distribuitor

02638762/04/$30.00+0.00# 2004 Institution of Chemical Engineers

www.ingentaselect.com=titles=02638762.htm Trans IChemE, Part A, January 2004Chemical Engineering Research and Design, 82(A1): 2533

THE INFLUENCE OF THE DISTRIBUTOR PLATE ON THEBOTTOM ZONE OF A FLUIDIZED BED APPROACHING

THE TRANSITION FROM BUBBLING TOTURBULENT FLUIDIZATION

J. M. PAIVA1*, C. PINHO2 and R. FIGUEIREDO3

1Mechanical Engineering and Industrial Management Department, ESTV, Polytechnic of Viseu, Viseu, Portugal2CEFTDEMEGI, Faculty of Engineering, University of Oporto, Oporto, Portugal

3Mechanical Engineering Department, Faculty of Science and Technology, University of Coimbra, Coimbra, Portugal

The dynamics of the bottom zone of a narrow fluidized bed, from bubbling to turbulentregimes, was studied in a cold model of 0.1 m i.d. and 1.3 m high. Tested distributortypes were perforated perspex plates, with six different perforation grids, metallic mesh

and porous ceramic, with pressures drops ranging from 0.05 to 350 kPa, corresponding tosuperficial air velocities from 0.1 to 2.3 m s1. Group B silica ballotini, within the range 0.3550.425 mm, were used as bed material. The experimental data consisted of pressure drop andabsolute pressure fluctuating signals, together with visual observations. The bottom zonepresented a dynamic condition that produced higher pressure drop values than those expectedwith the incipient fluidization condition, particularly for the distributor plates with a higheropen area. A simple model is used to describe the axial solids distribution and the dynamicscharacteristics of the voids created in that zone, and an attempt is made to differentiate thebubble voidage from that of the dense phase, with a value of through flow estimated in asystematic way.

Keywords: fluidization regimes; bottom zone; distributor plate.

INTRODUCTION

The performance of a gas distributor is very important tothe success of the operation of a fluidized bed (Geldartand Baeyens, 1985). The present work analyses theinfluence of some distributor types (porous ceramics,metallic mesh and perforated perspex), under differentfluidization velocities, covering bubbling and turbulentregimes, on the bottom zone of a fluidized bed usingGroup B silica ballotini within the size range 0.3550.425 mm. Although the bottom zone represents only asmall part of the bed, it is of considerable importancein what concerns the reactions taking place in a fluidizedbed reactor, owing to the high solids concentration and theretention of the reactants (Schlichthaerle and Werther, 1999).Svensson et al. (1993), based on experimental values of thepressure profile, defined the height of the bottom zone asthe level at which that profile starts to deviate from a straightline. Schlichthaerle and Werther (1999) proposed that itshould be defined as the height of the column where thesolid concentration begins to decrease.

As for the hydrodynamic structure of that region of thebed, some authors consider it to be core-annulus type (Sunet al., 1999), others cluster-like (Brereton and Grace, 1993),bubbling at all times (Werther and Wein, 1994; Svenssonet al., 1996a), turbulent (Schnitzlein and Weinstein, 1988;Bolton and Davidson, 1988; Bai et al., 1996), or with aturbulent core region and a dense bubbling condition at thewall (Schlichthaerle and Werther, 1999). Finally, Malcus et al.(2000), using an electrical capacitance tomography method,found no significant change of the hydrodynamic behaviourof the gassolid suspension, measured between two differentheights inside the bed. As for the influence of the distributorplate, apart the work of Svensson et al. (1996b), no otherreferences were found to the way it might affect the bottom ofthe bed, either from the point of view of the different type ofused plates, or concerning the pressure drop across thedistributor. Several authors considered a constant voidageat the bottom bed, extrapolating the pressure drop profilebetween the lowest pressure tap and the distributor (Johnssonand Leckner, 1995; Zijerveld et al., 1997), or found noinflexions of voidage (Malcus et al., 2000).

