Werther - Fluidized Bed Reactors
Short Description
Werther - Fluidized Bed Reactors...
Description
Fluidized-Bed Reactors JOACHIM W ERTHER, Hamburg University of Technology, Hamburg, Germany
1. 1.1. 1. 1. 1.2. 1. 2. 1.3.. 1.3 2. 2.1.. 2.1 2.2.. 2.2 2.3.. 2.3 2.4. 2. 4. 2.5. 2. 5. 2.6. 2. 6. 2.7. 2. 7. 2.8. 2. 8. 2.9. 2. 9. 2.9. 2. 9.1. 1. 2.9.2. 2.9. 2. 2.9.3.. 2.9.3 2.10.. 2.10 2.11.. 2.11 3. 3.1. 3. 1. 3.2. 3. 2. 3.3. 3. 3. 3.4.. 3.4 3.5.. 3.5
320 0 Introduction . . . . . . . . . . . . . . . . . . . . . . . . 32 320 0 Thee Fl Th Flui uidi diza zati tion on Pr Prin inci cipl plee . . . . . . . . . . . . . 32 321 1 Form Fo rmss of Fl Flui uidi dize zed d Be Beds ds . . . . . . . . . . . . . . 32 Advant Adv antag ages es and Dis Disad advan vanta tages ges of the . . . . . . . . . . . . . . . . 32 322 2 Fluidized-Bed Fluidized-Be d Reactor 322 2 Flui Fl uidd-Me Mech chan anic ical al Pr Prin inci cipl ples es . . . . . . . . . . . 32 322 2 Minimu Min imum m Flu Fluidi idiza zatio tion n Vel Veloci ocity ty . . . . . . . . . 32 324 4 Expans Exp ansion ion of Liq Liquid uid–So –Solid lid Flu Fluidi idized zed Bed Bedss . 32 Fluidi Flu idiza zatio tion n Pr Prop oper erti ties es of Ty Typi pica call Be Bed d Sol Solid idss 324 325 5 Stat St atee Di Diag agra ram m of Fl Flui uidi dize zed d Be Bed d . . . . . . . . . 32 326 6 Gass Di Ga Dist stri ribu buti tion on . . . . . . . . . . . . . . . . . . . . 32 327 7 Gass Je Ga Jets ts in Fl Flui uidi dize zed d Be Beds ds . . . . . . . . . . . . . 32 328 8 Bubb Bu bble le De Deve velo lopm pmen entt . . . . . . . . . . . . . . . . . 32 329 9 Elu El utr tria iattio ion n . . . . . . . . . . . . . . . . . . . . . . . . . 32 330 0 Circ Ci rcul ulat atin ing g Fl Flui uidi dize zed d Be Beds ds.. . . . . . . . . . . . . 33 Hydr Hy drod odyn ynam amic ic Pri Princ ncip iple less . . . . . . . . . . . . . . . 33 330 0 Locall Flow Struc Loca Structure ture in Circu Circulatin lating g Fluid Fluidized ized Beds Be ds . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 33 333 3 Design Des ign of Soli Solids ds Recy Recycle cle Sys System tem . . . . . . . . . 33 334 4 Cocurren Cocu rrentt Down Downflow flow Circ Circulati ulating ng Fluid Fluidize ized d 334 4 Beds (Downers) . . . . . . . . . . . . . . . . . . . . . 33 335 5 Attr At trit itio ion n of So Soli lids ds . . . . . . . . . . . . . . . . . . . 33 337 7 Soli So lids ds Mi Mixi xing ng in Fl Flui uidi dize zedd-Be Bed d Re Reac acto tors rs . . 33 338 8 Mech Me chan anis isms ms of So Soli lids ds Mi Mixi xing ng . . . . . . . . . . 33 . . . . . . . . . . . . . . 33 338 8 Vert Ve rtic ical al Mi Mixi xing ng of So Soli lids ds . 339 9 Hori Ho rizo zont ntal al Mi Mixi xing ng of So Soli lids ds . . . . . . . . . . . . 33 340 0 Solids Sol ids Re Resid sidenc ence-T e-Time ime Pro Prope perti rties es . . . . . . . 34 Solids Sol ids Mi Mixin xing g in Cir Circu culat lating ing Flu Fluidi idized zed Bed Bedss 340
!
Principles of Chemical Symbols (see also Reaction Engineering and Model Reactors and Their Design Equations) a:
A0: Ar : At: b:
!
volume volu me-s -spe peci cific fic ma mass ss-t -tra rans nsfe ferr ar area ea be be-tween twe en bub bubble ble and sus suspen pensio sion n pha phases ses,, 1 m cros cr osss-se sect ctio iona nall ar area ea of or orifi ifice ce,, m2 Arch Ar chim imed edes es nu numb mber er,, de defin fined ed by Equation (5) cros cr osss-se sect ctio iona nall ar area ea of re reac acto tor, r, m2 para pa rame mete terr de def. f. by Eq Equa uati tio on (5 (54 4)
2012 20 12 Wil Wiley ey-V -VCH CH Ve Verl rlag ag Gm GmbH bH & Co Co.. KG KGaA aA,, We Wein inhe heim im
DOI: 10.1002/14356007.b04_23 10.1002/14356007.b04_239.pub2 9.pub2
4. 4.1.. 4.1 4.2.. 4.2 5. 6. 7. 8. 8.1.. 8.1 8.2. 8.2. 8.3.. 8.3 8.4. 8. 4. 8.5.. 8.5 9. 9.1.. 9.1 9.2.. 9.2 9.2.1. 9.2.1. 9.2.2. 9.2 .2. 9.3.. 9.3 9.3.1. 9.3.1. 9.3.2. 9.3 .2. 10.
cv: cb: cc: c j: C b: C d: d o:
340 0 Gas Mi Gas Mixi xing ng in Fl Flui uidi dize zedd-Be Bed d Re Reac acto tors rs . . . 34 341 1 Gas Mi Mixin xing g in Bu Bubbl bblin ing g Flu Fluidi idized zed Bed Bedss . . . 34 341 1 Gas Mi Mixin xing g in Cir Circu culat latin ing g Flu Fluidi idized zed Bed Bedss . 34 Heat He at an and d Ma Mass ss Tr Tran ansf sfer er in Fl Flui uidi dize zedd-Be Bed d 341 1 Reactors . . . . . . . . . . . . . . . . . . . . . . . . . . . 34 343 3 GasGa s-So Soli lid d Sep epa arat atiion . . . . . . . . . . . . . . . . . 34 Inject Inj ection ion of Liq Liquid uid Rea Reacta ctant ntss int into o Flu Fluidi idize zed d 343 3 Beds . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 34 344 4 Indu In dust stri rial al Ap Appl plic icat atio ions ns . . . . . . . . . . . . . . . 34 Heter Het eroge ogeneo neous us Cat Cataly alytic tic Ga Gas-P s-Pha hase se . . . . . . . . . . . . . . . . . . . . . . . . . . 34 344 4 Reactions . Reactions 347 7 Poly Po lyme meri riza zati tion on of Ol Olefi efins ns.. . . . . . . . . . . . . . 34 347 7 Homoge Hom ogeneo neous us Gas Gas-Ph -Phase ase Rea Reacti ction onss . . . . . . 34 348 8 Gas– Ga s–So Soli lid d Re Reac acti tion onss . . . . . . . . . . . . . . . . . . 34 . . . . . . . . . . 35 352 2 Applic App licati ation onss in Bio Biotec techn hnolo ology gy . 354 4 Mode Mo deli ling ng of Fl Flui uidi dize zedd-Be Bed d Re Reac acto tors rs . . . . . 35 Model Mo deling ing of Liq Liquid uid–S –Sol olid id Flu Fluidi idize zed-B d-Bed ed 354 4 Reactors . . . . . . . . . . . . . . . . . . . . . . . . . . . 35 Model Mo deling ing of Gas Gas–S –Soli olid d Flu Fluidi idize zed-B d-Bed ed 354 4 Reactors . . . . . . . . . . . . . . . . . . . . . . . . . . . 35 Bubbli Bub bling ng Flu Fluidi idize zed-B d-Bed ed Rea Reacto ctors rs . . . . . . . . . 35 355 5 Circu Cir culat lating ing Flu Fluidi idize zed-B d-Bed ed Rea Reacto ctors rs . . . . . . . 35 356 6 New Dev Develo elopm pmen ents ts in Mo Mode delin ling g Flu Fluidi idize zedd357 7 Bed Reactors . Reactors . . . . . . . . . . . . . . . . . . . . . . . 35 Comput Com putati ationa onall Fluid Fluid Dyn Dynami amics cs . . . . . . . . . . 35 357 7 Modeli Mod eling ng of Flu Fluidi idize zed-B d-Bed ed Sys System temss . . . . . . . 35 358 8 359 9 Scale-up . . . . . . . . . . . . . . . . . . . . . . . . . . . 35 361 1 References . . . . . . . . . . . . . . . . . . . . . . . . . 36
solids soli ds vo volu lume me co conc ncen entr trat atio ion n bubb bu bble le at attr triti ition on ra rate te co cons nsta tant, nt, de defin fined ed by 2 4 Equation (50), s /m cycclo cy lone ne att ttri riti tion on ra rate te co con nst stan antt de defi fine ned d by 2 3 Equation (51), s /m jett at je attr trit itio ion n ra rate te co cons nsta tant nt,, de defin fined ed by 2 3 Equation (52), s /m conc co ncen entra tratio tion n in bu bubb bble le ph phas ase, e, km kmol ol/m /m3 conc co ncen entra tratio tion n in su susp spen ensi sion on ph phas ase, e, 3 kmol/m orifice diamete terr, m
320
d p: d pi: d t: d v: d v0: D: Dsh: Dsv: Fr p: Gs: h: ho: hgs: hwb: H : H mf mf : k G: L : ma: matt: mb: ms: np: p: Per, c: Q3: r a: _
_
r j:
Fluidized-Bed Reactors
Sauter Saut er di diam amet eter er,, de defin fined ed by Eq Equa ua-tion (6), m diam di amet eter er of pa part rtic icle le si size ze cl clas asss i , m bed diameter, m loca lo call bu bubb bble le vo volu lume me eq equi uiva vale lent nt sp sphe here re diameter, m init in itia iall bu bubb bble le di diam amet eter er,, m coef co effic ficie ient nt of mo mole lecu cula larr di diff ffus usio ion, n, m2 /s late la tera rall so soli lids ds di disp sper ersi sion on co coef effic ficie ient nt,, m2 /s vertic ver tical al sol solids ids dis disper persio sion n coe coeffic fficien ient, t, 2 m /s Frou Fr oude de nu numb mber er,, de defin fined ed by Equation (29) soli so lids ds ma mass ss flo flow w ra rate te,, ba base sed d on re reac acto torr 2 1 cross-sectional area, kg m s heig he ight ht ab abov ovee di dist stri ribu buto torr le leve vel, l, m heig he ight ht ab abov ovee di distr strib ibuto utorr wh where ere bu bubb bbles les are forming, m gasga s-to to-s -sol olid id he heat at tr tran ansfe sferr co coef effic ficie ient nt,, W 2 1 m K wallwa ll-to to-b -bed ed he heat at tr tran ansf sfer er co coef effic ficie ient nt,, W 2 m K expanded bed heigh ghtt, m bed be d hei heigh ghtt at at min minimu imum m flui fluidi diza zatio tion, n, m mass ma ss-t -tra rans nsfe ferr co coef effic ficie ient nt,, m/ m/ss jet length, m mass ma ss of el elut utri riat ated ed so soli lids ds,, kg mass ma ss flo flow w du duee to at attr trit itio ion, n, kg kg/s /s bed mass, kg solids mass flow ow,, g/s numb nu mber er of pa pass ssag ages es th thro roug ugh h cy cycl clon onee pressure, Pa Pecl eclet et num number ber,, defi defined ned by Equ Equatio ation n (43) cumu cu mula lati tive ve ma mass ss di dist stri ribu buti tion on attr at trit itio ion n ra rate te,, de defin fined ed by Eq Equa uati tion on (3 (33) 3),, 1 s reac re acti tion on ra rate te,, ba base sed d on ca cata taly lyst st ma mass ss,, 1 1 kmol kg s
1. Introd Introductio uction n 1.1. The Fluidizatio Fluidization n Principle Principle In fluidization an initially stationary bed of solid particles is brought to a ‘‘fluidized’’ state by an upwardstreamofgasorliquidassoonasthevolume flowrateofthefluidexceedsacertainlimitingvalue V mf (where mf denotes minimum fluidization). In _
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Re: S v:
Rey eyn nolds number volu vo lume me-s -spe peci cific fic su surf rfac acee ar area ea of pa part rtii1 cles, m time, s transpo tran sport rt dis diseng engagi aging ng hei height ght,, m supe su perfi rfici cial al flu fluid idiz izin ing g ve velo loci city ty,, m/ m/ss loca lo call bu bubb bble le ri rise se ve velo loci city ty,, m/ m/ss velo ve loci city ty at cy cycl clon onee in inle let, t, m/ m/ss supe su perfi rfici cial al mi mini nimu mum m flu fluid idiz izin ing g ve velo loci city ty,, m/s jett vel je eloc ocit ity y at or orifi ifice ce,, m/ m/ss slip sl ip ve velo loci city ty,, de defin fined ed by Eq Equa uati tion on (2 (27) 7),, m/s sing si ngle le pa part rtic icle le te term rmin inal al ve velo loci city ty,, m/ m/ss visi vi sibl blee bu bubb bble le flo flow, w, ba base sed d on be bed d ar area ea,, 3 2 1 m m s mini mi nimu mum m flu fluid idiz izin ing g flo flow w ra rate te,, m3 /s flow rate of gas issuing from orifi ificce, m3 /s mass fra racction of particl clee size fractio ion n i in bedmaterial velo ve loci city ty ra rati tio, o, de defin fined ed by Eq Equa uati tion on (1 (14) 4) pres pr essu sure re dr drop op of th thee gas gas di dist stri ribu buto tor, r, Pa Pa bed porosity local bubble gas holdup poro po rosi sity ty of ca cata taly lyst st par arti ticl clee bed be d po poro rosit sity y at mi mini nimu mum m flu fluid idiz izati ation on 2 1 elut el utri riat atio ion n ra rate te co cons nsta tant nt,, kg m s aver av erag agee li life fe ti time me of a bub ubb ble le,, s soli so lidd-to to-g -gas as ma mass ss flo flow w ra rati tio o 2 kinematic viscosit ity y, m /s stoi st oich chio iome metr tric ic nu numb mber er of sp spec ecie iess i in reaction j fluid density, kg/m3 solids density, kg/m3 stre st ress ss hi hist stor ory y pa para rame mete ter, r, de defin fined ed by Equation (54) para pa rame mete ter, r, de defin fined ed by Eq Equa uati tion on (2 (23) 3) pres pr essu sure re ra rati tio, o, de defin fined ed by Eq Equa uati tion on (2 (28) 8)
t : TDH : u: ub: uc: umf : uo: usl: ut : V b: _
V mf : V o: xi: _
_
: D pd: e: eb: ei: emf : * k : l : m: n: nij: a
rf : rs:
q: qb: y :
the fluidized bed, the particles are held suspended by the fluid stream; the pressure drop D pfb of the fluidonpassingthroughthefluidizedbedisequalto theweight theweig ht of th thee so soli lids ds min minus us thebuoy thebuoyan ancy cy,, di divi viddedbythecross-sectionalarea At ofthefluidized-bed vessel (Fig. 1):
¼ A H ð1e AÞ ðr r Þ g
D pfb
t
s
t
f
ð1Þ
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321
Figure 1. Pressure drop in flow through packed and fluidized beds
