laboratory Prediction of Flow Properties of R O B E R T D I E K M A N A N D W. L. FORSYTHE,J R . RESEARCH DEPARTMENT, STANDARD O I L co. ( I N D I A N A ) W . H I T I N G ,I N D .
Successful circulation of fluidized solids as a dense phase depends on keeping t h e solidsaerated. If allowed t o deaerate, t h e solids rapidly losefluidity. A laboratory method has been developed with used cracking catalysts for predicting t h e dense-phase flow of fluidized solids in commercial-scale standpipes. T h e method involves measurements of t h e viscosity and t h e deaeration rateof afluidized bed. Solidsthat increase moreslowly i n viscosityduring deaeration have better flow quality. Coarse particles of high density adversely affect flow behavior. Wide size distributions are not needed t o maintain good flow; average particle size i s t h e controlling factor.
I
N THE fluidization of solids with gases, such as in catalytic
cracking, the mixture assumes many properties of a fluid. Use is made of the fluidity of the mixture to circulate large amounts of solids rapidly from one vessel to another. Ciroulation is usually accomplished with few mechanical parts by means of carrier lines and standpipes (5). In carrier lines, a low ratio of solids t o gas, or a “dilute phase,” is maintained; in standpipes, a high ratio, or “dense phase,” is used.
FLUIDIZED BED
Deaeration causes flow variations in standpipes. A tendency toward deaeration is inherent in any fluidized system (8), coarse particles of high density deaerating most readily. Even in a fully aerated bed, partial deaeration may occur between gas bubbles as the solids move away from an aeration point. Inability to maintain smooth dense-phase flow in standpipes may limit the throughput of commercial processing units. A laboratory test has been needed for evaluating deaeration characteristics so that flow variations in the plant can be better understood. Such a test might also predict the flow of solids before plant use. The authors have carried out a study to develop such a test and to determine the effects of physical properties on flow behavior. Cracking catalysts, both natural clay and synthetic silica-alumina, were used as the solids. DEVELOPMENT O F TEST
DISTRIBUTOR
ORlD
AERATION
CAS
GAS
CARRIER LINE
Y
Figure 1.
DENSE
PHASE
DlUllE
PCUSE
Standpipe for Circulation of Fluidized Solids
A diagram of a standpipe is shown in Figure 1. Solids enter the top through an enlarged opening, pass downward, and discharge through a slide valve. Aeration gas is injected a t intervals along the standpipe to keep the solids fluidized by counteracting the tendency to settle or deaerate. Flow quality in a standpipe is indicated by the magnitude and the steadiness of the pressure build-up above the slide valve. As circulation through the slide valve increases, the pressure build-up above the slide valve decreases. If it drops too low, the solids may be meeting added frictional resistance caused by partial deaeration. If it fluctuates, a surging flow is probably occurring because of temporary deaeration and “bridging” of solids; this condition is analogous to “stick and slip” flow ( 2 ) .
Deaeration i n a fluidized bed can be observed upon shutting off the flow of aeration gas. Deaeration rate can be determined by timing the settling of the bed, but preliminary studies have shown this rate to be inadequate for fully characterizing flow behavior. Changes in fluidity of the solids during deaeration should also be a factor. Of the published methods for characterizing the fluidity of solids (4, 5, 7, 9, IO), the viscosity concept (7) appeared most adaptable for correlation with deaeration behavior. Because deaeration proceeds rapidly, direct measurements of viscosity during deaeration are impracticable. However, the change in fluidity during deaeration can be obtained indirectly from separate measurements of viscosity and deaeration. This indirect method requires the assumption that the viscosities during deaeration are the same as those obtained during aeration. I n practice, viscosities are measured a t steady-state aeration rates that cover the transition from a fluidized bed to a partially settled bed. At these low rates, the greatest changes in fluidity occur ( 7 ) and the effects of physical properties of the solids are most pronounced. Viscosity and deaeration rate are more readily measured with a confined fluidiaed bed than with a bed having continuous external cirrulation of solids. TEST METHOD
The study was made with used cracking catalysts that showed differences in flow behavior in three commercial units. The particle densities of these catalysts varied from 1.3 to 1.6 grams per ml., as determined by mercury displacement. Particle sizes ranged from 30 to 160 microns, as measured by Roller analysis (6, 11) and screen analysis ( I ) . Average particle sizes were obtained from a logarithmic probability plot of cumulative weight
1174
per cent finer than a given micron size; each average was taken as the micron size at the 50% point. Prior to being tested, the catalysts were saturated with gaseous ammonia to minimize accumulation of electrostatic charges. Such charges cause abnormal agglomeration and loss of fluidity a t room temperature, as used in the test, but are not observed a t cracking temperatures.
