Ind. Eng. C h e m . Res. 1987,26, 958-966
958
Hydrodynamics of Turbulent Bed Contactors. 1. Operating Regimes and Liquid Holdup G. V. Vunjak-Novakovib and D. V. Vukovii! Faculty for Technology and Metallurgy, Department 11002 Belgrade, Yugoslavia
of
Chemical Engineering, Belgrade University,
H. Littman* Relzsselaer Polytechnic Institute, Troy, New k’ork 12180-3590
T h e hydrodynamics of a three-phase turbulent bed contactor with countercurrent flow of gas and liquid has been studied and its performance compared with that of a packed-bed column using a consistent set of data. T h e first paper of this series deals with the operating regimes and liquid holdup in a bed of spherical particles. Two types of fluidized-bed operation and conventional packed-bed operation were studied, and criteria for identifying each regime are given. Correlations for the liquid holdup are reported for each regime. Introduction Turbulent bed contacting is a relatively new and highly effective means of contacting gas and liquid streams. Beds of low-density particles are fluidized by upwardly flowing gas and irrigated by downwardly flowing liquid. The particles used were spheres ranging in diameter from 10 to 38 mm and in density from 182 to 980 kg/m3. The beds are supported by nonflooding grids, placed sufficiently far apart to permit bed expansion and turbulent motion in the beds. Turbulent bed contactors have been used for gas absorption and scrubbing for the last 20 years. Recently, they have been used for SOz absorption and for particulate removal from stack gases. They offer considerable advantages over conventional packed-bed columns because they operate a t higher gas velocities, they have higher mass-transfer rates, and the mobility of the packing prevents plugging. This makes the turbulent bed contactor suitable for handling streams containing particulate matter and precipitates. In addition, there is practically no channeling or bypassing. The high gas and liquid throughputs (below about 8 m/s and 60 kg/(m2 s)) are made possible by bed expansion which avoids flooding. Overall mass-transfer rates as much as 2 orders magnitude higher than that in packed columns have been reported by Douglas et al. (1963). Industrial experience has also pointed out some of the disadvantages of turbulent bed contactors. These include bed pulsations, backmixing in the liquid phase, and mechanical erosion of the packing spheres. However, the characteristics of high capacity and high rates of mass transfer and particulate removal have made for successful industrial application of these units. Fundamental studies of the hydrodynamics of mobile beds such as operating regimes, liquid holdup, pressure drop, minimum fluidization velocity, bed expansion, and axial mixing in the liquid phase have been conducted by a number of investigators. The data reported in most cases have been obtained by using small columns (less than 150 mm in diameter) operated over a narrow range of variables. Therefore, literature data are not always consistent, and the applicability of the proposed empirical correlations is generally limited to the specific experimental conditions studied. Available theory developed to explain the hydrodynamics of these contactors and to relate it to that of packed columns is also limited (O’Neill et al., 1972). This paper deals with the basic hydrodynamic characteristics of countercurrent contacting of gas and liquid in
a bed of spherical packing and addresses itself to operating regimes and liquid holdup in fluidized and packed beds. Other hydrodynamic data (pressure drop, bed expansion, minimum fluidization velocity, maximum capacity of the contactor) and mass transfer will be reported in subsequent papers in this series. The data (Vunjak-Novakovie,1980) were obtained over a wide range of operating conditions. The possible modes of operation first outlined by O’Neill et al. (1972) are analyzed, and a model is developed to relate the hydrodynamics of turbulent bed contactors to that in packed beds. Correlations for liquid holdup are also presented. Literature Review (a) Operating Regimes. In packed beds, the pressure drop across the bed remains less than the sum of the weight of particles and liquid holdup. When the bed pressure drop reaches the total weight of the particles and the liquid holdup, the bed will fluidize and expand. O’Neill et al. (1972) recognized that a turbulent bed contactor can operate in one of two regimes: (1) fluidization without flooding and (2) fluidization due to incipient flooding. Fluidization which occurs at flow rates below the flooding point of the corresponding packed bed was termed fluidization without flooding. In that case, the pressure drop required to fluidize the bed is lower than that for flooding in a packed bed and the bed fluidizes before flooding occurs. This regime depends primarily on particle density and to a lesser extent on liquid flow rate