Ind. Eng. Chem. Process Des. Dev.
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23 1
Gas-Liquid Contacting in Vertical Two-Phase Flow Nlgel N. Clark’ Particle Analysis Center, West Virginia UniversiQf,Morgantown West Virginia 26506
Rory L. Flemmer Department of Chemical Engineering, University of Natal, Durban, South Africa
Deep shaft reactors (DSR’s) employ vertical two-phase bubble flow to effect mass transfer between gases and liquids. Mass transfer coefficients in these devices are high, and the large pressures at the base of the reactor serve to promote gas dissolution. DSR design depends on both hydrodynamic and mass-transfer aspects of twwhase flow theory. Experimental work on a 50-mm diameter, 17 m high pilot plant apparatus has demonstrated that DSR’s operate in plug flow. By use of a plug-flow model the mass-transfer rate was extracted from data gained in the apparatus by the absorption of oxygen into catalyzed sodium sulfite solution. Mass-transfer rates were found to be an order of magnitude higher than those found in bubble columns and stirred tanks, thus proving that the DSR is suitable as a gas-liquid contacting device.
Introduction Mass transfer between gas and liquid phases is a capital and energy-intensive unit operation in the chemical industries. Gas-liquid contacting may be achieved in bubble columns or in autoclaves with mechanical agitators, but the compression of gas to the reactor pressure leads to high energy consumption. Moreover, mass transfer rates in industrial scale reactors are generally low, so that the reactor vessel must be large and costly to cope with the throughput. An alternative scheme, the deep shaft reactor (DSR), is suggested for mass-transfer applications. Large diameter (up to 3 m) DSRs were developed by IC1 for sewage treatment (Hines et al., 1975) and have some advantages over conventional aeration equipment (Cox et al., 1980). It is envisaged that narrower bore DSRs will have wide application in the chemical industry for gas-liquid mass transfer (Clark and Flemmer, 1983) and for three-phase mass-transfer applications, such as uranium ore leaching (Robinson et al., 1958). The DSR consists of two vertical conduits, either concentric or in parallel, joined at the base. Liquid flows down one conduit, termed the downcomer, and returns via the other, the riser, to a header tank. Gas is introduced near the top of the downcomer and is carried around the DSR by the liquid, to be discharged in the header tank. Large-bore DSRs may operate in natural circulation by an air-lift effect, provided the gas is introduced at a great enough depth in the downcomer (Kubota et al., 19781, but in a smaller bore devices, where frictional losses are high, the liquid must be circulated with a pump (see Figure 1). Bubble Dispersion in Vertical Flow The designer of a DSR must recongize the pattern in which the gas and liquid are interdispersed. Various “flow patterns” exist in vertical gas-liquid flow, such as the bubble flow pattern, where the gas is dispersed as small bubbles in a liquid continuum, the slug flow pattern, where alternate gas and liquid slugs occupy the whole pipe diameter, and annular flow, where a gas core flows within a liquid annulus (Govier and Aziz, 1972). The flow pattern depends on gas and liquid superficial velocities, pipe geometry, and fluid properties such as viscosity, surface tension, and density. ai~~-43asia511 i24-0231$01.soia
The bubble flow pattern is most desirable because it provides a large interfacial area for mass transfer. Bubble coalescence must be discouraged since in narrow-bore pipes this may lead to slug flow, with loss of gas-liquid interfacial area and the posibility of cyclic hammering in the pipe fittings. Little coalescence occurs at low gas voidages, and for air-water systems the onset of slug flow may be given in terms of a minimum air voidage. Flow pattern maps are available in the literature to describe the limits of the bubble flow pattern in upflow (Serizawa et al., 1975; Griffith and Wallis, 1961; Oshinowo and Charles, 1974; Spedding and Nguyen, 1980), but many of the maps disagree on the locus of the bubble-slug transition boundary, as may be seen from Figure 2. This disagreement may be attributed to differences in the methods of introducing gas into the liquid stream (Iida, 1980). Although fewer maps are available to predict the bubble to slug flow transition boundary in downflow (Oshinowo and Charles, 1974; Spedding and Nguyen, 1980; Barnea et al., 1982), a similar disagreement exists there. Although bubble coalescence is generally reduced when solute is present in the liquid (Keitel and Onken, 1982), Govier and Azia (1972) have argued that the positioning of the flow regime boundaries is not substantially affected by the properties of the liquid and gas. Govier and Aziz (1972) have proposed that air-water regime maps should be generalized to describe other gas-liquid systems by modifying the water superficial velocity by a factor
and by modifying the air superficial velocity by the factor
where the subscript w denotes water, a denotes air, and subscripts g and 1 refer to the properties of an alternate gas and liquid. For purposes of DSR design, the lowest prediction of voidage at which slugs will form must be taken as the maximum gas voidage permissible in the reactor. In addition, the method of gas introduction must not encourage bubble coalescence. In the deep shaft reactor, gas is added in the downcomer, so that a gas sparger suitable for gas 0 1985 American
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Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 2, 1985
L 0
1
2
3
TOTAL SUPERFICIAL VELOCITY, Wg+W,
4
(mise=
)
Figure 3. Drift-flux plot of the low voidage data of Borishanskiy et al. for the case of a vertical ll-mm diameter pipe. Figure 1. Deep shaft reactor with forced circulation.