tion can be taken at common void fractions and cross-plotted to give the correlation of viscosity with deaeration time. The cross plot from Figure 3 is shown in Figure 4. The deaeration time in seconds required to reach a viscosity of 400 can be taken as a measure of flow quality. This useful concept is labeled the fluidity index; the higher the index, the better the flow quality. DISCUSSION OF METHOD
BROOKFIELD VISCOMETER I
I
I
Y
u,
DISTRIBUTOR
Figure 2.
Equipment for Measurement of Viscosity and Deaeration Rate
Several features of the equipment provide a uniform fluidized bed and sensitive laboratory measurements of fluidization behavior. The RVF Brookfield viscometer is particularly suited for viscosity measurement in a fluidized bed because the spindle is not supported by any lower bearing that could become clogged with particles. The cylindrical-screen spindle minimizes interference with particle behavior. Any turbulence set up by the spindle would affect particle behavior a t the point of viscosity measurement, Viscosity as measured with this equipment is related in Figure 5 to the behavior of a typical catalyst bed. At aeration rates below 0.02 foot per second, the gas channels through the bed. Small bubbles then gradually appear a t the center of the bed, while channeling continues nearer the wall. Between 0.020 and 0.035 foot per second, channeling ceases as bubbling becomes more pronounced, but little mixing occurs. When the bubbles grow large enough, at 0.035 foot per second, the mixing becomes turbulent and the viscosity approaches a constant value.
The arrangement of the equipment for measuring viscosity and deaeration rate is shown in Figure 2. 0.6
A 4-inch glass pipe is used for the aeration column. A canvas distributor supports the catalyst bed and gives uniform distribution of the aeration gas. A Brookfield Model RVF viscometer with a modified screen spindle is used for viscosity measurement, The spindle is a cylinder of 10-mesh wire screen, open a t both ends and fastened to the shaft by three wires. The torque required for rotation of the spindle is measured by the displacement of a torsion spring and is translated directly to a dial reading on an arbitrary scale of 0 to 500. One scale unit is equivalent to about 1.3 centipoises. Sufficient catalyst is weighed into the glass pipe to give a bed height of 8 inches durinz aeration. The viscometer is clamped in level position. Viscosities are measured a t a series of aeration rates up to 0.04 foot per second with the bed level kept constant a t 8 inches by addition or withdrawal of catalyst. The weight of the catalyst in the bed a t each aeration rate is recorded. Air is used as the aeration gas. The viscometer is removed from the aeration column, and deaeration rate is determined by timing the settling of thd catalyst bed. The catalyst is aerated at a rate of 0.04 foot per second with a bed height of 8 inches. Aeration is then stopped instantaneously with a quick-action pinch clamp. Times are measured with four stopwatches as the bed settles past levels of 7.5, 7.0, 6.5, and 6.0 inches. At least three determinations are averaged. Deaeration times usually are repeatable to 0.2 second.
a
-
Viscosity and deaeration time can be related in terms of the void fraction of the catalyst bed, calculated as follows: total bed volume Void fraction =
of particles in - weightparticle density
bed
total bed volume
During the viscosity measurements, the total bed volume is kept constant, while the weight of particles varies with the aeration rate. During the measurements of deaeration rate, the total bed volume decreases as the bed settles past the given levels, but the weight stays constant. Viscosity and deaeration time are plotted separately as functions of void fraction, as shown in Figure 3 for a typical catalyst. Although the void fraction drops as the maximum aeration rate of 0.04 foot per second is reached, no such inflection is found during deaeration. The trend of viscosity during deaeration is therefore assumed to follow the dashed portion of the upper curve. Values of deaeration time and assumed viscosity during deaeraJune 1953
0
100
VISCOSITY 200 300
400
500
I
0000-FLOWINO CATALYST FROM UNIT A
AERATION VISCOSITY WITH INCREASINQ AERATION
OURINQ
0
DEAERATIOW
a >
t
-
-
"' -I DEAERATION TIME, SEGONDS FROM AERATION AT 0.04 FT./SLQ.