and particle size. O’Neill et al. (1972) have stressed the advantages of operating in the incipient flooding regime because it results in conditions of higher interfacial contact of gas and liquid. Industrial contactors, however, are currently operated in both regimes. If the bed fluidizes without flooding, the liquid holdup does not vary with gas velocity in either the packed bed or the fluidized bed, as shown by O’Neill et al. (1972),Chen and Douglas (1968), and Groeneveld (1967). If the bed fluidizes at flooding, O’Neill et al. (1972) have predicted that the liquid holdup should increase with both gas velocity and particle density. Experimental data presented by O’Neill et al. (1973), Kit0 et al. (1978), VunjakNovakoviE et al. (1980), and Vunjak-NovakoviE and Littman (1984) show that this prediction is incorrect with regard to gas velocity. The model of O’Neill et al. (1972) represents an important contribution to the understanding of the behavior of turbulent bed contactors. Much of the confusion in the
0888-588518712626-0958$01.50/0 0 1987 American Chemical Society
Ind. Eng. Chem. Res., Vol. 26, No. 5, 1987 959 literature can be eliminated by identifying two operating regimes for fluidized-bed operation. Some caution is necessary, however, in using their phase diagram for predicting the operating regimes in beds of spherical particles because the diagram is based on pressure gradients in conventional packed beds and uses the Chen and Douglas (1968) correlation for liquid holdup in fluizied beds. The assumed value for the pressure gradient at the flooding point was not examined for spherical particles, and the Chen and Douglas (1968) correlation is valid only for fluidization without flooding. Of greater significance is the fact that the model is valid only at minimum fluidization. The predictions are not correct when the gas velocity is increased above ud because the void fraction of gas in the bed varies, as will be shown in part 2 of this series (Vunjak-NovakoviE et al., 1987). (b) Liquid Holdup. Empirical correlations available in the literature for predicting the liquid holdup in turbulent bed contactors are given in Table I. The operating regime and range of variables studied are also given in the table. Analysis of data in the literature shows that the equations proposed by Chen and Douglas (1968), Ushida et al. (1977), and Kit0 et al. (1978) give the most reliable predictions for the range of variables they studied. All the correlations predict that the liquid holdup is independent of the gas velocity. They show that liquid holdup increases with liquid flow rate to the power of 0.6-1.0 and particle density to the power 0-0.18 and decreases with the particle diameter to the power 0.5-0.89 and static bed height to the power 0 . 4 . Kito et al. (1978) and Ushida et al. (1977) have also shown the effects of column diameter, liquid viscosity, and free cross-sectional area of the supporting grid. (c) Hydrodynamics of the Packed Beds. The hydrodynamics of packed beds were investigated in detail owing to its wide application in mass-transfer operations. Many authors give empirical correlations for the static and overall liquid holdup in the bed, pressure drop, and flooding velocity, but the available data and correlations concern primarily the common types of packing. Data for packed beds with spherical particles are limited because such packing is not generally used in packed columns. Spheres have minimal surface area and voidage when compared with other packings, which results in low flooding velocities, high pressure drops, and poor masstransfer performance. Pressure drop and liquid holdup in beds of spherical particles have been correlated by Kolar and Broz (l965,1968a,b) and Kolar et al. (1970), but their data are not useful in this study because the particle diameters are too small. Equipment and Procedure, Bed Properties, and Operating Variables The hydrodynamics of packed and fluidized beds were investigated by using three pilot plant systems of similar design. A schematic diagram of the apparatus is shown in Figure 1. Plexiglas columns with inside diameters of 140,194, and 290 mm and a height of about 2 m were used in this work. Two of the units, D,= 140 and 290 mm, were installed at the Technical University Clausthal in F.R.G. and the third one, D,= 194 mm, a t Belgrade University. A flow-straightening section was installed below the bed in order to obtain a reasonably flat inlet gas-velocity profile. The spheres were supported on nonflooding grids of different types, which are described in Table 11. The characteristics of the particles are shown in Table 111. They are hollow spheres except for the 10-mm size with nominal densities of 700 and 1000 kg/m3. A wide
FLUIDIZED BED LIQUID DISTRIBUTOR
@
SECTION
MEASURING POINTS
SHUTTER MEASURHG PNEUMPTIC W DRIVE INDER
RDTAMmR
@
@
PRESSURE TEMPERATURE
u Figure 1. Schematic diagram of absorption column.
