Pressure Drop and Flow Characteristics of Packed Fluidized Systems B. C. Pillai and M. Raja Rae* Department of Chemical Engineering, Indian Institute of Technology, Bombay, India
Hydrodynamic characteristics of packed fluidized beds were investigated using particulate materials of varying size, shape, and density, fluidized by air within the voids of stationary packings such as solid spheres, solid cylinders, and open-ended wire mesh screen cylinders. Owing to channeling, the experimental bed pressure drop in packed fluidized beds has been found to be less than that predicted on the basis of models proposed for packed fluidization. The slugging tendencies were, however, completely eliminated, and bed height fluctuations were reduced by the use of fixed packings in gas-fluidized beds. Results of minimum fluidization mass velocity and bed expansion have been discussed with respect to those of conventional plain fluidized beds.
A packed fluidized bed is a modification of the conventional plain fluidized bed, in which a particulate material is fluidized in the interstices of stationary packings. The addition of a fixed packing to a fluidized bed results in significant changes in the quality and performance of the fluidized bed. This is brought about by the facts that (i) the packing controls the size of the voids and (ii) the presence of the packing reduces the slugging tendencies by inhibiting the growth of the bubbles and also their coalescence. This, in turn, might improve the mixing characteristics of a fluidized bed and thus considerably influence wall-to-bed heat transfer rates. The limited published work in packed fluidized beds includes pressure-drop and bed expansion studies by Sutherland et al. (1963), Gabor (1966), Gabor, et al. (1964) Kang et al. (1967), and Ganapathy (1970). Gas mixing phenomena in packed fluidized beds have been studied by Gabor and Mecham (1964) and Chen and Osberg (1967), while solids mixing phenomena have been investigated by Gabor (1964) and Kang and Osberg (1966). Some studies on the lateral transport in packed fluidized beds have been carried out by Gabor (1965). I t has been reported that due to the random orientation of the fixed packing, the fluid passing through the bed takes a tortuous path resulting in preferential flow in the bed, thereby leading to severe channeling. Since the bubble growth is inhibited, the slugging tendencies are eliminated thereby resulting in reduced bed expansion. It has also been observed by several investigators that the minimum fluidization velocity in a packed fluidized bed is higher than in a plain fluidized bed. Theoretical Considerations The introduction of an internal fixed packing in a fluidized bed would alter the bed characteristics because of the variation in the filling up of internal voids of the packings and might hinder particle motion. Thus the minimum fluid bed voidage in a packed fluidized bed becomes an important parameter affecting the behavior of a packed fluidized bed. The variation in packed fluid bed voidage depends on the type of fixed packings such as solid packings or openended screen cylinder packings (Pillai, 1968). It is assumed that (i) all the voids of the stationary packings are completely filled with the particulate material, and (ii) a t the onset of fluidization the motion of the particles in the voids of stationary solid packings is completely restricted. On the other hand, the particles are assumed to be completely free to move about in the case of open-ended screen 250
Ind. Eng. Chem., Process Des. Dev., Vol. 15, No. 2, 1976
cylinder packings. Hence two different models have to be considered to describe the pressure drop characteristics of packed fluidized beds and the utility of these models can be checked with experimental results. A schematic representation of a packed fluidized bed is shown in Figure 1. It has been assumed that when the particles in the packed fluidized bed expand into the open tube, the total bed pressure drop would be equal to the sum of the pressure drops across the packed fluidized bed AP' and the plain-fluidized bed AP" above it. Thus
AP = (AP'+ AP")
(1)
(a) Packed Fluidized Bed with Solid Packing. The pressure drop contribution of the packed fluidized system is evaluated by assuming that the Blake-Kozeny equation for laminar flow through a fixed bed is valid for packed section and the bed voidage corresponding to G, after the onset of fluidization, is given by
1+€0
where t l = settled bed voidage, R = bed expansion ratio, and to = voidage of empty packed bed. Thus, the pressure drop per unit bed height will be (1- t )
12
Similarly, the bed pressure drop across the plain fluidized section would be
(5)
= (W"/LOS) = ( W - W')/SLO
-~ -
-
~ €i)(R ( 1 -1)~o R - 1)
(EO
+
(4)
(6)
The total bed pressure drop per unit length in a packed fluidized bed, with solid fixed packing will be given by the sum of AP'and AP".
