Scaling Considerations for Circulating Fluidized Bed Risers, Patience Et Al. 1992

8/12/2019 Scaling Considerations for Circulating Fluidized Bed Risers, Patience Et Al. 1992

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Powder Technology, 72 (1992) 31-37 31

Scaling considerations for circulating fluidized bed risersG. S. Patience*, J. ChaoukiDkpati ement de Gt nie Chimi que, Ecole Polytechnique de Monttial, Monheal, Que., H3C 3A7 (Canada)F. Berruti and R. WongDepartment of Chemical and Petroleum Engineeri ng, University of Calgay, Calgary, Alt a., T2N I N4 (Canada)(Received June 10, 1991; in revised form December 27, 1991)

AbstractThe ratio between actual gas velocity to particle velocity in the hydrodynamically fully developed region ofCirculating Fluidized Bed risers (CFB) may be approximated by cp= 1 5.6/Fr + 0.47Frp = U& VP. This ratio,termed the slip factor, is about 2 at operating conditions characteristic of industrial risers several meters indiameter and agrees with observations of J. M. Matsen (in D. L. Keairns (ed.), Fluidization Technology, Vol. 2,Hemisphere, 1976, p. 135). The proposed relationship between the gas and solids velocity is an adequate firstapproximation to estimate gas and solids residence times, blower capacity and standpipe length.

Introduction

Circulating Fluidized Beds (CFB) are being consid-ered as alternatives to more conventional FluidizedBed processes because of their apparent intrinsic ad-vantages, including short and controllable residencetimes for the gas and solids, high turn down ratios andflexibility. Industrial scale plants for coal combustion,aluminum oxide calcination, catalytic cracking,Fischer-Tropsch synthesis successfully employ this tech-nology. Contractor and Chaouki [l] and Gianetto etal. [2] have discussed a number of potential catalyticprocesses that are likely candidates for CFB technology.

The main difference between bubbling, turbulent bedsand CFB risers is gas velocity. Typical gas velocitiesin CFBs range from 2-10 m s-l. At these velocities,solids are readily entrained by the gas and are carriedto the top of the vessel. Cyclones and rough cut sep-arators separate the solids from the gas phase. Thesolids are returned to the riser bottom by a standpipe.

The longitudinal solids hold-up in the riser, discussedby Yerushalmi et al. [3] and by Li and Kwauk [4],exhibits a relatively dense region at the solids entrypoint and a dilute phase above it. A number of modelshave been proposed to characterize the riser hydro-dynamics. Rhodes and Geldart [5] used the entrainmentmodel, developed for fluidized beds by Wen and Chen[6], to describe the dilute phase and adapted a classical

*Presently at E.I. du Pont de Nemours & Co., Wilmington,DE 19880. USA.

two phase model of a bubbling fluidized bed to thedense phase. Kunii and Levenspiel[7] adopted a similarapproach and correlated decay constants based on anumber of experimental investigations to model thedecrease in solids hold-up along the riser. Also, mea-surements of the internal flow structure of the riserby Hartge et al. [ES], ader et al. [ 9], and Bolton andDavidson [lo] indicated that large radial gradients exist,with significantly higher concentrations of solids nearthe wall.

A complete description of the hydrodynamics of sucha flow structure is difficult. The gas-solid flow is typicallycharacterized by large relative velocities between thetwo phases. Two mechanisms have been proposed toaccount for the difference in velocity: Yerushalmi etal. [3] suggested that particles agglomerate into clusterswhose void fraction approaches E,.,,~,whereas Rhodeset al. [ll] and Berruti and Kalogerakis [12] postulatedthat particles ascend in the core in a dilute phase anddescend along the wall as a dense annulus. Ishii et aE.[13] have recently incorporated both mechanisms intoa clustering annular flow model.