range of the bed voidage and packing factor could be obtained by varying the column diameter, static bed height, and particle diameter. Room air was supplied to the bottom of the column by the blower. Recirculation of gas was required in some experimental runs, according to the specific demands of mass-transfer experiments which were carried out simultaneously. The gas velocity was measured by using orifice plates connected in parallel. Water was admitted to the bed through a distributor which gave a uniform liquid distribution over the top of the bed. The liquid flow rate was measured by using rotameters, and a separator for liquid drops prevented liquid carry-over. The experiments were conducted in a one-stage column with static bed heights to 300 mm (to 1m for packed beds). The initial voidage was fixed by compacting the bed in all runs in which the loading, flooding, and fluidization velocities were determined. These bed heights were later reproduced in the same way by making sure that the same number of spheres packed to the same bed height. Superficial gas velocities were varied from 0 to 4 m/s and liquid mass flow rates from 0 to 34 kg/(m2 s). The pressure drop and liquid holdup were experimentally determined in fluidized and packed beds over the whole range of gas flow rates. Constant voidage in the packed bed was maintained at velocities above loading by using a screen. Pressures were measured below the grid and above the bed with pressure transducers whose outputs were time-averaged. At each level, the averaged pressure represented the input of four taps located symmetrically around the column. Pressure drops corrected for the drop in the empty column were determined with an accuracy of 0.5%. The liquid holdup was measured directly by using the pneumatically driven fast shutter shown in Figure 2. The shutter and valve on the liquid distributor were closed simultaneously, and a high velocity air flow, introduced into the column above the packing, prevented the liquid in the inlet line from being drawn into the bed. In the 194-mm column, the liquid which accumulated in the bed was collected at the exit of the overflow tube at the bottom of the bed after the magnetic switch on the water line was closed. The portion of the static liquid holdup, hh, representing the liquid held by adhesive forces on the particles was determined by weighing the wetted and dry packing. Appropriate corrections were then made. Results (a) Operating Regimes. In order to understand the results, it is necessary to discuss the subject of operating
960 Ind. Eng. Chem. Res., Vol. 26, No. 5, 1987
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Ind. Eng. Chem. Res., Vol. 26,No. 5, 1987 961
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962 Ind. Eng. Chem. Res., Vol. 26, No. 5, 1987 Table 11. Characteristics of the Grids Used in This Study free crosswidth of the slot spacing of the type of or mesh, mm slots, mm grid sectional area plate 0.36 7.78 7.15 slotted 0.52 3.95 6.35 screen 0.78 7.89 1.03 Table 111. ProDerties of the Particles Used in This Study
10
20 25
38
320 379 683 980 195 413 203 407 182 374
0.37-0.40
5625-6170
0.39-0.41
2568-3731
0.39-0.43
1720-2468
0.41-0.46
876-1352
Figure 2. Schematic diagram of the shutter used in measurements of liquid holdup.
u3
regimes. This is best done by examining the traditional loading-flooding curves in Figure 3 which show the bed pressure drop and liquid holdup plotted against the gas velocity for a fixed liquid mass flow rate. Up to the loading point (point B), there is essentially no influence of the gas velocity on the liquid holdup and the pressure drop is linear with gas velocity on a logarithmic plot. If the particle density is low enough (p, < 170 kg/m3), the bed will fluidize before the loading point is reached. Figure 3 shows this case using the data of Chen and Douglas (1968), Tichy et al. (1972), and Tichy and Douglas (1972) to illustrate the point. Note that the pressure drop becomes constant and the liquid holdup remains the same as that in the packed bed a t gas velocities in excess of the minimum. This is the fluidizing regime which O'Neill et al. (1972) referred to as "fluidization without flooding" or, in our terminology, "type-I operation". Fluidized beds of this type are visually similar to aggregative gas fluidized beds with large bubbles and slugs, especially in deep beds (Ho/D, > 1or Ho > 300 mm). The minimum fluidizing velocity increases with particle density as predicted by O'Neill et al. (1972), and the liquid holdup is constant over the whole range of gas velocities. For the heavier particles, p = 379 and 683 kg/m3, the beds fluidize after the packedPbed floods. This fluidized
Ind. Eng. Chem. Res., Vol. 26, No. 5, 1987 963 I
IO
5
I
I
lor
pp: 6 8 3 kg/m3
Cm3lh
I
~ ~ ~ $ ~ ’ ’
E
. a“
I
h
D?