Table I. Properties of the Particulate Materials Used Particulate material Aluminum powder Iron powder
Y
Y
I
of
CONDITONS
I BEFORE I 1 AFTER
u
Sand
11
PACKED-FLUIDIZED
Figure 1. Conditions of packed fluidized bed before and after the onset of fluidization.
Ps9
Mesh B.S.S.
g/cm3
181 128 90 181 128 90 59 253 181 128 90
(-72 + 100) (-100 + 150) (-150 + 2 0 0 ) (-72 + 100) (-100 + 150) (-150 + 200) (-240 + 300) (-52 + 72) (-72 + 100) (-100 + 150) (-150 + 200)
2.60
BED
ONSET OF FLUICIZATION, G L G ! m f ONSET OF FLUlDIZATlON AND EXPANSlON STATIONARY PACKING, G P G a
OVER
Av
do, p
6.78
2.65
@S
0.537 0.586 0.571 0.463 0.506 0.433 0.433 0.654 0.564 0.602 0.472
mm), (iii) brass cylinders of 6.4 X 6.4 mm size, and (iv) open-ended wire mesh screen cylinders of 12 X 12 mm and 6.4 X 6.4 mm size, made out of 12-mesh stainless steel screen nettings. The particulate materials used include aluminum powder, iron powder, and sand in the average particle size range of 59 w (-240 300 B.S.S.) to 253 w (-52 72 B.S.S.), and their properties such as density and shape factor are listed in Table I. Air at essentially atmospheric pressure was used as the fluidizing medium. The experiments were carried out by filling the test section with a weighed amount of fixed packings, introduced slowly through a 25 mm diameter plastic pipe, so as to avoid any damage to the supporting wire mesh, and later filling the entire voids of the fixed bed by the particulate material and allowing it to settle down within the packing voids by slowly cutting the air supply. The bed pressure drop was measured in terms of millimeters of water column for different air mass velocities in fixed and fluidized bed conditions (Pillai, 1968).
+
[
150
-
1' -
1
+ ( R - l)/to
Psdp2ds2
+ (1 -( t ot l+) ( RR -- 11)) P s t o
(7)
(b) Open-Ended Screen Cylinder Packed Fluidized Bed. In a screen cylinder packed fluidized bed, where there is complete freedom for particle motion, the bed pressure drop across the packed section would be AP'
(1 - d
P s
- to(1 - 6 1 ) P S (to + R - 1)
(8)
Thus from eq 6 for pressure drop across a plain fluidized section, and eq 8, the total bed pressure drop per unit length will be
(E)
=
to(1 - t (cg
i l ~ s
+ R - 1)
+
(1 - t i ) @ - 1)tops (tl R - 1)
+
(9)
Since the voidages of static beds of screen packings are in the range of 0.95 to 0.98, the particle motion within such a bed is almost as free as in a fluidized bed. The above equation can be modified by incorporating a channeling factor Cf, characteristic of the particulate material, which can be obtained from the experimental pressure drop studies in plain fluidized beds. Hence the pressure drop across a screen cylinder packed fluidized bed is
Experimental Work The schematic flow diagram of the setup used in the present work was shown in Figure 2. The apparatus essentially consists of a Pyrex glass column of 73.6 mm i.d. and 600 mm length, fitted between two calming sections by means of the tie rods. The bed was supported by means of a 300-mesh phosphor bronze screen, resting on a 3 mm thick perforated plate. The fixed packings used were (i) steel spheres (6.4 mm diameter), (ii) spherical glass beads (11.1
+