In Eq Equa uati tion on (1 (1), ), th thee po poro rosi sity ty e of th thee flu fluid idiz ized ed be bed d is the void void vol volume ume of the flui fluidize dized d bed (volum (volumee in interstices between grains, not including any pore volume in the interior of the particles) divided by the total bed volume; rs is the solids apparent density; and H is the height of the fluidized bed. In many respects, the fluidized bed behaves like a liquid. The bed can be stirred like a liquid; objects of greater specific gravity sink, whereas those of lower specific gravity float; if the vessel is tilted, the bed surface resumes a horizontal posit po sitio ion; n; if tw two o ad adja jace cent nt flu fluid idiz ized ed be beds ds wi with th diff di ffere erent nt be bed d he heig ight htss ar aree co conn nnec ecte ted d to ea each ch ot othe her, r, the heights become equal; and the fluidized bed flows out like a liquid through a lateral opening. Particularly advantageous features of the fluidized bed for use as a reactor are excellent gas– solid contact in the bed, good gas–particle heat and mass transfer, and high bed–wall and bed– internals heat-transfer coefficients. The fluidization principle was first used on an indus ind ustri trial al sca scale le in 192 1922 2 for the gas gasifi ificat cation ion of fine fine-grained coal [1]. Since then, fluidized beds have been applied in many industrially important processe ces ses. s. The pre prese sent nt spe spectr ctrum um of app appli licat catio ions ns extend ext endss fro from m a num numberof berof phy physic sical al pro proces cesses ses,, suc such h as cooling–h cooling–heating, eating, drying, sublimation–desublisublimation–desublimation, matio n, adso adsorpti rption–d on–desor esorptio ption, n, coat coating, ing,and and gran gran-ulation, ulat ion, to many heterogeneou heterogeneouss cata catalyti lyticc gasphase reactions as well as noncatalytic reactions. What follows is a survey of the fluid mechanicall pri ica princi nciple pless of flui fluidiz dizati ation on tec techno hnolog logy, y, gas and solid mixing, gas–solid contact in the fluidized bed, be d, ty typi pica call in indu dustr stria iall ap appli plica catio tions ns,, an and d ap ap-proach pro aches es to mod modelin eling g flui fluidiz dizeded-bed bed rea reacto ctors. rs. Further information is given in textbooks (e.g.,
[2]) and mon [2]) monogr ograph aphss (e. (e.g., g., [3– [3–8]). 8]). Sum Summary mary treatments can also be found in [9–19]. Other useful use ful lite literat rature ure inc includ ludes es rep report ortss of the Eng Engine ineerering Fou Founda ndatio tion n Con Confere ference ncess on Flu Fluidiz idizati ation on [20–22], the Circulating Fluidized Bed Conference en cess (e (e.g .g., ., [2 [23– 3–25 25], ], an and d – fo forr us usee of th thee flu fluid idiz ized ed bed in energy technology – the Fluidized Bed Combustion Conferences (e.g., [26–28]).
1.2. Forms of of Fluidized Fluidized Beds As th thee vo volu lume me flo flow w ra rate te V or or the sup superfi erficia ciall velocity u V / At of the fluid increases beyond the value V mf or umf (Fig. 2 A) corresponding to thee mi th mini nimu mum m flu fluid idiz izati ation on po poin int, t, on onee of two thingss happe thing happens: ns: in fluidization with a liquid , the bed begins to expand uniformly; in fluidization proc oces esss of gr grea eate terr in indu dustr stria iall with wi th a ga gass – a pr importance impor tance and the one discus discussed sed almost exclu exclu-sive si vely ly in th thee fo follo llowi wing ng ma mate teri rial al – vi virtu rtual ally ly solids-free solid s-free gas bubbles bubbles begin begin to form (Fig. 2 B). Thee loc Th local al mea mean n bub bubble ble siz sizee inc increa reases ses rap rapidly idly with increasing height above the grid because of coalescence of the bubbles. If the bed vessel is suffi su ffici cien entl tly y na narr rrow ow an and d hi high gh,, th thee bu bubb bble less ulti ul tima mate tely ly fil filll th thee en enti tire re cr cros osss se sect ctio ion n an and d pass through the bed as a series of gas slugs (Fig. 2 C). As the gas veloc velocity ity increases further, further, more and more solids are carried out of the bed, the original, sharply defined surface of the bed disappears, and the solids concentration comes to decrease continuously with increasing height. To ac achie hieve ve st stea eady dy-s -stat tatee op oper erat ation ion of su such ch a ‘‘tu ‘turbu rbulen lent’ t’’’ flui fluidiz dized ed bed (Fi (Fig. g. 2 D), sol solids ids entrained in the fluidizing gas must be collected _
¼ _
_
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Figure 2. Forms of gas–solids fluidized beds
and returned to the bed. The simplest way to do this is with a cyclone integrated into the bed vessel and a standpipe dipping into the bed. A further increase in gas velocity finally leads to the circulating fluidized bed (Fig. 2 E), which is characterized by a much lower average solids concen con centra tratio tion n tha than n the pre previo vious us sys system tems. s. The high hig h sol solids ids ent entrain rainmen mentt req requir uires es an effi efficie cient nt external solids recycle system with a specially design des igned ed pre pressu ssure re sea seall (sh (shown own as a sip siphon hon in Fig. 2 E).
1.3. Adva Advantag ntages es and Disa Disadvan dvantages tages of the Fluidized-Bed Reactor The maj major or adv advant antage agess of the (ga (gas–s s–solid olid)) flui fluidiz dized ed bed as a reaction system include 1. Easy handling handling and transport transport of solids solids due to liquid-like behavior of the fluidized bed 2. Uni Unifor form m tem temper peratu ature re dis distri tributi bution on due to intensive solids mixing (no hot spots even with strongly exothermic reactions) 3. Lar Large ge solid–gas solid–gas exchang exchangee are areaa by virtue of small solids grain size 4. Hi High gh he heat at-t -tra rans nsfe ferr co coef effic ficie ient ntss be betw twee een n bed an and d im imme mers rsed ed he heat atin ing g or co cool olin ing g surfaces 5. Unifo Uniform rm (solid) product product in batchwise batchwise process process because of intensive solids mixing
Set ag Set agai ains nstt th thes esee ad adva vant ntag ages es ar aree th thee fo foll llow owin ing g disadvantages: 1. Expensive Expensive solids separation separation or gas purificapurification tio n equ equipm ipment ent req requir uired ed bec becaus ausee of sol solids ids entrainment by fluidizing gas 2. As a consequence consequence of high solids solids mixing mixing rate, nonuniform residence time of solids, backmixing mix ing of gas gas,, and res result ulting ing lower lower conve conversio rsion n 3. In cata catalyt lytic ic rea reactio ctions, ns, und undesir esired ed byp bypass ass or broadening of residence-time distribution for reaction gas due to bubble development 4. Ero Erosio sion n of int intern ernals als and attrition attrition of sol solids ids (especially significant with catalysts), resulting from high solids velocities 5. Possib Possibility ility of defluidizati defluidization on due to agglomeragglomeration of solids 6. Gas Gas–so –solid lid cou counte ntercu rcurre rrent nt mot motion ion pos possib sible le only in multistage equipment 7. Diffic Difficulty ulty in scaling-up scaling-up Table 1 compares the fluidized-bed reactor with altern alt ernativ ativee gas gas–so –solid lid rea reacti ction on sys system tems: s: fixe fixeddbed, moving-bed, and entrained-flow reactors.
2. Fluid Fluid-Mecha -Mechanical nical Principle Principless 2.1. Mini Minimum mum Fluidizatio Fluidization n Velocity The minimu Themin imum m flui fluidiz dizatio ation n poi point, nt, whi which ch mar marks ks the boundarybetweenthefixed-andthefluidized-bed
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323
Table 1. Comparison of gas–solid reaction systems [2, 18]
conditions, can be determined by measuring the pressure drop D p across the bed as a function of volume flow rate V (Fig. 1). Measurement should always be performed with decreasing gas velocity, by starting in the fluidized condition. Only for very narrow particle-size distributions, however, does a sharply defined minimum fluidization point occur. The broad size distributions commonly encountered in practice exhibit a blurred range; conventionally, the minimum fluidization point is defined as the intersection of the extrapolated fixed-bed characteristic with the line of constant bed pressure drop typical of the fluidized bed (Fig. 1). The measurement technique already contains the possibility of calculating the minimum fluidization velocity umf : The pressure drop in flow _
through the polydisperse fixed bed at the point u umf , given, for example, by the Ergun relation [29] ( Fluid Mechanics), is set equal to the fluidized-bed pressure drop given by Equation (1). From the Ergun relation
¼
D p h
!
2
¼ 4:17 S ð1e eÞ 2 v
3
hu 0:29S v
þ
1ee r u 3
2
f
it follows
¼ 7:14 ð1e Þ n S
umf
mf
v
2v 3 ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi u ð Þ 75 64t þ ð Þ 1 0:067
e3mf
1 emf
rs
2
rf
rf n2
g
1
S 3v
1
ð2Þ
Accordingly, to calculate umf , the characteristics of the gas (rf , n), the density rs of the particles,
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Fluidized-Bed Reactors
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the porosity emf of the bed at minimum fluidization, and the volume-specific surface area S v of the solids must be known. The specific surface area defined by S v
area of all particles in the bed ¼ surface volume of all particles in the bed
(this takes into account only the external area, which governs hydraulic resistance, not the pore surface area as in porous catalysts) cannot be determined very exactly in practice. Hence umf should not be calculated on the basis of the measured particle-size distribution of a representative sample of the bed solids; instead, it is better measured directly. Equation (2) can be employed advantageously to calculate umf in an industrial-scale process on the basis of minimum fluidization velocities measured in the laboratory under ambient conditions [30]. An equation from WEN and YU [31] can be used for approximate calculations:
¼ 33:7
Remf
where
p ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ð þ Þ 1 3:6
105 Ar 1
¼ u nd
mf p
Remf
3 p 2
ð4Þ
r r ¼ gd n r
Ar
ð3Þ
s
f
f
ð5Þ
Here the surface mean or Sauter diameter calculated from the mass–density distribution q3 (d ) of the particle diameters
¼
d p
1 d max
R
d min
d 1 q3 d d d
ðÞðÞ
ð6Þ
u ut
n
¼e
ð7Þ
according to RICHARDSON andZAKI [33]. Here ut is the terminal velocity of isolated single particles; the exponent n is given as follows, provided the particle diameter is much smaller than that of the vessel:
n
8>< ¼> :
4:65 0:03 4:4 Re t 0:1 4:4 Re t 2:4
0 < Ret 0:2 0:2 < Ret 1 1 < Ret 500 500 < Ret
ð8Þ
The Reynolds number used above is calculated via the single-particle terminal velocity ut: Ret
¼ u n d t
p
ð9Þ
2.3. Fluidization Properties of Typical Bed Solids In fluidization with gases, solids display characteristic differences in behavior that can also affect the operating characteristics of fluidizedbed reactors. GELDART has proposed an empirically based classification of solids into four groups (A to D) by fluidization behavior [34]. The parameters employed are those crucial for fluidization properties: the mean particle diamerf ) ter (d p) and the density difference (rs between solid and fluid. Figure 3 shows the Geldart diagram with the interclass boundaries theoretically established by MOLERUS [35].
should be used for the characteristic particle diameter d p. Both the Ergun approach and the Wen and Yu simplification have been confirmed experimentally over a wide range of parameters. More recently, Vogt et al. [32] found that Equations (2) and (3) are also applicable to high-pressure fluidized beds in which the fluid is under supercritical conditions
2.2. Expansion of Liquid–Solid Fluidized Beds The uniform expansion of a bed on fluidization with a liquid can be described by
Figure 3. Geldart diagram (boundaries according to MOLERUS [35]) For explanation see text
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Fluidized-Bed Reactors
Solids of Group C are very fine-grained, cohesive powders (e.g., flour, fines from cyclones and electrostatic filters) that virtually cannot be fluidized without fluidization aids. The adhesion forces between particles are stronger than the forces that the fluid can exert on the particles. Gas flow through the bed forms channels extending from the grid to the top of the bed, and the pressure drop across the bed is lower than the value from Equation (1). Fluidization properties can be improved by the use of mechanical equipment (agitators, vibrators) or flowability additives, e.g., Aerosil. Solids of Group A have small particle diameters (ca. 0.1 mm) or low bulk densities; this class includes catalysts used e.g., in the fluidizedbed catalytic cracker. As the gas velocity u increases beyond the minimum fluidization point, the bed of such a solid first expands uniformly until bubble formation sets in at u umb > umf . The bubbles grow by coalescence but break up again after passing a certain size. At a considerable height above the gas distributor grid, a dynamic equilibrium is formed between bubble growth and breakup. If the gas flow is cut off abruptly, the gas storage capacity of the fluidized suspension causes the bed to collapse rather slowly.