Some experimental results concerning the changes occur-ring in the fraction of solids with height are also examined.Different velocities and eight different types of distributorplates were used in the experiments. A technique to estimatethe porosity fluctuations of the dense phase, in the bottom ofthe bed is also proposed.

25

*Correspondence to: Dr J. M. Paiva, Departmento de Engenharia Mecanicae Gestao Industrial, Escola Superior de Tecnologia, Campus Politecnico,3504-510 Viseu, Portugal.E-mail: [email protected]

EXPERIMENTAL

The experimental work was done in a cold, atmosphericfluidized bed, with a 100 mm internal diameter perspex tube,using air at ambient conditions as the fluidizing medium(Figure 1).

The fluidized bed was established by fluidizing Group Bsilica ballotini, 355425 mm, previously screened frombulk samples and then statistically assessed with t-tests toevaluate the consistency of the means, using SPSS1

software package, as well as a Coulter LS 230 ParticleSize Analyser. The airflow rate was measured with orificeplate flowmeters, equipped with a Furness Control FCO18differential pressure transducer, connected to a data acquisi-tion system.

Instead of measuring the evolution of the whole bedpressure drop, a different approach was followed. The300 mm bed static height was divided into horizontalslices. The pressure measurements were made at each oneof the six horizontal slices of 50 mm high the bed wasdivided in. The experimental conditions are summarized inTable 1.

Three pressure probes were used at each level, placed atangular intervals of 120 (Figure 2). The data acquiredrepresented cumulative values for each layer, i.e. thevalues of the first layer represent the slice 1 pressuredrop, the values of the second layer the slice 1 plus 2

pressure drop, and so on until the sixth layer, correspondingto the sum of the measured pressure drops from slices 1 to 6.Each probe was calibrated against U-tube water pressuremanometers. Values for the pressure drop of each slice wererecorded for consecutively decreasing flowrates, usingFurness Control differential transducers FCO15 and 16and Klay Peramic pressure transmitters CER8000,connected to a data acquisition system. Problems withpressure taps plugging were inexistent as the porosity ofthe particles avoided it and a set of two consecutive U-bends(one up and one down) was used after each tap to preventdamage to the pressure transmitters membranes. Figure 3shows a typical example of the pressure measured by aprobe located in the fifth layer and of the differentialpressure signal measured between two probes (second andthird, slice 2).

The data acquisition system used a sampling frequency of5 Hz. To ensure sufficient accuracy in the statistical analysis,an average of 1000 samples were taken for each spectrum;those readings were then weighed in order to output valuescorresponding to arithmetically averaged 1 s intervals. Lateron, using a suitable program, these were determined for eachposition of the flowmeter and the dubious points eliminatedaccording to Chauvenets criterion (Holman, 1994).

The plates pressure drop evolution for the tested gas flowrange for each distributor type was measured with pressureprobes calibrated against U-tube mercury pressuremanometers. Eight different distributors were used: sixperforated perspex plates, one metallic mesh and oneporous ceramic, with pressures drops ranging from 60 to300 kPa, and superficial gas velocities going from 0.1 to2.3 m s1 at standard pressure and temperature. The relevantdata for the distributors is given in Table 2. A moreextensive description of the experimental setup can befound in Paiva et al. (2000).

The uncertainty calculations follow the Coleman andSteele (1999) systematic approach. It consists in using adetailed analysis by means of an evaluation of the fixed orsystematic, and random or precision contributions, for eachof the measured variables, thus obtaining values of the fixed

Figure 1. Photo where the location of the several pressure taps can be observed and schematics of the experimental setup.

Table 1. Experimental conditions.