Most of the studies on the hydrodynamics of CFBsystems reported in the literature have been conductedusing laboratory scale units (i.e. relatively short andnarrow). Scale-up to industrial reactors several metersin diameter and tens of meters in height is uncertainat best. Experimental rigs. are not only limited bydiameter and height constraints but also by the maximumcirculation rates attainable. Matsen [14] reported thattypical industrial FCC units operate at solids fluxes

0032-5910/92/$5.00 0 1992 - Elsevier Sequoia. All rights reserved


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between 500 and 1500 kg m- s-l. The majority ofexperimental rigs employ non-mechanical devices andGeldart group B particles, which facilitate circulationrate control but may ultimately limit the maximumsolids fluxes attainable. Although CFB boilers generallyoperate at solids circulation rates less than 100 kg m-S -I catalytic reactors require different operating con-ditions.

Despite the growing body of literature, more fun-damental information on the hydrodynamics of largescale CFB reactors is needed to assess the potentialof this technology and to establish design criteria. Scale-up parameters are useful for the design of industrialCFB units. These parameters should not only estimatethe overall pressure drop for a given gas velocity andcirculation rate, necessary to size compressors and thestandpipe, but also adequately predict reactor per-formance including gas-solids contact efficiency andheat transfer characteristics. The internal flow structureof small experimental units is well understood, but Dryand La Nauze [15] suggest that the symmetry of theradial solids distribution measured in small units maynot apply to large units. However, experiments con-ducted in a commercial FCC riser by Saxton and Worley[16] using a radiation attenuation technique, indicatethat a two-phase type flow pattern might adequatelydescribe the flow phenomenon. In addition to theuncertainties in scaling-up the riser diameter, few studiesaddress the effect of height on the hydrodynamics. Intall risers, differences in the gas velocity between thetop and bottom of 50 are conceivable at high cir-culation rates. Grace [17] indicates that further com-plications may arise from the exit and entrance effects,wall intrusions or roughness and the coefficient ofrestitution of the particles. Glicksman et al. [18] doc-umented the increase in the solids void fraction bychanging the geometry of the exit from a smooth elbowto a sharp one and presented data suggesting thatobjects intruding into the riser may significantly influencethe local solids hold-up.

Glicksman et al. [18], Ishii et al. [13] and Ishii andMurakami [19] proposed scaling laws to predict thebehaviour in large scale units. Scale-up criteria werederived based on the principles of fluid-particle systems.The criteria were then verified in geometrically similarsmall scale lab units. Axial solids hold-up and pressurefluctuations were generally used as the basis for com-parison. Despite the differences in derivations, Glicks-man et al . [18] maintain that the scaling laws proposedin the literature are not dissimilar. They examined twounits with a 34 mm and 152 mm square cross section.Ishii et al . [13] developed scaling parameters based onthe Clustering Annular Flow Model and validated thetheory experimentally in two geometrically similar units200 mm and 50 mm in diameter. However, only very

low circulation rates and gas velocities were considered(U,


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However, work by Patience et al. [23] clearly showsthat the latter assumption is not true and better criteriaare required to evaluate the average solids velocity.

In large industrial scale FCC risers Matsen [14]reported that the slip factor cp, is approximately equalto 2 and hence the particle velocity equals Ug/2e.Comparisons between large scale industrial units andexperimental units is complicated not only because ofthe differences in geometry but also because of thedifferences in operating conditions; high circulationrates, high temperatures and pressures. Sternerding [24]showed that the riser pressure drop was independentof the transport gas properties in the range,1.4 lo-


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that operate at elevated circulation rates with GroupA powders as reported by Matsen [14]. At the entrance,the slip factor increases with mass flux, which may beattributable, in part, to the overprediction of the solidsfraction because the acceleration contribution to thepressure drop was neglected. Also, the slip factor ap-pears to be greater at higher gas velocities at the topof the riser.