¼
325
2.4. State Diagram of Fluidized Bed Whereas the onset of the fluidized state can be described by the minimum fluidization velocity, the bed operating range and the gas velocity needed to create a given fluidized state can be estimated with the help of the fluidized-bed state diagram (Fig. 4) devised by R EH [36]. This plot shows the fluid mechanical resistance characteristics of the fixed bed, fluidized bed, and pneumatic transport. The ordinate is the quantity 3 u2 4 g d p
r ðr r Þ f
s
f
and the abscissa is the Reynolds number Rep formed with the fluidization velocity u and the particle 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 an auxiliary grid with lines of constant M and constant Archimedes number. While the dimensionless groups plotted as ordinate and abscissa each contain both the particle diameter and the fluidization velocity, this is not the case with the parameters Ar and M defined by 3 p
¼ g n d ðr rr Þ
Ar
2
s
f
f
ð10Þ
Group B Solids have moderate particle sizes and densities. Typical representatives of this group are sands with mean particle diameters between ca. 0.06 and 0.5 mm. Bubble formation begins immediately above the minimum fluidization point. The bubbles grow by coalescence, and growth is not limited by bubble splitting. When the gas flow is cut off abruptly, the bed collapses quickly. Group D includes solids with large particle diameters or high bulk densities; examples are sands with average particle diameters > 0.5 mm. Bubbles begin to form just above the minimum fluidization point, but the character of bubble flow is markedly different from that in group B solids: group D solids are characterized by the formation of ‘‘slow’’ bubbles (Section 2.7). On sudden stoppage of the gas flow, the bed also collapses suddenly.
Figure 4. Reh status diagram with status points S and S 1–S4 (for explanation, see text)
326
Fluidized-Bed Reactors 3
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¼ gu n ðr rr Þ f
M
s
ð11Þ
f
The Reh status diagram can answer a number of practical questions. If, for example, the properties of the gas (rf , v) and the solid (dp, rs ) and the fluidization velocity u are given, the calculation of Ar and Re p yields, via the status point S in the diagram (Fig. 4), the average voidage e in the fluidized bed. Taking the line M const. through S at the intersection with the line e 1 a t S1 gives information on the particle size which is just elutriatedwhen a particles with a size distribution are fluidized, and the intersection of the same line with the fixed-bed limit e 0.4 (S2) indicates the particle size at which fluidization will break down if agglomeration occurs. The line Ar const. through S can be used to find the minimum fluidization velocity at S3 or – as a measure of the upper limit of fluidization – the maximum fluidizing velocity at S4. An important practical point is that the state diagram implies a classification scheme that
¼ !
¼
¼
relates various fluidized-bed systems to one another [37, 38] (Fig. 5). When a new fluidized-bed process is being designed, the position of the state point in the diagram will identify related fluidized-bed systems with potentially similar operating problems.
2.5. Gas Distribution The gas distribution device must satisfy the following requirements: 1. Ensure uniform fluidization over the entire cross section of the bed (especially important for shallow beds) 2. Provide complete fluidization of the bed without dead spots where, for example, deposits can form 3. Maintain a constant pressure drop over long operation periods (outlet holes must not become clogged)
Figure 5. Reh’s fluidized-bed state diagram with operating regions of different reaction systems a) Circulating fluidized bed; b) Fluidized-bed roaster; c) Bubbling fluidized bed; d) Shaft furnace; e) Moving bed
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Fluidized-Bed Reactors
Often, the gas distributor design must also prevent solids from raining through the grid both during operation and after the bed has been shut off. Porous plates of glass, ceramics, metal, or plastic are commonly used as gas distributors in laboratory apparatus; a variety of designs are used in pilot-plant and full-scale fluidized-bed reactors (see Fig. 6). Many more designs can be found, for example, in [2] and [39]. The principal requirement – uniform distribution of fluidizing gas over the bed cross section – can be met if the pressure drop D pd across the gas distribution grid is large enough. Suggested values for the ratio D pd / D pfb are 0.1–0.3 (with a minimum D pd of 3.5 kPa) [40], 0.2–0.4 [41], and > 0.3 [42]. For a given pressure drop D pd the gas velocity in the nozzle u o can be calculated from
¼ r2 C u o
D pd
D
2 o
where ro is the gas density in the orifice and C D is the drag coefficient. Applying the continuity equation :
V
¼ N A u o
o
o
327
erosion, and back-flow of solids. Erosion may occur at the distributor plate and at neighboring nozzles or walls due to gas jets as well as at the nozzle itself. Back-flow of solids into the windbox is caused by pressure fluctuations. In order to prevent this either the design pressure drop has to be larger than the pressure fluctuations or – if this is not feasible for economic reasons – a design must be chosen which tolerates short periods of gas flow reversal without permitting the solids to penetrate into the windbox. For the latter case the bubble cap design has turned out to be advantageous [43]. In the operation of fluidized-bed reactors, the quadratic response (D pd u2) of industrial gasdistributor designs must be kept in mind, because even if the fluidization velocity is lowered only slightly, an unacceptably low pressure drop across the gas distributor may occur. Industrial experience with different distributor designs, practical design rules, and a discussion of distributor-related problems, such as weepage into the windbox and erosion by grid jets and at grid nozzles, has been compiled in [44].
2.6. Gas Jets in Fluidized Beds
either the number of nozzles N o or the crosssectional area of the individual nozzle A o can be calculated for a given gas flow rate V . Problems related to the design of gas distributors are attrition of solids (see Section 2.11), _
Figure 6. Industrial gas distributors A) Perforated plate; B) Nozzle plate; C) Bubble-cap plate
Gas jets can form at the outlet openings of industrial gas distributors and also where gaseous reactants are admitted directly into the fluidized bed. A knowledge of the geometry of such jets, in
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particular the depth of penetration, is important for the implementation of chemical operations in fluidized-bed reactors, and not just from the standpoint of reaction engineering. It is also vital for reasons of design: the strongly erosive action of these jets means that internals, such as heatexchanger tubes, must not be located within their range. The literature contains many empirical correlations for estimating the mean depth of jet penetration L (e.g., [2–4]); these must, however, be used with care and, whenever possible, only within the range of parameter values for which they were derived. By way of example, M ERRY gives the following correlations for vertical gas jets [45]: L d o
¼ 5:2
0:2
" # 0:3
rf d o
1:3
rs d p
u2o
1
gd o
ð12Þ
and for horizontal jets [46]: L d o
"
ro u2o
¼ 5:25 ð1eÞ r g d s
0:2
ðÞ d p d o
p
#
0:4
rf
ðÞ rs
4:5
the motion of solids that they cause, however, also have some drawbacks: attrition of solid particles, erosion of internals, and increased solids entrainment by bubbles bursting at the bed surface. The existence of bubbles is particularly detrimental in the case of a heterogeneous catalytic gas-phase reaction, because the bypass of reactant gas in the bubble phase limits the conversion achieved in the fluidized bed. The ultimate cause of bubble formation is the universal tendency of gas–solid flows to segregate. Many studies on the theory of stability (e.g., [3, 4]) have shown that disturbances induced in an initially homogeneous gas–solid suspension do not decay but always lead to the formation of voids. The bubbles formed in this way exhibit a characteristic flow pattern whose basic properties can be calculated with the model of DAVIDSON and HARRISON [47]. Figure 7 shows the streamlines of the gas flow relative to a bubble rising in a fluidized bed at minimum fluidization conditions (e emf ). The characteristic parameter is the ratio a of the bubble’s upward velocity u b to the interstitial velocity of the gas in the suspension surrounding the bubble:
¼
0:2
ð13Þ
Here d o is the diameter of the outlet opening, uo is the outflow velocity, and r o is the density of the jet gas.
2.7. Bubble Development For many applications, especially physical operations and noncatalytic reactions, the state of a fluidized bed can adequately be described in terms of a single quantity averaged over the entire bed, such as the mean bed porosity e. In contrast, the design of fluidized-bed catalytic reactors requires that local fluid-flow conditions also be taken into account. The local fluid mechanics of gas–solid fluidized beds are determined by the existence of bubbles, which influence the performance of fluidized-bed equipment in several ways: the stirring action and convective solids transport by the rising bubbles are helpful; the resulting intensive solids motion produces a uniform temperature throughout the fluidized bed and rapid heat transfer between the bed and the heating or cooling tubes submerged in it. The bubbles and
a
¼u
ub
mf = emf
ð14Þ
The case a > 1 is typical for solids of Geldart groups A and B. The gas rising in the bubble flows downward again in a thin layer of suspension (‘‘cloud’’) surrounding the bubble. An important point for heterogeneous catalytic gas-phase reactions is that the presence of a boundary between bubble gas and suspension gas leads to the existence of two distinct phases (bubble phase and suspension phase) with drastically different gas–solid contact.
Figure 7. Gas flow for isolated rising bubbles in the Davidson model [47]
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If a < 1, some of the gas in the suspension phase undergoes short-circuit flow through the bubble, while only part of the bubble gas recirculates through the suspension. This type of flow is typical for fluidized beds of coarse particles (Geldart group D). Under the real operating conditions of a fluidized-bed reactor, a number of interacting bubbles occur in the interior of the fluidized bed. As a rule, the interaction leads to coalescence. As detailed studies have shown, this process is quite different from that between gas bubbles in liquids because of the absence of surface-tension effects in the fluidized bed [48, 49]. For predicting mean bubble sizes in freely bubbling fluidized beds, a differential equation for bubble growth should be used in the case of Geldart group A and B solids [50]: d d v dh
2eb 9p
1 3
¼
dv
ð15Þ
3l ub
with the following boundary condition at h
8>< ¼> :
d v0
m
1=3
0:008eb 1:3
:
o
0:2
industrial gas distributor
g
ð16Þ
where h o is the height above the grid where the bubbles form (for a porous plate, ho 0; for a perforated plate, h o L ; for a nozzle plate, h o is the height of the outlet opening above the plate; and for a bubble-cap plate, ho is the height of the lower edge of the cap above the plate). V0 is the volume flow rate of gas through the individual grid opening. The local volume fraction of bubble gas eb is given by
¼
_
¼ V:
eb
b
329
d 1 m; Geldart group A d 1 m; Geldart group B
ð20Þ
3:2 d t 0:33 0:05 3:2 d t 0:5 0:1
t
t
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 righthand side [51]). The crucial parameter here is the mean bubble lifetime l : l
280 ug
mf
ð21Þ
In practice, bubble growth is limited not only by the splitting mechanism based on the particlesize distribution of the bed solids, but also by internals (screens, tube bundles, and the like) that cause bubbles to break up. Computational techniques for estimating this process are given in [52, 53]. HILLIGARDT and WERTHER have derived a corresponding bubble-growth model for coarse-particle fluidized beds (Geldart group D) [50]. An example of a measured and calculated bubble-growth curve is presented in Figure 8.
porous plate 2 o
ð Þ V
¼ h :
¼
Fluidized-Bed Reactors
b =ub
2.8. Elutriation When bubbles burst at the surface of the fluidized bed, solid material carried along in their wake is ejected into the freeboard space above the bed. The solids are classified in the freeboard; particles whose settling velocity u t is greater than the gas velocity fall back into the bed, whereas particles with ut < u are elutriated by the gas
ð17Þ
and the visible bubble flow V b is _
:
0:8 ðuu Þ
Vb
ð18Þ
mf
The upward velocity u b of bubbles depends not only on the bubble size but also on the diameter d t of the fluidized bed: where :
p ffiffi ffi ffi
¼ V þ0:71s s
ub
b
_
b
_
gd v
ð19Þ
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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Fluidized-Bed Reactors
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Figure 9. Schematic drawing of fluidized bed and freeboard
stream. As a result, both the volume concentration of solids cv and the mass flow rate of entrained solids in the freeboard show a characteristic exponential decay (Fig. 9). With increasing height above the bed surface, the ‘‘transport disengaging height’’ (TDH ) is finally reached. Here the increased local gas velocities due to bubble eruptions have decayed, and the gas stream contains only particles with ut < u. When the TDH can be reached in a fluidized-bed reactor, this is associated with minimum entrained mass flow rates and solids concentrations, and hence with minimum loading on downstream dust collection equipment. Design of the dust collection system requires knowledge of the entrained mass flow rate Gs and the particle-size distribution of the entrained solids. For the design of the fluidized-bed reactor, the distribution cv (h) of the solids volume concentration and, for gas– solid reactions, the local particle-size distribution as 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 in Figure 10 [55]. The following relation is given for the TDH of Geldart group B solids as a function of the size d v of bubbles bursting at the bed surface [56]:
¼ 18:2 d
TDH
v
ð22Þ
Equation (25) was, however, derived for a bench-scale unit and may not scale to plant-size equipment. The mass flow rate G s of entrained solids per unit area leaving the fluidized-bed reactor is the
Figure 10. Estimation of transport disengaging height (TDH ), according to [55] umb Fluidization velocity at which bubble development begins
¼
sum of contributions from the entrainable particle size fractions (ut < u):
¼
Gs
X
xi k*i
i
ð23Þ
Here xi is the mass fraction of particle-size fraction i in the bed material and k*i is the elutriation rate constant for this fraction. The literature contains a number of empirical correlations for estimating k*i (e.g., [2–4]). More physical-based are the elutriation models of WEN and CHEN [57] and of K UNII and LEVENSPIEL [2, 58], which enable not only calculation of the exiting mass flow rate but also estimation of the concentration versus height cv (h) in the freeboard. The model by SMOLDERS and BAEYENS additionally takes the effect of variable freeboard geometry into account [59]. A literature survey on the factors affecting elutriation and the available modeling tools is given in [60].