Data on the experimental setup Value

Internal bed diameter, D (m) 0.10Height of the unit, L (m) 1.8Local atmospheric pressure, P (Pa) 97500Bed temperature, T (C) 15Fluidization velocity, U0 (m s

1) 0.12.3Bed material Silica ballotiniBed material density, rp (kg m

3) 2485Average particle diameter dp 1=(Sxi=dpi) (mm) 372Minimum fluidization velocity, Umf (m s

1) 0.1Static bed height, H (m) 0.3

Trans IChemE, Part A, Chemical Engineering Research and Design, 2004, 82(A1): 2533

26 PAIVA et al.

uncertainty IFr, and of the random uncertainty IAr,concerning the experimental result. As the focus is on theaxial evolution of the bed of solids, it was important toanalyse the pressure evolution with height as an indirectmeasure of the solids fraction. Therefore, the uncertaintyanalysis carried out was based on the acquired values ofpressure and the calculation of the pressure drop ofthe successive layers, as well as on the measurementof the other variable, the superficial velocity, whichwas evaluated for increasing values within the range fromincipient fluidization to the maximum experimentallyallowed value for that setup.

Table 3 presents the values for the standard deviation s,fixed IFDPi, random IADPi , and global IDPi uncertainties,associated to each layer pressure drop. In order to obtain

the uncertainties, three fixed errors were considered: thecalibration process error for the pressure transducers, theanalog-to-digital conversion and the hardware precisionindicated by the manufacturer. The final value is thesquare root of the sum of the squares of each contribution.As for the random contribution, errors arising from calibra-tion and from statistical reading of the acquisition datasystem were considered, the final result being calculatedas described previously for the fixed uncertainty. Table 4shows the uncertainty values regarding the fixed IF, randomIA, and global I uncertainties, for the range of superficialvelocities, from Umf to maximum U0, taking into considera-tion variables such as the thin plate orifice pressure drop, thegas temperature, the gas constant and the measurementprecision of the duct diameter.

Figure 2. Radial distribution of pressure tap for each slice.

Figure 3. Time sequence of, for example, pressure measured by a probe located on the top of (a) the fifth layer and (b) a differential transducer measuring thepressure drop in the second slice.

Table 2. Data on distributors=orifices.

TypePlate

reference

Nr oforifices

Nor

Openarea (%)

Orificediameterdor (mm)

Pressure dropratio, DPdist=DPt,

at Umf (%)MaximumU0 (m s

1)

Pressure dropratio, DPdist=DPt, at

maximum U0 (%)

Perforated plate withtriangular pitch (p#x)

p0x 50 0.045 0.3 946 0.6 9138p1x 109 0.098 0.3 418 1.0 8603p2x 199 0.179 0.3 180 1.2 8241p3x 300 0.270 0.3 54 1.7 7145p4x 386 0.347 0.3 46 2.0 5649p9x 948 0.853 0.3 18 2.2 1400

Metallic mesh (dyn) dyn 0.675 O 0.3 23 2.2 2318Porous ceramic (ker) ker 8.482 O 0.3 1 2.3 42


INFLUENCE OF THE DISTRIBUTOR PLATE 27

In both tables the global value is the square root of thesum of the squares of the fixed and random values and,beyond the indication of the absolute values, a percentageratio between that absolute value and the range of measure-ment is also given.

RESULTS AND ANALYSIS

Solids Concentration at the Bottom of the Bed

In order to calculate the evolution of the bed axial solidsconcentration es, the accumulated values of the experimentalpressure drop per slice, DPsj, were used so that the sum ofsuccessive values will determine the pressure drop of eachlayer under analysis, from the first to the sixth DPi:

DPi Xij1

DPsj (1)

Such values of DPi are shown in Figure 4 for twodifferent types of distributor, p9x and dyn, for increasedvalues of U0=Umf, from the first to the sixth layer. The

curves follow the expected hierarchy among the involvedlayers. In fact, the more slices are considered, the higher thepressure drop of the corresponding layer, whereas with theincrease of U0=Umf expected discrepancies between layersbecome progressively reduced with the increase of in-bedmixing.

The approach that is followed is based on the axialevolution of the pressure drop as an indirect step to calculatethe bed solids fraction. Following the basic idea of Svenssonet al. (1993), who made use of the axial pressure profile todefine the height of the bottom zone, the axial values of thesolids concentration are calculated from the pressure gradi-ent, considering that the pressure drop is due to the hydro-static head of solids minus buoyancy, between twodifferential pressure taps, when the bed is operated underincipient fluidization conditions:

dP

dH (rp rf )g(1 emf ) (2)

with rp and rf being the particles and the fluid density, andemf the dense phase porosity at minimum fluidization

Table 3. Standard deviation (s), fixed (IFDPi), random (IADPi) and global uncertainties (IDPi), for the successive layers pressuredrop DPi.