Results and Discussion Pressure Drop. Typical experimental bed pressure drop results obtained in packed fluidized beds for the iron powder-air system using fixed packings such as 6.4 mm diameter steel spheres, 11.1 mm glass beads, 6.4 X 6.4 mm brass cylinders and open-ended screen cylinders of 6.4 X 6.4 and 12 X 12 mm size have been presented in Figures 3 to 7 and compared with the results predicted by using eq 7 and 9. It is seen from the Figures 3 to 7 that the bed pressure drop under settled bed conditions shows linear variation with mass velocity. The experimental values of settled bed pressure drop are consistently lower than the calculated values for packed fluidized beds of solid packings. This corroborates the assumption that the voidage of an empty packed bed is not utilized completely by the particulate material and hence the settled bed voidage should be higher than in empty tube (e,) as shown in Table 11. This difference would lead to lower pressure drop per unit length, compared to the values calculated on the basis of complete filling of the voids. I t is obvious that the extent of variation in the bed voidage depends on the nature of the packing, and during the course of this work it was observed that complete filling of the voids of brass cylinders in the packed bed was more difficult than with other types of packed beds. Moreover, the ratio of the particle size to packing size appears to influence the extent of utilization of internal voids of the packing and thereby the bed pressure drop, the effect being more pronounced when this ratio is large. If, on the other hand, the ratio (dp/D,) is decreased as, for instance, the fluidization of iron powder of 90 p average particle size in large sized (11.1mm diameter) glass balls, uniform consolidation of particles in the voids of the packing is facilitated Ind. Eng. Chem., Process Des. Dev., Vol. 15, No. 2, 1976
251
Table 11. Comparison of Minimum Fluidization Velocity in Plain and Packed Fluidized Beds (Dt = 73.6 mm)
6.4-mm steel spheresiron powder
181 128 90 59
0.419 0.424 0.421 0.418
0.506 0.536 '0.552 0.587
0.595 0.617 0.652 0.659
136 79 65 38
191 152 87 34
325 193 159 93
1.74 1.27 1.83 2.72
2.43 2.22 2.74 2.13
6.4-mm steel spheres-sand
181 128
0.388 0.383
0.503 0.516
0.591 0.602
92 48
113 67
235 123
2.08 1.83
2.67 2.20
11-mm glass balls-iron powder
181
0.409 0.401 0.407 0.413
0.506 0.536 0.552 0.587
0.555 0.566 0.560 0.622
155
128 90 59
56 22
191 152 87 34
381 246 138 54
1.99 1.62 1.59 1.59
1.58 1.27 1.06 1.40
6.4 x 6.4-mm brass cylindersiron powder
181 90
0.395 0.340
0.506 0.552
0.588 0.609
102 52
191 87
258 154
1.35 1.77
2.22 2.30
6.4 x 6.4-mm brass cylinderssand
253
0.340 0.342
0.468 0.503
0.556 0.560
125 62
220 113
368 182
1.67 1.61
1.75 1.82
6.4 x 6.4-mm screen cylindersiron powder
181
0.955 0.966 0.965 0.965
0.506 0.536 0.552 0.587
0.570 0.603 0.620 0.635
310 233 143 56
191 152 87 34
323 243 149 58
1.69 1.60 1.72
1.87 1.93 1.70 1.62
12.0 x 12.0-mm screen cylindersalum. powder
181
128 90
0.973 0.970 0.975
0.598 0.586 0.577
0.640 0.646 0.622
350 192 80
204 112 52
360 198 82
1.76 1.77 1.59
1.48 1.84 1.57
12.0 x 12.0-mm screen cylindersiron powder
181 128 90 59
0.975 0.975 0.974 0.972
0.506 0.536 0.552 0.587
0.532 0.564 0.596 0.635
232 21 5 135 55
191 152 87 34
239 221 139 56.6
1.25 1.45 1.60 1.66
1.28 1.40 1.75 1.63
12.0 x 12.0-mm screen cylinders-sand
253
0.975 0.975 0.973
0.468 0.503 0.516
0.535 0.577 0.575
380 230 130
220 113 67
392 237 134
1.78 2.10 2.0
1.97 2.07 1.78
181
128 90 59
181
128
100
kll
Figure 2. Schematic flow diagram of the apparatus for pressure drop studies in plain and packed fluidized beds. 252
Ind. Eng. Chem., Process Des. Dev., Vol. 15,No. 2, 1976
1.71
4
I
,
'
,
I
01 = 7 3 4 m m
l
l
,
l
,
PA*I1lIcYLIIlE M A T E R I A L
(0)
dp-181
---
PACKING-6
0
0
300
CALCULATED 7
FROM E a -
"
0
'
4 X6.4MM
1
"
100
"
CRom
E?- 7
0
PACKING-6 "
600
(ma)
CALCULATED
-_
l
MICRONS
500
400
G , Kg /(hr)
, POWDER
BRASS C K I N D E R S
4x6 4MM
200
100
l IRON
"
i
BRASS CYLINDERS
~200
300
G, Kg,/(6r)Wzj
Fgure 5. Pressure gradient-flow diagram for packed fluidized beds of 6.4 X 6.4-mm brass cylinders. G,
KO/ chr)
cm'>
Figure 3. Pressure gradient-flow diagram for packed fluidized beds of 11.1-mmglass balls.