2.9. Circulating Fluidized Beds 2.9.1. Hydrodynamic Principles
In REH’s state diagram of the fluidized bed [36], the circulating fluidized bed (CFB) is located above the single-particle suspension curve for
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Fluidized-Bed Reactors
¼
Fr p
u
ð26Þ
q ffiffi ffi ffi ffi ffi ffi ffi ffi ffi ffi ðr r Þ gd s
f
rf
331
p
The dimensionless pressure drop y is the ratio of the pressure drop D p along the flow path Dh to the maximum possible value for ascending flow (the value that would be attained if the pipe cross section were filled with solids corresponding to the concentration at the minimum fluidization point). The parameter of the family of curves is a volume flow rate ratio Figure 11. Fluidized-bed state diagram, according to [61]
Re < 102 and porosities e greater than about 0.8 (dashed line in Fig. 5). The shortcoming of this diagram is that it does not show an important parameter in the operation of a circulating fluidized bed: the circulating solids mass flow rate per unit area Gs. The diagram of Figure 11 [61] attempts to remedy this by plotting the mean slip velocity u sl between gas and solids
¼ ue ðG1=re Þ s
usl
s
ð24Þ
e, versus the mean solids concentration cv 1 with Gs as the parameter. The limiting conditions are high solids concentration (bed at minimum fluidization) and cv 0 with usl ut (isolated single particle). In the circulating fluidized-bed region, slip velocity increases with increasing Gs and can become much higher than the single-particle settling velocity (the physical justification for this statement comes from the formation of strands or clusters of particles). In the entrained-flow region the slip velocities again decrease with decreasing solids concentration. The fluidized-bed state diagrams discussed thus far, as well as others (e.g., [62, 63]), are suitable mainly for the qualitative interpretation of flow phenomena. A diagram proposed by WIRTH (e.g., [11, 64, 65]) also provides quantitative assistance in the design of circulating fluidized beds. The schematic in Figure 12 applies to a given gas–solid system described by a constant value of the Archimedes number Ar . The ordinate is the dimensionless pressure drop of the fluidized bed
¼
!
y
¼ ðr r Þ ðD1 pe Þ g Dh s
f
mf
the abscissa is the particle Froude number
¼
ð25Þ
rs
m rf 1 emf
ð27Þ
ð Þ
where m is the ratio of solid-to-gas mass flow rates. The limiting curve bounds the region of stable, vertically upward gas–solid flow on the low gas velocity side. Figure 13 shows how the state diagram of Figure 12 is constructed for a circulating fluidized bed with siphon recycle. If solids holdup in the recycle line and siphon is ignored, this case represents operation with a constant bed mass independent of velocity. At high gas velocities and if acceleration effects are neglected, the bed material is distributed uniformly over the total height H cfb of the fluidized bed (Fig. 13 C). The circulating fluidized bed then exhibits a single steady-state section with a constant pressure gradient (D p / Dh). This pressure gradient can be calculated from the bed mass as
Figure 12. State diagram for the circulating fluidized bed with siphon, according to W IRTH [64] Ar const.,parameter of family of curvesis the volume flow rate ratio m rf /(rs (1 e mf )); Fr p particle Froude number for superficial minimum fluidization velocity (pu mf ), singleparticle terminal velocity (pt), and transport velocity (pT), respectively
¼
¼
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Fluidized-Bed Reactors
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Figure 13. Pressure profile in the circulating fluidized bed with siphon, according to W A) Fr pumf < Fr p < Fr pt ; B) Fr pt < Fr pmax ; C) Fr p > Fr pmax
H ð1e Þ H ¼ ððrr rr ÞÞgg H ¼ ð1e Þ H
y hom
s
f
mf
mf
mf
s
f
cfb
mf
cfb
ð28Þ
where H mf is the bed height at minimum fluidization. The states identified by y hom to the right of the bounding curve in Figure 12 are accessible by increasing the gas velocity (corresponding to increasing Fr p). With increasing Fr p the volume flow ratio increases; that is, relatively more solids are elutriated (and thus circulated). If Fr p is allowed to drop below the limit Fr pmax (Fig. 13 B, Fig. 12) two steady-state sections appear in the riser tube: the one in the lower part is marked by a high pressure gradient, that in the upper part by a lower gradient. Figure 13 illustrates the physical significance of these two pressure gradients. In practice, the transition between the two linear regions takes place gradually. The height of the transition zone corresponds to the transport disengaging height (TDH). The picture changes further if the gas velocity declines to values lower than the settling velocity ut of a single isolated particle. In this case (for Fr p < Fr pt, Fig. 13 A, Fig. 12), no more solids can be elutriated, and the pressure gradient in the upper linear region vanishes. All the solid material is now in the form of a bubbling or turbulent fluidized bed. The solids concentrations averaged over the tube cross section (1 e) can be calculated from the dimensionless pressure drop:
IRTH [64]
Besides the pressure and solids concentration profile, the circulating mass flowrate of solids Gs At is important for the design of the circulating fluidized bed. In particular, the design of the solids collection and recycle system depends very much on this quantity. The mass flow rate of solids depends on the flowregime. At gas velocities such that twosteady-state sections arepresent in the bed vessel (i.e.,Fr pumf < Fr p < Fr pT), the mass flow rate of entrained solids depends on the physics of the gas–solid flow. Figure 14 plots the
1 e
¼ ð1e Þ y mf
ð29Þ
Figure 14. Elutriation diagram when the circulating fluidized bedcontains two steady-state sections, according to W IRTH [64]
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Fluidized-Bed Reactors
dimensionless solids mass flow rate versus Fr p, with the Archimedes number as parameter. For a given Ar , the flow rate tends to zero as Fr p Fr pt and reaches a maximum at Fr p Fr pT. The slope of the elutriation curve becomes greater with increasing Ar ; that is, the coarser the particles, the greater is the relative change in the circulating mass flow rate of solids with a change in gas velocity. At high gas velocities in the circulating fluidized bed (i.e., when a single steady-state section exists), the entrained mass flow rate depends on the particle Froude number and the solids holdup. More detailed information about the application ofWirth’stheoryinpracticemaybefoundin[11]. Whereas WIRTH’ s analysis of the circulating fluidized bed starts from the pneumatic transport condition, the models of RHODES and GELDART [66], as well as K UNII and LEVENSPIEL [2, 58], are based on the bubbling fluidized bed and describe the circulating fluidized bed as a limiting case of a bubbling bed with a very high rate of solids entrainment.
!
¼
2.9.2. Local Flow Structure in Circulating Fluidized Beds
The Wirth state diagram, as a first step toward the local characterization of flow regimes in a circulating fluidized bed, describes the vertical profile of the solids concentration. In the lower section of a circulating fluidized bed a dense region exits near the gas distributor. It has been observed that in this bottom zone bubble-like voids coexist with a surrounding dense suspension. The solids volume 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 the bottom zone to the upper dilute zone is characterized by violent gas–solid mixing. Many recent experimental studies with various measurement techniques (e.g., X-ray tomography [68], capacitance tomography [69] and fiber-optical probes [70]) have shown that the upper section of the circulating fluidized bed exhibits characteristic horizontal profiles, with the concentration cv, wall near the vessel wall always significantly higher than the value cv averaged over the vessel cross section; for example, c v, wall 2.3 cv [71]. Local measurements of the solids concentration and solids velocity show that upward-
333
¼
Figure 15. Schematic diagram of flow structure in a circulating fluidized bed
flowing regions of low solids concentration and downward-flowing aggregates of high solids concentration alternate in time at every point inside the fluidized bed, with downward-moving aggregates (strands, clusters) predominating near the wall and upward-moving regions of low suspension concentration predominating in the central zone. However, no significant downward flow of solids near the wall was observed in highdensity circulating fluidized beds, e.g., [72]. The picture 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 energyminimization multiscale (EMMS) model [73]. It considers the tendency of a fluid in a gas–solid two-phase flow to pass through the particulate layer with least resistance and the tendency of the solids to maintain least gravitational potential. Least resistance means that the volume-specific energy consumption for suspending and transporting solids is minimized, and minimization of the gravitational potential is equivalent to the requirement that the local mean voidage e attains a minimum. The model has been applied as a description of fluid-mechanical phenomena in CFB risers of different sizes [74, 75] but also for the prediction of flow patterns of gas and solids in industrial-scale units, such as a CFB boiler [76] and a petrochemical processing unit [77].Another promising line of development is the introduction of the EMMS concept into computational fluid dynamical calculations of multiphase flows; first results obtained with a drag model based on the EMMS model are encouraging [78].
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Fluidized-Bed Reactors
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the mass flow rate of the solids can be regulated by varying the gas supplied to the standpipe. Because the solids path does not contain any sortof mechanical closure, the characteristic pressure distribution plotted in Figure 17 is obtained. The distribution of solids between the fluidized bed and the recycle line is directly related to this pressure distribution. Operating properties differ from one recycle design to another [79].
Figure 16. Design options for solids recycle A) Siphon; B) L-valve
2.9.3. Design of Solids Recycle System
Solids carried over with the fluidized gas are generally collected in cyclones. In the case of bubbling beds, the solids can easily be returned to the 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 cyclones arranged beside and outside the bed, with special ‘‘valves’’ needed to connect the standpipe to the bed vessel. Figure 16 shows two design options, the siphon and the L-valve. With the siphon, the solids are fluidized (i.e., enabled to flow back into the reactor). In the L-valve design,
2.10. Cocurrent Downflow Circulating Fluidized Beds (Downers) A certain drawback of circulating and bubbling fluidized beds when applied for gas-phase reactions is the backmixing which inevitably occurs in the gas phase. In bubbling fluidized beds it is the bubble-induced solids circulation, and in circulating fluidized beds the downflow of solids in the wall zone, which entrains gas in the upstream direction and thus lowers the yield of a catalytic reaction or gives rise to undesired consecutive or side reactions. These disadvantages caused by the hydrodynamic effects of both gas and solids flowing against gravity could be overcome in the so-called downer reactor, in which the flow directions of both gas and solids are downward, i.e., in the same direction as gravity [80]. Another incentive is the possibility
Figure 17. Pressure distribution in solids recycle system of a circulating fluidized bed a) Fluidized bed; b) Return leg
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of realizing short contact times between gas and solids 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 investigated both experimentally and by numerical simulation [86]. It has been found that the hydrodynamics of the downer are also characterized by a wall zone of increased solids concentration. However, axial and radial gas-solids flow structures are much more uniform than in conventional fluidized beds. Another result is that the length of the flow development zone is much shorter for the downer than for the riser, which means that reactions with very short contact times can be carried out under near-plug-flow conditions. However, the solids feeding process and the geometry of the entrance region are critical points that deserve special attention [87]. The patent and open literature suggest various applications for downer reactors, e.g., residual oil cracking [88], coal pyrolysis [89], and biomass pyrolysis [90]. The catalytic pyrolysis of heavy feeds for the production of light olefins has been investigated on the laboratory scale with promising results [88]. However, no large-scale industrial process has emerged yet.
2.11. Attrition of Solids The attrition of solid particles is an unavoidable consequence of the intensive solids motion in the fluidized bed. The attrition problem is especially critical in processes where the bed material needs to remain unaltered for the longest possible time, as in fluidized-bed reactors for heterogeneous catalytic gas-phase reactions. Catalyst attrition is important in the economics of such processes and may even become the critical factor. Catalyst attrition in fluidized-bed reactors occurs normally as surface abrasion (Fig. 18) which means that surface asperities are abraded and edges of the catalyst particles are rounded off. Fragmentation may also play a role, especially for some fresh catalyst particles which on entering the reactor may simply be crushed into pieces. If in an industrial process extraordinarily high catalyst losses are observed it is advisable to examine catalyst samples under the scanning electron microscope. If the sample contains many fragments this could be an indication of
Fluidized-Bed Reactors
335
Figure 18. Attrition modes and their effects on the particle size distribution ( q3 mass density distribution of particle sizes d p)
¼
a wrong design (e.g., too high a velocity at the cyclone inlet or at the distributor). When designing catalytic fluidized-bed processes, the attrition performance of candidate catalysts should be tested under standardized conditions in the process development stage. This test can be performed in a small laboratory apparatus; it consists essentially of an extended fluidization test in which the mass of solids carried out of the bed is recorded as a function of time. Figure 19 presents a typical test result: during the first hours of testing, both the attrited material and the fine fraction of the bed material are elutriated. Only after a relatively long operating period is a quasi-steady state attained. The
Figure 19. Result of an attrition measurement
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Fluidized-Bed Reactors
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attrition rate r a in this steady state can be defined as
¼ m1 DDmt
a
r a
ð30Þ
b
where ma is the elutriated mass and mb the bed mass. Usually r a is expressed as percentage per day; for relatively attrition-resistant, fluidizedbed catalysts,it isof the order of0.2 % per day [9]. Many standard test apparatuses have been proposed for comparative attrition tests (e.g., [91, 92]), but all such equipment has been suitable only for comparative studies of different catalysts under consideration for the same process. The attrition measured in large-scale equipment can be far different from the values measured in a test apparatus. A number of sources can be identified for catalyst attrition in industrial fluidized-bed reactors:
Figure 20. Experimental measurement
apparatus
for
attrition
1. Jet attrition at gas distribution grid openings and nozzles where gaseous reactants are Attrition due to the bubble induced solids movement is given by [98] admitted to the bed 2. Bubble attrition in the bed due to solids 3 m att; bubble;i ¼ cb d pi DQ3i mb ðuumf Þ ð32Þ motion caused by bubbles where mb denotes the bed mass which contains 3. Attrition in cyclones 4. Attrition in pneumatic conveying lines, such bubbles (i.e., which is located outside the as those between reactor and regenerator beds jet-dominated grid region). Equation (32) also denotes the mass production of attrited fines Empirical correlations are available for the which is resulting from the size fraction d pi in attriting action of a gas jet in the fluidized bed the bed. The stress on the catalyst particles will be [93] and for the size reduction effect of solids different in contact with a gas jet, in the bulk of motion in the bed [94, 95]. WERTHER and coworkers [96] employ the the bubbling fluidized bed, and during its passage laboratory apparatus shown schematically in through a cyclone. Recent investigations of Figure 20 which enables separate study of the cyclone-induced catalyst attrition [99–101] have attrition due to jets from nozzles of various shown that the mass flow of attrited fines which is produced by attrition inside the cyclone when a diameters and that due to bubbles. Under steady-state conditions the jet-attrition- solids mass flow mc DQ3ci of particles of the size related mass production of fines per unit time for fraction d pi enters the cyclone is given by a gas distributor with a number no of orifices from u2c mother particles with diameter d p,i which are m c m Q d D ¼ ð33Þ att; c ;i c c 3ci pi p mc present in the catalyst inventory with a mass fraction D Q3i is proportional to the particle size where u is the gas velocity at the cyclone inlet, c d pi, the mass fraction D Q3i, the density r o of the and m the solids loading of the incoming gas flow c gas issuing from the orifice, the square of the orifice diameter d o, and to the cube of the jet exit m ¼ m c ð34Þ c rc uc Ac velocity uo [97, 98]: _
ffiffi ffi
m att; jet ;i
2 3 3i r0 d 0 u0
¼ c n d DQ j 0 pi
ð31Þ
where rc is the density of the inflowing gas, and Ac the cross-sectional area of the cyclone inlet.