s
9
- SLUG FLCW PATTERN BOUNDARIES OF : -
BUBBLE
20
b
wg/t = C,(W, + W,) + u,,
A : SPEDDING 6 NGLNEN (1980)
B: GRIPPITH 6 WALLIS (1961) C: OSHINOWO 6 CHARLES (1974)
g
EE
B
e
(1)
where t is the gas phase voidage, W, and W, are the liquid and gas superficial velocities, Up is a drift velocity accounting for the local slip, and C, is the constant accounting for interaction of the profiles. Values of C, from 1.1to 1.3 have been proposed in the literature for bubble flow (Nassos and Bankoff, 1967; Govier and Aziz, 1972; Clark and Flemmer, 1984a,b), and U,, is given by the formula of Harmathy (1960)
4
E
differs from the volume fraction of gas flowing for two distinct reasons. First, the difference in gas and liquid density causes the gas bubbles to rise within the fluid, so that there is a local slip velocity between the two phases. Secondly, since both mixture velocity and bubble population profiies (distributions) across the pipe diameter have a maximum at the pipe center, the gas phase is more concentrated in the faster region of flow, thus reducing the average gas void fraction present in the pipe. The driftflux model is presented as the equation
1.0
a
z
2
0 lW0
4
PHASE PROUDE MMBER
6
a
= 0.25 m/s for an air-water system 10
Wwg+W1)/(gD)ii
Figure 2. Disagreement over the positioning of the bubble to slug flow transition boundary in vertical gas-liquid flow.
introduction in downflow must be designed. In the authors’ experience introducing the air through 0.5-1-mm holes drilled in a long (about 10 pipe diameters) thin tube placed along the pipe axis produces a satisfactory bubble flow. Introducing the gas over a short length of pipe (less than 3 diameters) may lead to a “falling film”flow pattern, described by Oshinowo and Charles (1974). Gas Voidage i n Vertical Flow The gas voidage in the pipe must be predicted in order to evaluate the hydrostatic head and determine the surface area available for mass transfer. Although many authors have proposed models to predict gas voidage present in the pipe, Zuber and Findlay’s (1964) drift-flux model has gained the widest acceptance and has been used to describe air-water, steam-water, and boiling organic liquid systems (Ardron and Hall, 1980; Zuber et al., 1967). Zuber and Findlay (1964) argue that the voidage of gas in the pipe
The drift-flux model is represented as a plot of average gas velocity, Wg/t, vs. total superficial velocity, W, + WI, so that a lilie with a slope of Co and intercept U,, is generated. Two such plots are presented here in support of the drift-flux model. The raw data of Borishanskiy et al. (1977) for air-water upflow in an ll-mm pipe have been plotted by the authors in Figure 3. The data agree closely with a line for Co = 1.187 and U,, = 0.25 m/s. The Zuber and Findlay model also may be applied to bubble downflow. In the case of downflow, the slip velocity, U p , will change in sign, because the bubble slip due to gravity opposes the direction of flow. A drift-flux plot showing results of the authors (Clark and Flemmer, 1984a) for upward and downward air-water bubble flow in a 50mm pipe is given in Figure 4. Close agreement between these data and the best fit lines (C, = 1.16, U , = 0.25 and -0.25 m/s) serves to vindicate eq 1. In designing a DSR the values of Co = 1.2 and U,, = 0.25 m/s should prove sufficiently accurate for the prediction of gas voidage in an air-water system. In other gas-liquid systems, the value of U,, should be predicted by the equation of Harmathy (1960), eq 2.
