HOLDUP AND MASS TRANSFER IN BUBBLE COLUMNS CONTAINING SCREEN CYLINDERS B.
H .
CHEN
A N D
RADHE
VALLABH
Department of Chemical Engineering, N o v a Scotia Technical College, H a l i f a x , Nova Scotia, Canada
The effect of cylindrical screen packings on gas holdup and rate of mass transfer in countercurrent gas-liquid bubble columns was studied qualitatively over a reasonable range of gas and liquid flow rates. The packings were made from wire cloths of 8 to 14 meshes per inch and their size varied from !/2 x $2 inch to 1 x 1 inch. Absorption of CO1 by water in two columns, 23/< and 5 ; ~ inches in i.d., was used for this investigation. The addition of screen cylinders to a conventional bubble column greatly increased the gas holdup, t ~ and , the volumetric mass transfer coefficient, KLa, significantly reduced liquid surface fluctuation, and considerably enlarged the range of gas flow over which bubble columns operate normally with high efficiency. Both tc and KLa decrease with increasing packing size and mesh number. KLa also increases with liquid flow, but is essentially independent of column diameter. CG is independent of both column diameter and liquid flow rate over the range of variables studied.
THEgas bubble column in which a gas is bubbled through a deep liquid without agitation is often used as a gasabsorption device or a chemical reactor. Shulman and Molstad (1950) and Houghton et al. (1957) found that bubble columns are particularly useful for mass transfer processes in which the liquid-film resistance is controlling. However, the bubble column has the deficiencies of bubble coalescence and the formation of gaseous slugs a t high gas flows; both may reduce significantly its contacting efficiency. Sutherland et al. (1963), Ishii and Osberg (1965), and Chen and Osberg (1967b) have reported that on addition of cylindrical screen packings to a gas-solid fluidized bed, the growth of bubbles in the bed becomes restricted. As a result, the modified fluidized bed has smaller and more uniformly dispersed gas bubbles within it, and hence less bypassing of the gases and less bed surface fluctuations than in a normally fluidized bed. I t appears likely that the screen packing may also limit the bubble growth in bubble columns, because bubble phenomena in both bubble columns and fluidized beds have been reported to be similar (Davidson and Harrison, 1963). The purpose of this study is to investigate mass transfer characteristics of a packed bubble column in terms of the volumetric mass transfer coefficient and gas holdup, and to coinpare these data with those obtained in a conventional bubble column. Experimental
Equipment. Figure 1 schematically represents the experimental apparatus. Two Plexiglas columns, 5% and 23/1 inches in i.d., were used, with the gas distributor made of twelve %-inch i.d. copper tubes for the larger column and four Xe-inch i.d. stainless steel tubes for the smaller one. Air containing a known proportion of COZ gas was fed Ind. Eng. Chem. Process Des. Develop., Vol. 9, No. 1, January 1970
1. Water
s Figure 1. Flow diagram D. Air distributor R. Rotameters S. Sampling points P. Pocked section
to the distributor a t the bottom of the column and discharged to the atmosphere a t the top. Water was taken from the city mains and introduced into the top of the packing from a shower-head liquid distributor. The water temperature was not controlled, but its variation was small throughout the experiment. Rotameters were used to measure the flows of the inlet air and COz and inlet and outlet liquid. The flows were regulated by needle valves. The test section of the column was carefully packed with screen cylinders up to a desired height. The packings were made from wire screen cloths of different mesh numbers. Their physical properties are listed in Table I. Procedure and Data Treatment. The column was filled with water to a predetermined height, and the screen packing was then dropped, one a t a time, into the column. 121
