Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 1, 1978
Subscript q = gaseous quench, gas phase i = feed conditions 1 = liquid quench, liquid phase m = metals s = sulfur 1 , 2 , 3 = corresponds to first, second, and third section of the reactor. Thus, q r n 2 is the dimensionless metal concentration is the dimensionless in the second section of the reactor.
33
concentration of sulfur, C$C,i, in the second section of the reactor after the first quench is added, etc.
Literature Cited Shah, Y. T., Mhaskar, R. D., Paraskos, J. A,, lnd. Eng. Chern. Process Des. Dev., 15,400 (1976).
Received for review September 27,1976 Accepted August 1,1977
Bed Expansion in Three-phase Fluidization Wen Y . Soung Hydrocarbon Research, Inc., Subsidiary of Dynalecfron Go., Lawrenceville, New Jersey 08648
Bed expansion data for three-phase fluidation have been determined for beds of commercial Go-Mo extrudate catalysts in n-heptane and nitrogen in Lucite tubes of 12.70 and 15.24-cm diameter. Gas and liquid velocities have been varied from 0 to 25.9 and 0.86 to 9.40 cm/s, respectively. Three cylindrical catalysts sizes were used, 0.0635, 0.1270, and 0.1600 cm diameter, and all were 0.4763 cm in length. Contraction of the bed due to gas injection was observed in the 0.0635-cm catalyst bed. With 0.1270-cm particles this phenomenon is much less perceptible, and no contraction at all with gas injection is observed in a bed of 0.1600-cm catalyst. An attempt has been made to isolate the gas injection effect on bed expansion from the effect of liquid velocity. Correlations are presented for the effect of gas velocity on bed expansion, based on particle Reynolds number, sphericity of the particle, and the liquid-to-gas velocity ratio. The results show that the catalyst bed will expand substantially upon gas injection if the liquid-to-gas velocity ratio is kept below 0.25, 0.55, and 0.65 for 0.0635, 0.127, and 0.1600-cm diameter catalysts, respectively.
Introduction Three-phase fluidization is a process used to contact gas, liquid, and solids. The solid particles are fluidized by an upward concurrent flow of gas and liquid. The industrial applications of three-phase fluidization have been reviewed by 0stergaard (1971). The most important industrial application of the process is the heterogeneous catalytic hydrogenation of residual oils or coal slurry for the removal of sulfur and the production of hydrocarbon distillates by hydrocracking (such as the H-Oil and H-Coal processes). An important property of the three-phase fluidized bed is bed expansion as it relates to the volume of the reactor and the mean residence time of the liquid phase. The height of an expanded catalyst bed is controlled by liquid velocity in the industrial applications of a three-phase system. Hence, it is important to be able to predict the catalyst bed expansion at various liquid and gas velocities. The goal of the present work was to study the effects of catalyst particle size and liquid gas velocities on catalyst bed expansion in a three-phase fluidization system with commercial Cc-Mo extrudate catalysts. An attempt was made to isolate the gas injection effect on bed expansion from the effect of liquid velocity. Previous Work Richardson and Zaki (1954) proposed that the liquid voidage of a particulate fluidization system can be correlated by =
(ullu,)n
similar observations. Ostergaard and Theison (1966) studied the effect of particle size and bed height on bed contraction. They concluded that bed contraction is greater in beds of small particles than in those with large particles, and contraction increases with increasing bed height. Experimental Section The columns were Lucite tubes with 12.70 and 15.24-cm i.d., and both were of 5.18 m height. The tubes each consisted of four 1.22 m long Lucite sections, flanged and bolted together as shown in Figure 1.The gas distributor, which was also used as a bed support, was made up of 40 mesh screens supported by an expanded metal grid. The liquid and gas used in this study were n-heptane and nitrogen. The solid particles were commercial Co-Mo extrudate catalyst (American Cyanamid Co.). Three different sized samples of catalyst (0.0635, 0.1270, and 0.1600 cm diameter) were used. The temperature of the system was carefully controlled a t 12.8 "C. Table I gives the physical properties of the experimental solids. The height of the dead catalyst bed was recorded at the beginning of each run. The initial static bed heights were 243.8 and 172.7 cm for the 12.70 and 15.24-cm columns, respectively. Two sets of conditions were tested, one being particulate fluidization and the other three-phase fluidization. The height of the expanded bed was recorded when steady-state conditions were reached. If conditions were such that slugging existed at the top of the bed, an average of the maximum and minimum heights was taken.
