werther - fluidized bed reactors

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Fluidized-Bed Reactors

JOACHIM WERTHER, Hamburg University of Technology, Hamburg, Germany

1. Introduction. . . . . . . . . . . . . . . . . . . . . . . . 320

1.1. The Fluidization Principle. . . . . . . . . . . . . 320

1.2. Forms of Fluidized Beds . . . . . . . . . . . . . . 321

1.3. Advantages and Disadvantages of the

Fluidized-Bed Reactor . . . . . . . . . . . . . . . . 322

2. Fluid-Mechanical Principles . . . . . . . . . . . 322

2.1. Minimum Fluidization Velocity . . . . . . . . . 322

2.2. Expansion of LiquidSolid Fluidized Beds. 3242.3. Fluidization Properties of Typical Bed Solids 324

2.4. State Diagram of Fluidized Bed . . . . . . . . . 325

2.5. Gas Distribution . . . . . . . . . . . . . . . . . . . . 326

2.6. Gas Jets in Fluidized Beds . . . . . . . . . . . . . 327

2.7. Bubble Development . . . . . . . . . . . . . . . . . 328

2.8. Elutriation. . . . . . . . . . . . . . . . . . . . . . . . . 329

2.9. Circulating Fluidized Beds. . . . . . . . . . . . . 330

2.9.1. Hydrodynamic Principles . . . . . . . . . . . . . . . 330

2.9.2. Local Flow Structure in Circulating Fluidized

Beds . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 333

2.9.3. Design of Solids Recycle System . . . . . . . . . 334

2.10. Cocurrent Downflow Circulating FluidizedBeds (Downers) . . . . . . . . . . . . . . . . . . . . . 334

2.11. Attrition of Solids . . . . . . . . . . . . . . . . . . . 335

3. Solids Mixing in Fluidized-Bed Reactors . . 337

3.1. Mechanisms of Solids Mixing . . . . . . . . . . 338

3.2. Vertical Mixing of Solids. . . . . . . . . . . . . . 338

3.3. Horizontal Mixing of Solids . . . . . . . . . . . . 339

3.4. Solids Residence-Time Properties . . . . . . . 340

3.5. Solids Mixing in Circulating Fluidized Beds 340

4. Gas Mixing in Fluidized-Bed Reactors . . . 340

4.1. Gas Mixing in Bubbling Fluidized Beds. . . 341

4.2. Gas Mixing in Circulating Fluidized Beds . 341

5. Heat and Mass Transfer in Fluidized-Bed

Reactors . . . . . . . . . . . . . . . . . . . . . . . . . . . 341

6. Gas-Solid Separation. . . . . . . . . . . . . . . . . 343

7. Injection of Liquid Reactants into Fluidized

Beds . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 3438. Industrial Applications . . . . . . . . . . . . . . . 344

8.1. Heterogeneous Catalytic Gas-Phase

Reactions. . . . . . . . . . . . . . . . . . . . . . . . . . 344

8.2. Polymerization of Olefins. . . . . . . . . . . . . . 347

8.3. Homogeneous Gas-Phase Reactions . . . . . . 347

8.4. GasSolid Reactions. . . . . . . . . . . . . . . . . . 348

8.5. Applications in Biotechnology . . . . . . . . . . 352

9. Modeling of Fluidized-Bed Reactors . . . . . 354

9.1. Modeling of LiquidSolid Fluidized-Bed

Reactors . . . . . . . . . . . . . . . . . . . . . . . . . . . 354

9.2. Modeling of GasSolid Fluidized-Bed

Reactors . . . . . . . . . . . . . . . . . . . . . . . . . . . 3549.2.1. Bubbling Fluidized-Bed Reactors . . . . . . . . . 355

9.2.2. Circulating Fluidized-Bed Reactors . . . . . . . 356

9.3. New Developments in Modeling Fluidized-

Bed Reactors. . . . . . . . . . . . . . . . . . . . . . . 357

9.3.1. Computational Fluid Dynamics . . . . . . . . . . 357

9.3.2. Modeling of Fluidized-Bed Systems . . . . . . . 358

10. Scale-up . . . . . . . . . . . . . . . . . . . . . . . . . . . 359

References . . . . . . . . . . . . . . . . . . . . . . . . . 361

Symbols (see also ! Principles of ChemicalReaction Engineering and ! Model Reactorsand Their Design Equations)

a: volume-specific mass-transfer area be-

tween bubble and suspension phases,

m1

A0: cross-sectional area of orifice, m2

Ar: Archimedes number, defined by

Equation (5)At: cross-sectional area of reactor, m2

b: parameter def. by Equation (54)

cv: solids volume concentration

cb: bubble attrition rate constant, defined by

Equation (50), s2/m4

cc: cyclone attrition rate constant defined by


cj: jet attrition rate constant, defined by


Cb: concentration in bubble phase, kmol/m3

Cd: concentration in suspension phase,

kmol/m3

do: orifice diameter, m

DOI: 10.1002/14356007.b04_239.pub2


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1. Introduction

1.1. The Fluidization Principle

In fluidization an initially stationary bed of solidparticles is brought to a fluidized state by an

upwardstreamofgasorliquidassoonasthevolumeflowrateofthefluidexceedsacertainlimitingvalueV_mf(where mf denotes minimum fluidization). In

the fluidized bed, the particles are held suspendedby the fluid stream; the pressure drop Dpfbof thefluidonpassingthroughthefluidizedbedisequaltotheweight of the solids minus thebuoyancy, divid-edbythecross-sectionalareaAt ofthefluidized-bedvessel (Fig. 1):

Dpfb At H 1e rsrf gAt

1

dp: Sauter diameter, defined by Equa-

tion (6), m

dpi: diameter of particle size class i, m

dt: bed diameter, m

dv: local bubble volume equivalent sphere

diameter, mdv0: initial bubble diameter, m

D: coefficient of molecular diffusion, m2/s

Dsh: lateral solids dispersion coefficient, m2/s

Dsv: vertical solids dispersion coefficient,

m2/s

Frp: Froude number, defined by

Equation (29)

Gs: solids mass flow rate, based on reactor

cross-sectional area, kg m2 s1

h: height above distributor level, m

ho: height above distributor where bubbles

are forming, m

hgs: gas-to-solid heat transfer coefficient, W

m2K1

hwb: wall-to-bed heat transfer coefficient, W

m2K

H: expanded bed height, m

Hmf: bed height at minimum fluidization, m

kG: mass-transfer coefficient, m/s

L: jet length, m

ma: mass of elutriated solids, kg

m_att: mass flow due to attrition, kg/s

mb: bed mass, kg

m_s: solids mass flow, g/s

np: number of passages through cyclone

p: pressure, Pa

Per, c: Peclet number, defined by Equation (43)

Q3: cumulative mass distribution

ra: attrition rate, defined by Equation (33),

s1

rj: reaction rate, based on catalyst mass,

kmol kg1 s1

Re: Reynolds number

Sv: volume-specific surface area of parti-

cles, m1

t: time, s

TDH: transport disengaging height, m

u: superficial fluidizing velocity, m/sub: local bubble rise velocity, m/s

uc: velocity at cyclone inlet, m/s

umf: superficial minimum fluidizing velocity,

m/s

uo: jet velocity at orifice, m/s

usl: slip velocity, defined by Equation (27),

m/s

ut: single particle terminal velocity, m/s

V_b: visible bubble flow, based on bed area,

m3 m2 s1

V_mf: minimum fluidizing flow rate, m3/s

V_o: flow rate of gas issuing from orifice, m3/s

xi: mass fraction of particle size fraction i in

bedmaterial

a: velocity ratio, defined by Equation (14)

Dpd: pressure drop of the gas distributor, Pae: bed porosityeb: local bubble gas holdupei: porosity of catalyst particleemf: bed porosity at minimum fluidizationk

*: elutriation rate constant, kg m2 s1

l: average life time of a bubble, sm: solid-to-gas mass flow ration: kinematic viscosity, m2/snij: stoichiometric number of species i in

reaction j

rf: fluid density, kg/m3

rs: solids density, kg/m3

q: stress history parameter, defined byEquation (54)

qb: parameter, defined by Equation (23)y: pressure ratio, defined by Equation (28)

320 Fluidized-Bed Reactors Vol. 15


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In Equation (1), the porosity e of the fluidized bedis the void volume of the fluidized bed (volume ininterstices between grains, not including any porevolume in the interior of the particles) divided bythe total bed volume; rs is the solids apparentdensity; andHis the height of the fluidized bed.

In many respects, the fluidized bed behaveslike a liquid. The bed can be stirred like a liquid;objects of greater specific gravity sink, whereasthose of lower specific gravity float; if the vesselis tilted, the bed surface resumes a horizontal

position; if two adjacent fluidized beds withdifferent bed heights are connected to each other,the heights become equal; and the fluidized bedflows out like a liquid through a lateral opening.Particularly advantageous features of the fluid-ized bed for use as a reactor are excellent gassolid contact in the bed, good gasparticle heatand mass transfer, and high bedwall and bedinternals heat-transfer coefficients.

The fluidization principle was first used on anindustrial scale in 1922 for the gasification of fine-

grained coal [1]. Since then, fluidized beds havebeen applied in many industrially important pro-cesses. The present spectrum of applicationsextends from a numberof physical processes, suchas coolingheating, drying, sublimationdesubli-mation, adsorptiondesorption, coating,and gran-ulation, to many heterogeneous catalytic gas-phase reactions as well as noncatalytic reactions.

What follows is a survey of the fluid mechan-ical principles of fluidization technology, gas andsolid mixing, gassolid contact in the fluidized

bed, typical industrial applications, and ap-proaches to modeling fluidized-bed reactors.Further information is given in textbooks (e.g.,

[2]) and monographs (e.g., [38]). Summarytreatments can also be found in [919]. Otheruseful literature includes reports of the Engineer-ing Foundation Conferences on Fluidization[2022], the Circulating Fluidized Bed Confer-ences (e.g., [2325], and for use of the fluidizedbed in energy technology the Fluidized BedCombustion Conferences (e.g., [2628]).