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 2, 1985 233
predict. Thus, for an air-water flow with a velocity of 2 m/s in a DSR, the turning losses will be approximately 3 kPa, a pressure loss small in comparison to with the total pressure gradient in a deep reactor. However, in the design of a short DSR, or in the design of natural circulation reactors, these losses may be considered. Pressure losses in concentric riser and downcomer DSRs have been discussed by Kubota et al. (1978), and are generally smaller than eq 5 would predict. In addition to the turning losses, there are losses associated with the introduction of gas into the liquid stream, due to the acceleration of the liquid phase across the sparger. From a momentum balance, the pressure drop at the sparger is given by f l s = PlWl[W1/(1 - 4 - Wll (6)
-
0
y
1 2 3 TOTAL SUPERFICIAL VELOCITY, W g + W ~ (m/sec.)
Figure 4. Drift-flux plot of both upward and downward bubble flow data. The profile constant, C,, is the same in the two configurations.
Pressure Loss in Bubble Flow In the DSR, three separate effects contribute to the overall pressure gradient, namely the hydrostatic head, frictional pressure loss, and acceleration effects. In twophase bubble flow, acceleration effects constitute less than 1% of the total pressure gradient and are generally neglected (Hughmark and Presburg, 1961; Oshinowo and Charles, 1974), so that the pressure gradient is the s u m of the hydrostatic head and frictional loss dP/dx = p,g(l - t) + Ftp (3) where P is the pressure, x is the reactor length, PI is the liquid density, and Ftpis the two-phase flow frictional pressure loss, and where gas density is neglected. The frictional pressure loss, Ftp,may be found from a graphical correlation, such as that of Lockhart and Martinelli (1949) or from a formula such as
Ftp= Fsp(l- ty
(4)
where Fspis the irreversible pressure loss which would occur if only the liquid were flowing in the pipe; Fspmay be evaluated from a friction factor diagram and the D'Arcy equation (Govier and Aziz, 1972). Values of n in eq 3 vary from 1.8 to 2.4 (Butterworth and Hewitt, 1977; Hsu and Graham, 1976; Orkizewski, 1967), but for the low values of gas voidage found in DSRs small variations in n are not significant, and a value of 2 will be sufficiently accurate. Besides the overall pressure gradient in the riser and downcomer, there are pressure losses associated with pipe fittings, the turning of the flow at the base of the reactor, and the introduction of the gas phase into the liquid flow. However, these pressure loss terms are small with respect to the total pressure gradient in the reactor. Turning losses in a 180° bend in single-phase liquid flow are given by = 1.5plw:/2 (5) and in two-phase bubble flow, where gas voidages are low, losses will not be significantly larger than eq 5 would hpt
where t is the void fraction of gas in the flow immediately after the sparger. In a 2 m/s velocity flow, with a gas void fraction of 0.2, this loss would be as small as 1 kPa. Residence Times Keitel and Onken (1981) have indicated that knowledge of the liquid phase dispersion is important in modeling gas-liquid contacting devices. Residence times of the liquid in two-phase bubble flow were tested on a 50-mm diameter, 17 m tall DSR apparatus. A salt solution was injected into the flow 2 m above the gas sparger, and the conductivity was monitored at the top of the DSR riser, with a conductivity probe connected to a flatbed recorder. Tests with only liquid flowing showed the apparatus to approximate plug flow, and the addition of gas up to voidages of 15% caused no significant change in the residence time distribution. A plug-flow model would accordingly provide a good description of residence time in a bubble flow DSR. Mass Transfer The interfacial area available for mass transfer is expressed as a, the surface area per unit volume. This area may be calculated from the formula u = 6~/db (7) where db is the Sauter diameter, or area-volume mean bubble diameter, which is calculated for a population of bubbles according to the formula
db =Cd3/Cd2 (8) where the bubbles in the population have diameters d. Little quantitative information on mass transfer in two-phase bubble flow is available in the literature, although mass-transfer rates in gas-liquid flow are known to be high (Hsu and Dudukovic, 1980). Several correlations are available for the annular flow pattern (Wales, 1966; Jagota et al., 1973; Shilimkan and Stepanek, 1977), but these are not readily extended to the bubble flow pattern. Scott and Hayduk (1966) and Shah and Sharma (1975) investigated bubble flow in horizontal pipes of less than 25-mm diameter and presented a correlation for kla,the product of liquid side mass transfer coefficient and interfacial area, in terms of flow rates, pipe diameter, and fluid properties. This correlation may not be extended to predict kl alone for larger pipes, as it does not account for variation in bubble diameter. Lamont and Scott (1966) used data gained in an 8-mm pipe to obtain the dimensionally inconsistent equation k1 = 5 x 10-'jRe0.49m/s (9) However, Lamont and Scott's work was conducted by use of a stratified stream of bubbles flowing along a horizontal tube, so that 9 may not be applied directly to DSR design
234
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 2, 1985