Layer i s (Pa) IFDPi (Pa) IFDPi (%) IADPi (Pa) IADPi (%) IDPi (Pa) IDPi (%)

1 0.609.68 15.11 3.08 5.11 1.04 15.95 3.252 0.579.05 21.38 2.18 7.22 0.74 22.57 2.303 1.0110.13 26.19 1.78 8.85 0.60 27.65 1.884 0.9114.97 149.58 7.62 10.21 0.52 149.92 7.645 0.6011.13 154.08 6.28 11.42 0.47 154.50 6.306 2.1134.01 249.54 8.48 12.51 0.43 249.85 8.49

Total 1.6425.46 196.29 6.67 5.11 0.17 196.36 6.67

Table 4. Fixed (IF), random (IA) and global correlated (I) uncertainties of U0 and Umf range of superficialvelocities.

IF (m s1) IF (%) IA (m s

1) IA (%) I (m s1) I (%)

U0 (2.1 m s1) 0.00825 0.38 0.00179 0.08 0.00845 0.39

U0 (1.5 m s1) 0.01290 0.61 0.00150 0.09 0.01039 0.62

U0 (1.0 m s1) 0.01472 1.27 0.00137 0.12 0.01479 1.28

Umf (0.45 m s1) 0.02461 3.55 0.00180 0.26 0.02467 3.56

Umf (0.28 m s1) 0.03182 5.92 0.00267 0.42 0.03190 5.93

Umf (0.1 m s1) 0.03418 6.82 0.00243 0.48 0.03427 6.84

Figure 4. Typical evolution of the bed pressure drop using (a) p9x and (b) dyn distributor plates, for increasing values of U0/Umf (u, first layer;e, second; , third; , fourth; --, fifth; s, sixth).


28 PAIVA et al.

conditions. As the objective is the analysis of the bedbehaviour beyond the point of minimum fluidization condi-tions, and the measurement system is made between fixedprobes, the reduction of the amount of solids must beaccounted for by means of the introduction of the bubblesvoid fraction d, so that the resulting pressure drop is

dP

dH (rp rf )g(1 emf )(1 d) (3)

For the high gas velocities found in the experimentalconditions, ca 20Umf for these group B particles, it isnecessary to account for the friction effects resulting fromthe collisions between particles and between particles andconfining walls. These effects will be equated as a fractionof the dynamic force exerted, following a similar approachfrom Yerushalmi et al. (1978), so that the pressure dropowing to the drag exerted by the gas upon those particlescan be expressed as

dP 12l

Umf

emf Up

2rf dNp (4)

where Umf is the minimum fluidization velocity, Up theparticles velocity and Np the number of particles in a bedsection of thickness dH, so that

dP

dH 3

4l

Umf

emf Up

2rf (1 emf )(1 d)

D2

d3p(5)

with D the bed diameter and dp and the average diameter ofthe particles. Finally, l accounts for the friction effectsbetween the particles and the particles and the walls, andas a result of multiple set of experiments performed, shall becorrelated with

l 1:5 104 U0Umf

1,5(6)

This factor increases with increasing superficial velocity,as beyond stationary conditions the bed will produce agrowing frequency of collisions, and the obtained valuesare within the range 102104 defined by Barths (1954)model. In addition, results are in good agreement withformer experimental results (Paiva et al., 2000).