1.2. Forms of Fluidized Beds

As the volume flow rate V_ or the superficialvelocity u V_/At of the fluid increases beyondthe valueV_mforumf(Fig. 2 A) corresponding tothe minimum fluidization point, one of twothings happens: influidization with a liquid, thebed begins to expand uniformly; in fluidizationwith a gas a process of greater industrialimportance and the one discussed almost exclu-sively in the following material virtuallysolids-free gas bubbles begin to form (Fig. 2 B).

The local mean bubble size increases rapidlywith increasing height above the grid because ofcoalescence of the bubbles. If the bed vessel issufficiently narrow and high, the bubblesultimately fill the entire cross section andpass through the bed as a series of gas slugs(Fig. 2 C). As the gas velocity increases further,more and more solids are carried out of the bed,the original, sharply defined surface of the beddisappears, and the solids concentration comesto decrease continuously with increasing height.

To achieve steady-state operation of such aturbulent fluidized bed (Fig. 2 D), solidsentrained in the fluidizing gas must be collected

Figure 1. Pressure drop in flow through packed and fluidized beds

Vol. 15 Fluidized-Bed Reactors 321


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and returned to the bed. The simplest way to dothis is with a cyclone integrated into the bedvessel and a standpipe dipping into the bed. Afurther increase in gas velocity finally leads tothe circulating fluidized bed (Fig. 2 E), which ischaracterized by a much lower average solidsconcentration than the previous systems. The

high solids entrainment requires an efficientexternal solids recycle system with a speciallydesigned pressure seal (shown as a siphon inFig. 2 E).

1.3. Advantages and Disadvantages of

the Fluidized-Bed Reactor

The major advantages of the (gassolid) fluidized

bed as a reaction system include

1. Easy handling and transport of solids due toliquid-like behavior of the fluidized bed

2. Uniform temperature distribution due to in-tensive solids mixing (no hot spots even withstrongly exothermic reactions)

3. Large solidgas exchange area by virtue ofsmall solids grain size

4. High heat-transfer coefficients betweenbed and immersed heating or cooling

surfaces5. Uniform (solid) product in batchwise process

because of intensive solids mixing

Set against these advantages are the followingdisadvantages:

1. Expensive solids separation or gas purifica-tion equipment required because of solidsentrainment by fluidizing gas

2. As a consequence of high solids mixing rate,

nonuniform residence time of solids, back-mixing of gas, and resulting lower conversion

3. In catalytic reactions, undesired bypass orbroadening of residence-time distribution forreaction gas due to bubble development

4. Erosion of internals and attrition of solids(especially significant with catalysts), result-ing from high solids velocities

5. Possibility of defluidization due to agglomer-ation of solids

6. Gassolid countercurrent motion possible on-ly in multistage equipment

7. Difficulty in scaling-up

Table 1 compares the fluidized-bed reactor withalternative gassolid reaction systems: fixed-bed, moving-bed, and entrained-flow reactors.

2. Fluid-Mechanical Principles

2.1. Minimum Fluidization Velocity

Theminimum fluidization point, which marks theboundarybetweenthefixed-andthefluidized-bed

Figure 2. Forms of gassolids fluidized beds



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conditions, can be determined by measuring thepressure dropDpacross the bed as a function ofvolume flow rate V_(Fig. 1). Measurement should

always be performed with decreasing gas veloci-ty, by starting in the fluidized condition.

Only for very narrow particle-size distribu-tions, however, does a sharply defined minimumfluidization point occur. The broad size distribu-tions commonly encountered in practice exhibit ablurred range; conventionally, the minimumfluidization point is defined as the intersectionof the extrapolated fixed-bed characteristic withthe line of constant bed pressure drop typical ofthe fluidized bed (Fig. 1).

The measurement technique already containsthe possibility of calculating the minimum flu-idization velocityumf: The pressure drop in flow

through the polydisperse fixed bed at the pointu umf, given, for example, by the Ergun rela-tion [29] (

!Fluid Mechanics), is set equal to the

fluidized-bed pressure drop given by Equa-tion (1). From the Ergun relation

Dp

h 4:17 S2v

1e2e3

hu0:29Sv 1ee3

rfu2

it follows

umf 7:14 1emfn Sv

ffiffiffiffiffiffiffiffiffiffi ffiffiffiffiffiffiffiffiffiffiffiffiffi ffiffiffiffiffiffiffiffiffiffiffiffi ffiffiffiffiffiffiffiffiffiffiffiffiffi ffiffiffiffiffiffiffiffiffiffiffi ffiffiffiffiffiffiffiffiffiffiffiffiffi

10:067 e3mf

1emf2rsrf g

rfn2

1S3v

vuut

1

2

64

3

752

Accordingly, to calculateumf, the characteristicsof the gas (rf, n), the density rsof the particles,

Table 1. Comparison of gassolid reaction systems [2, 18]



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the porosityemfof the bed at minimum fluidiza-tion, and the volume-specific surface area Sv ofthe solids must be known. The specific surfacearea defined by

Sv

surface area of all particles in the bed

volume of all particles in the bed

(this takes into account only the external area,which governs hydraulic resistance, not the poresurface area as in porous catalysts) cannot bedetermined very exactly in practice. Hence umfshould not be calculated on the basis of themeasured particle-size distribution of a represen-tative sample of the bed solids; instead, it is bettermeasured directly. Equation (2) can be em-ployed advantageously to calculate umf in anindustrial-scale process on the basis of minimumfluidization velocities measured in the laboratoryunder ambient conditions [30].

An equation from WEN and YU [31] can beused for approximate calculations:

Remf 33:7 ffiffiffiffiffiffiffiffiffiffiffi ffiffiffiffiffiffiffiffiffiffiffiffiffi ffiffiffiffiffiffiffiffiffiffiffiffi

13:6 105 Arp

1 3where

Remf umfdpn

4

Ar gd3

p

n2 rsrf

rf5

Here the surface mean or Sauter diameter calcu-lated from the massdensity distribution q3 (d) ofthe particle diameters

dp 1Rdmaxdmin

d1 q3ddd6

should be used for the characteristic particlediameterdp.

Both the Ergun approach and the Wen and Yusimplification have been confirmed experimen-tally over a wide range of parameters. Morerecently, Vogt et al. [32] found that Equations(2) and (3) are also applicable to high-pressurefluidized beds in which the fluid is under super-critical conditions

2.2. Expansion of LiquidSolid

Fluidized Beds

The uniform expansion of a bed on fluidizationwith a liquid can be described by

u

ut en 7

according to RICHARDSON andZAKI [33]. Here ut isthe terminal velocity of isolated single particles;the exponentn is given as follows, provided theparticle diameter is much smaller than that of the

vessel:

n 4:65 0 < Ret 0:24:4 Re0:03t 0:2 < Ret 14:4 Re0:1t 1 < Ret 5002:4 500 < Ret

8>>>: 8

The Reynolds number used above is calculatedvia the single-particle terminal velocity ut:

Ret ut dpn

9

2.3. Fluidization Properties of Typical

Bed Solids

In fluidization with gases, solids display charac-teristic differences in behavior that can alsoaffect the operating characteristics of fluidized-bed reactors. GELDART has proposed an empiri-cally based classification of solids into fourgroups (A to D) by fluidization behavior [34].The parameters employed are those crucial forfluidization properties: the mean particle diame-ter (dp) and the density difference (rs rf)between solid and fluid. Figure 3 shows theGeldart diagram with the interclass boundariestheoretically established by MOLERUS [35].

Figure 3. Geldart diagram (boundaries according toMOLERUS [35])For explanation see text



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Solids of Group C are very fine-grained,cohesive powders (e.g., flour, fines from cyclonesand electrostatic filters) that virtually cannot befluidized without fluidization aids. The adhesionforces between particles are stronger than theforces that the fluid can exert on the particles. Gas

flow through the bed forms channels extendingfrom the grid to the top of the bed, and thepressure drop across the bed is lower than thevalue from Equation (1). Fluidization propertiescan be improved by the use of mechanical equip-ment (agitators, vibrators) or flowability addi-tives, e.g., Aerosil.

Solids of Group A have small particle dia-meters (ca. 0.1 mm) or low bulk densities; this

class includes catalysts used e.g., in the fluidized-bed catalytic cracker. As the gas velocity uincreases beyond the minimum fluidizationpoint, the bed of such a solid first expandsuniformly until bubble formation sets in at uumb> umf. The bubbles grow by coalescence butbreak up again after passing a certain size. At aconsiderable height above the gas distributorgrid, a dynamic equilibrium is formed betweenbubble growth and breakup. If the gas flow is cutoff abruptly, the gas storage capacity of the

fluidized suspension causes the bed to collapserather slowly.

Group B Solids have moderate particlesizes and densities. Typical representatives ofthis group are sands with mean particle diametersbetween ca. 0.06 and 0.5 mm. Bubble formationbegins immediately above the minimum fluidi-zation point. The bubbles grow by coalescence,and growth is not limited by bubble splitting.

When the gas flow is cut off abruptly, the bedcollapses quickly.

Group D includes solids with large particlediameters or high bulk densities; examples aresands with average particle diameters > 0.5 mm.Bubbles begin to form just above the minimumfluidization point, but the character of bubbleflow is markedly different from that in group Bsolids: group D solids are characterized by the

formation of slow bubbles (Section 2.7). Onsudden stoppage of the gas flow, the bed alsocollapses suddenly.

