bubbles u p through the medium; volatile products, water vapor, and unreacted hydrocarbon leave via line C, equipped with a valve, H, which serves to reduce the pressure to atmospheric. Then the exit gases are passed into a n appropriate trapping VOL. 6
NO. 3
JULY 1967
327