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Figure 21. Dependence of attrition on time (bubble- and jetinduced attrition) and number of passages np through a cyclone.
337
Figure 22. Dimensionless attrition rate of FCC catalyst as a function of stress history.
A variety of approaches exist for reducing Equations (31)–(34) describe the catalyst attrition under conditions of steady state, i.e., attrition in industrial fluidized-bed reactors. The when the particles are more or less rounded off jet attrition action can be controlled with special (Fig. 18). To describe also the initial breakage gas distributor designs ([9]; e.g., by the use of and attrition of fresh catalyst particles, it is bubble caps, Fig. 6) such that gas jets do not issue necessary to follow the fate of the particles on directly into the bed at high velocity. Attrition their introduction into the reactor, which is due to bubbles can be lowered by limiting bubble possible with population balance models (cf. growth (avoiding high gas velocities and large Section 9.3.2). Klett et al. [102] and Hartge bed heights; use of fine catalysts with low umf , as et al. [103] have defined a stress history implied by Eqs. 18 and 24). Attrition in cyclones can be prevented, in the simplest case, by replaparameter cing the cyclones with devices such as filters. * Attrition can also be minimized by cutting back t =t j for jet--induced attrition load on the cyclone, for example, by placing ð35Þ the ¼ t =t b* for in--bed attrition the cyclones above the TDH. Relatively high np =n*p for attrition in cyclones catalyst attrition also occurs in circulating fluidwhere the definition of the characteristic ized beds where very large quantities of solids parameters t *, t *, and n * can be taken from must be collected in the cyclones.
8>< >:
j
b
p
Figure 21, np is the number of passages of a given particle through the cyclone, and t b and t j are the time periods during which the particle is sub jected to bubble and jet stress, respectively. If it is assumed that the effects of the different stress mechanisms on the catalyst particles are additive, a uniform treatment of the overall stress history for all three attrition mechanisms is given by m att
ð Þ¼
m att;¥
1:1 b 1
1:11=b > 1 :11=b
ð36Þ
The parameter b is characteristic of a given catalyst. Figure 22 shows measurements with FCC catalyst [103] which lead to b 1.16. Equation (36) allows the description of the stresshistory-dependent attrition rate and can be used for the simulation of fluidized bed reactors (see Section 9.3.2).
¼
3. Solids Mixing in Fluidized-Bed Reactors The intensive solids mixing typical of fluidizedbed reactors has several effects on performance. In catalytic reactions, the large-scale vertical solids mixing results in a transport of the gas components, adsorbed to the catalyst, so that the gas phase is backmixed and the conversion and selectivity are impaired. In noncatalytic gas– solid reactions, the mean solids residence time and residence-time distribution, as well as the propagation behavior of the solids from individual feed points, play a role. In general, fast and strongly exothermic reactions require fairly vigorous solids mixing to prevent temperature peaks near the reactant inlet.
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Figure 23. Solids mixing in bubbling fluidized beds due to particle drift (A) and wake transport (B) a) Cloud; b) Wake
3.1. Mechanisms of Solids Mixing The wake of the rising bubbles produces a rather slight upward and lateral drift of the particles (Fig. 23 A) [104]. In addition, solid particles are drawn upward in the wake, portions of the wake are shed at irregular intervals during bubble motion, and new portions of solids are taken into the wake (Fig. 23 B). Solids transport in the wake is essentially the reason that vertical solids mixing is from one to two orders of magnitude better than horizontal mixing. For reasons of continuity, the upward transport of particles by bubbles is coupled with a downward movement in the suspension phase that surrounds the bubbles. Measurements of the local bubble-gas flow have shown that the rising bubbles are not distributed evenly over the bed cross section. As a typical example, Figure 24 A gives a plot of the radial distribution of the bubble-gas flow at three heights above the grid in a fluidized bed 1 m in diameter. The profile is comparatively flat in the bottom zone but exhibits a steeper slope as the height increases, with an annular zone of preferentially rising bubbles. The resulting circulation of the solids also features an annular region of upward transport in the wakes with predominantly downward motion of the solids in the center and at the periphery of the bed (Fig. 24 B). The large-scale solids circulation can be reinforced by uneven distribution of the fluidized gas over the distributor cross section [106]. Figure 25 presents examples of industrial fluid-
Figure 24. A) Radial distribution of bubble-gas flow; B) Relationship between bubble distribution and solids circulation [105] 1 m, quartz sand, umf 0.013 m/s, u 0.2 m/s, d t H mf 0.5 m, V b visible bubble flow
¼
¼
_
¼
¼
¼
ized-bed furnaces in which forced circulation of the solids is employed to improve coal burnup.
3.2. Vertical Mixing of Solids The propagation behavior of the solids in a fluidized bed can be described by a number of models (e.g., [2, 109]). Most commonly used is the dispersion model, in which solids transport is described by a diffusion law. The numerical value of the dispersion coefficient D sv for solids mixing in the vertical direction increases with increasing gas velocity because of the growth in the number and size of bubbles. The following simple empirical correlation is given for fine particles (Geldart groups A and B) [2]:
¼ Dsv
m2 =s
0:06 0:1
þ
u
m=s
ð37Þ
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339
Figure 25. Fluidized-bed furnaces with forced circulation of solids A) According to [107]; B) According to [108]
For a plant-scale fluidized bed (0.9 1.26 m2 in plan, bed height 4 m) equipped with a bundle of horizontal tubes, a very similar relation was derived for a solid of Geldart group B [110]:
¼ Dsv
m2 =s
0:056
u umf
m=s
ð38Þ
Because solids circulation becomes more marked in larger-diameter fluidized beds, the dispersion coefficient increases rapidly with increasing bed diameter d t (Fig. 26). For this case the following expression is found [2]:
¼ Dsv
m2 =s
0:030
d t
coworkers model the horizontal propagation of coal in a fluidized-bed furnace, describing the carbon conversion in terms of a simple first-order reaction (rate constant k with dimension s1) [113]. The crucial parameter is the ratio k d2t / Dsh between the rate of the chemical reaction and the rate of dispersive mass transport. For high values of k (fast reaction), large reactor diameters d t,and low values of the dispersion coefficient D sh, the
0:65
m
ð39Þ
The above correlations can provide only rough values. Other effects observed in practice include, in particular, a significant effect of particle-size range [111, 112].
3.3. Horizontal Mixing of Solids In gas–solid reactions, the propagation behavior of the solids in the horizontal direction is important if, for example, the solid material is fed into the bed at isolated feed points. WERTHER and
Figure 26. Vertical solids dispersion in fluidized beds of fine particles (Geldart groups A and B) [2]
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Fluidized-Bed Reactors
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local carbon concentration in the bed exhibits a rather steep horizontal profile, resulting in a significantly nonuniform distribution of gas emissions over the bed cross section. On the basis of K UNII and LEVENSPIEL’ s model of bubble-induced solids mixing [114], an expression has been derived for calculating the horizontal dispersion coefficient Dsh averaged over the bed height H , given local bubble properties (bubble diameter d v, bubble-gas holdup e b) [115]: H
Dsh ¼ 0:67 103 þ0:023
1 H
Z 0
eb 1 eb
ffi ffi ffi ffi ffi ffi q
g d v3 dh
ð40Þ
This correlation holds for solids of Geldart groups B and D with Archimedes numbers between 8 600 and 58 000.
3.5. Solids Mixing in Circulating Fluidized Beds The circulating fluidized bed exhibits a complex gas–solid flow pattern as discussed in Section 2.9. Different regions can be discriminated with respect to the prevailing mechanisms of solids motion and mixing. An extensive survey on experimental findings in solids mixing is given in [116]. In the upper diluted zone of the circulating fluidized bed, clusters are formed with mainly upward flow in the core and predominantly downwards motion near the wall. While the wall region can be modeled by a plug-flow approach, the core region exhibits radial gradients. The Peclet number characterizing radial solids mixing in the core region
Per; s
*
¼ u D 2; R c
r s
3.4. Solids Residence-Time Properties Many applications of fluidization technology involve continuous processing of solids. Important considerations in such cases are not only the mean solids residence time but also the residence-time distribution. Whereas all elements have the same residence time in a plug-flow system, a stirred tank exhibits a broad distribution of residence times. To a good approximation, the residence-time properties of the fluidized bed with respect to the solids are the same as those of a stirred tank. The mean residence time t is the ratio of the solids mass m b in the reactor to the solids throughput m s: _
t
¼ mm:
b
ð41Þ
s
The mass fraction dms / mb of solids having a residence time between t and t dt is
þ
dms mb
¼ 1t e dt t
t
ð42Þ
Similarly, the fraction f of solids having a residence time less than t in the bed is calculated as
¼ 1e =t
f
t
ð43Þ
The residence-time distribution can be narrowed by placing a number of fluidized beds in series. Multistage systems of this type are used, for example, in fluidized-bed drying [18].
increases from 150 to 300 with increasing solids volume concentrations [117]. A recent investigation of solids mixing in the bottom zone with solid carbon dioxide as a tracer showed that in this zone solids are almost ideally mixed in the vertical direction but lateral mixing is limited with dispersion coefficients of about 0.1 m 2 /s which corresponds to Peclet numbers of around 40, [118]. Counteracting to solids mixing, segregation occurs in applications using particles of a broad size distribution and/or different densities. Easily fluidized particles tend to be elutriated while others tend to sink. A dynamic equilibrium between solids mixing and segregation is established, causing a spatial distribution of particles with significantly different solids properties, as was shown in an experimental study with a mixture 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 of the gas are particularly important for catalytic reactions but are also significant for gas–solid reactions when gaseous reactants are to be converted to the greatest possible extent in fluidized beds (e.g., reduction of fine-grained iron ores to sponge iron with gaseous reductants [18]). Gas
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mixing is closely linked to the motion and mixing of the solids in the bed.
4.1. Gas Mixing in Bubbling Fluidized Beds If the flow and mixing of gas in the bubbling fluidized bed are described by a simple one-phase dispersion model, the coefficients Dgv and Dgh of gas dispersion in the vertical andhorizontal directions have similar numerical values and follow trends similar to those of the solids dispersion coefficients. By way of example, Figure 27 shows the effect of fluidized-bed diameter d t on vertical gas dispersion. The increase in dispersion coefficient withvesseldiameter might be attributable to the formation of large-scale solids circulation patterns, which becomes more marked in larger equipment. As in the solids case, the coefficients of horizontal gas dispersion are a factor of 10–100 lower than those of vertical gas dispersion. A single-phase dispersion model gives only a rough description of gas mixing in bubbling fluidized beds. A more exact description comes from models that take account of local flow conditions in the bed, especially the presence of bubbles (see Chap. 9).
Fluidized-Bed Reactors
which are summarized in [120]. The bubbles in a bubbling fluidized bed influence the gas residence-time distribution and mixing directly through the bypass action of the bubble-gas flow and gas exchange between the bubbles and the surrounding suspension phase, and also indirectly through the solids motion that they induce. In the circulating fluidized bed, on the other hand, the gas-mixing properties are controlled by segregation due to the formation of solid aggregates (jets, clusters) and the rapid downward movement of solids strands predominantly near the wall. GRACE and coworkers, for example, show that a single-phase dispersion model cannot describe the tracer gas residence-time distributions that they measured [121]. They propose instead a two-phase model featuring exchange between a wall zone with stagnant gas and a core zone with plug flow. For the case of horizontal gas mixing, WERTHER and coworkers [122, 123] have shown that, for the bed solids they used (quartz sand, d p 0.13 mm, Geldart group B), horizontal gas mixing in the top part of the circulating fluidized bed in the core zone can be described by the model of turbulent single-phase flow [124]. The Peclet number
¼
Per; c
4.2. Gas Mixing in Circulating Fluidized Beds Only a few detailed studies of gas mixing in circulating fluidized beds have been published,
341
*
¼ u D 2; R c
r c
ð44Þ
(defined in terms of uc, the superficial velocity in the core zone; R*, the radius of the core zone; and Dr, c, the horizontal dispersion coefficient in the core zone) has a value of 465, which is in fairly good agreement with values measured in singlephase flows [125]. This value is independent of the solids circulation rate Gs. The circulating fluidized bed thus exhibits no especially intensive horizontal gas mixing, at least in the upper section where solids concentrations are relatively low.