2.4. State Diagram of Fluidized Bed

Whereas the onset of the fluidized state can bedescribed by the minimum fluidization velocity,the bed operating range and the gas velocityneeded to create a given fluidized state can be

estimated with the help of the fluidized-bed statediagram (Fig. 4) devised by REH[36]. This plotshows the fluid mechanical resistance character-istics of the fixed bed, fluidized bed, and pneu-matic transport. The ordinate is the quantity

3

4

u2

g dp rf

rsrfand the abscissa is the Reynolds number Repformed with the fluidization velocity u and theparticle diameter d

p. The state parameter in the

fluidized-bed region is the mean bed porosity e.The use of the diagram is facilitated by anauxiliary grid with lines of constant Mand con-stant Archimedes number. While the dimension-less groups plotted as ordinate and abscissa eachcontain both the particle diameter and the fluidi-zation velocity, this is not the case with theparameters Arand Mdefined by

Ar g d3p

n2 rsrf

rf

10

Figure 4. Reh status diagram with status points S and S1S4(for explanation, see text)



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M u3

g n rf

rsrf 11

The Reh status diagram can answer a number ofpractical questions. If, for example, the proper-ties of the gas (rf, v) and the solid (dp, rs ) and thefluidization velocityu are given, the calculationofArandRepyields, via the status point S in thediagram (Fig. 4), the average voidage e in thefluidized bed. Taking the lineM const. throughS at the intersection with the line e! 1 a t S1 givesinformation on the particle size which is justelutriatedwhen a particles with a size distributionare fluidized, and the intersection of the same linewith the fixed-bed limit e 0.4 (S2) indicates theparticle size at which fluidization will breakdown if agglomeration occurs. The line Arconst. through S can be used to find the minimumfluidization velocity at S3 or as a measure of theupper limit of fluidization the maximum fluid-izing velocity at S4.

An important practical point is that the statediagram implies a classification scheme that

relates various fluidized-bed systems to one an-other [37, 38] (Fig. 5). When a new fluidized-bedprocess is being designed, the position of the statepoint in the diagram will identify related fluid-ized-bed systems with potentially similar oper-ating problems.

2.5. Gas Distribution

The gas distribution device must satisfy thefollowing requirements:

1. Ensure uniform fluidization over the entirecross section of the bed (especially importantfor shallow beds)

2. Provide complete fluidization of the bed with-out dead spots where, for example, depositscan form

3. Maintain a constant pressure drop over longoperation periods (outlet holes must not be-come clogged)

Figure 5. Rehs fluidized-bed state diagram with operating regions of different reaction systemsa) Circulating fluidized bed; b) Fluidized-bed roaster; c) Bubbling fluidized bed; d) Shaft furnace; e) Moving bed



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Often, the gas distributor design must also pre-vent solids from raining through the grid bothduring operation and after the bed has been shutoff.

Porous plates of glass, ceramics, metal, orplastic are commonly used as gas distributors in

laboratory apparatus; a variety of designs areused in pilot-plant and full-scale fluidized-bedreactors (see Fig. 6). Many more designs can befound, for example, in [2] and [39].

The principal requirement uniform distribu-tion of fluidizing gas over the bed cross section can be met if the pressure dropDpd across the gasdistribution grid is large enough. Suggestedvalues for the ratio Dpd/Dpfb are 0.10.3 (witha minimum Dpd of 3.5 kPa) [40], 0.20.4 [41],and> 0.3 [42].

For a given pressure drop Dpd the gas velocityin the nozzle uocan be calculated from

Dpd ro2

CD u2o

where ro is the gas density in the orifice and CD isthe drag coefficient. Applying the continuityequation

V: No Ao uo

either the number of nozzles No

or the cross-sectional area of the individual nozzle Aocan becalculated for a given gas flow rate V_.

Problems related to the design of gas distri-butors are attrition of solids (see Section 2.11),

erosion, and back-flow of solids. Erosion mayoccur at the distributor plate and at neighboringnozzles or walls due to gas jets as well as at thenozzle itself. Back-flow of solids into the wind-box is caused by pressure fluctuations. In order toprevent this either the design pressure drop has to

be larger than the pressure fluctuations or if thisis not feasible for economic reasons a designmust be chosen which tolerates short periods ofgas flow reversal without permitting the solids topenetrate into the windbox. For the latter case thebubble cap design has turned out to be advanta-geous [43].

In the operation of fluidized-bed reactors, thequadratic response (Dpd u2) of industrial gas-distributor designs must be kept in mind, becauseeven if the fluidization velocity is lowered onlyslightly, an unacceptably low pressure dropacross the gas distributor may occur. Industrialexperience with different distributor designs,practical design rules, and a discussion of dis-tributor-related problems, such as weepage intothe windbox and erosion by grid jets and at gridnozzles, has been compiled in [44].

2.6. Gas Jets in Fluidized Beds

Gas jets can form at the outlet openings ofindustrial gas distributors and also where gaseousreactants are admitted directly into the fluidizedbed. A knowledge of the geometry of such jets, in

Figure 6. Industrial gas distributorsA) Perforated plate; B) Nozzle plate; C) Bubble-cap plate



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Ifa < 1, some of the gas in the suspensionphase undergoes short-circuit flow through thebubble, while only part of the bubble gas recir-culates through the suspension. This type of flowis typical for fluidized beds of coarse particles(Geldart group D).

Under the real operating conditions of a flu-idized-bed reactor, a number of interacting bub-bles occur in the interior of the fluidized bed. As arule, the interaction leads to coalescence. Asdetailed studies have shown, this process is quitedifferent from that between gas bubbles in liquidsbecause of the absence of surface-tension effectsin the fluidized bed [48, 49].

For predicting mean bubble sizes in freelybubbling fluidized beds, a differential equationfor bubble growth should be used in the case ofGeldart group A and B solids [50]:

d

dhdv 2eb

9p

13

dv3lub

15

with the following boundary condition at h ho:

dv0

m

0:008e1=3b porous plate

1:3

V: 2

o

g

0:2

industrial gas distributor 16

8>>>:where ho is the height above the grid where thebubbles form (for a porous plate, ho 0; for aperforated plate,ho L; for a nozzle plate,hoisthe height of the outlet opening above the plate;and for a bubble-cap plate,hois the height of thelower edge of the cap above the plate). V_0is thevolume flow rate of gas through the individualgrid opening.

The local volume fraction of bubble gasebisgiven by

eb V:

b=ub 17

and the visible bubble flow V_b is

V:

b 0:8 uumf 18

The upward velocity ub of bubbles depends notonly on the bubble size but also on the diameter dtof the fluidized bed:

where

ub V:

b0:71_sb _sffiffiffiffiffiffiffi

gdvp

19

b 3:2d0:33t 0:05 dt 1 m; Geldart group A

3:2 d0:5t 0:1 dt 1 m; Geldart group B

20

Outside these limits,b is taken as constant.The differential equation (Eq. 15) describes

not only bubble growth by coalescence but also

the splitting of bubbles (second term on the right-hand side [51]). The crucial parameter here is themean bubble lifetimel:

l 280 umfg

21

In practice, bubble growth is limited not only bythe splitting mechanism based on the particle-size distribution of the bed solids, but also byinternals (screens, tube bundles, and the like) thatcause bubbles to break up. Computational tech-

niques for estimating this process are given in[52, 53].

HILLIGARDT and WERTHERhave derived a cor-responding bubble-growth model for coarse-par-ticle fluidized beds (Geldart group D) [50].

An example of a measured and calculatedbubble-growth curve is presented in Figure 8.

2.8. Elutriation

When bubbles burst at the surface of the fluidizedbed, solid material carried along in their wake isejected into the freeboard space above the bed.The solids are classified in the freeboard; parti-cles whose settling velocityutis greater than thegas velocity fall back into the bed, whereasparticles with ut < u are elutriated by the gas

Figure 8. Bubble growth in a fluidized bed of fine particles(Geldart group A; data points from [54], calculation from[50])



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stream. As a result, both the volume concentra-tion of solids cv and the mass flow rate ofentrained solids in the freeboard show a charac-teristic exponential decay (Fig. 9). With increas-ing height above the bed surface, the transportdisengaging height (TDH) is finally reached.Here the increased local gas velocities due tobubble eruptions have decayed, and the gasstream contains only particles with ut < u. WhentheTDHcan be reached in a fluidized-bed reac-

tor, this is associated with minimum entrainedmass flow rates and solids concentrations, andhence with minimum loading on downstreamdust collection equipment. Design of the dustcollection system requires knowledge of the en-trained mass flow rate Gs and the particle-sizedistribution of the entrained solids. For the designof the fluidized-bed reactor, the distribution cv (h)of the solids volume concentration and, for gassolid reactions, the local particle-size distributionas a function of height in the freeboard must be

known.For solids of Geldart group A, the TDH

can be estimated with the diagram shown inFigure 10 [55]. The following relation is givenfor the TDH of Geldart group B solids as afunction of the size dvof bubbles bursting at thebed surface [56]:

TDH 18:2 dv 22Equation (25) was, however, derived for a

bench-scale unit and may not scale to plant-size

equipment.The mass flow rateGsof entrained solids per

unit area leaving the fluidized-bed reactor is the

sum of contributions from the entrainable parti-cle size fractions (ut < u):

GsX

i

xi k*i 23

Here xi is the mass fraction of particle-sizefraction i in the bed material and k*i is theelutriation rate constant for this fraction. Theliterature contains a number of empirical corre-lations for estimating k*i (e.g., [24]). Morephysical-based are the elutriation models ofWENand CHEN[57] and of KUNIIand LEVENSPIEL[2, 58], which enable not only calculation of theexiting mass flow rate but also estimation of theconcentration versus height cv (h) in the free-board. The model by SMOLDERS and BAEYENS

additionally takes the effect of variable freeboardgeometry into account [59].

A literature survey on the factors affectingelutriation and the available modeling tools isgiven in [60].