In equation (5), the difference between the interstitialvelocity, taken as Umf=emf (supposing that the dense phase

voidage remains constant above Umf) and the particlevelocity Up, is its terminal velocity Ut,

Uf Up Ut 4

3

rpgdpCDcrf

s(7)

According to the previous equation, CDc the drag coeffi-cient for a cloud of particles, is calculated using the Wen andYu (1966) correlation, quoted by Klinzing (1981):

CDc CDse4,7mf (8)Back to equation (5), es equals the product of the fraction

of solids in the dense phase by the remainder of the fractionof bubbles regarding the whole bed,

es (1 emf )(1 d) (9)Finally, the result of summing both hydrostatic head and

friction loss contribution, equations (2) and (5), is:

dPcalcdH

(rp rf )g 3

4l(Ut Up)2rf

D2

d3p

" #es (10)

where Ut is calculated by means of equation (7). Conse-quently, the values of es are those that make the dimension-less pressure drop DP, a ratio between the experimentalpressure drop values DPexp, and the calculated ones, equa-tion (10), equal to unity:

DP DPexpDPcalc

(11)

The calculation procedure was then, by means of applyingequations (2)(10) to the experimental results of the slicedbeds pressure drop, determining es so that equation (11) issatisfied. The average relative contribution of the frictionalcomponent of dP=dH [equation (10)] is in the range of 0 to4560%, for U0 from 0.1 to 2.1 m s

1, growing with thereduction of the number of distributor orifices, the exceptionbeing the dyn distributor plate.

Because the bed was divided into slices inside which thepressure drop evolution was studied, instead of measuringthe bed pressure drop as if it was a whole entity, someparticularities concerning the different regions along the bedheight could be detected. These local characteristics are

Figure 5. Evolution of the bed axial solids concentration, es, for values of U0=Umf (u, solid line, U0=Umf 1; e, dashed line, 2; n, solid line, 3; , dashed line,4; s, solid line, 5; u, dashed line, 7; e, solid line, 10; n, dashed line, 13; , solid line, 17; s, dashed line, 20).



usually hidden when the analysis considers the bed as a soleentity.

Figure 5 shows the results obtained for the evolution ofthe fraction of solids with height, for several values of theratio U0=Umf and distributor plates p9x and dyn. It mustbe stressed that at higher velocities the bed expands and,consequently, particles are displaced towards higher slicesand thus cannot be accounted for in the lower slices of thebed, the ones under analysis.

Therefore, as the fluidizing velocity increases, the localbottom solids concentration broadly decreases as expected.Nonetheless, at heights around 0.15 m, mainly for the lowervelocities, it reaches a maximum, and then decreases again.The trend described becomes evident in Figure 6, where theseveral tested distributors are compared for increasingvelocity values, from 1 to 20 U0=Umf. At low fluidizingvelocities, differences between the several distributors areminor. Nevertheless, the curves appear to have a similar

Figure 6. Axial solids concentration profiles, es, in the fluidized bed for various fluidization ratios of U0=Umf (u, ker; s, p9x; , dyn; *, p4x; , p3x;e, p2x; n, p1x; --, p0x).


30 PAIVA et al.

behaviour to those described before, presenting a maximumsolids concentration for values of H in the interval0.150.20 m. Near the distributor orifices, where the influ-ence of emerging gas jets strongly disrupts the uniformdistribution of the solid phase, there is no distinction of thepattern of behaviour as a function of the used distributor.However, differences become increasingly apparent forhigher fluidization velocities. Unlike the authors mentionedabove (Svensson et al., 1993; Johnsson and Leckner, 1995;Zijerveld et al., 1997; Schlichthaerle and Werther, 1999;Malcus et al., 2000), who consistently report either adecrease or a constant value of the pressure drop in thebottom zone of the bed, the experimental results presentedshow the existence of a non-linear evolution for the group Bparticles. In particular, the lower pressure drop plates(namely ker and p9x) reveal a distinct maximum valuedue to the presence of solids in the lower portion of thatregion of the bed, while the other plates keep showing thesame typical pattern of behaviour. The higher the velocity ofthe jets, the more accentuated becomes the establishment ofa local pressure reduction. This classical energy compensa-tion has the effect of dragging the surrounding particles tothis region, enhancing the solids downcoming flow, result-ing in a growing accumulation closer to the gas distributorplate as the gas mass flow rate increases.