In Fig. 4, data measured by Wong [27] in a 3 m tallriser 50 mm in diameter is shown. The apparent slipfactor in the acceleration region is plotted togetherwith the actual slip factor in which the accelerationcontribution of the particles is taken into account. Thecontribution of particle acceleration to pressure drop,hence density, was calculated based on the work byWeinstein and Li [30]. The figure indicates that, althoughthe actual slip factor in the acceleration zone is greaterthan 2, ignoring the acceleration effect greatly over-estimates the slip factor, hence total solids hold-up.

The slip factor calculated in the hydrodynamicallydeveloped region, based on data reported by a numberof researchers, is plotted against the gas velocity inFig. 5. Table 1 summarizes the particle characteristicsand riser geometry of each study. Both Geldart A andB powders were used in the experiments for which theparticle terminal velocities vary between 0.2 m s-l and2 m s-l. A slip factor of 2 correlates the data reasonablyat gas velocities between 6 and 12 m s-l. This agreeswith the value reported by Matsen [14] for industrialrisers, which typically operate at velocities greater than8 m s-l.

To account for the increase in q with decreasing gasvelocity, as shown in the figure, the following relationshipis proposed:cp= 1 5.6jFr + 0.47Frp4* 5)

6- I I I0 U =

GE7.9 ms

= 57 kg/ms5- d; = 174 / . un

0 (p experimental0 0 0 corrected for acceleration9. 4-

i9 00 3- 0 0.? 0Ci 2 D CJ 0 00

1 .oHeight (m)

Fig. 4. Slip factors in the hydrodynamically developing region(acceleration) of a riser, Wong [27].

10 I I I 1 I I I_ - - (P = 1+5~3/Fr+0.47Fr,~.~~

060 2 4 6 8 10 12 14 16upb/s

Fig. 5. Slip factors in the hydrodynamically fully developed regionat different gas velocities. Data referenced in Table 1.

Agreement between predicted and experimental cir-culation rates of Van Swaaijs data [20] using thiscorrelation is good, as shown in Figs. 1 and 5. The fitis superior compared to the single parameter estimateof rp=2.

Equation (5) suggests that, at gas velocities muchgreater than single particle terminal velocities, the solidshold-up increases with diameter, i.e.4(1- = (1 +5.6(Dg)o.5/U, + 0.47Fr,0-41)cG,lUg 6)The effect of riser diameter on suspension density hasnot been fully explored. Arena et al. [31] studied tworisers 41 mm and 120 mm in diameter and concludedthat the density increased with diameter at the sameoperating conditions. Kato et al. [34] reported that, forsmall tubes, the density increased with diameter to the0.4 power, whereas a power of 0.2 fit data collectedby Findlay and Knowlton [38] better. Larger riserdiameters were used in the latter study. The correlationpredicts that the solids hold-up is relatively independentof particle characteristics as long as the superficial gasvelocity is much greater than the particle terminalvelocity. Moreover, it suggests that the solids hold-upis relatively insensitive to gas properties.

At high gas velocities, typical of pneumatic conveying,5.6/Fr tends to zero and cp approaches 1+0.47Frp41.Typically, FCC risers operate at slip factors near two;Govier and Aziz [37] suggest that in pneumatic conveying1 < cp< 2, which agrees with the proposed correlation.

Brereton [26] reports large differences in the slipfactor between a smooth exit and an abrupt geometry.The slip varies between 1.88 and 2.32 for sand particles(open squares in Fig. 5) in a CFB with a smooth exit.The slip factor for sand in the same unit with an abruptexit geometry (filled squares) varies from 8.2 at lowgas velocities to 3.6. Most of the data reported in the


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TABLE 1. References and experimental conditions for the data in Fig. 5Key Study Riser geometry Particle properties Remarks

V Arena et al. [LB]0 Brereton [26]A0 Patience [25]+ Rhodes and Geldart [5]v Stemerding [24]0 Van Swaaij et al. [20]A Wong [27]0Estimates.