2.9. Circulating Fluidized Beds

2.9.1. Hydrodynamic Principles

In REHs state diagram of the fluidized bed [36],the circulating fluidized bed (CFB) is locatedabove the single-particle suspension curve for

Figure 10. Estimation of transport disengaging height(TDH), according to [55]umb Fluidization velocity at which bubble developmentbegins

Figure 9. Schematic drawing of fluidized bed and freeboard



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Re < 102 and porositiesegreater than about 0.8(dashed line in Fig. 5). The shortcoming of thisdiagram is that it does not show an importantparameter in the operation of a circulating fluid-

ized bed: the circulating solids mass flow rate perunit area Gs. The diagram of Figure 11 [61]attempts to remedy this by plotting the mean slipvelocity usl between gas and solids

usl ueGs=rs

1e 24

versus the mean solids concentration cv 1e,with Gs as the parameter. The limiting condi-tions are high solids concentration (bed at mini-mum fluidization) and cv

! 0 with usl

ut

(isolated single particle). In the circulating flu-idized-bed region, slip velocity increases withincreasing Gsand can become much higher thanthe single-particle settling velocity (the physical

justification for this statement comes from theformation of strands or clusters of particles). Inthe entrained-flow region the slip velocitiesagain decrease with decreasing solids concen-tration.

The fluidized-bed state diagrams discussedthus far, as well as others (e.g., [62, 63]), aresuitable mainly for the qualitative interpretationof flow phenomena. A diagram proposed byWIRTH(e.g., [11, 64, 65]) also provides quantita-tive assistance in the design of circulating fluid-ized beds. The schematic in Figure 12 applies to agiven gassolid system described by a constantvalue of the Archimedes numberAr. The ordinateis the dimensionless pressure drop of the fluid-ized bed

y Dp

rsrf 1emf g Dh 25

the abscissa is the particle Froude number

Frp uffiffiffiffiffiffiffiffiffiffi ffiffiffiffiffiffiffiffiffirsrfrf

gdp

q 26The dimensionless pressure drop yis the ratio ofthe pressure dropDp along the flow pathDh to themaximum possible value for ascending flow (the

value that would be attained if the pipe crosssection were filled with solids corresponding tothe concentration at the minimum fluidizationpoint). The parameter of the family of curves is avolume flow rate ratio

m rfrs 1emf

27

where m is the ratio of solid-to-gas mass flowrates. The limiting curve bounds the region ofstable, vertically upward gassolid flow on the

low gas velocity side.Figure 13 shows how the state diagram of

Figure 12 is constructed for a circulating fluid-ized bed with siphon recycle. If solids holdup inthe recycle line and siphon is ignored, this caserepresents operation with a constant bed massindependent of velocity. At high gas velocitiesand if acceleration effects are neglected, the bedmaterial is distributed uniformly over the totalheightHcfbof the fluidized bed (Fig. 13 C). Thecirculating fluidized bed then exhibits a single

steady-state section with a constant pressuregradient (Dp/Dh). This pressure gradient can becalculated from the bed mass as

Figure 12. State diagram for the circulating fluidized bedwith siphon, according to WIRTH [64]Ar const.,parameter of family of curvesis the volume flowrate ratiom rf/(rs (1emf));Frp particle Froude numberfor superficial minimum fluidization velocity (pumf), single-particle terminal velocity (pt), and transport velocity (pT),respectively

Figure 11. Fluidized-bed state diagram, according to [61]



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yhomrsrfgHmf1emfrsrfgHcfb1emf

HmfHcfb

28

where Hmf is the bed height at minimumfluidization.

The states identified byyhom to the right of thebounding curve in Figure 12 are accessible byincreasing the gas velocity (corresponding toincreasingFrp). With increasingFrpthe volumeflow ratio increases; that is, relatively more solidsare elutriated (and thus circulated).

IfFrp is allowed to drop below the limitFrpmax(Fig. 13 B, Fig. 12) two steady-state sectionsappear in the riser tube: the one in the lower partis marked by a high pressure gradient, that in theupper part by a lower gradient. Figure 13 illus-trates the physical significance of these twopressure gradients. In practice, the transitionbetween the two linear regions takes place grad-ually. The height of the transition zone corre-sponds to the transport disengaging height

(TDH).The picture changes further if the gas velocitydeclines to values lower than the settling velocityut of a single isolated particle. In this case (forFrp < Frpt, Fig. 13 A, Fig. 12), no more solidscan be elutriated, and the pressure gradient in theupper linear region vanishes. All the solid mate-rial is now in the form of a bubbling or turbulentfluidized bed.

The solids concentrations averaged over thetube cross section (1

e) can be calculated from

the dimensionless pressure drop:

1e 1emf y 29

Besides the pressure and solids concentration pro-file, the circulating mass flow rate of solids Gs Atis important for the design of the circulating fluid-ized bed. In particular, the design of the solidscollection and recycle system depends very muchon this quantity. The mass flow rate of solidsdepends on the flowregime. At gas velocities suchthat twosteady-state sections arepresent in the bedvessel (i.e.,Frpumf < Frp< FrpT), the mass flowrate of entrained solids depends on the physicsof the gassolid flow. Figure 14 plots the

Figure 13. Pressure profile in the circulating fluidized bed with siphon, according to WIRTH[64]A) Frpumf < Frp < Frpt; B) Frpt < Frpmax ; C)Frp > Frpmax

Figure 14. Elutriation diagram when the circulating fluidizedbedcontains two steady-state sections, according to WIRTH [64]



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dimensionless solids mass flow rate versus Frp,with the Archimedes number as parameter. For agivenAr, the flow rate tends to zero as Frp! Frptand reaches a maximum at Frp FrpT. The slopeof the elutriation curve becomes greater withincreasingAr; that is, the coarser the particles, the

greater is the relative change in the circulatingmass flow rate of solids with a change in gasvelocity.

At high gas velocities in the circulating fluid-ized bed (i.e., when a single steady-state sectionexists), the entrained mass flow rate depends onthe particle Froude number and the solids holdup.More detailed information about the applicationofWirthstheoryinpracticemaybefoundin[11].

Whereas WIRTHs analysis of the circulatingfluidized bed starts from the pneumatic transportcondition, the models of RHODES and GELDART[66], as well as KUNIIand LEVENSPIEL[2, 58], arebased on the bubbling fluidized bed and describethe circulating fluidized bed as a limiting case ofa bubbling bed with a very high rate of solidsentrainment.

2.9.2. Local Flow Structure in Circulating

Fluidized Beds

The Wirth state diagram, as a first step toward thelocal characterization of flow regimes in a circu-lating fluidized bed, describes the vertical profileof the solids concentration. In the lower sectionof a circulating fluidized bed a dense region exitsnear the gas distributor. It has been observed thatin this bottom zone bubble-like voids coexistwith a surrounding dense suspension. The solidsvolume concentration is higher at the wall(cv

0.4) then in the center (cv

0.15) of the

bottom zone [67]. The splash zone which links thebottom zone to the upper dilute zone is charac-terized by violent gassolid mixing. Many recentexperimental studies with various measurementtechniques (e.g., X-ray tomography [68], capaci-tance tomography [69] and fiber-optical probes[70]) have shown that the upper section of thecirculating fluidized bed exhibits characteristichorizontal profiles, with the concentrationcv, wallnear the vessel wall always significantly higherthan the value cv averaged over the vessel cross

section; for example,cv, wall2.3cv[71].Local measurements of the solids concentra-

tion and solids velocity show that upward-

flowing regions of low solids concentration anddownward-flowing aggregates of high solidsconcentration alternate in time at every pointinside the fluidized bed, with downward-movingaggregates (strands, clusters) predominating nearthe wall and upward-moving regions of lowsuspension concentration predominating in thecentral zone. However, no significant downwardflow of solids near the wall was observed in high-density circulating fluidized beds, e.g., [72]. Thepicture of the local flow structure in a circulating

fluidized bed, as derived from these observations,is shown schematically in Figure 15.A modeling approach which is based on the

local flow structure of the CFB is the energy-minimization multiscale (EMMS) model [73]. Itconsiders the tendency of a fluid in a gassolidtwo-phase flow to pass through the particulatelayer with least resistance and the tendency of thesolids to maintain least gravitational potential.Least resistance means that the volume-specificenergy consumption for suspending and trans-

porting solids is minimized, and minimization ofthe gravitational potential is equivalent to therequirement that the local mean voidage eattainsa minimum. The model has been applied as adescription of fluid-mechanical phenomena inCFB risers of different sizes [74, 75] but also forthe prediction of flow patterns of gas and solids inindustrial-scale units, such as a CFB boiler [76]and a petrochemical processing unit [77].Anotherpromising line of development is the introductionof the EMMS concept into computational fluid

dynamical calculations of multiphase flows; firstresults obtained with a drag model based on theEMMS model are encouraging [78].

Figure 15. Schematic diagram of flow structure in a circu-lating fluidized bed



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2.9.3. Design of Solids Recycle System

Solids carried over with the fluidized gas aregenerally collected in cyclones. In the case ofbubbling beds, the solids can easily be returned tothe bed through the standpipe of the cyclone,which dips directly into the bed.

Due to the large amounts of circulating solids,circulating fluidized beds require very large cy-clones arranged beside and outside the bed, withspecial valves needed to connect the standpipeto the bed vessel. Figure 16 shows two design

options, the siphon and the L-valve. With thesiphon, the solids are fluidized (i.e., enabled toflow back into the reactor). In the L-valve design,

the mass flow rate of the solids can be regulatedby varying the gas supplied to the standpipe.

Because the solids path does not contain anysortof mechanical closure, the characteristic pres-sure distribution plotted in Figure 17 is obtained.The distribution of solids between the fluidized

bed and the recycle line is directly related to thispressure distribution. Operating properties differfrom one recycle design to another [79].