in the computation of true KLU.However, since the present study was concerned with the effect of the screen packing on the mass transfer coefficient rather than the true values of the coefficient, it was felt acceptable to use the logmean driving force. The use of the log-mean driving force always results in the most conservative value of KLU. The gas holdup, t ~ ,was determined by directly measuring the height of the aerated liquid, 2, and that of the clear liquid without aeration, 2,. The average gas holdup was then evaluated from the relation:
Table 1. Propwtiesof CylindricalWire-Screen Packing d., Inch x Inch
Meshes per Inch
d,, Inch
Voidage
%X% % X %
8 10 14 10 10
0.025 0.025 0.025 0.025 0.025
0.97 0.98 0.98 0.97 0.97
YZ
x 95
%X% 1x1
The gas and liquid flows were turned on at the required rates; the liquid drain valve was adjusted so that the aerated liquid height was maintained a t a level equal to the packed height. After steady-state conditions had been reached, samples of gas and liquid were taken. The inlet and outlet gas and liquid flows were then shut off simultaneously, and the height of the clear liquid was noted. The rate of COa absorption was determined by pipetting 20 ml. of the liquid sample quickly into a known amount of 0.08N Ba(OH)2solution. The solution was always present in excess, and the excess amount was back-titrated using 0.05N HC1. A similar procedure was followed for the analysis of gaseous samples. Values of the volumetric coefficient KLUwere obtained based on liquid-phase composition changes using the equation:
€G
z- 2,
=-
Z
Since the fluctuation of the liquid surface was minimized because of the presence of screen cylinders, the value of t~ obtained was believed to involve only negligible error. Material balances of COZ were frequently made and found to be always within &11%.Data were reproducible with a maximum deviation of *8%. Typical experimental data and calculated results are shown in Table 11. Detailed tabulations of data are given by Vallabh (1968). Results and Discussion
Effect of Gas Flow Rate. The effect of the superficial gas velocity, U G , on the gas holdup, cG, is shown in Figure 2 for 10-mesh-per-inch screen packing. The variation of C G with U G is seen to depend on the range of gas flow rate. At low flows, C G increases with increasing U G but becomes relatively independent of U G at high flows. Also shown for comparison are the data of Fair et al. (1962) for unpacked bubble columns and those obtained in the present study. The addition of screen packings has resulted in a two- to threefold increase of gas holdup, compared with that generated in an unpacked column. I n gas bubble columns much of the expansion is caused by the formation of bubbles which rise through the con-
in which ~ X isL the logarithmic-mean concentration driving force. The use of such a driving force in the present work is arbitrary. The flow pattern in liquid-gas bubble contactors is intermediate between the extremes of no backmixing and complete mixing, and hence the effect of nonideal flow should have been taken into account
~~
~~
~
Table II. Typical Experimental Data and Calculated Results (dl = 2% inches, d, = % inch, d, = 8 mesh per inch, 2 = 4.0 feet)
Run No
G
ZO
CG
102yi
148 149 150 151 152 153
27 42 64 90 110 130
2.70 2.50 2.08 1.88 1.72 1.64
32.5 37.5 48.0 53.1 57.0 59.0
13.33 13.2 13.2 13.95 13.41 12.66
154 155 156 157 158 159
27 42 64 90 110 130
2.60 2.48 2.12 2.02 1.68
35.0 38.0 47.0 49.5 55.1 58.0
13.33 13.20 13.20 13.95 13.41 12.66
160 161 162 163 164 165
27 42 64 90 110 130
2.70 2.58 2.24 2.00 1.76 1.72
32.5 35.8 44.b 50.0 56.0 57.1
13.23 13.20 13.20 13.41 13.95 12.66
1OZY2
1o5x1
lojX*l
IOjX*2
10jAXL
KLa
12.86 12.81 12.90 13.03 14.06 13.31
2.01 4.52 6.96 8.00 8.94 8.66
3.12 3.36 3.60 3.42 3.30 3.00
330 350 380 440 460 470
12.71 12.60 12.78 13.20 14.43 12.93
2.93 4.23 5.52 6.36 6.90 6.45
2.76 2.92 2.91 3.00 2.82 2.69
310 340 360 370 420 430
12.46 12.20 12.90 12.93 14.06 12.71
5.20 6.00 6.51 6.90 7.90 7.58
2.50 2.61 2.50 2.40 2.50 2.28
238 246 262 280 300 293
L = 30,800 Lb./(Hr.) (Sq. Ft.) 2.01 4.52 6.96 8.00 8.44 8.60
7.34 8.39 10.04 10.94 12.38 11.91
L = 23,000 Lb./(Hr.)(Sq. Ft.)
1.80
4.21 6.23 8.28 9.10 9.89 9.20
9.59 10.79 11.47 12.10 13.67 12.16
L = 13,000 Lb./(Hr.)(Sq. Ft.)