LOADING POINT
S
I
+-- /
D,= 140mm $ =0 . 5 2 H, = 3 0 0 mm dp = IO mm L : 5.56kg/m2s
I
””*”I””””’ p o : 156 kg/m3
Ho ~ 3 0 mm 0 20.52 dp=lOmm ~~237 kqlm’ 9
+
0
I
1I
0.5
FLUIDIZED BED
I 002
ii II /I
005
01
02
05
I
2
3
5 u h/s)
Figure 4. Effect of liquid flow rate on the operating liquid holdup.
(b) Pressure Drop, Liquid Holdup, and Flooding Velocity in Packed Beds. The pressure drop below the loading point was correlated by using Leva’s (1953) equation to give
0.I 0.5
m
E
\
m
E.
0.1
1
J 0
c
0.05 UMF < U F
FOR 156 kg/m3
0.05
CHEN AND DOUGLAS I19681 UMF
=
UF
0.1 0.5 1.0 GAS VELOCITY u ( m / s l
2.0
Figure 3. Pressure drop and operating liquid holdup in the packed and fluidized beds.
bed regime which O’Neill et al. (1972) referred to as “fluidization due to incipient flooding” we denote as “type-I1 operation”. If the packed bed does not fluidize first, the pressure drop will continue to increase with gas velocity until the bed floods. Between points B and C in Figure 3, the pressure drop increases rapidly with gas velocity, as expected. The holdup also increases and waves are observed on the sheets of fluid flowing downward and on the fluid surface coating the particle. At the flooding point, point C, liquid begins to accumulate in the bed. Depending on the particle density, a pressure drop is reached a t which the bed begins to fluidize and expand at point E. Due to the steep increase of pressure drop and liquid holdup with gas velocity (region CE in Figure 3), the effect of particle density is opposite to that in type-I fluidized beds. When the particle density is increased, there is an increase in liquid holdup, but the minimum fluidizing velocity remains equal to the flooding velocity. At point F, the top of the bed is fluidized and, as the gas velocity is increased, the bed continues t o fluidize to lower and lower levels in the bed until, a t point G, the entire bed is fully fluidized. In type-I1 operation, the bed has the appearance of a liquid fluidized bed with small bubbles. It is these small bubbles and the mixing in the liquid phase that leads to good contacting both in gas and liquid-film controlling mass-transfer operations. Finally, if the bed does not fluidize before the pressure drop per unit of bed length reaches point D where Ap = p L g H , , it will remain a flooded packed bed. For the parameters specified in Figure 3, a bed with a particle density greater than 1300 kg/m3 will never fluidize. Surprisingly, type-I1 operation is possible only over a limited range of particle density.
The pressure drop per unit of bed height a t the flooding point was found to be about 2 kPa/m. The maximum pressure drop which can be generated in a packed bed that is completely flooded (point D in Figure 3) was obtained from a large number of runs (needed because the bed height fluctuates) and found to be that of the column filled with water, that is, A p / H o = p ~ =g 9.81 kPa/m (2) The static liquid holdup was independent of the particle diameter
hLa= 0.02 m3/m3
(3) The same value was obtained by Chen and Douglas (1968). The operating liquid holdup at gas velocities below the loading point was correlated by a regression analysis of 186 experimental values as hLo= 6.49ReL-0~139FrL0~429(HO/Dc)-0~567 u = f27% (4) In terms of the quantities varied, the total liquid holdup, h~~+ hLs, was hL = 2.48 X 10-3(H0/Dc)-0~567dp-00.568L0.719 + 0.02 (5) The flooding velocity was correlated by a regression analysis of 60 experimental values to give log [(uF2/gc)Fp(pG/pL)I = 0.247 - 1.615(L/G)0~25(p~/pL)0.125 (6) (c) Liquid Holdup in Fluidized Beds. The experimental data for liquid holdup, bed pressure drop, and bed expansion as a function of gas velocity follow the packed-bed curves below the minimum fluidizing velocity as mentioned earlier. Thereafter the liquid holdup depends on the fluidizing regime (Figure 3). Figures 4 and 5 show the influence of liquid mass flow rate and particle diameter on the operating liquid holdup in type-I1 operation. An increase in liquid flow rate or a decrease in particle diameter increases the liquid holdup. The effect of static bed height on liquid holdup is shown in Figure 6. We confirm the work of previous investigators that hLdecreases with bed height in packed and fluidized beds. Note in Figure 6 that when the bed is near flooding and not fully fluidized, the effect of bed height is small. The liquid holdup data were taken under the same conditions in different columns. Some differences in the liquid holdup values were observed at the lowest liquid
964
Ind. Eng. Chem. Res., Vol. 26, No. 5 , 1987
Table IV. Comparison of Experimental Data for Liquid Holdup, h L ~ ,with the Predictions of Literature Correlations (Hoecker, 1980) hr- (calcd), m3/m3 L, HO, hdexptl), Kat0 et al., Handl, Chen and Douglas, kg/(m2 s) kgpjq63 m m3/m3 1978 1976 eq 7 1968 5.56