2.10. Cocurrent Downflow Circulating

Fluidized Beds (Downers)

A certain drawback of circulating and bubblingfluidized beds when applied for gas-phase reac-

tions is the backmixing which inevitably occursin the gas phase. In bubbling fluidized beds it isthe bubble-induced solids circulation, and incirculating fluidized beds the downflow of solidsin the wall zone, which entrains gas in theupstream direction and thus lowers the yield ofa catalytic reaction or gives rise to undesiredconsecutive or side reactions. These disadvan-tages caused by the hydrodynamic effects of bothgas and solids flowing against gravity could beovercome in the so-called downer reactor, in

which the flow directions of both gas and solidsare downward, i.e., in the same direction asgravity [80]. Another incentive is the possibility

Figure 17. Pressure distribution in solids recycle system of a circulating fluidized beda) Fluidized bed; b) Return leg

Figure 16. Design options for solids recycleA) Siphon; B) L-valve



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of realizing short contact times between gas andsolids of around or even below one second.

Downer systems have been intensely studied[80]. Hydrodynamics [81, 82], gas mixing [83],and solids mixing [84, 85] have been investigatedboth experimentally and by numerical simulation

[86]. It has been found that the hydrodynamics ofthe downer are also characterized by a wall zoneof increased solids concentration. However, axialand radial gas-solids flow structures are muchmore uniform than in conventional fluidizedbeds. Another result is that the length of the flowdevelopment zone is much shorter for the downerthan for the riser, which means that reactions withvery short contact times can be carried out undernear-plug-flow conditions. However, the solidsfeeding process and the geometry of the entranceregion are critical points that deserve specialattention [87].

The patent and open literature suggest variousapplications for downer reactors, e.g., residual oilcracking [88], coal pyrolysis [89], and biomasspyrolysis [90]. The catalytic pyrolysis of heavyfeeds for the production of light olefins has beeninvestigated on the laboratory scale with prom-ising results [88]. However, no large-scale indus-trial process has emerged yet.

2.11. Attrition of Solids

The attrition of solid particles is an unavoidableconsequence of the intensive solids motion in thefluidized bed. The attrition problem is especiallycritical in processes where the bed material needsto remain unaltered for the longest possible time,as in fluidized-bed reactors for heterogeneouscatalytic gas-phase reactions. Catalyst attrition

is important in the economics of such processesand may even become the critical factor.

Catalyst attrition in fluidized-bed reactorsoccurs normally as surface abrasion (Fig. 18)which means that surface asperities are abradedand edges of the catalyst particles are roundedoff. Fragmentation may also play a role, espe-cially for some fresh catalyst particles which onentering the reactor may simply be crushed intopieces. If in an industrial process extraordinarilyhigh catalyst losses are observed it is advisable to

examine catalyst samples under the scanningelectron microscope. If the sample containsmany fragments this could be an indication of

a wrong design (e.g., too high a velocity at thecyclone inlet or at the distributor).

When designing catalytic fluidized-bed pro-cesses, the attrition performance of candidatecatalysts should be tested under standardizedconditions in the process development stage.This test can be performed in a small laboratoryapparatus; it consists essentially of an extendedfluidization test in which the mass of solids

carried out of the bed is recorded as a functionof time. Figure 19 presents a typical test result:during the first hours of testing, both the attritedmaterial and the fine fraction of the bed materialare elutriated. Only after a relatively long oper-ating period is a quasi-steady state attained. The

Figure 19. Result of an attrition measurement

Figure 18. Attrition modes and their effects on the particlesize distribution (q3 mass density distribution ofparticle sizes dp)



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attrition rate ra in this steady state can bedefined as

ra 1mb

DmaDt

30

where ma is the elutriated mass and mb the bed

mass. Usually ra is expressed as percentage perday; for relatively attrition-resistant, fluidized-bed catalysts,it isof the order of0.2 % per day [9].

Many standard test apparatuses have beenproposed for comparative attrition tests (e.g.,[91, 92]), but all such equipment has been suit-able only for comparative studies of differentcatalysts under consideration for the sameprocess. The attrition measured in large-scaleequipment can be far different from the values

measured in a test apparatus.A number of sources can be identified forcatalyst attrition in industrial fluidized-bedreactors:

1. Jet attrition at gas distribution grid openingsand nozzles where gaseous reactants areadmitted to the bed

2. Bubble attrition in the bed due to solidsmotion caused by bubbles

3. Attrition in cyclones

4. Attrition in pneumatic conveying lines, suchas those between reactor and regenerator beds

Empirical correlations are available for theattriting action of a gas jet in the fluidized bed[93] and for the size reduction effect of solidsmotion in the bed [94, 95].

WERTHER and coworkers [96] employ thelaboratory apparatus shown schematically inFigure 20 which enables separate study of theattrition due to jets from nozzles of various

diameters and that due to bubbles.Under steady-state conditions the jet-attrition-

related mass production of fines per unit time fora gas distributor with a number no of orifices frommother particles with diameter dp,i which arepresent in the catalyst inventory with a massfractionDQ3i is proportional to the particle sizedpi, the mass fractionDQ3i, the densityroof thegas issuing from the orifice, the square of theorifice diameterdo, and to the cube of the jet exitvelocity u

o[97, 98]:

m att; jet;i cjn0dpiDQ3ir0d20 u30 31

Attrition due to the bubble induced solids move-ment is given by [98]

m att; bubble;i cbdpiDQ3imbuumf3 32where mb denotes the bed mass which contains

bubbles (i.e., which is located outside thejet-dominated grid region). Equation (32) alsodenotes the mass production of attrited fineswhich is resulting from the size fraction dpi inthe bed.

The stress on the catalyst particles will bedifferent in contact with a gas jet, in the bulk ofthe bubbling fluidized bed, and during its passagethrough a cyclone. Recent investigations ofcyclone-induced catalyst attrition [99101] haveshown that the mass flow of attrited fines which is

produced by attrition inside the cyclone when asolids mass flow m_cDQ3ciof particles of the sizefraction dpi enters the cyclone is given by

m att; c;i cc m c DQ3ci dpi u2cffiffiffiffiffimc

p 33

whereucis the gas velocity at the cyclone inlet,andmc the solids loading of the incoming gas flow

mc m crc

uc

Ac

34

wherercis the density of the inflowing gas, andActhe cross-sectional area of the cyclone inlet.

Figure 20. Experimental apparatus for attritionmeasurement



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Equations (31)(34) describe the catalystattrition under conditions of steady state, i.e.,when the particles are more or less rounded off(Fig. 18). To describe also the initial breakageand attrition of fresh catalyst particles, it isnecessary to follow the fate of the particles ontheir introduction into the reactor, which ispossible with population balance models (cf.Section 9.3.2). Klett et al. [102] and Hartgeet al. [103] have defined a stress historyparameter

t=t*jfor jet--induced attrition

t=t*bfor in--bed attrition

np=n*pfor attrition in cyclones

8>>>: 35

where the definition of the characteristicparameters tj

*, tb*, and np

* can be taken fromFigure 21, np is the number of passages of a givenparticle through the cyclone, andtbandtjare thetime periods during which the particle is sub-

jected to bubble and jet stress, respectively. If it isassumed that the effects of the different stress

mechanisms on the catalyst particles are additive,a uniform treatment of the overall stress historyfor all three attrition mechanisms is given by

m attm att;

1:1 b 1:11=b

1 > 1:11=b

36

The parameter b is characteristic of a givencatalyst. Figure 22 shows measurements withFCC catalyst [103] which lead to b 1.16.Equation (36) allows the description of the stress-

history-dependent attrition rate and can be usedfor the simulation of fluidized bed reactors(see Section 9.3.2).

A variety of approaches exist for reducingattrition in industrial fluidized-bed reactors. The

jet attrition action can be controlled with specialgas distributor designs ([9]; e.g., by the use ofbubble caps, Fig. 6) such that gas jets do not issuedirectly into the bed at high velocity. Attritiondue to bubbles can be lowered by limiting bubblegrowth (avoiding high gas velocities and largebed heights; use of fine catalysts with lowumf, asimplied by Eqs. 18 and 24). Attrition in cyclonescan be prevented, in the simplest case, by repla-cing the cyclones with devices such as filters.Attrition can also be minimized by cutting back

the load on the cyclone, for example, by placingthe cyclones above the TDH. Relatively highcatalyst attrition also occurs in circulating fluid-ized beds where very large quantities of solidsmust be collected in the cyclones.

3. Solids Mixing in Fluidized-Bed

Reactors

The intensive solids mixing typical of fluidized-bed reactors has several effects on performance.In catalytic reactions, the large-scale verticalsolids mixing results in a transport of the gascomponents, adsorbed to the catalyst, so that thegas phase is backmixed and the conversion andselectivity are impaired. In noncatalytic gassolid reactions, the mean solids residence timeand residence-time distribution, as well as thepropagation behavior of the solids from individ-ual feed points, play a role. In general, fast and

strongly exothermic reactions require fairly vig-orous solids mixing to prevent temperature peaksnear the reactant inlet.

Figure 22. Dimensionless attrition rate of FCC catalyst as afunction of stress history.

Figure 21. Dependence of attrition on time (bubble- and jet-induced attrition) and number of passages np through acyclone.



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3.1. Mechanisms of Solids Mixing

The wake of the rising bubbles produces a ratherslight upward and lateral drift of the particles(Fig. 23 A) [104]. In addition, solid particles aredrawn upward in the wake, portions of the wakeare shed at irregular intervals during bubblemotion, and new portions of solids are taken intothe wake (Fig. 23 B). Solids transport in the

wake is essentially the reason that vertical solidsmixing is from one to two orders of magnitudebetter than horizontal mixing.