without verification. Further mass-transfer data for vertical bubble flow are presented in the experimental section below. Overall Model Provided that a mass-transfer coefficient and a representative bubble radius are known, the theory presented above permits the design of a narrow-bore DSR using a numerical solution based on a plug-flow model. A computer program is used to iterate around the DSR in discrete intervals of length to predict total mass transfer in the DSR. A plug-flow model of this type is justified by the residence time tests conducted above. An initial estimate of the pressure drop around the DSR permits the calculation of an approximate value for P1, the mean pressure in the first incremental length. Given the mass flow of gas to the DSR, the gas superficial velocity in the first increment can be calculated by using the assumed pressure, P1. Hence, the gas void fraction, surface area, and mass transfer rate in the first increment are computed by using the theory developed above. The pressure for the second increment is then calculated from the hydrostatic head and frictional loss terms. By adjusting the bubble diameter, and the gas superficial velocity, to account for the change in pressure and loss of gas due to dissolution over the first increment, the same technique is applied to the second increment, and so on, until the whole DSR length has been treated in this fashion, and the overall mass transfer rate is predicted. If the final pressure at the top of the riser predicted by the program is significantly different from that required in the design, then the initial pressure, PI,must be adjusted by trial and error, and the mass transfer rate be recalculated. The computer plug-flow model provides an accurate means for the design of a DSR. Although a closed form equation for the hydrodynamic design of vertical air lift pumps is available (Clark, 1984), it is not readily extended to account for the loss of gas due to mass transfer, and it cannot be used to predict the overall performance of a DSR. Use of the Plug-Flow Model The plug-flow model provides for the rapid design of a DSR by use of the hydrodynamic and mass-transfer theory developed above. However, no data have been found in the literature to predict the mass-transfer coefficient in vertical gas-liquid bubble flow. To obtain the mass transfer data essential for DSR design, the plug-flow model was used to interpret data for the transfer of oxygen from air bubbles into a catalyzed sulfite solution, in a 50-mm diameter 17 m tall DSR apparatus, which had been used previously in holdup studies (Clark and Flemmer, 1984a). Mass transfer was measured by monitoring the dissolution of oxygen in a solution of sodium sulfite, a method recommended by Danckwerts (1970). Oxygen absorbed from the bubbles into the liquid is consumed by the sulfite to form sulfate, so that extended experimental runs are possible. Physical absorption of oxygen was not suitable for characterization of the mass transfer rate, because the liquid phase would have become saturated with oxygen over a small fraction of the total reactor length. Use of the oxygen-sulfite absorption method relies on adjusting the reaction rate so that it is sufficiently high to hold the dissolved oxygen in the liquid bulk at a low concentration, yet not so high that it significantly enhances mass transfer by promoting reaction in the film at the gas-liquid interface (Danckwerts, 1970; Linek and Vacek, 1981). To adjust the reaction rate to suit these criteria, a 0.8 M sulfite solution was catalyzed with M cobalt and adjusted to a pH of 7.5. All copper or brass fittings
were removed from the apparatus and replaced with stainless steel components,since dissolved copper is known to affect the cobalt catalysis. However, results may still be in slight error due to the presence of iron dissolved from the pipes and due to the fact that commercial grade sulfite was used. Linek and Vacek (1981) have reported up to 30% error in the use of commercial grade sulfite solutions. Nevertheless, the scale of the experiment prohibited the use of reagent quality chemicals and stainless steel piping for economic reasons. Experimental runs were carried out at a temperature of 33 "C, which was maintained with cooling coils to remove the heat of reaction and the frictional heat generated by the flow. The pH was corrected to 7.5 throughout the run. Since the reaction rate in the sulfite system is unaffected by the sulfite concentration over the range 0.4 to 0.8 M sulfite, an experimental run was stopped only when the sulfite concentration dropped below 0.5 M. For each data point, the flow rate of air and the flow rate of the circulating solution was set, and the molarity of the sulfite solution was tested by adding a sample of sulfite solution to a solution of iodine and back-titrating with sodium thiosulfate (Vogel, 1961). After 30 to 60 min of operation a second sample was taken, and the change in sulfite molarity was noted. From this change in molarity, and the total volume of the circulating sulfite solution, an overall rate of sulfite consumption was calculated. From stochiometry, this implied a corresponding oxygen consumption rate in the DSR apparatus. A plug-flow computer program, as described above, was written in Basic language and used to interpret the mass transfer results. A value of mass-transfer coefficient and bubble diameter was assumed, and the plug-flow model was used to predict overall