Figure 7 shows this trend as a function of superficial gasvelocity. With the rise of gas velocity at the orifices,corresponding to larger jets with stronger dragging capabili-ties, the maximum solids concentration gets closer to thedistributor plate with the increase of fluidizing gas flow.Although a decrease in the open area of the distributorscauses a higher velocity at the orifices, their reduction innumber also means a greater distance between them, so theeffect of pressure compensation does not have such a strongspatial influence as in the case of the lower pressure dropplates. The drag over the surrounding particles no longerprevails, and as the jets high velocity gas flow cannot passeasily through the bed, the overall gas velocity decreases andconsequently a pressure build-up takes place just above thegas distributor. This situation is hydrodynamically solved inan intermittent way (Svensson et al., 1993) when thebubbles are formed, or more adequately, where the bubblesare formed. A path is created through which the temporarilyretained gas flow will pass (van der Schaaf et al., 1999), thepressure reduces and the solids concentration increaseshigher up the bed. This way, considering the whole bed,the traditional gas flow split between the dense phase flow,the visible bubble flow and the invisible bubble flow, isre-established.

Dense Phase Voidage at the Bottom of the Bed

There is another aspect that can be taken into considera-tion from the analysis of the pressure drop evolution insidethe bed, the fluctuation of the emulsion porosity. Severalreferences in the literature agree that calculated values ofbubbles diameters and rising velocities using the model ofWerther and Wein (1994) were in fairly good agreementwith the void sizes found experimentally. Especially underconditions of superficial velocities typical for combustors(and for the group B particles widely used in thoseprocesses). Similar good results for the whole bed wereobtained in the present work, with the range of particles

used in these experiments. The first layers revealed,however, some differences when using the lower pressuredrop plates, as can be seen in Figure 8 (where those resultsare presented for the first and the sixth layer).

If the bottom bed is modeled according to the model ofWerther and Wein (1994), by calculating the bubblesvoidage d, with equation (12),

d 1:45Ar0:18 U0 UmfUba

p (12)

a good adjustment to the experimental results of the wholebed can be achieved, as can be observed in Figure 8, sixthlayer. In equation (12), the bubble absolute velocity Uba, is(Werther and Wein, 1994)

Uba 1:45Ar0:18(U0 Umf ) 0:71ug0:5db (13)where u takes the value 2

pD for 0.1

porosity ee (Figure 9), to borrow an expression from Kuniiand Levenspiel (1969), can be estimated from

ee emf De 11 d

1 d (1 emf ) (15)

where d represents the values of the fraction of bubblescalculated from equations (12)(14) and d the values thatmake DP equal to 1; De represents the range of variation ofthe dense phase voidage around the incipient conditionsvalue.

The results of applying this expression to the experimentaldata reveal a change in the porosity of the emulsion, awayfrom the postulated constant dense phase porosity emf,especially in the lower part of the bed, as the fluidizationvelocity increases and reaches the transition from bubbling toturbulent regime. The difficulty in determining experimentallythe dense phase flow is the reason why the usual practice hasbeen to calculate its value based on the two-phase flowpremises, forcing the invisible bubbles flow to accommodatediscrepancies between expected and detected experimentalbehavior. This would explain, on the other hand, why theexperimentally based calculated values of the bubbles voidfraction d, were different from the Werther and Wein (1994)model results for the first layers of the bed, meaning that the

need to allow the passage of the excess of air was notcompletely performed by the bubbles phase (either visible orinvisible), but that the dense phase widens to accommodatethe extra gas flow, a dynamic phenomenon described byWilhelm and Kwauk (1948), Lewis et al. (1949) and Wenand Yu (1966), especially when moving nearer the feedingpoint of the fluidizing medium. As for the upper layers, valuesof porosity are almost constant, approaching 0.45, which, forthese 0.3550.425 mm group B silica ballotini is the expectedvalues, meaning that the air flow has a more stable pathcondition.