For reasons of continuity, the upward trans-port of particles by bubbles is coupled with adownward movement in the suspension phasethat surrounds the bubbles. Measurements of thelocal bubble-gas flow have shown that the risingbubbles are not distributed evenly over the bedcross section. As a typical example, Figure 24 Agives a plot of the radial distribution of the

bubble-gas flow at three heights above the gridin a fluidized bed 1 m in diameter. The profile iscomparatively flat in the bottom zone but exhibitsa steeper slope as the height increases, with anannular zone of preferentially rising bubbles. Theresulting circulation of the solids also features anannular region of upward transport in the wakeswith predominantly downward motion of thesolids in the center and at the periphery of thebed (Fig. 24 B).

The large-scale solids circulation can be

reinforced by uneven distribution of the fluidizedgas over the distributor cross section [106].Figure 25 presents examples of industrial fluid-

ized-bed furnaces in which forced circulation ofthe solids is employed to improve coal burnup.

3.2. Vertical Mixing of Solids

The propagation behavior of the solids in afluidized bed can be described by a number ofmodels (e.g., [2, 109]). Most commonly used isthedispersion model, in which solids transport isdescribed by a diffusion law. The numericalvalue of the dispersion coefficient Dsvfor solidsmixing in the vertical direction increases withincreasing gas velocity because of the growth inthe number and size of bubbles. The followingsimple empirical correlation is given for fineparticles (Geldart groups A and B) [2]:

Dsv

m2=s

0:060:1 u

m=s

37

Figure 24. A) Radial distribution of bubble-gas flow; B)Relationship between bubble distribution and solids circula-tion [105]

dt 1 m, quartz sand, umf 0.013 m/s, u 0.2 m/s,Hmf 0.5 m, V_b visible bubble flow

Figure 23. Solids mixing in bubbling fluidized beds due toparticle drift (A) and wake transport (B)a) Cloud; b) Wake



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For aplant-scale fluidized bed(0.9 1.26 m2 inplan, bed height 4 m) equipped with a bundle ofhorizontal tubes, a very similar relation wasderived for a solid of Geldart group B [110]:

Dsv

m2=s

0:056 uumf

m=s

38

Because solids circulation becomes moremarked in larger-diameter fluidized beds, thedispersion coefficient increases rapidly with in-creasing bed diameterdt(Fig. 26). For this casethe following expression is found [2]:

Dsv

m2=s 0:030 dt

m 0:65

39

The above correlations can provide only roughvalues. Other effects observed in practiceinclude, in particular, a significant effect of par-ticle-size range [111, 112].

3.3. Horizontal Mixing of Solids

In gassolid reactions, the propagation behavior

of the solids in the horizontal direction is impor-tant if, for example, the solid material is fed intothe bed at isolated feed points. WERTHER and

coworkers model the horizontal propagation ofcoal in a fluidized-bed furnace, describing thecarbon conversion in terms of a simple first-orderreaction (rate constant k with dimension s1)[113]. The crucial parameter is the ratio k d2t/Dshbetween the rate of the chemical reaction and therate of dispersive mass transport. For high valuesofk(fast reaction), large reactor diameters dt,andlow values of the dispersion coefficient Dsh, the

Figure 25. Fluidized-bed furnaces with forced circulation of solidsA) According to [107]; B) According to [108]

Figure 26. Vertical solids dispersion in fluidized beds of fineparticles (Geldart groups A and B) [2]



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local carbon concentration in the bed exhibits arather steep horizontal profile, resulting in asignificantly nonuniform distribution of gasemissions over the bed cross section.

On the basis of KUNII and LEVENSPIELsmodel of bubble-induced solids mixing [114],

an expression has been derived for calculatingthe horizontal dispersion coefficient Dsh aver-aged over the bed height H, given local bubbleproperties (bubble diameter dv, bubble-gasholdupeb) [115]:

Dsh 0:67 1030:023 1H

ZH0

eb1eb

ffiffiffiffiffiffiffiffiffiffig d3v

q dh 40

This correlation holds for solids of Geldartgroups B and D with Archimedes numbers be-tween 8 600 and 58 000.

3.4. Solids Residence-Time Properties

Many applications of fluidization technologyinvolve continuous processing of solids. Impor-tant considerations in such cases are not only themean solids residence time but also the resi-dence-time distribution. Whereas all elements

have the same residence time in a plug-flowsystem, a stirred tank exhibits a broad distribu-tion of residence times. To a good approxima-tion, the residence-time properties of the fluid-ized bed with respect to the solids are the same asthose of a stirred tank. The mean residence time tis the ratio of the solids massmbin the reactor tothe solids throughput m_s:

t mbm:

s

41

The mass fraction dms/mb of solids having aresidence time between tand t dtisdms

mb 1

t e

tt dt 42

Similarly, the fraction f of solids having aresidence time less than tin the bed is calculatedas

f 1et=t 43The residence-time distribution can be nar-

rowed by placing a number of fluidized beds inseries. Multistage systems of this type are used,for example, in fluidized-bed drying [18].

3.5. Solids Mixing in Circulating

Fluidized Beds

The circulating fluidized bed exhibits a complexgassolid flow pattern as discussed in Section2.9. Different regions can be discriminated with

respect to the prevailing mechanisms of solidsmotion and mixing. An extensive survey onexperimental findings in solids mixing is givenin [116]. In the upper diluted zone of the circu-lating fluidized bed, clusters are formed withmainly upward flow in the core and predominant-ly downwards motion near the wall. While thewall region can be modeled by a plug-flowapproach, the core region exhibits radial gradi-ents. The Peclet number characterizing radial

solids mixing in the core region

Per;s uc 2R*

Dr;s

increases from 150 to 300 with increasing solidsvolume concentrations [117]. A recent investiga-tion of solids mixing in the bottom zone with solidcarbon dioxide as a tracer showed that in this zonesolids are almost ideally mixed in the verticaldirection but lateral mixing is limited with dis-persion coefficients of about 0.1 m2/s which cor-

responds to Peclet numbers of around 40, [118].

Counteracting to solids mixing, segregationoccurs in applications using particles of a broadsize distribution and/or different densities. Easilyfluidized particles tend to be elutriated whileothers tend to sink. A dynamic equilibrium be-tween solids mixing and segregation is estab-lished, causing a spatial distribution of particleswith significantly different solids properties, aswas shown in an experimental study with amixture of iron powder and quartz sand with a

broad particle size distribution [119].

4. Gas Mixing in Fluidized-Bed

Reactors

The mixing and residence-time distribution ofthe gas are particularly important for catalyticreactions but are also significant for gassolidreactions when gaseous reactants are to be con-

verted to the greatest possible extent in fluidizedbeds (e.g., reduction of fine-grained iron ores tosponge iron with gaseous reductants [18]). Gas



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mixing is closely linked to the motion and mixingof the solids in the bed.

4.1. Gas Mixing in Bubbling Fluidized

Beds

If the flow and mixing of gas in the bubblingfluidized bed are described by a simple one-phasedispersion model, the coefficientsDgv andDgh ofgas dispersion in the vertical andhorizontal direc-tions have similar numerical values and followtrends similar to those of the solids dispersioncoefficients. By way of example, Figure 27 showsthe effect of fluidized-bed diameterdton verticalgas dispersion. The increase in dispersion coeffi-

cient withvesseldiameter might be attributable tothe formation of large-scale solids circulationpatterns, which becomes more marked in largerequipment. As in the solids case, the coefficientsof horizontal gas dispersion are a factor of 10100lower than those of vertical gas dispersion.

A single-phase dispersion model gives only arough description of gas mixing in bubblingfluidized beds. A more exact description comesfrom models that take account of local flowconditions in the bed, especially the presence of

bubbles (see Chap. 9).

4.2. Gas Mixing in Circulating

Fluidized Beds

Only a few detailed studies of gas mixing incirculating fluidized beds have been published,

which are summarized in [120]. The bubbles in abubbling fluidized bed influence the gas resi-dence-time distribution and mixing directlythrough the bypass action of the bubble-gas flowand gas exchange between the bubbles and thesurrounding suspension phase, and also indirect-

ly through the solids motion that they induce. Inthe circulating fluidized bed, on the other hand,the gas-mixing properties are controlled by seg-regation due to the formation of solid aggregates(jets, clusters) and the rapid downward move-ment of solids strands predominantly near thewall. GRACE and coworkers, for example, showthat a single-phase dispersion model cannot de-scribe the tracer gas residence-time distributionsthat they measured [121]. They propose instead atwo-phase model featuring exchange between awall zone with stagnant gas and a core zone withplug flow.

For the case of horizontal gas mixing,WERTHERand coworkers [122, 123] have shownthat, for the bed solids they used (quartz sand,dp 0.13 mm, Geldart group B), horizontal gasmixing in the top part of the circulating fluidizedbed in the core zone can be described by themodel of turbulent single-phase flow [124]. ThePeclet number

Per;c uc 2R*

Dr;c44

(defined in terms ofuc, the superficial velocity inthe core zone;R*, the radius of the core zone; and

Dr, c, the horizontal dispersion coefficient in thecore zone) has a value of 465, which is in fairlygood agreement with values measured in single-phase flows [125]. This value is independent ofthe solids circulation rate Gs. The circulating

fluidized bed thus exhibits no especially inten-sive horizontal gas mixing, at least in the uppersection where solids concentrations are relativelylow.

5. Heat and Mass Transfer in

Fluidized-Bed Reactors

Fluidized-bed reactors exhibit a uniform temper-ature distribution even in case of highly exother-

mic or endothermic reactions. Approximationsof the heat transfer rates are necessary for thedesign and control of fluidized-bed reactors in

Figure 27. Vertical gas dispersion in a fluidized bed of solidsof Geldart group A (measurements by various workers; [2])



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order to determine the appropriate design ofinternals for cooling or heating and to estimatethe changes in the performance with changingoperating conditions.