PACX/NO
1 100
- 6 4 nm
ST€€L SPIICRCS
P
I
zoo
300
G. K 9 / t n r l
*oO
cm5
Figure 6. Pressure gradient-flow diagram for packed fluidized beds of 6.4 X 6.4-mm screen cylinders. PACK/NG I
100
?6
4
mm
srccL sPncnE3
206
300
400
G, U g l ( h r ) (m2J
Figure 4. Pressure gradient-flow diagram for packed fluidized beds of 6.4-mm steel spheres. and thereby the bed pressure drop approaches closely the ideal value as seen in Figure 3a. From the onset of fluidization, it is seen that the observed bed pressure drop falls rapidly and attains a consistently uniform value for reduced mass velocity of 1.5 and I' is reduced by 40 to 70% from the above, although this A ideal fluid bed AP.However, the values of AP calculated by using eq 7 , on the assumption of severe restriction to parti-
cle motion in the voids of solid-fixed packings, shown in Figures 3 and 4, exhibit a decreasing trend with the increasing mass velocity. This might be explained as being due to the assumption of complete restriction to particle motion, which does not include the variation of bed voidage with mass velocity. Since the voids in a packed fluidized bed with solid fixed packings would always cause severe hindrance to particle motion, the assumption of complete restriction to particle motion would hold good and the pressure drop predicted on the basis of the assumption would remain consistently uniform, if the channeling factor is incorporated in eq 7 . I n the case of screen cylinder packed fluidized beds, the results of experimental bed pressure drop, as shown in FigInd. Eng. Chem., Process Des. Dev., Vol. 15, No. 2, 1976
253
PARNCYLAIE
3 rn
. IRON
POWDER
_ ___ - _ _ _ _ _ _Et-_ _ _ _- _ - _ _ _ _-__ 1.10 9
CALCULATED F R O M
p---
MATUIAL
-
A
300
0.
108
gr =
73 6 m m
(a)
PACKING--12.0 "
X 12.0MM l
"
PACKIN6-
6 ' 4 mm
S T E E L SPHLRdS
SCREEN C Y L I N D E R S
'
l
'
'
'
'
J
1
2
3
4
5
6
a
7
GlGmf
$Ld:D12CD
SOLID-
IRON
AP-5)
PACKING-12
=I
POWDER
MICRONS
0 X 1 2 0 MM
i
,."e
mm
I 1
1
SCREEN CYLINDERS
2
3
4
5
6
7
0
G I Qmf 0
100
200
300
400
G , K ~ / W im?
Figure 7. Pressure gradient-flow diagram for packed fluidized beds of 12 X 12-mm screen cylinders. ures 6 and 7, were found to remain constant with increase in mass velocity and were consistently lower by about 10 to 25% than the values predicted on the assumption of unhindered particle motion. For instance, for fluidization of iron powder of the smallest size used, d, = 90 w , in large-sized screen cylinders of 12 mm X 12 mm, the calculated and observed values of aP agreed well within 10%as seen in Figure 7. This shows that even in the case of screen cylinder packed fluidized beds, there is some degree of channeling due to maldistribution of the internal voids of the packing, especially with increased ratio of particle size to the packing size. Channeling. As could be seen from Figures 3 to 5, the experimental bed pressure drop immediately after the onset of fluidization is reduced by about 40 to 60% from the ideal bed pressure drop. This brings out the fact that the presence of solid packings, because of the nature of their random orientation in the bed, would cause preferential flow paths and thus lead to channeling tendencies in the bed. However, an examination of Figures 6 and 7 shows that the channeling tendencies are reduced considerably in a screen cylinder packed fluidized bed, compared to fluidized bed with solid packings. Minimum Fluidizing Mass Velocity. It is obvious that the changes in the voidage of the particulate material within the interstices of the stationary packings would greatly influence the value of the minimum fluidization velocity, and in the present work it was observed that the minimum fluidization velocity has shifted to higher values by about 5 to 45%depending upon the nature of the packing used. The increase in the minimum fluidization velocity can be expressed in terms of the bed voidage group obtained from the Blake-Kozeny equation. Thus for the same value of pressure drop, the ratio of the observed value of the minimum fluidization velocity to the theoretical value, based on complete filling of the packing voids, would be
The values of these two ratios for typical systems are listed in Table I1 and it could be seen that the expected value of 254
Ind. Eng. Chem.,Process Des. Dev., Vol. 15, No. 2, 1976
Figure 8. Variation of bed expansion ratio with mass velocity in a packed fluidized bed: (a) 6.4-mm steel spheres; (b) 6.4 X 6.4-mm brass cylinders.