379
0.100 0.200 0.300 0.200 0.300 0.200 0.300 0.200 0.300 0.100 0.200 0.300 0.300 0.300
13.9 25.0 33.3 5.56
683
13.9 25.0
0.203 0.154 0.131 0.277 0.235 0.403 0.343 0.485 0.412 0.226 0.171 0.145 0.262 0.381
0.213 0.146 0.109 0.252 0.230 0.404 0.355 0.522 0.480 0.223 0.171 0.152 0.277 0.515
0.093 0.093 0.093 0.188 0.188 0.294 0.294 0.365 0.365 0.093 0.093 0.093 0.188 0.294
0.182 0.135 0.113 0.285 0.239 0.459 0.385 0.580 0.486 0.192 0.142 0.119 0.252 0.406
0.024 0.024 0.024 0.027 0.027 0.030 0.030 0.032 0.032 0.024 0.024 0.024 0.027 0.030
RICKED BED OPERATION
pp(kgI:
d p =38 (mm)
I0 '
5.102
1 I
01
0.2
03
0.5
2
1 u
TYPE
I1 FLUIDIZED BED OPERATION
1I
3
(m/s)
Figure 5. Effect of particle diameter on the liquid holdup.
mass flow rates, probably as a result of change in the liquid distribution. When D J d > 10, wall effects can be neglected and the liquid shoufd be uniformly distributed over the entire cross section of the column. Liquid holdup data were correlated for both types of fluidized-bed operation. Although experimental data for type-I operation were obtained for a limited number of packings and low liquid flow rates, they were sufficient to 3 5 10 20 50 L(kg/mzs) show, as other investigators have found (Chen and Douglas, 1968; Gelperin et al., 1968;Kupriyanov et al., 1969; Ushida Figure 6. Effect of initial bed height on liquid holdup in type-I1 et al., 1977; Groeneveld, 1967; that there is no difference fluidized beds. in liquid holdup between type-I fluidized beds and packed beds. We found that eq 4 fits all available literature data Table V. Predicted Effect of Static Bed Height, Particle for type-I operation. Diameter Liquid Mass Flow Rate, and Particle Density on Liquid Holdup in Type-I1 Fluidized Beds hL for type-I1 operation was correlated by a regression analysis of 243 experimental values to give eq proposed by eq proposed by eq 7 Kit0 et al. (1978) eq 7 Kit0 et al. (1978) hL = 7.33ReL4,059Fr L0~435(H0/Dc)4~433(pp/pL)0~090 + 0.02 H0-3).433 H0-0'.40 LO.812 L0.64 (7) d0.495 d-0.44 po.090 po.09 For the air-water system, eq 7 can be simplified to surmise that the enhanced effect of liquid flow rate prehL = dicted by eq 7 includes the effect of the support grid, which 4.43 x 10-3dP4.494L0.812(pp/p~)0.0g0(Ho/Dc)4.433 + 0.02 is expressed separately in Kito's equation. (8) Chen and Douglas' (1968) equation hL = 2.4 X 10-3d,4,5(L/pL)0,6+ 0.02 (9) Discussion The predictions of eq 7 and of the equations proposed by Kit0 et al. (1978) and Handl (1976) are compared in Table IV with values of the liquid holdup obtained in type-I1 fluidized beds. Both Kito's equation and eq 7 fit the experimental data quite well with the maximum deviations of +19% to -14% and +13.7% to -21.6%, respectively. Both of these equations predict almost the same variation of the four main parameters affecting liquid holdup in fully fluidized beds as shown in Table V. We
should be valid for type-I operation. However, eq 9 was obtained in a square column (305 X 305 mm) using a single = 305 mm) with liquid flow rates varying bed height (Ho between 2.8 and 13.8 kg/(m2 s) and particle diameters ranging from 12.7 to 38 mm. When the equivalent column diameter and the static bed height are replaced in eq 5, we obtain the relation hL = 2.7 X 10-3dp4.568(L/pL)0.719 + 0.02 (10)
Ind. Eng. Chem. Res., Vol. 26, No. 5, 1987 965 which is remarkably similar to Chen and Douglas’ (1968) equation. The most uniform and smooth fluidization was observed in the region of partial fluidization. Slugs and bubbles first appear in fully fluidized beds. It is worth noting that the height of a transfer unit for mass transfer has a flat minimum in the region of partial fluidization, essentially equal to that a t the flooding point of the packed bed (VunjakNovakoviC, 1980). Figure 6 shows that the liquid holdup decreases both in packed beds near flooding and in partially fluidized beds when the static bed height is increased. In packed beds near flooding and in partially fluidized beds, the static bed height has a significant effect on liquid holdup. Since nonuniform liquid distribution in deep beds causes a decrease in the total liquid holdup, except near flooding, it was presumed that partially fluidized beds also operate under flooding conditions. This assumption was later verified using bed expansion data and will be discussed in detail in part 2 of this series (Vunjak-NovakoviE e t al., 1987). The total volume of liquid in the fluidized bed increases when the static height is increased, but the rate of increase is negative. There is, therefore, an optimum packing volume for a single stage. Since the number of