Figure 3. Viscosity and Deaeration Measurements
The rapid decrease in viscosity a t the lower aeration rates is related to the greater free path of the particles as they are lifted and separated. Because of this rapid decrease, viscosities in the range of 300 to 500 are most sensitive to differences in flow quality. At these high viscosities and low aeration rates, the bed has the same appearance as during deaeration. The assumption of identical viscosities during aeration and deaeration is therefore reasonable in this range. The precision of the viscosity readings varies with the condition of the aerated bed and is best a t intermediate viscosities when the bed is quiescent. A t higher viscosities, the bed is less stable and less reproducible; a t lower values, the spindle is affected by bed turbulence. The deviations found a t the measured viscosity levels are: f l 0 a t 400, f 5 a t 250, and f 1 3 a t 50. Although measurement of deaeration time is started at a relatively low aeration rate, the results are applicable to deaeration in plant operation where higher aeration rates prevail. Measurements starting a t an aeration rate of 1.0 foot per second show that deaeration takes place in two stages: an instantaneous ini-
INDUSTRIAL AND ENGINEERING CHEMISTRY
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Table I .
Physical Properties and Fluidity of Used Cracking Cata Iysts Flow i n Commercial Unit Unit A Vnit B
-
Particle-Size Good Distribution, % Roller analysis 0-20 microns 0.3 0.0 20-40 microns 19.7 9.7 40-80 microns 5 2 . 0 57.1 A4bove80microns 2 8 . 0 33.2
Unit C
Poor
Good Poor
Good Poor
0.0 9.9 59.4 30.7
0.0 0.0 14.8 6.1 51.7 57.3 33.5 36.6
0.0 18.2 43.4 38.4
0.0 13.3 43.9 42.8
been determined, although particle shape affects the results of Roller analysis. B quantitative distinction between particle size and particle shape by microscopic examination should be of value and warrants futLre study. EFFECT OF PARTICLE SIZE ON FLUIDITY
Of the physical properties of a catalyst that may affect fluidity, only particle size has been studied in the present work. To improve dense-phase flow in most con-imercial fluid cracking units, particle size can be changed more easily than particle density. 500
Average particle size, microns Particledensitg,g./ml. Fluidity index
66 1.60 2 . 9 2 . 4 1 . 7
59 1.58
67 1.51
65 71 1.36 1.41 3 . 9 2 . 8
67 73 1.43 1.49 6 . 7 2 . 3
GOOD-FLOWING CATALYS FROM UNIT A
0
In \a
400
- 300
BUBBLING
VIGOROUS BUBBLING
LITTLE MIXING
TURBULENT MIXING
v)
0
8
0
5001
d
'" 200 >
/
40
01 0
-
GOOD FLOWING CATALYST FROM UNIT A
-
0 0
1
2.
3
4
DEAERATION TIME, SECONDS Figure 4.
Correlation of Viscosity with Deaeration T i m e
tial settling when the bubbles of aeration gas escape from the bed, and a slower settling as the dispersed aeration gas flows out of the spaces around the particles. The second stage appears to be controlling in dense-phase flow and is covered in the test by deaeration from 0.04 foot per second. APPLICATION OF METHOD
The flow properties of fluid cracking catalysts predicted by the laboratory test are consistent with trends of flow behavior in commercial fluid cracking units. Table I gives the physical properties and the flow quality of seven catalysts from three commercial units, These units were not equipped with Cottrell precipitators; hence, the samples do not contain particles finer than about 20 microns. For each unit, the fluidity index of the catalysts correlates with observed flow quality. The fluidity index of the good-flowing catalysts from unit A is about the same as that of the poor-flowing catalyst from unit B. Unit A differs in design and size from unit B and apparently can tolerate a catalyst Kith poorer flow properties. In unit A, the poor-flowing catalyst has a higher particle density than one good-flowing catalyst and a coarser particle size than the other, In units B and C, the poor-flowing catalysts are both denser and coarser. Higher particle density causes higher viscosity a t a given aeration rate, because the particles are closer together and the spindle meets increased friction. Coarser particle size causes faster deaeration and hence faster loss of fluidity. The effect of particle shape for these catalysts has not
1176
Figure 5.