(1)
Daksinamurty e t al. (1971, 1972) suggested correlations in terms of dimensionless groups, U J U t and g l U g / o . The phenomenon of bed contraction in three-phase fluidization was reported by Turner (1963). Wolk (1962) reported
Results a n d Discussion Figures 2, 3, and 4 give the experimental data for three catalyst sizes: 0.0635, 0.1270, and 0.1600-cm diameter, respectively. An attempt was made to correlate the data and to
0019-7882/78/1117-0033$01.00/0 0 1978 American Chemical Society
34
Ind. Eng. Chern. Process Des. Dev., Vol. 17, No. 1, 1978
t I
4I
/
II II
I
I-T
!I r l c
I
U I
0
0 1
0 2
0 1
0 4
0 5
0 6
UllUt
Figure 1. Fluidization apparatus with 12.7 or 15.2-cm i.d. tubes: a, 12.7 or 15.2-cmi.d. tube; b, c, gas rotameters; d, HZsupply;e, check valve; f, thermometer;g, h, orifice runs; i, centrifugal pump; j, liquid supply; k, vent. separate the effect of gas velocity on bed expansion from the effect of liquid velocity. The data correlation was based on sections (a), (b), and (c) below. (a) Correlating the P a r t i c u l a t e Fluidization Data in t e r m s of H - Ho/Hvs. U , / Ut as Shown in F i g u r e 5. The data are divided in two groups; one of them with a low bed expansion corresponding to a packed fluidization state and the other with a high bed expansion corresponding to a true fluidization state. The break point in Figure 5 or Figure 6 corresponds approximately to the minimum fluidization state. The data of particulate fluidization have been correlated as eq 2. The residual variance indicates that (Ho/H)ugf0can be predicted within &3 % and f15% (95% confidence interval) by eq 2a and 2b, respectively, over the range of the data.
(b) Evaluating the Effect of Gas Velacity on Bed Expansion. A factor called the F value is defined as the ratio of (Ho/H) in a three-phase system to (Ho/H)u,=o (calculated
Figure 2. Ho/H vs. U1/Ut;0.0635-cm diameter X 0.4763 cm length catalyst. Numbers on the curves are U,/Ut: 0 = 0.00; o = 0.12; 0 = 0.40; A = 0.45; = 0.59; = 1.01; E = 1.73; e = 2.26. with a zero gas velocity). The value of (Ho/H)u,=o can be calculated from eq 2 a t the same liquid velocity as in the comparable three-phase system. Hence
(3)
Figure 6 shows the F curves for 0.0635,0.1270, and 0.1600-cm catalyst. The F value indicates the effect of gas velocity on bed expansian. A value greater than 1.0 implies that the bed will contract or slump when the gas is introduced into a particulate fluidization bed. A value less than 1.0 means the bed will expand further when the gas is introduced into a particulate fluidization bed. For the bed of 0.0635-cm catalyst particles, the bed contracted substantially upon gas injection and stayed that way for a fairly wide range of the gas velocity. When the value of Ul/U, reached about 0.8 the bed started to expand gradually. However, the bed height was still less than would be found in a comparable particulate fluidization system. Beyond this point, the height of the catalyst bed then continually increased with increasing gas velocity. When UIIU,was about 0.3 or less, the bed height was greater than that of a particulate fluid-
Table I. Physical Properties of the Experimental Solids
Solid U.0635-cm diameter cylindrical extrudates 0.1270-cm diameter cylindrical extrudates 0.1600-cm diameter cylindrical extrudates
Effectiveu diameter, d, Sphericity,b True density,? cm 4% g/cm"
Terminal velocity in heptane,d Ut, cm/s
Diameter, D, cm
Length, L, cm
0.0635 f 0.0051
0.4763 f 0.3175
0.2785
0.628
1.37
11.44
0.1270 f 0.0076
0.4763 f 0.3175
0.3509
0.744
1.35
16.35
0.1600 f 0.0076
0.4763 f 0.3175
0.379
0.780
1.39
18.50
Diameter of sphere having the volume of the particle; d, = ( 3 D 2 L / 2 ) 1 /for 3 a cylindrical extrudated catalyst. Defined as (surface of sphere having same volume as particle)/(surface of the particle); & = dP2/[(D2/2) + D L ] for a cylindrical shape catalyst; D = diameter of a cylindrical catalyst; L = length of a cylindrical catalyst. Heptane saturated particle density. Calculated from information in Brown et al. (1950).