CONCLUSIONS

An investigation of the bottom zone of a fluidized bed,using three types of distributor plates, was performed usingpressure drop measurements. The results show that the flowchanges with variation of the operating conditions. Themeasurements at several different equal heights showchanges in the hydrodynamic behaviour of the gassolidsuspension. For the tested group B particles and static bedheight, there is no linear decrease, or even a sustaineduniform trend towards a decrease of the solids concentrationwith height in that region of the bed. Furthermore, using amodel for the calculation of the fraction of bubbles, theresults include the detection of a fluctuation in the values ofthe porosity of the emulsion, with height and with thefluidization velocity.

NOMENCLATURE

A bed cross-sectional area, m2

db bubble diameter, mdor orifice diameter, mdp average particle diameter, mD internal bed diameter, mF drag force, Ng gravitational acceleration, m s2

H vertical distance from gas distributor to bed free surface, mIr global uncertainty associated to a result or to a variable rL height of the unit, mNor number of orifices of the distributor plateP local atmospheric pressure, PaT ambient temperature, CU0 superficial fluidization velocity, m s

1

Uba mean bubble absolute rise velocity, m s1

Umf minimum fluidization velocity, m s1

Ut terminal velocity, m s1

Figure 8. Evolution of the bed bubble fraction, dmod, equation (8), and the resulting values for d, with different distributors, for several values of U0=Umf

(u, ker; s, p9x; , dyn; *, p4x; , p3x; e, p2x; n, p1x; --, p0x).

Figure 9. Evolution of the porosity of the emulsion, ee, equation (11), forseveral values of U0=Umf (u, layer 1; e, layer 2; n, layer 3; , layer 4;--, layer 5; s, layer 6).


32 PAIVA et al.

Up particle velocity, m s1

xi mass fraction in the i interval

Greek symbolsd bubble fraction, equation (12)d bubble fraction value that yields to equation (11)DP pressure drop, PaDP dimensionless pressure drop, equation (11)DPdist distributor pressure drop, PaDPsj slice pressure drop, PaDPt total bed pressure drop, PaDe variation of the dense phase voidage around the incipient

conditions valueee emulsion void fractionemf void fraction at minimum fluidization conditionses fraction of solidsl impact and attrition factorr density, kg m3

s standard mean deviationu parameter in equation (13)

Dimensionless numbersAr Archimedes number, rf(rp7rf)gdp

3=m2

CD drag coefficient, FD=rU2A

CDc corrected drag coefficient, CDsemf4,7

CDs drag coefficient for a single particle

Subscripts0 relative to the bed free surfaceb bubblec clustercalc calculatedD dragdist distributorexp experimentalf fluidi layerj slicel bedmf relative to minimum fluidization conditionsor relative to the number of orifices in the distributor platep relative to the particless singlet total

Superscripts relative to the dimensionless bed pressure drop, equation (11)0 relative to the normalization of the standard mean deviation

using the maximum value for a specific run

REFERENCES

Bai, D., Shibuya, E., Masuda, Y., Nakawaga, N. and Kato, K., 1996, Flowstructure in a fast fluidized bed, Chem Eng Sci, 51: 997966.

Barth, W., 1954, Chem Ing Tech, 20(1): 2932, in Pneumatic Conveying ofSolids, Marcus, R.D., Leung, L.S., Klinzing, G.E. and Rizk, F. (eds)(Chapman and Hall, London, UK).

Bolton, L.H. and Davidson, J.F., 1988, Recirculation of particles in fastfluidized risers, in Circulating Fluidized Bed Technology II, Basu, P. andLarge, J.F. (eds) (Pergamon Press, Oxford, UK), pp 139146.

Brereton, C.M.H. and Grace, J.R., 1993, Microstructural aspects of thebehavior of circulating fluidized beds, Chem Eng Sci, 48: 25652572.

Coleman, H.W. and Steele, W.G., 1999, Experimentation and UncertaintyAnalysis for Engineers (Wiley, New York, USA).

Geldart, D. and Baeyens, J., 1985, The design of distributors for gas-fluidized beds, Powder Technol, 42(1): 6778.

Holman, J.P., 1994, Experimental Methods for Engineers, 6th edition(McGraw-Hill, New York, USA).