masstransfer units reaches an asymptote as the bed height is raised (Vunjak-NovakoviC,1980, Vunjak-NovakoviEet al., 1984), staging is necessary in order to achieve optimum performance for the given volume of packing with the available pressure drop. By use of the experimental results for liquid holdup and pressure drop in packed and fluidized beds, a phase diagram was constructed for countercurrent flow of gas and liquid in a bed of spherical particles. This plot, presented in Figure 7, shows the range of particle densities for each regime of fluidized bed operation as a function of liquid mass flow rate and particle diameter. The diagram is qualitatively similar to the first of the two-phase diagrams proposed by O’Neill et al. (1972). Quantitative differences are apparent since O’Neill et al. (1972) had no relevant experimental pressure drop and liquid holdup data available to them. The regime boundary lines seen in Figure 7 were obtained by using the force balance at minimum fluidization (Ap/HO)mF = (1- 6J)(pp - PG)g + hLoPLg
(11)
and our experimental results and correlations for liquid holdup and pressure drop. From Figure 3, it is apparent that (AP/H~< ) ~(Ap/HOIc ~ for type-I fluidized beds (12) (@/Ho)c
< (Ap/Ho)m~< (A~/Ho)Dfor type-I1 fluidized beds (13)
(Ap/H0)&
> (Ap/Ho)D
for packed beds
(14)
The criterion for the transition from a type-I to a type-I1 operation was taken to be (AP/HO)mF = (Ap/HO)C
(15)
Combining eq 11 and 15, we obtain (1 - d ( P p - P G k
+ hLdLg = (AP/Ho)c
Noting that only the operational liquid holdup contributes to the pressure drop, i.e., hLo = hL - 0.02, and solving for the particle density, we find that bP)t
= PG + (Ap/ffo)c/P/(l - &I [ ( h -~ 0.02)/(1 - d l I (16)
( p )t is the density of a particle which fluidizes a t the floo&ng point (point C in Figure 3) and, therefore, represents the density a t which there is a transition from type4 to type-I1 operation. (Ap/Ho)c and hL are obtained from experimental data of the type given in Figure 3. Since both of these quantities vary with d, and L , the regime boundary line in Figure 7 depends on those quantities. Alternatively, ( p may be obtained by using correlations for hLand (Ap/ o)c. In that case, hL is calculated from eq 7 and (AP/H,)~from eq 1 after replacing G with ~ G Equations 7 and 16 are then solved iteratively for ( P , ) ~ We have mentioned earlier that point D in Figure 3 gives the highest pressure drop per unit of bed height in a particular system for which fluidized operation is possible. If the bed does not fluidize before point D is reached, it will remain a flooded packed bed. The maximum density of particles which can be fluidized in the countercurrent flow of gas and liquid was calculated, therefore, from (@/HO)mF = (Ap/HO)D (17)
Hp)
so that b p ) m = (Ap/HO)D[l/(l - d g l
+ [(hL)D/1 - €01
(18)
In order to solve for (pp)mby using eq 15, (hL)Dis obtained graphically by using the experimental loadingflooding curves for packed beds (Figure 3), and eq 2 shows that (AplHo)~ is 9.81 kPa/m. Equation 18 then gives Cop),, as a function of d, and L. The regime boundary line showing this relationship is plotted in Figure 7. The operating regime is controlled by conditions at the minimum fluidizing point. As a result, the actual gas velocity is missing from the phase diagram. Figure 7 is the regime map for predicting the operating regimes of a turbulent bed contactor. It can also be used for selecting the operating conditions in a column. For example, if a certain liquid flow rate is required by mass-transfer considerations and packing of a standard size is selected, the diagram shows the range of particle densities for which typed or type-I1 operation will occur. Note that, when the particle density exceeds about 300 kg/m3, type-I operation is impossible; for particle densities above about 1300 kg/m3, only packed-bed operation is possible. Several experiments were conducted to illustrate the surprising observation that high-density particles cannot be fluidized in countercurrent flow of gas and liquid. A bed of carbonate particles (d, = 10 mm, p = 2600 kg/m3) was fluidized with gas and then irrigate& The bed collapsed instantaneously, and liquid accumulated above the bed. The liquid could not flow through the bed. When the liquid flow was shut off, the bed did not fluidize until the packing had been completely dried by the gas stream. Countercurrent operation in the packed bed was possible at low gas velocities. With the most dense particles used in this study (d, = 10 mm, p, = 980 kg/m3), fluidized bed operation was possible a t liquid flow rates less than 28 kg/(m2 s). When this value was exceeded, the same situation we had observed using carbonate particles occurred again.