D Height Type dPCm) Cm) (pm) Zg mm3) (m SC)0.041 6.4 Glass 88 2600 0.46 Smooth exit0.152 9.3 Sand 148 2650 0.99 Smooth exit0.152 9.3 Sand 148 2650 0.99 Abrupt exit0.152 9.3 Alumina 65 3500 0.36 Abrupt exit0.083 5 Sand 275 2630 1.9 Abrupt exit, 20 < T< 2500.152 5 FCC 64 1800 0.2 Abrupt exit0.051 2-10 FCC 65 1600 0.18 , Developed region, 15 < T < 5000.18 FCC 65 1400 0.16 Developed region0.05 3 Sand 93 2500 0.48 Abrupt exit0.05 3 Sand 174 2500 1.2 Abrupt exit

CFB with the abrupt exit geometry varies between 4and 6 The slip factor with the alumina particles andan abrupt exit geometry (open triangles) lies between2.1 and 4.7. The large slip velocities reported by Brereton[36] may be a result of the contribution of two differentphenomena. Firstly, the abrupt exit configuration, inthis unit, might affect the solids behaviour significantly.Secondly, secondary gas was supplied to the columnwhich extends the acceleration region of the riser. Thecombination of the abrupt exit and extended accel-eration zone may prevent the establishment of a fullydeveloped flow region. However, experiments by Pa-tience [25] and Wong [27], conducted in short riserswith abrupt right angle exits, for which the accelerationzone would presumably affect the solids loading morethan in taller units, gave much lower slip factors.

Whereas values of rp greater than those predictedby the proposed correlation are evident in the workof Brereton [26], smaller values of cp are calculatedfrom data obtained in the experimental unit of Arenaet al. [28] as reported by Louge and Chang [29] andYang [22]. In this case, the slip velocity approachesthe single particle terminal velocity. At gas velocitiesof 5 m s-l and circulation rates ranging from 80 to390 kg rnp2 s-l the slip varies between 1.15 and 1.055.Inaccuracies in the measurement of the circulation ratecould explain the differences in slip factor as reportedby Brereton [26] and Arena et al. [28]. The slip factoris inversely proportional to the mass flux; hence, under-estimating the circulation rate will result in large cal-culated slip factors and overestimating the rate giveslow values of q. Patience and Chaouki [33] show thatusing the particle velocity along the downcomer wall,the technique used by Brereton [26] in the experimentswith sand, can underestimate the circulation rate byup to 40 . Brereton [26] used a butterfly valve technique

in the experiments with alumina and eqn. (5) fits thesedata well compared to the experiments with sand wherethe solids circulation rates were made by tracking thewall velocity in the downcomer.

The general agreement between slip factors reportedfor industrial FCC reactors and experimental riserssuggests that pressure drop predictions may be possiblewithout having to develop large and costly pilot plants.That is, given the desired residence time of the gasand solids, the blower requirements and height anddiameter of the reactor may be calculated. The insistenceof geometrical similarity, as suggested by Ishii et al.[13] and Ishii and Murakami [19], appears to be toorestrictive. In the present work, the proposed slip factormodel is shown to provide a reasonable estimate ofthe average gas and solids residence times. Particlecharacteristics do not affect the slip factor as long asthe operating gas velocity is significantly greater thanthe particle terminal velocity (in the fully developedflow regime of the riser). Van Swaaij et al. [20] andStemerding [24] used Geldart Group A powders,whereas Patience [25], Brereton [26] and Wong [27]used Geldart Group B particles. Zhang et al. [35] suggestthat particle density and particle size distribution ofGeldart A materials do not affect the radial voidageprofile when comparing systems operating at the samesuspension density.