5. Heat and Mass Transfer in Fluidized-Bed Reactors
Figure 27. Vertical gas dispersion in a fluidized bed of solids of Geldart group A (measurements by various workers; [2])
Fluidized-bed reactors exhibit a uniform temperature distribution even in case of highly exothermic or endothermic reactions. Approximations of the heat transfer rates are necessary for the design and control of fluidized-bed reactors in
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Fluidized-Bed Reactors
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order to determine the appropriate design of internals for cooling or heating and to estimate the changes in the performance with changing operating conditions. However, up to now there is no general theory on heat and mass transfer in fluidized beds. Numerous correlations for the calculation of heat and mass transfer coefficients are reported in the literature. Since these correlations are mainly based on experimental investigations performed under laboratory conditions, they may be different to the situation in large-scale reactors. Details on models of heat and mass transfer with their respective range of application are given in related surveys, e.g., [14–17, 126, 130].
magnitude larger than for gases alone [126]. For single phase flow a stagnant gas layer is established at the wall causing a hindered heat transfer. This layer is disrupted by solids transported at the wall. The solids adsorb heat and are mixed into the fluidized bed [9]. An example of the time-averaged local heat transfer along the circumference of a tube immersed horizontally in a fluidized bed is given in Figure 28. It exhibits lower values of the heattransfer coefficient below the tube due to a gas gap caused by bubbles and lower values on top of the tube because of solids being at rest. With intensified mixing this effect becomes less significant. The dependence of the heat-transfer coeffiFluid-to-Particle Heat and Mass Transfer. cient on the superficial gas velocity is illustrated Since the particle surface area is very large, fluid- in Figure 29. Fluidized beds of fine particles yield to-particle heat and mass transfer is rarely a limit- a larger heat-transfer coefficient than coarse ing factor in the design and operation of fluidized particles. According to MOLERUS and WIRTH bed reactors. The heat-transfer coefficients of [126], different transfer mechanisms can be idenfluidized-beds range between characteristic values tified. In case of fine particles, solids act as agents for flow through a fixed bed and flow around a transporting heat between walls and bed, whereas single particle [127]. gas convective transport is the mechanism dominating the heat transfer of coarse particles. The Fixed bed ( Rep > 80) heat-transfer coefficient of particles of intermehgs d p 0:5 0:33 diate sizes exhibits a maximum due to the super Nu ¼ ¼ 2þ1:8 Rep Pr l g position of these two transport mechanisms. Heat-transfer rates in circulating fluidized beds are lower than in bubbling fluidized beds due to Single particle Nu
¼ h l d ¼ 2þ0:6 Re : gs
p
g
05 p
Pr 0:33
where hgs is the gas–solid heat-transfer coefficient, d p the particle size, and l g the thermal conductivity of gas. The mass transfer coefficient can be determined applying the analogy of heat and mass transfer by replacing in the above formulas the Nusselt number Nu by the Sherwood number Sh and the Prandtl number Pr by the Schmidt number Sc. For particle Reynolds numbers below 100 and for fine particles, the transfer coefficients are significantly lower than estimated by the above formulas. If necessary, the effect of adsorption in mass transfer and of radiation in heat transfer needs to be taken into account additionally. Heat Transfer to Submerged Surfaces. Heat-transfer coefficients between fluidized bed and submerged surfaces are one or two orders of
Figure 28. Local heat-transfer coefficient around a 35 mm diameter tube immersed horizontally in a fluidized bed of 0.37 mm alumina particles operated at a superficial gas velocity of 0.8 m/s and a temperature of 500 C, adapted from [128]
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343
Figure 29. Heat-transfer coefficients determined with a tube immersed vertically in a fluidized bed of glass beads of different size operated at ambient conditions, adapted from [129, 126]
reduced solids volume concentrations and are dominated by clusters and strands [130]. The heat-transfer coefficient increases with increasing pressure [131] and temperature. The effect of radiation has to be considered for temperatures above 500 C, but opaque particles can form an effective radiation shield [132].
6. Gas-Solid Separation The fluidizing gas inevitably carries fine catalyst particles by entrainment to the reactor exit. Not only for environmental reasons (i.e., to minimize emissions) is it necessary to separate the solids from the gas. It may also be necessary to stop the main reaction and to avoid unwanted side or consecutive reactions or to protect following process steps or machines from particle-laden streams. In fluidized-bed technology cyclones are mostly used for this purpose. K NOWLTON [133] has given a survey on the state of the art of cyclone design and application in fluidizedbed reactors. The cyclone should not be considered as a separate apparatus following the fluidized bed but should be seen as an integral part of the fluidized-bed process. The reason is that, not only in circulating fluidized beds but also in bubbling or turbulent fluidized beds, the catalyst particles which are recovered in the cyclone are recycled to the fluidized bed. The collection efficiency of the cyclone is thus responsible for maintaining the particle size distribution in the bed inventory, which in turn determines the
fluidized-bed fluid mechanics and the chemical performance of the bed as a reactor. The interrelation between fluidized bed and cyclone is discussed in Section 9.3.2. The influence of cyclone performance on the overall process performance is increasingly considered. For example, PULUPULA et al. [134] investigated the role of cyclones in the regenerator system of a commercial FCC unit. ARNOLD et al. [135] were able to trace the deterioration of plant performance in theALMA maleicprocess back to problemswithcycloneefficiency.Achangeofthe cyclone design improved the particle size distribution of the bed inventory and consequently bed hydrodynamics and chemical conversion. S MIT et al. [136] report on cyclone performance in turbulent fluidized-bed Synthol reactors for Fischer–Tropsch synthesis. Carbon deposition on the catalyst particles influences the bed hydrodynamics, which in turn, via the elutriation mechanism, influence cyclone performance.
7. Injection of Liquid Reactants into Fluidized Beds The injection of reactants in liquid form into the bed was already an essential part of the first fluidized-bed catalytic process. In the FCC process (Section 8.1) crude oil is injected at the base of the reactor and evaporated in contact with the hot catalyst particles. The direct heat transfer is very efficient and avoids a separate evaporator for the feed. The cooling action of the evaporating reactant is a further advantage in the case of an
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exothermal reaction. Liquid-feed injection is therefore practiced not only in the FCC process but also, for example, in the syntheses of aniline ( Aniline, Section 3.2.), caprolactam ( Caprolactam, Section4.1.4.), andmelamine ( Melamine and Guanamines, Section 4.1.) and in BP Chemicals’ Inovene process [137] for the gasphase production of low-density polyethylene. Despite its industrial significance knowledge about the mechanisms of liquid mixing and evaporation in the fluidized bed is relatively scarce. Investigations with nonvaporizing horizontal gas–liquid spray jets have shown that with proper design of the injection nozzle it is possible to penetrate over several decimeters into the bed before the jet breaks up [138, 139]. On the other hand, it was found that under vaporizing conditions for atomizer nozzles with spray angles between 20 and 120 the injected liquid wetted the bed particles and subsequently evaporated from their surface while the particles were mixing in the bulk of the bed [140, 141]. This latter mechanism helps to transport the reactant away from the location of the nozzle and thus contributes to equalization of the feed distribution inside the reactor. The special case that a large oil droplet impinges on a smaller hot catalyst particle was recently investigated in a 3D direct numerical simulation to analyze droplet–particle collisions in the Leidenfrost regime [142]. The calculations were carried out for conditions prevailing near the feed nozzle in an FCC riser. Vapor layer pressure induced by evaporation and the droplet surface tension are the driving forces for droplet recoiling and rebounding. The contact time for a FCC particle and an oil droplet turned out to be about 140 m s.
!
! !
8. Industrial Applications In this chapter the industrial uses of fluidized-bed reactors are classified as follows: 1. 2. 3. 4. 5.
Heterogeneous catalytic gas-phase reactions Polymerization of olefins Homogeneous gas-phase reactions Gas–solid reactions Biotechnology applications
In each of these areas, the most important applications are listed and a few typical examples
are analyzed in more detail. For further descriptions of processes, the reader is referred to relevant articles in the A series. Complete descriptions of industrial uses of the fluidized-bed reactor can also be found in [2, 10, 18, 19].
8.1. Heterogeneous Catalytic Gas-Phase Reactions The fluidized-bed reactor offers the following principal advantages over the fixed-bed reactor for heterogeneous catalytic gas-phase reactions: 1. High temperature homogeneity, even with strongly exothermic reactions. 2. Easy solids handling, permitting continuous withdrawal of spent catalyst and addition of fresh if the catalyst rapidly loses its activity. 3. Ability to operate in the explosion range, provided the reactants are not mixed until they are inlet to the fluidized bed. This is because the high heat capacity of the bed solids, together with intensive solids mixing, prevents the propagation of explosions.
!
Catalytic Cracking. ( Oil Refining, Section 3.2.). The ease of solids handling was the basic reason for the success of catalytic cracking of long-chain hydrocarbons in the fluidized bed (Fig. 30). The cracking reaction is endothermic and involves the deposition of carbon on the catalyst surface, which quickly renders the catalyst inactive. Accordingly, the catalyst must be continuously discharged from the reactor and regenerated in an air-fluidized regenerator bed (b), where its carbon loading is lowered from 1–2 to 0.4–0.8 wt %. The combustion in this bed simultaneously furnishes the heat required for the cracking reactor; the catalyst acts as a heat carrier. The temperature in the regenerator is 570–590 C and in the reactor, 480–540 C [2]. In a stripper, steam is admitted to remove hydrocarbons adhering to the catalyst before it is forwarded to the regenerator. With the advent of high-activity zeolite catalysts in the 1960s, the bubbling fluidized bed, operated at gas velocities between 0.31 and 0.76 m/s [2], was replaced by the riser cracker (Fig. 31), in which the oil fed in at the bottom of the riser (c) is vaporized in contact with the hot catalyst and the mixture of oil vapors and
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345
of the order of a few seconds and the narrow gas residence-time distribution, the high activity of the zeolite catalyst is optimally utilized and a higher gasoline yield is achieved [2, 10]. Synthesis of Acrylonitrile. The crucial factor in the successful use of the fluidized-bed reactor for the synthesis of acrylonitrile by the ammonoxidation of propene (Sohio process) ( Acrylonitrile) was reliable control of this strongly exothermic reaction:
!
C3 H6 NH3 3=2 O2
þ
þ
! C H Nþ3 H O D 3
3
2
H r
¼ 515 kJ=mol of acrylonitrile :
Figure 30. Fluid catalytic cracking process (Kellogg-Orthoflow system; according to [143, 144]) a) Reactor; b) Regenerator
The reaction is carried out at a bed temperature of 400–500 C and gas contact time of 1– 15 s [145] or 5–20 s [2]. Figure 32 is a schematic of the reactor. Air is fed to the bottom of the fluidized-bed vessel. The reactants ammonia and propene are fed in through a separate distributor (b). Catalyst regeneration by carbon burnoff occurs in the space between the air distributor and the feed-gas distributor. The heat of reaction is removed by bundles of vertical tubes (a) inside the bed (horizontal tubes are used in other designs [146]).
cracking gas transports the catalyst up through the riser. In the reactor bed (a), solids are collected before passing through the stripper (b) to the regenerator (f). By virtue of the short contact time
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
346
Fluidized-Bed Reactors
Fischer–Tropsch Synthesis. The Fischer– Tropsch synthesis of hydrocarbons is used on a large scale for fuel production in the Republic of South Africa [149]. Synthesis gas generated from coal in Lurgi fixed-bed gasifiers enters the Synthol reactor (Fig. 33), where it is reacted over an iron catalyst at ca. 340 C. The reactor works on the principle of the circulating fluidized bed. The mean porosity in the riser is 85 %, and the gas velocity varies between 3 and 12 m/s [2]. Reaction heat is removed by way of heat-exchanger tube bundles placed inside the riser. However, experience has shown that this reactor is costly, relatively expensive to operate and maintain, and scale-up to the size of the reactors in operation is probably close to the maximum achievable for operation at 350 C and 2.5 MPa. Therefore, in the 1990s the 16 circulating fluidized-bed reactors operating at Sasol’s Secunda site were replaced by eight turbulent fluidized-bed reactors each of 10.7 m diameter, which achieve a higher per-pass syngas conversion [150].
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Different process routes have been developed for the synthesis of maleic anhydride. The Mitsubishi process [152, 153] used the naphtha cracker C4 fraction. The ALMA process uses n-butane as feedstock [154, 155]. A more recent development is the Du Pont process, which is also based on n-butane but uses a circulating fluidized bed as reactor (Fig. 34) [156]. It is based on a vanadium phosphorus oxide (VPO) catalyst which oxidizes n-butane to maleic anhydride by a redox mechanism on its surface layers [157]. In the riser n -butane is selectively oxidized by the oxidized catalyst. In the fluidized-bed regenerator the spent catalyst is reoxidized. Since 1996 a commercial plant has been operating in Asturias, Spain [158]. Other Processes. Other catalytic reactions carried out in fluidized-bed reactors are the oxidation of naphthalene to phthalic anhydride ( Phthalic Acid and Derivatives) [2, 10, 151]; the ammoxidation of isobutane to methacrylonitrile [2]; the reaction of acetylene with acetic acid to vinyl acetate [2]; the oxychlorination of ethylene to 1,2-dichloroethane ( Chlorinated Hydrocarbons) [2, 10, 159, 160]; the chlorination of meth-
!
!
Figure 33. Fischer–Tropsch synthesis in the Synthol reactor [2, 147] a) Hopper; b) Standpipe; c) Riser; d) Cooler(coil); e) Reactor; f) Gooseneck
Figure 34. The Du Pont maleic anhydride process [158].
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Fluidized-Bed Reactors
ane [2]; the reaction of phenol with methanol to cresol and 2,6-xylenol [2, 161]; the reaction of methanol to gasoline [162, 163]; the synthesis of phthalonitrile by ammoxidation of o -xylene ( Phthalic Acid and Derivatives) [164]; the synthesis of aniline by gas-phase hydrogenation of nitrobenzene ( Aniline, Section 3.2.) [165]; and the low-pressure synthesis of melamine from urea ( Melamine and Guanamines) [166]. An overview on the various fluidized-bed catalytic processes has been given [167].
!
!
!