oxygen consumption, for comparison with the experimental value. By trial and error, the program was used to extract the correct value of the product of the mass-transfer coefficient and average specific surface area, klu. It was not possible to extract k p in a more simple fashion from the raw data, because a third to two thirds of the oxygen in the air bubbles was consumed over the reactor length, and the gradient of oxygen partial pressure around the DSR was not simply described. Results for kla were largely independent of liquid flow rate (see Figure 5) because of two opposing influences on the overall mass-transfer rate. Although the mass-transfer coefficient increases with increasing mixture velocity, due to the increasing turbulence, the residence time in the reactor is reduced. An alternative way of viewing the decrease in residence time is to note that a higher liquid flow rate will reduce the void fraction of gas present in the flow, even if the gas flow rate is held constant, thus reducing the interfacial area available for mass transfer. Results were also gained to determine the mass-transfer coefficient, kl, alone, using the plug flow model. For these later results, the overall oxygen consumption in the apparatus was found by measuring the concentration of oxygen in the gas leaving the top of the riser, a rapid technique used on mass-transfer apparatus by Danckwerts and Rizvi (1971). Sauter mean bubble diameter was determined from photographs of the two-phase flow 2 m below the sparger. The use of low gas flow rates, and the fact that the solution contained a high sulfite concentration, supported the assumption that there was no bubble coalescence over the length of the apparatus. The plug-flow program was used to extract the mass-transfer coefficient, kl, for each run. A regression on the points yielded the following correlation kl = 1.205 X 10-6Re0.628 m/s
(10)
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 2, 1985 235
0
0.5
0.4
0
d
d
7
a
0.3
01 u v
;-
E
E
d
0.2
LIQUID SUPERFICIAL VELOCITY (m/sec.) 1.12 o 1.50
m
."
2 0.1
A 1.81
0
0
0
2.23 3.01
0
50
100
AIR FLOWRATE TO D . S . R . (litres/min. at S . T . P . )
Figure 5. Mass transfer data for a 50-mm diameter deep shaft reactor, for absorption of oxygen into catalyzed sulfite solution.
Values of the mass transfer coefficient, klwere on average 20% higher than would be predicted by the equation of Lamont and Scott (1966), eq 9, over the range of flow rates used, despite the considerable differences between Lamont and Scott's apparatus and the DSR used in this study. Nevertheless, the effect of pipe diameter has not been assessed in either study, so that scaleup of reactors will require further mass-transfer investigation. Experimental work has demonstrated that mass-transfer rates in the DSR apparatus were high. The rates shown in Figure 5 were an order of magnitude higher than those found in bubble columns (Mangartz and Pilhofer, 1981),
thus supporting the use of DSRs for gas-liquid contacting in the chemical industry. Numerical Example of Plug-Flow Model The plug-flow model is best demonstrated by using a numerical solution to predict the variation of pressure, void fraction, bubble diameter, and gas composition over the length of a DSR. A simple DSR system was chosen, and the plug-flow computer model was used to predict the gradients of these variables for the following operating conditions: reactor diameter, 50 mm i.d.; reactor length, 40 m (riser and downcomer 20 m each, neglecting the turning distance); sparger position, at top of downcomer; gas inlet composition, 20% oxygen; reaction system, sulfite oxidation, assume no oxygen dissolved in liquid bulk; bubble diameter, 2 mm at atmospheric pressure and 20 "C, at sparger; assume no coalescence; reactor temperature, 20 O C ; liquid superficial velocity, 2 m/s; free gas superficial felocity, 0.5 m/s; pressure at sparger, 150 kPa abs. The plug-flow model solution is shown in Figure 6. A discontinuity exists in the gas void fraction plot at the base of the reactor (20 m length), because the bubble rise velocity changes direction abruptly at this point and causes a change in the solution to the Zuber and Findlay (1964) drift-flux equation. The first derivatives of pressure and bubble diameter are also discontinuous at this point, due to the reversal of the hydrostatic head. Figure 6 serves to illustrate the use of a plug-flow numerical solution in DSR design. Solution is rapid and accurate and can be readily extended to include more complex reaction systems and reactor geometries. Conclusion Hydrodynamic theory and a plug-flow model have been presented for the description of mass transfer in deep shaft reactors. Mass transfer rates and coefficients, determined with this plug flow model, are an order of magnitude higher than those found in bubble columns and stirred tanks, thus supporting the use of DSRs for gas-liquid contacting. DSRs are superior to bubble columns and stirred tanks by virtue of their excellent mass transfer properties and simplicity of construction. Existing theory presented above is sufficient for the design of a forced circulation DSR. Future work would concentrate on assessing the DSR for m
"
J I
a
; 3
m
-
-1.5
-1.0
Figure 6. Predicted gradients of pressure, bubble size, gas holdup, and gas bubble composition for the case of oxygen-sulfite absorption in a 20 m high, 50 mm diameter deep shaft reactor.