Conclusions It has been shown that there are two fluidized-bed operating regimes (denoted type I and type 11) whose boundaries depend on particle density, particle size, and liquid mass flow rate. A criterion has been developed to predict the regime boundary based on liquid holdup and pressure drop data. It is also shown both theoretically and experimentally that fluidized-bed operation is impossible if the particle densities are too great. The boundary line
U
~
966
Ind. Eng. Chem. Res., Vol. 26, No. 5, 1987 p~
0.1
02
D, = 1 4 0 h m )
pp=
d p : lO(mm)
LS
03
0 4 05
1
pG
pL = density of liquid, kg/m3 pp = particle density, kg/m3 (pp)m = maximum density of particle
2
3 u(mW
Figure 7. Phase diagram for countercurrent flow of air and water in beds of spherical particles.
on the regime map between t y p e 4 operation and a flooded packed bed is given. Experiments have been carried out in three pilot columns with the diameters of 140,194, and 290 mm to obtain liquid holdup data in fluidized beds and packed beds and flooding velocities and pressure drop in packed beds. This paper describes operating regimes and liquid holdup in the different operating regimes. Correlations for liquid holdup in packed beds and type-I and type-I1 fluidized beds are presented in eq 5 and 7 . Acknowledgment
A part of the experimental work was performed at Technical University Clausthal in F.R.G., in the Institute for Thermal Processes. G. Vunjak-NovakoviE is grateful to Prof. A. Vogelpohl, Director of the Institute, for placing on disposal the available experimental units. The help of Dipl. Ing. A. Obermayer in the experimental work is also gratefully acknowledged. This work was supported by the Research Council of Serbia and National Science Foundation through Grant F5F035Y provided by the Yugoslav-US Joint Board on Scientific and Technological Cooperation. Nomenclature A = cross-sectional area of column, m2 a = dry surface of packing, per unit of static bed volume, m2/m3 D, = column diameter, m do = hole diameter in the supporting grid, m d = particle diameter, m = packing factor, Fp= a/c3;for spherical packing = (6/dp)(1
E',
-
€1~3)
FrL = Frounde number for the li uid phase, (L/pL)2/gdp g = gravitational acceleration, m/s G = gas mass flow rate, kg/(m2 s) Ga = Galileo number, d;p2g/pL2 Ho = static bed height, m h L = liquid holdup per unit of static bed volume, m3/m3 hh = static liquid holdup per unit of static bed volume, m3/m3 h h = operational liquid holdup per unit of static bed volume,
9
m3/m3 L = liquid mass flow rate, kg/m2s Ap = pressure drop, Pa ReL = Reynolds number for the liquid phase, d&L/pL u = superficial gas velocity, m/s U,F = minimum fluidization velocity, m/s U F = flooding velocity, m/s URF = gas velocity at which bed is fully fluidized, m/s V = volume, m3 WeL = Weber number, dA2/pLuL Greek Symbols t
co
= bed voidage = initial bed voidage
= viscosity of the liquid phase, kg/ms = density of gas, kg/m3