347
recirculating gas. The catalysts used have such a high activity that more than 105 parts by volume of polymer can be produced per unit weight of active substance in the catalyst [2]. Because of the high degree of catalyst dilution in the granular polymer, the catalyst need not be removed from the product. In the process developed by BP Chemicals, prepolymers with a diameter from 0.2 to 0.25 mm rather than catalyst particles are fed into the fluidized bed [169]. Mitsui Petrochemical Industries has developed a process for the gas-phase fluidized-bed polymerization of propene ( Polyolefins); a plant using this process came on stream in 1984 [170]. The Unipol–Shell process was jointly developed by Union Carbide and Shell and commissioned in 1986. Burdett et al. [171] have given a broad overview on this still-developing technology, which presents many challenges for the engineer. One of the biggest problems is the stickiness of the particles under the operating conditions of the process, which has often led to particle sintering with subsequent defluidization of the bed. Seville et al. [172] monitored the motion of particles in a scaled polymer reactor and studied the sintering kinetics in order to determine a safe operating window. Cai and Burdett [173] developed a model of single-particle polymerization in the fluidized bed to simulate particle growth and particle-temperature evolution with the residence time of a catalyst particle in the reactor.
!
8.2. Polymerization of Olefins The gas-phase polymerization of ethylene in the fluidized bed was developed by Union Carbide (Unipol process [168]; see Fig. 35) ( Polyolefins). The reaction gas (ethylene and its comonomers butene or hexene) fluidizes the bed at 75– 115 C and 20–30 bar. Extremely fine-grained catalyst is metered into the bed. Polymerization occurs on the catalyst surface and yields a granular product with diameter ranging from 0.25 to 1 mm. Ethylene conversion is comparatively low, 2 % per pass; so the reaction gas is recycled. The heat of reaction is removed by cooling the
!
8.3. Homogeneous Gas-Phase Reactions
Figure 35. Gas-phase polymerization of ethylene (Unipol process) [2] a) Compressor;b) Cooler; c) Catalyst feed hopper; d) Reactor; e) Separator
The decisive advantage of the fluidized bed for homogeneous gas-phase reactions is the ability to carry large quantities of heat into or out of the reactor by using direct heat exchange via the bed solids. An example is the Exxon fluid coking process (Fig. 36; [2, 18, 174, 175]), which converts heavy residual oils to petroleum coke and gas-oil. The reactor (d) and heater (e) beds are connected in a single solids loop. The bed material is coke generated in coking at 480–570 C, which grows to spherical particles 0.1–1 mm in diameter in the reactor. The coke is discharged continuously from the reactor and heated to 500– 690 C by partial combustion in the heater. The hot coke stream then transports the heat needed
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sions can be reduced. If limestone is added to the bed, the calcination reaction CaCO3
! CaOþCO
2
yields CaO, which can bind in situ the SO2 produced in combustion: SO2 CaO 1=2 O2
þ
Figure 36. Fluid coking process [2, 18] a) Slurry recycle; b) Stripper; c) Scrubber; d) Reactor; e) Heater; f) Quench elutriator
for the endothermic coking reaction into the reactor. Excess coke is removed as a coarse fraction in a classifier connected to the heater. Fluid coking is used, e.g., for refining bitumen from the Athabasca tar sands in Canada. To make efficient use of the product coke, Exxon combined the fluid coking process with a fluidizedbed gasification reactor [2, 175]. This FlexiCoking process was first implemented in 1976 in Japan; the daily capacity of one plant is ca. 3400 t of vacuum residue. The bed solids also find use as heat-transfer agents in the thermal cracking of naphtha, a process carried out in the Lurgi sand cracker [2, 18, 176]. The solids circulating between the reactor and the heater consist of coarse sand particles (ca. 1 mm in diameter). When the coke deposit produced in cracking is burned off the particle surface with air, the solids are heated to 800–850 C and can thus deliver the heat required for endothermic cracking. The temperature in the reactor is ca. 700–750 C. Other thermal cracking processes include the BASF Wirbelfliess process [2, 18, 177], and the Kunii–Kunugi process [2].
8.4. Gas–Solid Reactions Coal Combustion. The high heat capacity of the fluidized bed permits stable combustion at low temperature (ca. 850 C), so that the formation of thermal and prompt nitrogen oxides [178] can be suppressed and total nitrogen oxide emis-
þ
! CaSO
4
During the 1980s the fluidized bed was established in power-plant engineering. The unit size rapidly increased from 5 MWe in 1970 to about 350 MWe during this time [179]. Meanwhile (ca. 2006) some 500 power plants are in operation worldwide. By far the majority of these plants operate with circulating fluidized beds. As an example, Figure 37 shows a Lurgi design. The staged admission of the combustion air minimizes NO production from nitrogen in the fuel in the lower part of the combustion chamber. The admission of secondary air completes the combustion in the upper part of the chamber by oxidizing most of the CO. Some of the circulating solids are led through the external fluidized-bed cooler, which enhances the flexibility of control and permits load variation over a wide range. More recent developments aim at even larger capacities with a further enlargement of the combustion chamber and making use of supercritical steam conditions and once-through boiler design. One problem associated with the size enlargement is the distribution of both the coal and the secondary air from the sidewalls over the cross section of the combustion chamber. Since the lateral mixing of gas and solids in a circulating fluidized bed is quite slow, sufficient numbers of feed ports for the coal and air injection nozzles have to be arranged on the sidewalls. The ‘‘pantsleg’’ design shown in Figure 38 is one possibility to provide sufficient lateral mixing at the bottom of the combustion chamber. A first 450 MWe unit is being built in Lagisza/Poland [181] and 600 MWe CFB combustors are in the design phase [182]. If a fluidized-bed furnace running under a pressure of 12–16 bar is linked to a gas turbine, the efficiency of the power plant can be markedly enhanced [183]. At the same time, however, this concept imposes severe requirements on gas cleaning [184]. Pressurized fluidized-bed combustion has been tested in several large experimental plants (e.g., [186]). In the meantime
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Figure 37. Power plant with circulating fluidized-bed furnace (Lurgi process) [180] a) Circulating fluidized-bed reactor; b) Recycling cyclone; c) Siphon; d) Fluidized-bed heat exchanger; e) Convective pass; f) Dust filter; g) Turbine; h) Stack
several plants of the 80 MW range are in commercial operation. Recently, a pressurized fluidized-bed combustor with an electrical power of 360 MW was erected by ABB Carbon [185] (Fig. 39); these units employ bubbling fluidized beds [187, 188]. Pressurized fluidized-bed boi-
lers employing circulating fluidized beds are still under development (e.g., [189]). For further details on fluidized-bed combustion systems see the proceedings of the Fluidized Bed Combustion Conferences [25] and a monograph [190].
Figure 38. Furnace cross section of a large CFB combustor (after [179])
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Figure 39. Power plant with pressurized fluidized-bed combustor (ABB design) [187] a) Pressurized fluidized-bed boiler; b) Cyclone; c) Gasturbine; d) Economizer; e) Ashremoval; f) Fuel feed; g) Feed-water tank; h) Steam turbine; i) Condenser; j) Bed material hopper
Waste Incineration. The incineration of municipal sewage sludge in fluidized-bed furnaces is now practiced in many facilities [191]. The total amount of sludge thermally treated in Germany in 1996 was 513 000 t (dry basis) [192]. Figure 40 is a diagram of an incinerator. Moist, centrifuged sludge is fed from above to the fluidized bed by means of piston pumps. The fluidizing air is preheated in an oil-fired muffle before reaching the furnace. Fuel oil, used as a supplemental fuel, is metered directly into the bed. Developments in sludge incineration have achieved energy autarky by recovering waste heat and utilizing it to predry the sludge so that self-sustaining combustion is possible [194–196]. Later, the more stringent emission limits set forth in the 17th Bundesimmissionschutzverordnung (BImSchV, regulation in the Federal Republic of Germany concerning the limitation of immissions) may necessitate staged combustion (as in power generation), particularly to control NO x emissions [197, 198]. For the incineration of municipal waste, a furnace with a ‘‘rotating’’ fluidized bed has been developed. The inclined distribution grid in this design generates two rollerlike flows of circulat-
Figure 40. Fluidized-bed furnace for municipal sludge incineration (Uhde system) [193]
ing bed solids, leading to rapid and efficient mixing of the waste in the bed [199]. Coal Gasification. A number of fluidizedbed processes have been developed for gasifying
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Figure 41. Concept for cogeneration power plant based on high-pressure gasification in circulating fluidized bed [185] a) High-pressure CFB gasification; b) Gas cleaning; c) CFB combustion; d) Waste-heat boiler; e) Gas turbine; f) Steam turbine
!
coal ( Gas Production; e.g., [2]). Interest in these processes for cogeneration power plants has recently become more intense. In the cogeneration system shown in Figure 41, high-pressure gasification is combined with combustion in a circulating fluidized bed; efficiencies of more than 40 % are expected, depending on the available gas turbine technology [180]. The Rheinische Braunkohlenwerke company has developed a high-temperature Winkler (HTW) process based on Winkler gasification (Fig. 42) [200]. The pressure (ca. 10 bar) and temperature (ca. 1100 C) are higher than in the Winkler process; coal is gasified with oxygen and steam. Recycling of solids from the cyclone to the fluidized bed results in a much higher carbon conversion than in the Winkler process. Gasification of Solid Waste. In comparison with incineration, the gasification of solid waste offers the advantage of a smaller volume of offgas, so the cleaning system can be made smaller. In the Japanese Pyrox process, heat required by the gasification reactor is supplied by sand heated in a fluidized-bed furnace, so that a high-Btu-gas can be generated [2]. A cement kiln plant at Ruedersdorf in Germany is operated with a biomass gasification reactor. This circulating fluidized-bed reactor designed by Lurgi supplies fuel gas for the calciners [201].
Calcination. The calcination of aluminum hydroxide in the Vereinigte Aluminiumwerke (VAW)/Lurgi circulating fluidized-bed process features an overall thermal efficiency of more than 70 %, which is achieved through downstream heat recovery from the calcined alumina and the off-gas [18, 202–204]. The circulating fluidized bed proper (c) is coupled to two venturi fluidized beds (a), in which the moist hydroxide is first dried and heated by direct contact with the off-gas before it is forwarded to the calcination furnace (see Fig. 43). The five-stage fluidizedbed cooler (d) downstream of the furnace serves to preheat the combustion air. A furnace 3.8 m in diameter and 20 m tall, with a fluidization velocity of 3 m/s and mean particle diameter of 0.04– 0.05 mm, produces more than 500 t/d of Al 2O3 [18]. Other applications include the calcination of limestone (in multistage fluidized-bed furnaces), lime muds, and crude phosphates [2, 18]. Roasting Processes. Fluidized-bed roasting follows the general reaction equation
þ !Metal oxideþSulfur dioxide
Metal sulfide Atmospheric oxygen
This is one of the earliest industrial uses of fluidization. Many such processes are used in the roasting of pyrite, zinc blende, and other sulfide
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Figure 42. Flow sheet of HTW demonstration plant [200] a) Coal lock hopper system; b) Gasifier; c) Cyclone; d) Ash lock hopper system; e) Raw gas cooler; f) Wet dust separator; g) Carbon monoxide conversion
ores. Bubbling fluidized beds with gas velocities between 0.5 and 2.3 m/s [2] are employed; heat generated by the exothermic roasting reaction is removed by tube banks immersed in the bed, via a solid heat-transfer agent, or by simple water injection. Roasting furnaces are available in very large sizes (bed diameters up to 11 m) with capacities of several hundred tonnes of ore per day [2, 18].
!
Iron Ore Direct Reduction. ( Iron). For the direct reduction of iron ore, Lurgi has developed two processes [205]. Hydrogen is applied as reductant in the Circored process, coal gas is used in the Circofer process. A plant for the Circored process has been built in Trinidad with a capacity of 500 000 t iron briquette per year [206]. Applying two stages, a circulating fluidized-bed reactor reduces the preheated iron ore (800 C)at 630–650 C to a degree of metallization of 65– 85 % and a bubbling fluidized-bed reactor proceeds at temperatures up to 680 C to achieve a degree of metallization of 93–95 % (see Fig. 44). The metallized product is then transported to the hot briquetting unit. At a pressure of 4 bar, the process gas is recycled to a gas cleaning unit and made up with hydrogen.
Other Processes. Fluidized-bed processes for the production of high-purity silicon and activated carbon and the chlorination and fluorination of metal oxides are described in [2]. A detailed description of TiO2 synthesis in a fluidized-bed reactor and a survey of the use of fluidized-bed processes in the production of nuclear fuels are given in [10].
8.5. Applications in Biotechnology A comprehensive survey of the use of fluidizedbed reactors in biotechnology is given in [10, 207]. Liquid–Solid and Liquid–Gas–Solid Systems. are used in aerobic and anaerobic wastewater treatment (nitrification and denitrification); the microorganisms are grown as a biofilm on particulate supports to prevent their entrainment from the reactor with the fluidizing medium. The advantages of the fluidized-bed reactor over the fixed-bed reactor include higher capacity per unit volume and less susceptibility to plugging [208]. A study showing the potential of liquid–solid circulating fluidized beds in
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Figure 43. Fluidized-bed calcination of aluminum hydroxide (VAW/Lurgi system) [18] a) Venturi fluidized bed; b) Cyclones; c) Fluidized-bed furnace; d) Fluidized-bed cooler; e) Recycle cylone; f) Electrostatic precipitator
biotechnological processes such as fermentation has been published recently [209]. Full-scale fluidized-bed biogas production reactors have come on stream since 1984 [210, 211]. The process consists of two stages, acidification and methanation. Sand particles 0.1–0.3 mm in diameter serve as support for the microorganisms; at fluidization velocities of 8– 20 m/h, the biofilm grows to a thickness of 0.06– 0.2 mm on these particles. The reactors are large devices with diameters of 4.6 m and bed heights of 21 m. Gas–Solid Systems. Gas–solid fluidizedbed fermenters have been investigated on a pilot scale for the growth of Saccharomyces cerevisiae [212], the production of ethanol with S . cerevisiae [213, 214], and the enrichment of glutathione in yeast by S. cerevisiae [214, 215]. In these
applications, the substrate is metered into the bed in liquid form. A process used in Japan for the culture of Aspergillus sojae on wheat groats employs a solid substrate [216, 217]. The latter process is in service on a plant scale (bed mass 500 kg, bed diameter 1.5 m). The reactor (Fig. 45) contains an agitator (c) just above the distributor (d), as well as a rotating separator (a) in the top of the vessel. Water is sprayed onto the bed from above to maintain the proper moisture level; electrodes (b) dipping into the bed measure this parameter. The moisture content of the solids is generally a critical parameter for the fluidizedbed fermenter; the bioreactions extinguish if it becomes too low, whereas the particles agglomerate and fluidization is disrupted if it becomes too high. Sterilized air is used for fluidization. Seed spores of the microorganisms are fed into the bed via the ejector (e). This system achieves a
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Figure 44. Flow sheet of Lurgi Circored process [206] a) Preheater; b) Cyclone; c) First stage reactor; d) Second stage reactor; e) Briquetting unit; f) Gas cleaning unit
considerable gain in cell yield and an enrichment of certain enzymes by a factor of 5–15 over conventional fixed-bed cultures. The generated biomass forms the basis for soy sauce production.