Ind. Eng. Chem. Process Des. Dev. 1985,2 4 , 236-238
236
use in three-phase systems and on the development of a comprehensive model to predict mass transfer in vertical bubble flow.
BorWMnskiy, V. V.; Andreyevskly, A. A,; Bykov, G. S.; Zaletnev, A. F., Fokin, B. S.; Voiukhova, T. G. FluklMech.-SovletRes. 1977, 6 , 51. Butterworth, D.; Hewtlt, G. F. "Two Phase Flow and Heat Transfer"; Oxford university Press; Oxford, 1977. Clark, N. N. "Air Lift Pumps for the Hydraulic Transport of Solids"; Powder Advisory Centre: London, 1984. Clark, N. N.; Fiemmer, R. L. C. CHEMSA 1983, 9 , 62. Clark, N. N.; Flemmer, R. L. C. Chem. Eng. Sci. 1984a, 3 9 , 170. Clark, N. N.; Fiemmer, R. L. C. AIChE J . 1964b, In press. Cox, G. C.; Lewin, V. H.; West, J. T.; Brlgnal, W. J.; Redhead, D. L.; Roberts, J. G.; Shah, N. K.; Waller, C. B. Water Pollut. Control 1960, 79, 70. Danckwerts, P. V. "Gas-Liquid Reactions"; McGraw-HIII; New York, 1970. Danckwerts, P. V.; Rlzvl, S. F. Trans. Inst. Chem. Eng. 1971, 49, 124. Govter, G. W.; Azlz, K. "The Flow 01 Complex Mixtures in Pipes"; Van Nostrand-Reinhoid; New York, 1972. Grifflth, P.; Wallis, G. B. J . Heat Transfer 1981, 3 5 , 58. Harmathy, T. 2. AIChE J . 1960, 6 , 281. Hines, D. A.; Bailey, M.; Ousby, J. C.; Roesler, F. C. 1.Ch.E. Conference, York, England, April 1975. Hsu, Y.; Dudukovic, M. P. I n "MuRlphase Transport"; Veziroglu, T. N., Ed.; Hemisphere: Washington, 1980. Hsu, Y.; Graham, R. W. "Transport Processes in Boiling and Two Phase Systems"; McGraw-Hill; New York, 1976. Hughmark, G. A.; Pressburg, B. S.AIChE J . 1961, 7 , 677. Iida, Y. Bull. Soc. Mech. Eng. Jpn. 1980, 3 , 247. Jagota, A. K.; Rhodes, E.; Scott, D. S. Chem. Eng. J . (Lausanne) 1973, 5 , 23. Keitel, G.; Onken, V. Chem. Eng. Sci. 1961, 3 6 , 1927. Keitel, G.; Onken, V. Chem. Eng. Sci. 1982, 37, 1635. Kubota, H.; Honsono, Y.; Fujle, K. J . Chem. Eng. Jpn. 1978, 7 7 , 319. Lamont, J. C.; Scott, D. S. Can. J . Chem. Eng. 1966, 4 4 , 201. Linek, V.; Vacek, V. Chem. Eng. Sci. 1981, 36, 1747. Lockhart, R. W.; Martinelll, R. C. Chem. Eng. frog. 1949, 45, 39. Mangartz, K. H.; Pilhofer, T. H. Cbem. Eng. Sci. 1961, 36, 1069. Nassos, G. P.; Bankoff, S. G. Chem. Eng. Sci. 1967, 2 2 , 661. Orklzewski, J. J . Per. Technoi. 1967, 19, 829. Oshinowo, T.; Charles, M. E. Can. J . Cbem. Eng. 1974, 5 2 , 25. Robinson, R. E.; James, G. S.; Van Zyl, P. C. N.; Marsden, D. D.; Bosman. D. J. I n t . Conf. Peaceful Uses At. Energy, 1955. Scott, D. S.; Hayduk, W. Can. J . Chem. Eng. 1966, 44, 130. Serizawa, A,; Kataoka, I.; Michiyoshi, I. Int. J . Multiphase Flow 1975, 2, 235. Shah, A. K.; Sharma, M. M. Can. J . Chem. Eng. 1975, 5 3 , 572. Shlllmkan, R. V.; Stepanek, J. B. Chem. Eng. Sci. 1977, 3 2 , 1397. Spedding, P. L.; Nguyen. van T. Chem. Eng. Sci. 1980, 3 5 , 779. Wales, C. E. A I C M J . 1968. 72, 1166. Zuber, N.; Findlay, J. A. General Electric Co. Report GEAP-4592 (1964): for less detailed account, see J . Heat Transfer 1965, 8 7 , 453. Zuber, N.; Staub, F. W.; Bijwaard, G.; Kroeger, P. G. General Electric Co. Report GEAP 5417 (1967).