However, up to now there is no general theoryon heat and mass transfer in fluidized beds.

Numerous correlations for the calculation of heatand mass transfer coefficients are reported in theliterature. Since these correlations are mainlybased on experimental investigations performedunder laboratory conditions, they may be differ-ent to the situation in large-scale reactors. Detailson models of heat and mass transfer with theirrespective range of application are given in re-lated surveys, e.g., [1417, 126, 130].

Fluid-to-Particle Heat and Mass Transfer.

Since the particle surface area is very large, fluid-to-particle heat and mass transfer is rarely a limit-ing factor in the design and operation of fluidizedbed reactors. The heat-transfer coefficients offluidized-beds range between characteristic valuesfor flow through a fixed bed and flow around asingle particle [127].

Fixed bed(Rep> 80)

Nu hgs dplg

21:8Re0:5p Pr0:33

Single particle

Nu hgs dplg

20:6Re0:5p Pr0:33

where hgs is the gassolid heat-transfer coeffi-cient, dp the particle size, and lg the thermalconductivity of gas. The mass transfer coefficientcan be determined applying the analogy of heat

and mass transfer by replacing in the aboveformulas the Nusselt number Nu by the Sher-wood numberSh and the Prandtl number Prbythe Schmidt number Sc. For particle Reynoldsnumbers below 100 and for fine particles, thetransfer coefficients are significantly lower thanestimated by the above formulas. If necessary,the effect of adsorption in mass transfer and ofradiation in heat transfer needs to be taken intoaccount additionally.

Heat Transfer to Submerged Surfaces.Heat-transfer coefficients between fluidized bedand submerged surfaces are one or two orders of

magnitude larger than for gases alone [126]. Forsingle phase flow a stagnant gas layer is establishedat the wall causing a hindered heat transfer. Thislayer is disrupted by solids transported at the wall.The solids adsorb heat and are mixed into thefluidized bed [9].

An example of the time-averaged local heattransfer along the circumference of a tube im-mersed horizontally in a fluidized bed is given inFigure 28. It exhibits lower values of the heat-transfer coefficient below the tube due to a gasgap caused by bubbles and lower values on top ofthe tube because of solids being at rest. Withintensified mixing this effect becomes lesssignificant.

The dependence of the heat-transfer coeffi-cient on the superficial gas velocity is illustratedin Figure 29. Fluidized beds of fine particles yielda larger heat-transfer coefficient than coarseparticles. According to MOLERUS and WIRTH[126], different transfer mechanisms can be iden-tified. In case of fine particles, solids act as agentstransporting heat between walls and bed, whereasgas convective transport is the mechanism domi-nating the heat transfer of coarse particles. Theheat-transfer coefficient of particles of interme-diate sizes exhibits a maximum due to the super-

position of these two transport mechanisms.Heat-transfer rates in circulating fluidized bedsare lower than in bubbling fluidized beds due to

Figure 28. Local heat-transfer coefficient around a 35 mm

diameter tube immersed horizontally in a fluidized bed of0.37 mm alumina particles operated at a superficial gasvelocity of 0.8 m/s and a temperature of 500 C, adaptedfrom [128]



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reduced solids volume concentrations and aredominated by clusters and strands [130].

The heat-transfer coefficient increases withincreasing pressure [131] and temperature. Theeffect of radiation has to be considered for tem-peratures above 500C, but opaque particles canform an effective radiation shield [132].

6. Gas-Solid Separation

The fluidizing gas inevitably carries fine catalystparticles by entrainment to the reactor exit. Notonly for environmental reasons (i.e., to minimizeemissions) is it necessary to separate the solidsfrom the gas. It may also be necessary to stop themain reaction and to avoid unwanted side orconsecutive reactions or to protect followingprocess steps or machines from particle-ladenstreams. In fluidized-bed technology cyclonesare mostly used for this purpose. KNOWLTON

[133] has given a survey on the state of the artof cyclone design and application in fluidized-bed reactors.

The cyclone should not be considered as aseparate apparatus following the fluidized bedbut should be seen as an integral part of thefluidized-bed process. The reason is that, notonly in circulating fluidized beds but also inbubbling or turbulent fluidized beds, the catalystparticles which are recovered in the cyclone arerecycled to the fluidized bed. The collection

efficiency of the cyclone is thus responsible formaintaining the particle size distribution in thebed inventory, which in turn determines the

fluidized-bed fluid mechanics and the chemicalperformance of the bed as a reactor. The interre-lation between fluidized bed and cyclone is dis-cussed in Section 9.3.2.

The influence of cyclone performance on theoverall process performance is increasingly con-sidered. For example, PULUPULA et al. [134] in-vestigated the role of cyclones in the regeneratorsystem of a commercial FCC unit. ARNOLDet al.[135] were able to trace the deterioration of plantperformance in theALMA maleicprocess back toproblemswithcycloneefficiency.Achangeofthecyclone design improved the particle size distri-bution of the bed inventory and consequently bedhydrodynamics and chemical conversion. SMITet al. [136] report on cyclone performance inturbulent fluidized-bed Synthol reactors forFischerTropsch synthesis. Carbon deposition onthe catalyst particles influences the bed hydrody-namics, which in turn, via the elutriation mecha-nism, influence cyclone performance.

7. Injection of Liquid Reactants into

Fluidized Beds

The injection of reactants in liquid form into thebed was already an essential part of the firstfluidized-bed catalytic process. In the FCC pro-cess (Section 8.1) crude oil is injected at the baseof the reactor and evaporated in contact with thehot catalyst particles. The direct heat transfer is

very efficient and avoids a separate evaporator forthe feed. The cooling action of the evaporatingreactant is a further advantage in the case of an

Figure 29. Heat-transfer coefficients determined with a tube immersed vertically in a fluidized bed of glass beads of differentsize operated at ambient conditions, adapted from [129, 126]



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exothermal reaction. Liquid-feed injection istherefore practiced not only in the FCC processbut also, for example, in the syntheses of aniline(! Aniline, Section 3.2.), caprolactam (! Cap-rolactam, Section4.1.4.), andmelamine (! Mel-amine and Guanamines, Section 4.1.) and in BP

Chemicals Inovene process [137] for the gas-phase production of low-density polyethylene.

Despite its industrial significance knowledgeabout the mechanisms of liquid mixing andevaporation in the fluidized bed is relativelyscarce. Investigations with nonvaporizing hori-zontal gasliquid spray jets have shown that withproper design of the injection nozzle it is possibleto penetrate over several decimeters into the bedbefore the jet breaks up [138, 139]. On the otherhand, it was found that under vaporizing condi-tions for atomizer nozzles with spray anglesbetween 20 and 120 the injected liquid wettedthe bed particles and subsequently evaporatedfrom their surface while the particles were mix-ing in the bulk of the bed [140, 141]. This lattermechanism helps to transport the reactant awayfrom the location of the nozzle and thus con-tributes to equalization of the feed distributioninside the reactor. The special case that a large oildroplet impinges on a smaller hot catalyst parti-

cle was recently investigated in a 3D directnumerical simulation to analyze dropletparticlecollisions in the Leidenfrost regime [142]. Thecalculations were carried out for conditions pre-vailing near the feed nozzle in an FCC riser.Vapor layer pressure induced by evaporation andthe droplet surface tension are the driving forcesfor droplet recoiling and rebounding. The contacttime for a FCC particle and an oil droplet turnedout to be about 140 ms.

8. Industrial Applications

In this chapter the industrial uses of fluidized-bedreactors are classified as follows:

1. Heterogeneous catalytic gas-phase reactions2. Polymerization of olefins3. Homogeneous gas-phase reactions4. Gassolid reactions5. Biotechnology applications

In each of these areas, the most importantapplications are listed and a few typical examples

are analyzed in more detail. For further descrip-tions of processes, the reader is referred to rele-vant articles in the A series. Complete descrip-tions of industrial uses of the fluidized-bed reac-tor can also be found in [2, 10, 18, 19].

8.1. Heterogeneous Catalytic

Gas-Phase Reactions

The fluidized-bed reactor offers the followingprincipal advantages over the fixed-bed reactorfor heterogeneous catalytic gas-phase reactions:

1. High temperature homogeneity, even withstrongly exothermic reactions.

2. Easy solids handling, permitting continuouswithdrawal of spent catalyst and addition offresh if the catalyst rapidly loses its activity.

3. Ability to operate in the explosion range,provided the reactants are not mixed untilthey are inlet to the fluidized bed. This isbecause the high heat capacity of the bedsolids, together with intensive solids mixing,prevents the propagation of explosions.

Catalytic Cracking. (! Oil Refining, Sec-tion 3.2.). The ease of solids handling was thebasic reason for the success of catalytic crackingof long-chain hydrocarbons in the fluidized bed(Fig. 30). The cracking reaction is endothermicand involves the deposition of carbon on thecatalyst surface, which quickly renders the cata-lyst inactive. Accordingly, the catalyst must becontinuously discharged from the reactor andregenerated in an air-fluidized regenerator bed(b), where its carbon loading is lowered from 12to 0.40.8 wt %. The combustion in this bed

simultaneously furnishes the heat required forthe cracking reactor; the catalyst acts as a heatcarrier. The temperature in the regenerator is570590C and in the reactor, 480540C[2]. In a stripper, steam is admitted to removehydrocarbons adhering to the catalyst before it isforwarded to the regenerator.

With the advent of high-activity zeolite cat-alysts in the 1960s, the bubbling fluidized bed,operated at gas velocities between 0.31 and0.76 m/s [2], was replaced by the riser cracker

(Fig. 31), in which the oil fed in at the bottom ofthe riser (c) is vaporized in contact with the hotcatalyst and the mixture of oil vapors and



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cracking gas transports the catalyst up throughthe riser. In the reactor bed (a), solids are collect-ed before passing through the stripper (b) to theregenerator (f). By virtue of the short contact time

of the order of a few seconds and the narrow gasresidence-time distribution, the high activity ofthe zeolite catalyst is optimally utilized and ahigher gasoline yield is achieved [2, 10].