Johnsson, F. and Leckner, B., 1995, Vertical distribution of solids in acirculating fluidized bed in a circulating fluidized bed furnace, inProceedings of FBC 13, Heinschel, K.J. (ed), 710 May, Orlando, FL,p 671.

Klinzing, G.E., 1981, GasSolid Transport (McGraw-Hill, New York,USA).

Kunii, D. and Levenspiel, O., 1969, Fluidization Engineering (Krieger,New York, USA).

Lewis, W.K., Gilliland, E.R. and Bauer, W.C., 1949, Characteristics offluidized particles, Ind Eng Chem, 41: 11041117.

Malcus, S., Chaplin, G. and Pugley, T., 2000, The hydrodynamics of thehigh-density bottom zone in a CFB riser analyzed by means of electricalcapacitance tomography (ECT), Chem Eng Sci, 55: 41294183.

Paiva, J.M., Pinho, C. and Figueiredo, R., 2000, Distributor plate influenceon fluidization quality, measuring pressure drop evolution in adjacentslices through the bed, in Proceedings of the FEDSM2000ASME FluidsEngineering Division Summer Meeting, 1115 June, Boston, MA(Multi-media CD-ROM FEDSM2000-1026.ps) (ASME, New York, USA).

Schlichthaerle, P. and Werther, J., 1999, Axial pressure profiles and solidconcentration distributions in the CFB bottom zone, in CirculatingFluidized Bed Technology VI, Werther, J. (ed.) (Dechema, Germany),pp 185190.

Schnitzlein, M.G. and Weinstein, H., 1988, Design parameters determiningsolids hold-up in fast fluidized bed system, in Circulating Fluidized BedTechnology II, Basu, P. and Large, J.F. (eds) (Pergamon Press, Oxford,UK), pp 205211.

Sun, G., Chao, Z, Fan, Y. and Shi, M., 1999, Hydrodynamic behavior inthe bottom region of a cold FCC riser, in Circulating Fluidized BedTechnology VI Werther, J. (ed.) (Dechema, Germany), pp 179184.

Svensson, A., Johnsson, F. and Leckner, B., 1993, Fluid-dynamics of thebottom bed of circulating fluidized bed boilers, in Proceedings of the12th International Conference on Fluidized Bed Combustion, Rubow, L.(ed) (ASME, New York, USA), pp 887897.

Svensson, A., Johnsson, F. and Leckner, B., 1996a, Bottom bed regimes in acirculating fluidized bed boiler, Int J Multiphase Flow, 22(6): 11871204.

Svensson, A., Johnsson, F. and Leckner, B., 1996b, Fluidization regimes innon-slugging fluidized beds: the influence of pressure drop across the airdistributor, Powder Technol, 86: 299312.

van der Schaaf, J., Schouten, J.C., Johnsson, F. and van den Bleek, C.M.,1999, Bypassing of gas through bubble chains and jets in circulatingfluidized beds, in Proceedings of 6th International Conference onCirculating Fluidized Beds, Werther, J. (ed) (Dechema, Frankfurt amMain, Germany), pp 4752.

Wen, C.Y. and Yu, Y.H., 1966, Mechanics of fluidization, Chem Eng ProgSymp Ser, 62(62): 100111.

Werther, J. and Wein, J., 1994, Expansion behavior of gas fluidized beds inthe turbulent regime, AIChE Symp Ser, 301(90): 3144.

Wilhelm, R.H. and Kwauk, M., 1948, Fluidization of solid particles, ChemEng Prog, 44: 201218.

Yerushalmi, J., Cankurt, N.T., Geldart, D. and Liss, B., 1978, Flowregimes in vertical gassolid contact systems, AIChE Symp Ser,74(176): 113.

Zijerveld, R.C., Koniuta, A., Johnsson, F., Marzocchella, A., Schouten, J.C.and van den Bleek, C.M., 1997, Axial solids distribution and bottom beddynamics for circulating fluidized bed combustor application, AIChESymp Ser, 93(318): 97102.

The manuscript was received 11 June 2002 and accepted for publicationafter revision 10 September 2003.



distribuitor

Documents