122
7.64 8.83 9.61 10.41 11.20 10.61
11.39 11.70 12.16 12.40 13.45 12.40
Ind. Eng. Chern. Process Des. Develop., Vol. 9,No. 1, January 1970
dt :
d, =
6 in.
d,=
10
1 =
8,764 Ib/ hr. sq.ft.
d,:
mesh/ in.
L
400
-
:
5 . 7 5 in. I O mesh in. 8 7 6 0 lb/hr.-sq.ft. d n , in.
0
.5
e
.75 1.0 none
v
0 -
300-
-2
c
I .
u)
-
c
Y
ZOO
-
1001
0
.I
.2
.3
'
uo, ft./sec.
Figure 2. Effect of packing size on gas holdup
tinuous phase at a rate dependent upon their diameter. The presence of screen packing has been observed to restrict the size of the bubble in the column, generating a more uniformly dispersed bubble phase than in a conventional bubble column. The net effect of this factor is to increase the amount of gas holdup. The effect of gas flow rates G on the volumetric mass transfer coefficient, KLa, is shown in Figures 3, 5, 7, and 9 for various packing diameters and mesh openings and various column diameters. Over the ranges of variables studied, the variation of KLa with G follows a common pattern for all cases. KLu first increases rapidly, then levels off and finally approaches a constant value as the gas flow rate is increased. The corresponding data for an unpacked bubble column obtained in this study are shown in Figure 3. They compared well with those reported by Shulman and Molstad (1950). A comparison of data (Figure 4) indicates that KLUis greatly increased by the addition of screen packing to the column. Furthermore, the range in which KLu increases with G has been widened and this may also be one of the advantages of using screen packing. The increase in KLUis obviously a direct consequence of the increase in gas holdup. Since the bubble size in a screen-packed bubble column remains fairly constant and is independent of gas flow at low gas flow rates, any increase in gas holdup in such a column means an increase in the interfacial area, a. KL is not dependent on G and hence KLa should increase in accordance with the increase in ec. However, a t high gas flows, a becomes almost independent of G, as indicated by Voyer and Miller Ind. Eng. Chem. Process Des. Develop., Vol. 9, No. 1, January 1970
I
I
40
2 0
.4
I
I
\
100
80
60
I20
I b / hr.-sq.ft.
G,
Figure 3. Effect of packing size on over-all mass transfer coefficient
d,:
8 rnesh/in.
dn
:
.S in. ,
d, : 3 in.
1, Ib/ hr.rq.ft. 30,800
4
A
13, o o o
4
e
13,000
4
30,800
185
5,160
1.85
A
I
0
.I
Z, ft.
0
I
I
I
I
.2
.3
.4
.5
uc, ft./sec.
Figure 4. Effect of gas flow, liquid flow, and bed height on gas holdup
123
column packed with 3/4 x 3/4 inch cylinders, or with 1 x 1 inch cylinders. Consequently, higher ec should be observed for smaller than for larger packings. Similar observations were reported by Sutherland et ul. (1963) and by Chen and Osberg (1967b) in connection with studies on the characteristics of a packed fluidized bed. The increase of C G with decreasing packing size ,has an obvious effect on the volumetric mass tranfer coefficient, KLU. Since bubbles in a screen-packed column are of uniform size, the increase in CG means an increase in the interfacial area. Therefore KLu would be expected to increase with decreasing packing size. Experimentally determined values of KLu were plotted in Figures 3 and 5 as a function of packing size. The data show the expected variation of KLu with packing size. Figure 5 is a plot of 1nKLu us. 1nL with packing size as parameter. The three straight lines can be represented, respectively, by
G : 3 0 Ib/hr.-sq.ft. d, : 5.75 in. :d, IO mesh/ in. d,, in.