Synthesis of Acrylonitrile. The crucial fac-

tor in the successful use of the fluidized-bedreactor for the synthesis of acrylonitrile bythe ammonoxidation of propene (Sohio process)(! Acrylonitrile) was reliable control of thisstrongly exothermic reaction:

C3H6NH33=2 O2 ! C3H3N3 H2O DHr 515 kJ=mol of acrylonitrile:

The reaction is carried out at a bed tempera-

ture of 400500C and gas contact time of 115 s [145] or 520 s [2]. Figure 32 is a schematicof the reactor. Air is fed to the bottom of thefluidized-bed vessel. The reactants ammonia andpropene are fed in through a separate distributor(b). Catalyst regeneration by carbon burnoffoccurs in the space between the air distributorand the feed-gas distributor. The heat of reactionis removed by bundles of vertical tubes (a) insidethe bed (horizontal tubes are used in other designs[146]).

Figure 30. Fluid catalytic cracking process (Kellogg-Ortho-flow system; according to [143, 144])a) Reactor; b) Regenerator

Figure 31. Riser cracking process (UOP system), [2]a) Reactor; b) Stripper; c) Riser; d) Slide valve; e) Air grid;f) Regenerator

Figure 32. Synthesis of acrylonitrile (Sohio process) [2]a) Cooler with internals; b) Distributor



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FischerTropsch Synthesis. The FischerTropsch synthesis of hydrocarbons is used on alarge scale for fuel production in the Republic ofSouth Africa [149]. Synthesis gas generated fromcoal in Lurgi fixed-bed gasifiers enters theSynthol reactor (Fig. 33), where it is reacted over

an iron catalyst at ca. 340C. The reactor workson the principle of the circulating fluidized bed.The mean porosity in the riser is 85 %, and the gasvelocity varies between 3 and 12 m/s [2]. Reac-tion heat is removed by way of heat-exchangertube bundles placed inside the riser.

However, experience has shown that thisreactor is costly, relatively expensive to operateand maintain, and scale-up to the size of thereactors in operation is probably close to themaximum achievable for operation at 350 Cand 2.5 MPa. Therefore, in the 1990s the 16circulating fluidized-bed reactors operating atSasols Secunda site were replaced by eightturbulent fluidized-bed reactors each of 10.7 mdiameter, which achieve a higher per-pass syngasconversion [150].

Different process routes have been developedfor the synthesis of maleic anhydride. The Mit-subishi process [152, 153] used the naphthacracker C4 fraction. The ALMA process usesn-butane as feedstock [154, 155]. A more recentdevelopment is the Du Pont process, which is

also based on n-butane but uses a circulatingfluidized bed as reactor (Fig. 34) [156]. It is basedon a vanadium phosphorus oxide (VPO) catalystwhich oxidizes n-butane to maleic anhydride by aredox mechanism on its surface layers [157]. Inthe riser n-butane is selectively oxidized by theoxidized catalyst. In the fluidized-bed regenera-tor the spent catalyst is reoxidized. Since 1996 acommercial plant has been operating in Asturias,Spain [158].

Other Processes. Other catalytic reactionscarried out in fluidized-bed reactors are the oxi-dation of naphthalene to phthalic anhydride (!Phthalic Acid and Derivatives) [2, 10, 151]; theammoxidation of isobutane to methacrylonitrile[2]; the reaction of acetylene with acetic acid tovinyl acetate [2]; the oxychlorination of ethyleneto 1,2-dichloroethane (! Chlorinated Hydrocar-bons) [2, 10, 159, 160]; the chlorination of meth-

Figure 34. The Du Pont maleic anhydride process [158].

Figure 33. FischerTropsch synthesis in the Synthol reactor[2, 147]a) Hopper; b) Standpipe; c) Riser; d) Cooler(coil); e) Reactor;f) Gooseneck



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ane [2]; the reaction of phenol with methanol tocresol and 2,6-xylenol [2, 161]; the reaction ofmethanol to gasoline [162, 163]; the synthesis ofphthalonitrile by ammoxidation ofo-xylene (!Phthalic Acid and Derivatives) [164]; the syn-thesis of aniline by gas-phase hydrogenation of

nitrobenzene (! Aniline, Section 3.2.) [165];and the low-pressure synthesis of melamine fromurea (! Melamine and Guanamines) [166].

An overview on the various fluidized-bedcatalytic processes has been given [167].

8.2. Polymerization of Olefins

The gas-phase polymerization ofethylenein thefluidized bed was developed by Union Carbide(Unipol process [168]; see Fig. 35) (! Polyole-fins). The reaction gas (ethylene and its como-nomers butene or hexene) fluidizes the bed at 75115C and 2030 bar. Extremely fine-grainedcatalyst is metered into the bed. Polymerizationoccurs on the catalyst surface and yields a gran-ular product with diameter ranging from 0.25 to1 mm. Ethylene conversion is comparativelylow, 2 % per pass; so the reaction gas is recycled.The heat of reaction is removed by cooling the

recirculating gas. The catalysts used have such ahigh activity that more than 10

5 parts by volumeof polymer can be produced per unit weight ofactive substance in the catalyst [2]. Because ofthe high degree of catalyst dilution in the granularpolymer, the catalyst need not be removed from

the product. In the process developed by BPChemicals, prepolymers with a diameter from0.2 to 0.25 mm rather than catalyst particles arefed into the fluidized bed [169].

Mitsui Petrochemical Industries has devel-oped a process for the gas-phase fluidized-bedpolymerization of propene (! Polyolefins); aplant using this process came on stream in1984 [170]. The UnipolShell process was joint-ly developed by Union Carbide and Shell andcommissioned in 1986.

Burdett et al. [171] have given a broad over-view on this still-developing technology, whichpresents many challenges for the engineer. Oneof the biggest problems is the stickiness of theparticles under the operating conditions of theprocess, which has often led to particle sinteringwith subsequent defluidization of the bed. Sevilleet al. [172] monitored the motion of particles in ascaled polymer reactor and studied the sinteringkinetics in order to determine a safe operating

window. Cai and Burdett [173] developed amodel of single-particle polymerization in thefluidized bed to simulate particle growth andparticle-temperature evolution with the resi-dence time of a catalyst particle in the reactor.

8.3. Homogeneous Gas-Phase

Reactions

The decisive advantage of the fluidized bed for

homogeneous gas-phase reactions is the ability tocarry large quantities of heat into or out of thereactor by using direct heat exchange via the bedsolids. An example is the Exxon fluid coking

process(Fig. 36; [2, 18, 174, 175]), which con-verts heavy residual oils to petroleum coke andgas-oil. The reactor (d) and heater (e) beds areconnected in a single solids loop. The bed mate-rial is coke generated in coking at 480570 C,which grows to spherical particles 0.11 mm indiameter in the reactor. The coke is discharged

continuously from the reactor and heated to 500690C by partial combustion in the heater. Thehot coke stream then transports the heat needed

Figure 35. Gas-phase polymerization of ethylene (Unipolprocess) [2]a) Compressor;b) Cooler; c) Catalyst feed hopper; d) Reactor;e) Separator



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for the endothermic coking reaction into thereactor. Excess coke is removed as a coarsefraction in a classifier connected to the heater.Fluid coking is used, e.g., for refining bitumenfrom the Athabasca tar sands in Canada. To makeefficient use of the product coke, Exxon com-bined the fluid coking process with a fluidized-bed gasification reactor [2, 175]. This Flexi-Coking process was first implemented in 1976in Japan; the daily capacity of one plant is ca.

3400 t of vacuum residue.The bed solids also find use as heat-trans-fer agents in the thermal cracking of naphtha,a process carried out in the Lurgi sand crack-er [2, 18, 176]. The solids circulating betweenthe reactor and the heater consist of coarsesand particles (ca. 1 mm in diameter). Whenthe coke deposit produced in cracking isburned off the particle surface with air, thesolids are heated to 800850 C and can thusdeliver the heat required for endothermic

cracking. The temperature in the reactor isca. 700750C.

Other thermal cracking processes include theBASF Wirbelfliess process [2, 18, 177], and theKuniiKunugi process [2].

8.4. GasSolid Reactions

Coal Combustion. The high heat capacityof the fluidized bed permits stable combustion at

low temperature (ca. 850C), so that the forma-tion of thermal and prompt nitrogen oxides [178]can be suppressed and total nitrogen oxide emis-

sions can be reduced. If limestone is added to thebed, the calcination reaction

CaCO3 ! CaOCO2

yields CaO, which can bind in situ the SO2

produced in combustion:

SO2CaO1=2 O2 ! CaSO4During the 1980s the fluidized bed was estab-lished in power-plant engineering. The unit sizerapidly increased from 5 MWein 1970 to about350 MWe during this time [179]. Meanwhile (ca.2006) some 500 power plants are in operationworldwide. By far the majority of these plantsoperate with circulating fluidized beds. As anexample, Figure 37 shows a Lurgi design.

The staged admission of the combustion airminimizes NO production from nitrogen in thefuel in the lower part of the combustion chamber.The admission of secondary air completes thecombustion in the upper part of the chamber byoxidizing most of the CO. Some of the circulatingsolids are led through the external fluidized-bedcooler, which enhances the flexibility of controland permits load variation over a wide range.

More recent developments aim at even larger

capacities with a further enlargement of thecombustion chamber and making use of super-critical steam conditions and once-through boilerdesign. One problem associated with the sizeenlargement is the distribution of both the coaland the secondary air fro

werther - fluidized bed reactors

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