that can be fluidized, k/m3 (p,)$ = transition particle density (type I to type 11), kg/m3 u = mean relative deviation uL = surface tension of liquid r$ = free cross-sectional area of the supporting grid, m2/m2 Subscripts G = gas phase L = liquid phase mF = minimum fluidization C = flooding conditions, point C in Figure 3 D = flooding conditions, point D in Figure 3 Literature Cited Akselrod, L. S.; Yakovenko, M. M. Theor. Found. Chem. Eng. (Engl. Transl.) 1969, 3, 124. Balabekov, 0. S.; Tarat, E. Ya.; Romankov, P. G.; Mikhalev, M. F. J. A p p l . Chem. USSR (Engl. Transl. 1971,44, 1068. Chen, B. H.; Douglas, W. J. M. Can. J . Chem. Eng. 1968, 46, 245. Douglas, H. R.; Snider, I. W. A.; Tomlinson, G. H., 11. Presented a t the 50th Annual Meeting AIChE, Buffalo, NY, 1963. Gelperin, N. I.; Kruglyakov, B. S. Sou. Chem. Ind. 1977, 53, 66. Gelperin, N. I.; Liferenko, V. A,; Grishko, V. Z.; Sokolov, V. I. Prom. Sanit. Ochistka Gazov 1976,3, 14. Gelperin, N. I.; Savchenko, V. I.; Grishko, V. Z. Theor. Found. Chem. Eng. (Engl. Transl.) 1968, 2, 65. Groeneveld, K. J. V. Ph.D. Dissertation, Technical University, Delft, The Netherlands, 1967. Handl, R. Ph.D. Thesis, Technical University Clausthal, F.R.G., 1976. Hoecker, J. Seminar work, Technical University Clausthal, F.R.G., 1980. Kito, M.; Tabei, K.; Murata, K. Znd. Eng. Chem. Process Des. Deu. 1978, 17, 568. Kolar, V.; Broz, Z. Collect. Czech. Chem. Commun. 1965,30, 2527. Kolar, V.; Broz, Z. Collect. Czech. Chem. Commun. 1968a, 33,2722. Kolar, V.; Broz, Z. Collect. Czech. Chem. Commun. 1968b, 33,3870. Kolar, V.; Broz, Z.; Tichy, J. Collect. Czech. Chem. Commun. 1970, 35, 3344. Kupriyanov, V. N.; Kan, S. V.; Yatskov, A. D., private communication, 1969. Leva, M. Tower Packings and Packed tower Design, 2nd ed.; US Stoneware: Akron, OH, 1953. O'Neill, B. K.; Nicklin, D. J.; Leung, L. S.Presented a t the International Conference on Fluidization and Its Application, Toulouse, Oct 1973. O'Neill, B. K.; Nicklin, D. J.; Morgan, N. J.; Leung, L. S. Can. J . Chem. Eng. 1972,50, 595. Petrov, P. M.; Tassev, Zh. Chem.-Ing. Technol. 1978, 50, 887. Tarat, E. Ya.; Burkat, V. S.; Dudorova, V. S. J . Appl. Chem. USSR (Engl. Transl.) 1974, 47, 105. Tichy, J.; Douglas, W. J. M. Can. J. Chem. Eng. 1972, 50, 702. Tichy, J.; Wong, A.; Douglas, W. J. M. Can. J . Chem. Eng. 1972,50, 215. Ushida, S.; Chang, C. S.; Wen, C. Y. Can. J. Chem. Eng. 1977, 55, 392. Vunjak-Novakovic, G. V. Ph.D. Dissertation, Belgrade University, Yugoslavia, 1980. Vunjak-NovakoviE, G. V.; VukoviE, D. V.; Vogelpohl, A.; Obermayer, A. In Fluidization, Proceedings of the Third Engineering Foundation Conference on Fluidization; Hennikar, N. H., Matsen, J. M., Grace, J. R., Eds.; Plenum: New York, 1980. Vunjak-NovakoviE, G . V.; Littman, H. Presented a t the 8th International Congress CHISA '84, Praha-Czechoslovakia, 1984; Lecture 07.4. Vunjak-NovakoviE, G. V.; VukoviE, D. V.; Littman, H. Presented a t the 8th International Congress CHISA '84, Praha-Czechoslovakia, 1984; Paper Y 8.23. Vunjak-NovakoviE, G. V.; VukoviE, D. V.; Littman, H. 2nd. Eng. Chem. Res. 1987, following paper in this issue. Receiued for review October 8, 1985 Revised manuscript received February 11, 1987