112
108 K
104 ( a ) P A C K l f f i - 6 ~ 4 k M SCREEN CYLINDEK
loot
2
3
4
5
6
7
8
9
G , ' G ~
K
1
G
/ Gmi
Figure 9. Variation of bed expansion ratio with mass velocity in a packed fluidized bed: (a) 6.4 X 6.4-mm screen cylinders; (b) 11.1mm glass balls.
(G,//co) based on the packing free cross-section has not been attained, due to nonuniform distribution of the voids of the particulate material in the bed. This is especially true of packed fluidized systems with sphere packings. However, screen cylinder packed fluidized beds gave close agreement between the ratio of experimental minimum fluidization velocities and the ratio of the voidage functions. In the case of brass cylinder packed fluidized beds, however, no such definite relationship could be observed, probably due to the more complex nature of gas flow paths resulting from greater maldistribution of voids in solid cylinder packed fluidized beds. Bed Expansion Ratio. The bed expansion characteristics of packed fluidized beds expressed as bed expansion ratio were shown in Figures 8 and 9 as a function of re-
duced mass velocity for iron powder fluidized in the voids of packings such as 6.4 mm size steel spheres, 6.4 X 6.4 mm brass cylinders, 11.1 mm glass beds, and 6.4 mm X 6.4 mm open screen cylinders. It was observed in the present work that the bed expansion is generally reduced in the case of packed fluidized beds with solid fixed packings as compared to those of conventional plain fluidized beds. This confirms the observations of Gabor et al. (1964) on bed expansion for fluidization of copper shot (-140 200 mesh) in packed fluidized beds. Such a reduction in bed expansion could be attributed to the restriction to particle motion and the interference with the growth of the bubbles in a packed fluidized bed. However, because of the greater degree of freedom for particle circulation within the voids of screen cylinder fixed packings, it was observed that the decrease in bed expansion with these packings was less marked than in fluidized beds with solid fixed packings.
+
R = bed expansion ratio S = cross-sectional area of the column, m2 W = total weight of the particulate material in a plain fluidized bed, kg
W‘ = weight of the particulate material in the packed section of a packed fluidized bed, kg
W” = weight of the particulate material in the fluidized section of a packed fluidized bed, kg
Greek Letters w = dynamic viscosity, kg/(m)(h) ps =
solids density, kg/m3
cf = voidage of expanded bed
emf = minimum fluid bed voidage ’
€0
= voidage of packed bed
= initial voidage of the bed within the voids of the packing 61‘ = expanded bed voidage within interstices of packing & = particle shape factor €1
Literature Cited Nomenclature d, = average diameter of the particulate material, w D, = diameter of the fixed packing Gm, - minimum fluidization velocity in plain fluidized beds, (kg)/(h)(m*) Gm; = minimum fluidization velocity in packed fluidized beds, (kg)/(h)(-m2) LO = initial height of the fixed bed hp = total bed pressure drop hp’ = pressure drop contribution in the packed section of a packed fluidized bed AP” = pressure drop contribution in the open fluidized section of a packed fluidized bed
Chen. B. H., Osberg. G. L.. Can. J. Chem. E r g , 45, 91 (1967). Gabor, J. D., A.I.Ch.€. J.. 10, 345 (1964). Gabor. J. D., Mecham, W. J.. Jonke. A. A,. Chem. Eng. Prog. Symp. Ser.. 60 (47), 96 (1964). Gabor. J. D., Mecham, W. J., Ind. Eng. Chem., Fundam., 3, 60 (1964). Gabor, J. D.. A.I.Ch.E. J., 11, (1965). Gabor, J. D., Chem. Eng. Prog. Symp. Ser., 62, 32 (1966). Ganapathy, K. K., M. Tech. Thesis, I.I.T., Bombay, 1970. Kang, W. K., Osberg. G. L., Can. J. Chem. Eng., 44, 142 (1966). Kang, W. K.. Sutherland. J. P., Osberg, G. L., Ind. Eng. Chem., Fundam.. 8 , 499 (1967). Pillai, B. C.. Ph.D. Thesis, I.I.T., Bombay, 1968. Plllai, 8 . C.. Raja Rao. M.. Ind. J. Tech., 9, 77 (1971). Sutherland, J. P.. Vassilatos. G., Kubota, H.. Osberg, G. L., A.I.Ch.E. J., 9, 437 (1963).
Received for review January 31, 1975 Accepted November 4, 1975
Ind. Eng. Chem., Process Des. Dev., Vol. 15, No. 2, 1976
255