0
.5
0
.7
KLa KLa KLa
4I
I
6
8
1
IO1
I
1
I
2
4
6
oc 0:
LOa7 Lo'' LO6'
for %--inch packing for %-inch packing for 1-inch packing
I
Io3
1 , Ib/hr.- s q , f t
Figure 5. Variation of over-all mass transfer coefficient with packing size and liquid flow rate
(1967) and by the present data (Figure 2); KLUshould therefore be relatively insensitive to G a t high flows. Effect of Liquid Flow Rate. The effect of the liquid flow rate on the gas holdup is shown in Figure 4. Although the data points are slightly scattered, C G is essentially independent of the liquid flow within the range studied, 5700 < L < 31,000 lb./(hr.)(sq. ft.). Similar results were reported by Hoogendoorn and Lips (1965), Webber (1960), and Hofman (1961). Figure 4 also shows that C G is not significantly affected by changing the column height from 1.85 to 4 feet. Figure 5 shows the variation of KLU with liquid flow rate L for runs with and without packing. For all cases, KLu increases with increasing L and the increase follows the general relation,
KLa a L' with 6 equal to 0.58 for the unpacked column and varying between 0.68 and 0.87 for the packed column. The stronger dependence of the packed-column data on L is believed to be the result of the screen-induced turbulence in the liquid phase. Furthermore, for the different types of packing used in this study, the degree of axial mixing is expected to be different, and this difference may also contribute to the change in exponent on L . Effect of Screen-Packing Size. The effect of the packing size on CG may be seen from Figure 2. The gas holdup decreases with increasing packing size for a given mesh opening and gas flow rate. In bubble columns, the size of the bubbles depends largely on the size of the screen packing-for example, bubbles are smaller in a column packed with % x Y2 inch screen cylinders than in a 124
oc
Thus decreasing packing size also increased the rate at which KLUincreases with L . Effect of Mesh Number. It has been reported that in a screen-packed column, the flow pattern of a fluid through it is strongly influenced by the mesh opening of the packing (Chen and Osberg, 1967a). The void spaces in the column due to the hollow structure of the packing and also the void spaces formed by neighboring cylinders offer relatively low resistance to flow. Therefore, with close-mesh packings, the upward-flowing fluids tends not to pass through the mesh openings but to follow the path created by the void spaces. This causes flow irregularity such as channeling within the bed. With packings of wideopen mesh, the effect is small. This flow irregularity was observed in the present investigation. When a rising bubble encountered a screen packing of 1 4 meshes per inch, instead of directly crossing the packing, it sidestepped the packing and followed the void space to the top of the column. To investigate the effect of screen mesh number, the gas holdup and mass transfer coefficient were measured for % x % inch packing made from wire cloths of three mesh numbers-8, 10, and 14 meshes per inch (Figures 6 and 7). Consistent with the theory and the visual observation that channeling becomes increasingly noticeable as mesh number is increased, the gas holdup and the volumetric coefficient decrease with increasing mesh number for a fixed gas and liquid flow rate and a given packing size. Although the mesh number has not been varied over a wide range, the present results on K I Uindicate a general agreement with the findings of Voyer and Miller (1967), who reported that 6-mesh-per-inch packing gives the highest interfacial area and consequently the highest volumetric mass transfer coefficient, KLa. Effect of Column Diameter. Gas holdup data were obtained for various gas and liquid flow rates in a 6-inch and a 3-inch column. Figure 8 indicates that although data scatter, C G is not sensitive to the change of column diameter from 6 to 3 inches. For unpacked bubble columns, a significant effect of Ind. Eng. Chem. Process Des. Develop., Vol. 9, No. 1, January 1970
z : 1.85 f t . ,d
dn:
in.
0.5
: 8 merh/in.
L,
Ib/hr. sq. ft.
dt, in.
8,769
6
t
836
6
0
5,749
3
1.c -
.8
8 IO
0
14
.5
Figure 8. Effect of column diameter and liquid flow on gas holdup
d,: 5.75 in. dn : . 5 in. 1 : 5 7 6 0 I b / hr.- rq. ft. d ,, mesh/ in.
0
.4
ft./sec.
Ug,
Figure 6. Variation of gas holdup with packing mesh number and gas flow rate
0
.3
.2
.I
0 uo, ft. rec
Ib/ hr.- r q . f t .
G
: 60
dn
: . 5 in.
d,
: 8 mesh/ in.
d,, in.
320-
0
5.75
0
2.75
280-
240-
-
I .
-2
200-
Y
I60
-
""1 I
I
20
I 40
I 60
I 80
I
100
1 I20
G, I b / hr.-sa.ft.
Figure 7. Variation of over-all mass transfer coefficient with packing mesh number and gas flow Ind. Eng. Chem. Process Des. Develop., Vol.
9,No. 1, January 1970
7
10'
2
5
IO'
2
3
5
1, Ib/ hr.-rq.ft.