Acknowledgment The authors are grateful to David Wright of Natal University for his invaluable assistance in determining residence times in the DSR.
Nomenclature a = interfacial area per unit volume, m2/m3 Co = profile constant in drift-flux model
d = diameter of bubble, m d b = Sauter mean bubble diameter, m F = frictional losses, Pa/m g = acceleration due to gravity, m/sz
k , = liquid side mass transfer coefficient, m/s n = exponent in pressure loss correlation P = pressure, Pa R e = Reynolds number U,, = bubble drift velocity, m/s V = incremental volume, m3 W = superficial velocity, m/s x = length along pipe, m X, Y = dimensionless regime map multipliers 6 = gas voidage p = density, kg/m3 u = surface tension, N/m Subscripts 0 = at atmospheric pressure 1, 2 = in first and second incremental length a = air property g = gas property
1 = liquid property sp = single phase tp = two phase w = water property
Literature Cited
Received for review June 24, 1983 Revised manuscript received April 20, 1984 Accepted April 30, 1984
Ardron, K. H.; Hall, P. C. J . Heat Transfer 1980, 702, 3. Barnea, D.; Shoham, 0.; Taitel, Y. Chem. Eng. Sci. 1962, 3 7 , 741.
Separation of Gold Cyanide Ion from Anion-Exchange Resins Henry H. Law,' Wllson L. Wilson,+ and N. Elise Gabriel* AT&T Bell Laboratories, Murray Hill, New Jersey 07974
The separation of gold cyanide ions from anlon-exchange resins is enhanced when organidwater mixtures are used as eluent solvent. Mixtures of water and dimethylformamide, dimethylacetamide, acetone, N-methyl-2pyrroliiinone, dimethyl sulfoxide, hexamethylphosphoramide, or tetrahydrofuran show better elution than aqueous solutions. The elution rate Is faster at 45 OC than at 23 O C with the preferred eluent (5 M KSCN in 50 vol % dimethylformamide/water).
Introduction It is desirable to recover gold even from very low concentration sources because of ita high price. Strong base anion-exchange resins have been used to remove gold Department of Chemistry, Stanford University, Stanford, CA 94305. t Department of Chemistry, Massachusetts Institute of Technology, Cambridge, MA 02139. 0196-4305/85/1124-0236$01.50/0
cyanide ion from plating rinse solutions. However, it is difficult to recover the gold from the resins by elution. The current method of recovering the precious metals from resins is by incineration, even though this is not environmentally sound and could result in substantial convective losses. Recently, Law (1982, 1983) found that a mixture of dimethylformamide (DMF) and water enhanced the separation of gold cyanide ions from Amberlite IRA-400 resins. These encouraging results prompted us to investigate what other solvents have properties similar to those 0 1985 American Chemical Society