35
Ind. Eng. Chem. Process Des. Dev., Vol. 17,No. 1, 1978 LOO-,
'
0 . 3 L
0
I 0. I
'
I
'
1
0.2
0.3
I
I
1
I 0.5
0.4
1
I 0.6
"I'Ut
Figure 3. HoIH vs. UI/Ut;0.1270 cm diameter
0.4763 cm length catalyst. Numbers on the curves are Ug/Ut: 0 = 0.00; = 0.10; 0 = 0.16; o = 0.19; A = 0.26; = 0.30; = 0.67; = 1.06; e = 1.40.
,025
X
I
I
0
0.2
,
1 , 0.1
1
1 0.4
1
.
0.5
1 .6
I
I .O
I
I 1.0
UI'Ut
Figure 5. Data correlation for particulate fluidization system of
0.0635, 0.1270, and 0.1600-cm diameter catalysts in 12.7 or 15.2-cm i.d. tube.
I
0.9
I \ :.66
For the catalyst bed of 0.1270-cm particles, the bed contraction is much less perceptible upon gas injection. For the bed of 0.1600-cm particles, the particulate fluidization bed always expands as gas is injected. However, the extent of expansion is not substantial until U J U , is reduced below 0.55 or 0.65 for 0.127 or 0.160-cm particles, respectively. A t this point, the turbulence again increases markedly when gas velocity increases, and the solids have top-to-bottom motion. The slopes of the sloped-line sections of the curves in Figure 6 become larger as the catalyst particle size becomes smaller. This indicates that an increased gas velocity causes the amount of increment of catalysts bed expansion to increase as the particle size decreases. Bstergaard and Theison (1966) reported similar observations. (c) Correlating the F Values in Terms of Particle
Reynolds Number, Rep,the Liquid to Gas Velocity Ratio, and Sphericity of the Particle, The resulting equations are
F = 1.50 + 0.16 In - - 0.065 In (&Re,)
);:(
(0.06 5 u 1< 0.6)
(4a)
u, and ization system. It was observed experimentally that when U1/U, was reduced below 0.25, the agitation of the system increased markedly and the solids had some top-to-bottom motion. This top-to-bottom mixing would be useful as a way to level out any reactor temperature gradients. It was observed experimentally that the gas bubbles are small and spherical in shape when gas velocity is low. Because of coalescence, the bubbles change from small spherical to large spherical caps and then to gas slugs when gas velocity increases. The particulate mixing was observed experimentally to be increased from a local scale to a full scale top-tobottom mixing when gas velocity increases.
F = 2.09 - 0.17 In (& Re,)
(0.6 I U1 < 5 )
u,
(4b)
The residual variance from this correlation indicates that F can be predicted within f20% (95%confidence interval) over the range of the data. Note that preliminary data of small columns (1.6,2.5, and 3.8 cm i.d.1 show that the wall effect appears to be important on bed expansion. However, no quantitative information can be given at this stage. No wall effect on bed expansion existed (or was negligible if any) for 12.70 and 15.24-cm i.d. columns, because no difference was observed experimentally between the results gathered.
Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 1, 1978
38
i.1)
c
1.2
-
P
-
-
00
-
1.0
v
0
0
-
-
0.0
-
-
0.6
1
1
0.03
I
0.05
I
I l l 1
I
I
0.07 0.10
I
i l l l l
I
0.40
0.20
I
0.60 0.00 i . 0
i
2.0
1
I
4.0
I
i I l l J 6.0 3 . 0 10.0
UI/Ug
Figure 6. F
vs.
U,/U,: 0,-,
0.0635-cm diameter catalyst; A ,
- - -, 0.1270-cm diameter catalyst; 0 ,- - -, 0.1600-cm diameter cata-
lyst.
0.9
0.0
1
Summary The bed expansion in a three-phase fluidized bed can be predicted by eq 2 , 3 , and 4.Their application has been illustrated in the "Application" section.
0.7
0.6 0.5
n-no H
0.4
Acknowledgment 0.3
The author wishes to thank Hydrocarbon Research, Inc., a subsidiary of Dynalectron Co., for granting permission to publish this paper.
Nomenclature D = diameter of a cylindrical extrudate catalyst, cm d, = effective catalyst diameter (see definition in the footnote of Table I), cm
F = definition in eq 3 Ho = height of dead catalyst bed, cm H = height of expanded catalyst bed, cm L = length of a cylindrical extrudate catalyst, cm n = constant in eq 1 U , = superficial gas velocity (based on empty cross section of tube), cmls
U I = superficial liquid velocity (based on empty cross section
O0.03 ' O I 0.025
I
I
0 . l ~
0.20
I
I
I
0.3
0.4
I
I 0.6
I
I 0.8
I
I 1.0
Ul'Ut
Figure 7. Convenient chart for 0.1600-cm diameter catalyst; numbers on the curves are U,/ Ut.
Application Example of the Correlation Results A convenient chart can be plotted for any catalyst of interest for a practical application. The chart can be prepared by the following steps. (a) Assume a series value of UlIUt. Get the value of (HolH)u,=o for each corresponding value of U1/Ut from eq 2 or Figure 5. (b) Assume a series value of Ug/Utfor each value of UIIUt. Calculate UI/U,, ( = UllUJU,lUt), and find an F value for eq 4.Compute HdH,namely (Ho/H)u,=o X F. (c) Calculate ( H - Ho)/H, and percent bed expansion, (H- Ho)/Ho X 100%. (d) Repeat steps (b) and (c) for another value of (UllUJ with Ug/Utas a parameter until all the values of (LJJlYt)are done. Figure 7 shows a plot of percent bed expansion vs. U1/Ut with U,/Ut as a parameter for the 0.1600-cm catalyst bed. The chart is ready for predicting bed expansion a t various gas and liquid flow rates.
of tube), cm/s
Ut = terminal velocity of the particle in the liquid phase, cmls Re, = particle Reynolds number, d p1 Ut/gl €1 = liquid voidage of a particulate ffuidization bed 4s = sphericity of catalyst particle; see definition in the footnote of Table I p1 = density of liquid, g/cm3 p1 = viscosity of liquid, P u = surface tension of liquid, dynlcm
Literature Cited Brown, G. G., et al., "Unit Operations", Wiley. New York, N.Y., 1950. Daksinamurty, P., Rao. K.. Subbaraju, R. V., Subrahmanyam,V., Ind. Eng. Chem. Process Des. Dev., 11, 316 (1972). Dakshinamutry P., Subrahmanyam, V., Rao, J. N., Ind. Eng. Chem. Process Des. Dev., 10, 322 (1971). Bstergaard, K., "Fluidization", Chapter 16, J. F. Davidson and D. Harrison, Ed., Academic Press, New York, N.Y., 1971. Bstergaard, K., Theisen, P. I., Chem. Eng. Sci., 20, 165 (1965). Richardson, J. F., Zaki, W. N., Trans. Inst. Chem. Eng., 39, T212 (1954). Turner, R.. "Fluidization Papers Joint Symposium, London", 1963. Wolk. R., Master's Thesis, Polytechnic Institute of Brooklyn, 1962.
Receiued for review December 23, 1976 Accepted June 30,1977