2
I
I
I
I
0.01
0.02
0.03
0.04
(
5
AERATDN RATE, FT/SEC. Bed Behavior with Increasing Aeration Rate
The improvement in fluidity with finer particle size is shown in Table I1 for four fractions of a typical catalyst. The fractions were obtained by elutriation as progressively coarser cuts. Comparison of the average particle size and the fluidity index shows the sensitivity with which the test detects small changes in particle size. For less than a threefold decrease in average particle size, the fluidity index increases almost seven times.
Table I I .
Effect of Particle Size on Fluidity
Particle-Size Distribution, % Roller analysis 0-20 microns 20-40 microns 60-80 mjcrons 40-60 microns Above 80 microns Screen analysis Through 150-mesh 100-1 50-mesh 80-100-mesh On 80-mesh Avefage partiole size, microns Fluidity index
Catalyst Fraction Total C
A
B
0.0 2.2
0.0 3.3
0.0 14.8
0.4
90.6 8.1 0.9 0.4 90 1.3
-D
0.0 8.3
j51.7 33.5
:;"2 0 . 7
0.4 63.4 }35..5
100.0 0.0 0.0 0.0
95.8 3.6 0.4 0.2
100.0 0.0 0.0 0.0
100.0 0.0 0.0 0.0
67 3.4
65 3.9
52 6.0
36 8.6
27:;
}24.0 73.8
0.7
Fraction B, with nearly 100% 40- to 80-micron particles, has nearly the same fluidity index and average particle size as the total catalyst with particles from 25 to 150 microns in size. Wide size distributions therefore are not required for maintenance of good flow. Average particle size controls fluidity. Much better flow in standpipes can be obtained by reducing average particle size by as little as 10 microns. However. a practical 10%-erlimit is imposed by the tendency of particles below 30 microns to agglomerate or to be lost from the unit. Severtheless, the ability of the test to show large differences in fluidity for small changes in particle size makes it helpful for predicting the dense-phase flow quality of solids prior t o plant use.
INDUSTRIAL A N D E N G I N E E R I N G C H E M I S T R Y
Vol. 45, No. 6
ACKNOWLEDGMENT
The authors appreciate the assistance of J. D. Nielsen and A. W. Ries in the experimental work.
(4) Kramers, H., Chem. Eng. Sci., 1, 35 (1951). (5) Leva, M., Grummer, M., Weintraub, AI., and Pollohik, M., Chem. Eng. Progr., 44, 619 (1948). (6) Matheson, G. L., Proc. Am. PetroleumInst., 27 [III],18 (1947). (7) Matheson, G. L., Herbst, W. A., and Holt, P. H., IND.ENG. CHEM.,41, 1099 (1949).
LITERATURE CITED
Fcrsythe, W. L., Jr., and Hertwig, W. R., IND.ENG.CHEM.,41, 1200 (1949). Gregory, s. J *A P p l . 2r SUPPI. Issue No* p* (1952). Gunness, R. C., Chem. Eng. Progr., 49, 113 (1953). *.I
Chern.l
( 8 ) Morse, R. D., Ibid., 41, 1117 (1949). (9) Morse, R. D., and Ballou, C. O., Chem. Eng. Progr., 47, 199
(1951). (10) Parent, J. D., Yagol, N., and Steiner, C. S., IFid., 43,429 (1947). (11) Roller, P. S., Trans. Am. SOC.Testing &futeriaZs,32, 607 (1932). ACCEPTEDApril 6, 1953 RECEIVED for review February 24, 1953.
Gas-Flow Patterns in Beds of Fluidized Solids E. R. GILLILAND, E.A. MASON,
AND R. C. OLIVER M A S S A C H U S E T T S INSTITUTE O F T E C H N O L O G Y , C A M B R I D G E , MASS.