Figure 9. Effect of column diameter and liquid flow on over-all mass transfer coefficient
125
column diameter on CG has been reported. Ellis and Jones’ data (1965) indicate that wall effect increases gas holdup a t diameters up to 3 inches and for diameters greater than 3 inches, gas holdup is independent of the diameter. Yoshida and Akita (1965) have reported that the effect of column diameter becomes insignificant only when the diameter is greater than 6 inches, while Hughmark’s data (1967) show this independence to occur at diameters greater than 4 inches. Significant effects of vessel diameter on e G were reported by Fair et al. (1962) for vessel diameters up to 18 inches. Therefore, on the basis of the present results it may be concluded that by using proper screen packings, the effect of column diameter on C G is minimized, probably because the screen packing disperses the gas phase uniformly across the column cross section. The effect of column diameter on KLU is shown in Figure 9. Although the data are not taken over the same range of liquid flow, the data in the overlapping region indicate that the effect of column diameter on KLU is slight. A comparison of the present results on KLUwith similar data in the literature suggests that the effect of column diameter on KLUhas been reduced again by the use of proper screen packings. Acknowledgment
The authors thank the National Research Council of Canada for financial support. B. H . Chen is grateful to the Atlantic Industrial Research Institute, Nova Scotia Technical College, for a Summer Fellowship. Nomenclature U
L NA
S
126
interfacial area per unit volume of packed tower, sq. ft./cu. ft. mesh number, meshiin. packing dimension, in. x in. column diameter, in. wire diameter, in. gas mass velocity, lb./ (hr.)(sq. ft.) over-all volumetric mass transfer coefficient, lb. moles/ (hr.) (cu. ft.) (lb. moles) / (cu. ft.) liquid mass velocity, 1b.i (hr.) (sq. ft.) rate of mass transfer of component A, lb. moles/ hr . column cross sectional area, sq. ft.
uG = gas velocity, ft./sec.
a x L = logarithmic-mean driving force, lb. moles CO*/ lb. moles liquid x1 = outlet liquid composition, lb. moles COz/lb. mole
liquid xz = inlet liquid composition, lb. moles COz/lb. mole
liquid yl = inlet gas composition, lb. moles COn/lb. mole gas
yz = outlet gas composition, lb. moles COz/lb. mole gas
CG
= gas holdup, fractional
p i = molal liquid density, lb. molesicu. ft.
* = equilibrium composition
Literature Cited
Chen, B. H., Osberg, G. L., Can. J . Chem. Eng. 45, 46 (1967a). Chen, B. H., Osberg, G. L., Can. J . Chem. Eng. 45, 90 (1967b). Davidson, J. F., Harrison, D., “Fluidized Particles,” Cambridge University Press,, London, 1963. Ellis, J. E., Jones, E. L., Two-Phase Flow Symposium, Exeter, England, June 1965. Fair, J. R., Lambright, A. J., Anderson, J. W., IND.ENG. CHEM.PROCESS DESIGNDEVELOP.1, 33 (1962). Hofman, H., Chern. Eng. Sci. 14, 193 (1961). Hoogendoorn, C. J., Lips, J., Can. J . Chem. Eng. 43, 125 (1965). Houghton, G., McLean, A. M., Ritchie, P. D., Chem. Eng. Sci. 7, 26 (1957). Hughmark, G. A,, IND.ENG. CHEM. PROCESS DESIGN 6,218 (1967). DEVELOP. Ishii, T., Osberg, G. L., A.1.Ch.E. J . 11, 279 (1965). Shulman, H . L., Molstad, M. C., Ind. Eng. Chem. 42, 1058 (1950). Sutherland, J. P., Vassilatos, G., Kubota, H., Osberg, G. L., A.1.Ch.E. J . 9, 427 (1963). Vallabh, R., M. Eng. thesis, Nova Scotia Technical College, Halifax, Nova Scotia, Canada, 1968. Voyer, R. D., Miller, A. L., 17th Canadian Chemical Engineering Conference, Niagara Falls, Ontario, Canada, Oct. 17, 1967. Webber, T. H., dissertation, Darmstadt, 1960. Yoshida, F., Akita, K., A.1.Ch.E. J . 11, 9 (1965). RECEIVED for review March 24, 1969 ACCEPTED September 22, 1969
Ind. Eng. Chem. Process Des. Develop., Vol. 9, No. 1, January 1970