ConclusionsCorrelations available in the literature do not seem

to predict the relationship between the gas and solidsvelocity adequately. An examination of a large pool ofdata from both experimental laboratory scale CFBs andindustrial units indicates that the ratio of interstitial


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gas velocity to particle velocity, or slip factor, is ap-proximately equal to 2 in the hydrodynamically fullydeveloped flow regime of risers at superficial gas ve-locities greater than 6 m s-. However, an improvedrelationship is also proposed,

which better describes the slip factor dependence ongas velocity, riser diameter, and particle properties.The suspension density increases with diameter anddecreases with increasing gas velocity. At high gasvelocities cp approaches 1 0.47Frto.41. This relationshipapplies to the region, above the acceleration zone atthe entrance and below the deceleration zone at theexit where values of C+Jre greater. Therefore, to estimategas and solids residence times, blower capacity, andstandpipe length requirements the entrance and exiteffects must be considered. However, entrance and exitlengths are typically much shorter than the developedregion, so errors in ignoring the pressure drop con-tribution in a 40 m tall riser would be less than 10 .The slip factor, cp, is reported to be independent ofgas properties and particle characteristics (at gas ve-locities much greater than the single particle terminalvelocity).

List of symbolsD riser diameterd, particle diameterdPld.z pressure gradientFr Froude number, U D). Fr, Froude number, V gD)~5GS solids flux in riserg gravitational constantP time-averaged pressureT temperatureutx superficial gas velocityv , particle velocityK I slip velocityv , particle terminal velocityz vertical co-ordinate

Greek l et t ersE void fractionEmf void fraction at minimum fluidizationPs gas density4 particle densitycp slip factor

References

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C. Y. Wen and L. H. Chen, AZChE J., 28 (1982) 117.K. Kunii and 0. Levenspiel, Powder Technol., 61 (1990) 193.E.-U. Hartge, Y. Li and J. Werther, in P. Basu (ed.), CirculatingFluidized Bed Technology, Pergamon, Toronto, 1986, p. 153.R. Bader, J. Findlay and T. M. Knowlton, in P. Basu andJ. F. Large (eds.), Cir culating Flui dized Bed Technology I I ,Pergamon, Oxford, 1988, p. 123.L. W. Bolton and J. F. Davidson, in P. Basu and J. F. Large(eds.), Cir culating Flu idized Bed Technology I I , Pergamon,Oxford, 1988, p. 139.M. 3. Rhodes, P. Laussmann, F. Villain and G. Geldart, inP. Basu and J. F. Large (eds.), Circulating Fiuidized BedTechnology ZZ, Pergamon, Oxford, 1988, p. 20.F. Berruti and N. Kalogerakis, Can. J. Chem. Eng., 67 (1989)1010.H. Ishii, T. Nakajima and M. Horio, J. Chem. Eng. Jpn., 22(5) (1989) 484.J. M. Matsen, in D. L. Keairns (ed.), Fluidization Technology,Vol. 2, Hemisphere, 1976, p. 135.R. J. Dry and R. D. La Nauze, Chem. Eng. Prog., 86 (7)(1990) 31.A. L. Sazton and A. C. Worley, Oil Gas J., 68 (20) (1970)84.

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31 U. Arena, A. Cammarota, L. Massimilla and D. Pirozzi, inP. Basu and J. F. Large (eds.), Circulating Fluidized BedTechnology II , Pergamon, Oxford, 1988, p. 223.

32 J. Findlay and T. M. Knowlton, Final Report: Pipeline Gasfrom Coal (IGT Hydrogasification Process), (1980) 306.

33 G. S. Patience and J. Chaouki, in P. Basu, M. Horio andM. Hasatani (eds.), Circulating Flu idized Bed Technology I I I ,Pergamon, Oxford, 1991, p. 627.

34 K. Kato, Y. Ozawa and H. Endo, in K. Ostergaard and A.Sorenson (eds.), Fluidization, Engineering Foundation, NewYork, 1986.

35 W. Zhang, Y. Tung and F. Johnsson, Chem. Eng. Sci., 46(12) (1991) 3045.

36 B. Jazayeri, Hydrocar bon Process., 70 (5) (1991) 93.37 G. W. Govier and K. A&, Z7zeFl ow of Complex M ixtzues inPipes, Van Nostrand Reinhold, Toronto, 1972.

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Scaling Considerations for Circulating Fluidized Bed Risers, Patience Et Al. 1992