T h e effect of t h e flow pattern of t h e gas is important in fluidized-solid reactors. I n previous studies on t h e residence-time history of t h e fluidizing gas, a technique was developed whereby t h e effect of nonuniformities in gas flow can be predicted for first-order homogeneous reactions. In order t o study t h e influence of gas flow patterns on chemical reactions of higher order, t h e oxidation of nitric oxide, a third-order reaction, has been carried o u t in small fluidized beds. T h e measured conversion of reactants has been compared t o t h a t predicted from known kinetics, residence-time data, and certain assumptions regarding t h e gas-flow pattern. T h e experiments were carried o u t in glass reactors using spherical glass beads (Scotchlite) as t h e fluidized solid. Preliminary work showed t h a t t h e presence of these beads had a negligible effect on t h e reaction under t h e conditions chosen. T h e progress of t h e reaction was measured photometrically. For comparison, t h e residence-time history of gas flowing through t h e same apparatus was determined. It was found t h a t knowledge of t h e residence-time distribution of t h e gases flowing through a reactor provides a good method for estimating t h e effect of gas-flow pattern on homogeneous reactions in fluidized beds.
F
1
LUIDIZED solid techniques have found widespread use in the chemical industry on both a laboratory and a plant scale. The solid mixing which takes place in fluidized beds promotes a degree of temperature control not readily obtainable in other types of chemical reactors. However, this solid motion causes mixing of the products of reaction with the entering reactants and permits some bypassing of reactants, so that the time for reaction is not uniform for all entering gas. Gas mixing and bypassing lead to a decreased reaction rate and nonuniformity of products, both of which are generally undesirable. The effect of these irregularities in gas flow becomes increasingly serious as the desired degree of conversion increases. Several experimental techniques have been employed in a n attempt to determine the nature and extent of the gas mixing in fluidized beds. Residence times of the gas flowing through small fluidized beds have been investigated by first supporting the solids with a mixture of tracer gas and air and then determining the composition of the gas leaving the bed as a function of time after the addition of tracer gas was discontinued. Studies concerning the back-mixing of the gases were made by injecting tracer gas into fluidized beds and analyzing gas samples taken from various points in the bed. The results of this work on the
Table I. Size No.
Inch
;mm.
0.24 0.0178 0.0061
11 13
0.0040
15
0.00275
June 1953
Properties of Glass Beads
Diameter Microns % Deviation 6000 452 155 I 02 70
6
7
6 12
Abs. Density, G./Co. 2.24
2.78,246 2.46 2.43 2.43
gas mixing in fluidized solids have been applied t o the prediction of the chemical conversion to be expected in reactors having similar flow patterns. I n a third experiment, a chemical reaction has been carried out in laboratory scale fluidized beds in order to determine directly the effect of gas mixing on chemical reaction and t o evaluate the methods that have been proposed for predicting the effect of known gas mixing on chemical conversion. GAS-M I X I N G STUD1 ES
I n the residence-time studies, a bed of finely divided solids was suspended in a mixture of tracer gas and air. After the composition in the bed had become uniform, the flow of tracer gas was abruptly halted and a series of samples was taken of the gas as it left the top of the bed, These gas samples were analyzed for tracer content. Helium and carbon dioxide were used as tracers and all samples were analyzed using a thermal conductivity apparatus. Studies were made in Lucite columns having internal diameters of 3 and 4.5 inches. The ratio of the height to the diameter of the fluidized bed, LID,was varied from 7.8 to 23:3. Various sizes of Scotchlite glass beads manufactured by the Minnesota Mining and Manufacturing Co. were used (Table I). Other data were taken using a silica-alumina cracking catalyst manufactured by the American Cyanamid Co. The method of size analysis has been reported (6). With the exception of the use of carbon dioxide as a tracer, details of the procedure and of the apparatus used have been reported (5). The data presented in Figure 1 were obtained with helium a8 a tracer, using glass beads in the 4.5-inch column a t a bed height of 35 inches, but in general respects are similar to those obtained under other operating conditions. The value of C/Co is plotted as a function of QO/Ve, where C is the tracer concentration at time 8, COis the initial tracer concentration, Q is the volumetric rate of flow, e is the time since the tracer flow was discontinued, V is the gross volume of the fluidized bed, and E is the fraction of voids in the bed. The group Q e p e represents the number of
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