9. Modeling of Fluidized-Bed Reactors 9.1. Modeling of Liquid–Solid Fluidized-Bed Reactors An expansion formula of the Richardson–Zaki type, Equation (7), describes the hydrodynamics of liquid–solid fluidized beds fairly well [218]. The difficulty in modeling this kind of reactor for bioreactions thus lies not so much in determining the flow and mixing conditions in the fluid as in describing the diffusion processes in the biofilm and the kinetics of the biological reactions [219]. In view of the small number of experimental studies reported thus far, no final judgment can bemadeonthesuitabilityofvariousmodels[208].
9.2. Modeling of Gas–Solid Fluidized-Bed Reactors Figure 45. Solid-state fermentation of Aspergillus sojae in the fluidized bed (adapted from [216]) a) Separator; b) Electrode; c) Agitator; d) Distributor; e) Ejector
Exhaustive literature surveys can be found in [2, 9, 10, 220]. [221]. Many models exist in the literature, which are classified in the cited
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references under various schemes. The available information can be summed up as follows: No generally accepted model of the fluidizedbed reactor exists; instead, many models have been proposed on the basis of more-or-less extensive experimental findings for various applications. Any fluidized-bed reactor model can be broken down into separate components that describe, with varying degrees of accuracy, the hydrodynamics (depending on solid properties, operating conditions, and geometry), gas–solid contact, and reaction kinetics. The essential point is that the reactor geometry effect, which is important for scale-up (Chap. 10), manifests itself in the flow conditions and must therefore be included in the hydrodynamic part of the model. Before a reactor model found in the literature can be applied to a given problem, the designer must determine whether numerical values are available for all model parameters, that is, whether the model is appropriate for design calculations or is a ‘‘learning model’’ [222] in which the numerical values of important parameters can be determined only after the model is adapted to actual test results. Reaction kinetics may be determined in a fixed-bed reactor, provided measurements are performed under conditions comparable to those that prevail in the fluidized-bed reactor (e.g., the same solids composition and particle-size distribution, the same activity state in the case of catalysts) [223]. However, the kinetic parameters can also be determined directly by measurements in a bench-scale fluidized-bed apparatus [224].
9.2.1. Bubbling Fluidized-Bed Reactors
By far the majority of fluidized-bed reactor models described in the literature deal with reactions in bubbling fluidized beds [2, 9, 10, 225, 226]. For a specific application, modeling depends on the bubble flow regime. For slow-bubble systems (Fig. 7, left), the short-circuit flow of gas through the bubbles must be taken into account [227]. For fast-bubble systems (Fig. 7, right), the species have to be balanced separately in the bubble and suspension phases. If models from the literature are employed, it should be taken into account that those devised in the past, when adequate computing hardware was
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not available, often sought to obtain analytical expressions for the degree of conversion of a single reaction (usually taken as first-order). The simplifying assumption of a single ‘‘effective’’ bubble size for the entire fluidized bed was therefore made [2], or the mass-transfer area between the bubble and suspension phases was taken as uniformly distributed over the height of the bed (HTU or NTU concept, where HTU denotes height of transfer unit and NTU denotes number of transfer units [228]. Today, in view of the computing power available at the PC level, the recommended procedure is to start from local mass-transfer relations, write balance equations for the differential volume element of the reactor, and then numerically integrate these equations. Figure 46 presents a model used by W ERTHER for a constant-volume reaction [224, 229]. Here the simplifying assumption is that flow through the suspension phase is at the minimum fluidization velocity umf . For a heterogeneous catalytic gas-phase reaction, the material balances for species i in the unsteady-state cases are as follows: Bubble phase
¼ ½uu ð1e Þ ddC h k ; aðC C Þ eb
dC bi dt Gi
bi
mf
bi
b
ð45Þ
di
Suspension phase
ð1e Þ½e þð1e Þe ddC t ¼u ð1e Þ ddC h þk ; aðC C Þ þð1e Þð1e Þr di
b
mf
mf
i
di
mf
b
Gi
bi
b
di
mf
s
M
X
ni; j r j
¼
j 1
ð46Þ
In Equations (45) and (46) the following simplifying assumptions have been made: 1. Plug flow through the suspension phase at an interstitial velocity (umf / emf ) 2. Bubble phase in plug flow, bubbles are solids free 3. Reaction in suspension phase only 4. Constant-volume reaction (see [224] for handling a change in number of moles) 5. Sorption effects neglected (see [229] for handling sorption)
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Figure 46. Two-phase model of the fluidized-bed reactor
Here ei is the porosity of the catalyst particles; of the few exceptions; here the model must take a is the local mass-transfer area per unit of account of the propagation of coal from the feed fluidized-bed volume, which can be calculated as point if the furnace emission behavior is to be described correctly [232, 234]. a
¼ 6d e
b
ð47Þ
v
for spherical bubbles; r j is the rate of partial reaction j per unit mass of catalyst; and nij is the stoichiometric number of species i in reaction j . The relation k G;i
¼
umf
3
r ffi ffi ffi ffi ffi ffi ffi ffi ffi ffi þ 4 Di emf ub pd v
ð48Þ
proposed by SIT and GRACE [230] has proved useful for describing the mass-transfer coefficient k Gi associated with component i in mass transfer between the bubble and suspension phases; D i is the molecular diffusion coefficient of species i . The freeboard space above the bubbling fluidized bed must be considered in the reactor model if the entrainment rate is high and the reactions in the freeboard are not quenched, for example, by cooling. Most fluidized-bed models include concentration profiles only for the vertical direction. This one-dimensional modeling is acceptable when the reactants are admitted uniformly over the bed cross section. If, however, reactants are metered into the bed at individual feed points, threedimensional modeling may become necessary. Such models have been devised for the combustion of coal in bubbling fluidized beds [232–234]. As a rule, the modeling of solids behavior in bubbling fluidized-bed reactors is based on that in stirred tanks. Fluidized-bed combustion is one
Temperature Homogeneity is a virtually fundamental property of the fluidized-bed reactor. Even so, one exception is industrially important: in high-pressure fluidized-bed furnaces, the high energy density can cause local hot spots near the fuel injection points [235]. Reactor models that take care of this have been described, e.g., in [236].
9.2.2. Circulating Fluidized-Bed Reactors
In the early days of circulating fluidized-bed reactor modeling, negligible axial dispersion and laterally uniform flow structure were believed to characterize these systems. Thus, simple plugflow models were used [237]. This approach was found to oversimplify the behavior of circulating fluidized-bed reactors, because a significant amount of axial dispersion was observed. As a result, the plug-flow model has often been modified by adding a dispersion term to the balance equations. Axial dispersion coefficients have been determined by many authors who measured the residence time distribution of tracer gases [238, 239]. Typical values of Peclet numbers they found are on the order of 10. By means of a model reaction it has been proved that in many cases circulating fluidizedbed reactors cannot be characterized by solely considering mixing phenomena [240]. Instead,
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the presence of mass-transfer limitations and bypassing was found to have a significant influence. In analogy to low-velocity fluidized beds a detailed description of the local flow structure within the reaction volume must serve as a basis for appropriate reactor modeling. The highly nonuniform flow structure of circulating fluidized beds described in Section 2.9.2 has led to reactor models which separately deal with different axial zones. The bottom zone–if it exists under the given operating conditions–can be described by models whose basic approaches were originally developed for modeling of bubbling fluidized beds [241]. Modeling of the upper section of the circulating fluidized bed is in most cases based on a proper description of the heterogeneous core–annulus flow structure [242–244]. These state-of-the-art models are one-dimensional and define two phases or zones which are present at every axial location: 1. A dense phase or annulus zone: high solids concentration, gas stagnant or moving downwards 2. A dilute phase or core zone: low solids concentration, gas flowing quickly upward. Similar to the situation in bubbling fluidized beds, the two phases exchange gas with each other and are modeled by separate equations which are obtained from mass balances for each component in each phase. Today’s models still suffer from the problem that not all fluid-mechanical variables can be predicted on the basis of the operating conditions. Instead, reasonable estimations or measurements in cold-flow models are used to obtain numerical values for many variables. A common feature of all models for the upper part of circulating fluidized beds is the description of the mass exchange between dense phase and dilute phase. In analogy to low-velocity fluidized beds the product of the local specific mass transfer area a and the mass-transfer coefficient k may be used for this purpose. Many different methods for determination of values for these important variables have been reported, such as tracer-gas backmixing experiments [241], non-steady-state tracer-gas experiments [245], model reactions [244] and theoretical calculations [243].
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Similar to the bubbling fluidized-bed reactor, the solids behavior of the circulating fluidizedbed reactor can usually be described as completely mixed. This does not hold for riser reactors with very high gas velocities, such as those used in FCC risers (u > 10 ms1). Here, better modeling results will be obtained by assuming dispersed plug flow of solids [239]. Like for bubbling fluidized beds, it can be assumed that circulating fluidized beds exhibit a high degree of temperature homogeneity even in the case of highly exothermic reactions. However, in the case of very large circulating fluidized beds for coal combustion, significant horizontal and vertical temperature profiles have been observed inside the combustion chambers [246]. Despite the many uncertainties, circulating fluidized bed reactors have been modeled successfully. For example, three-dimensional gas and solids concentration profiles were calculated in circulating fluidized-bed boilers with local injection of reactants [247] and coal feeding via discrete feeding points [248].
9.3. New Developments in Modeling Fluidized-Bed Reactors 9.3.1. Computational Fluid Dynamics
The models described above follow the ‘‘classical’’ chemical engineering approach which replaces the complex particle–fluid interaction in the fluidized bed by idealized configurations (plug flow, stirred tank, either valid overall or in regions) with mixing and mass-transfer coefficients describing the transport of matter. However, more recently, there has been a strong tendency to model the fluid mechanics of fluidized-bed reactors from first principles. The problem of computational fluid dynamics (CFD) modeling in this area is that the particle–particle and particle–fluid interaction must be considered on the particle scale, while the reactor performance must be described on a much larger scale, typically on the order of several meters. This leads to computational difficulties and problems with available computing capacities. At present (ca. 2006) there is no generally accepted CFD model of the fluidized-bed reactor available, but rapid progress can be seen in this area [249–253].
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A promising approach appears to be multiscale modeling strategy [254]. The idea essentially is that fundamental models which take into account the relevant details of fluid–particle (lattice Boltzmann model) and particle–particle (discrete-particle model) interactions are used to develop closure laws to feed continuum models which can then be used to simulate the flow structures on a larger scale. Figure 47 illustrates this approach, which finally leads to the discrete-bubble model and should be applicable to the large industrial scale of the bubbling fluidized-bed reactor. The multiscale methodology [255] still requires development work but provides a good chance to arrive at more realistic fluidized-bed reactor models in the not too far future.
9.3.2. Modeling of Fluidized-Bed Systems
Another line of development is the modeling of fluidized-bed reactor systems. Whereas previously the isolated fluidized bed was modeled,
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the focus now is on the coupling between the fluidized bed and the cyclone for catalyst recovery and recycle [256] or even on the coupling between two fluidized-bed reactors [257], e.g., reactor–regenerator systems as are used in the FCC and maleic anhydride processes. As an example, Figure 48 shows a fluidizedbed coupled with a cyclone and its translation into the model system. Attrition leads to a loss of material from the system, which requires the addition of fresh catalyst after some time (Fig. 49). A population balance model which considers the changes in the catalyst particle size intervals allows the change in the catalyst inventory with time to be followed. We see that it takes several weeks for the system to reach a quasisteady state. As a consequence of attrition and incomplete separation in the cyclone, the mean particle diameter in the bed increases with time, and this leads to larger bubbles and a reduced area of mass transfer between bubbles and the surrounding suspension in the bed. As a further consequence the conversion rate of a simple first-order reaction falls off with time. Finally, Figure 50
Figure 47. The multiscale approach for CFD modeling of fluidized-bed reactors [254].
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Figure 48. Fluidized-bed reactor model system [256]
Figure 49. Reactor behavior as a function of operating time [256]
shows that improvements in the efficiency of the solids-recovery system are able to increase the conversion rate again, which is in agreement with large-scale industrial experience [258, 259].
10. Scale-up Typical diameters of bench-scale fluidizedbed reactors are roughly 30–60 mm, and of
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Internals. Whereas the laboratory fluidized bed is generally operated with no internals, plant equipment often must contain bundles of heatexchanger tubes. Screens, baffles, or similar internals are frequently used to redisperse the bubble gas in industrial reactors. The mass-transfer area is thus increased relative to the fluidized bed without internals; the extra area can be utilized to partially offset the conversion-reducing effects of bed diameter and gas distributor [263]. Figure 50. Influence of the Sauter diameter on the chemical conversion of a simple first-order reaction [256].
pilot-scale units 450–600 mm, which should allow a reliable scale-up [9]. Full-scale fluidized-bed reactors used in the chemical industry have diameters up to ca. 10 m. Circulating fluidized bed combustors are even bigger with bed cross-sectional areas reaching 200 m2 [261]. As equipment size increases, characteristic changes take place in the gas–solid flow that can decisively affect reactor performance. Such changes result either directly from the geometry or indirectly from design changes made as the unit is enlarged. In particular, experience has shown that the following factors affect the performance of bubbling fluidized beds during scale-up [262]:
Catalyst Particle-Size Distribution. Bubble growth is influenced by the proportion of fines in the particle-size distribution of the bed (usually measured as the weight fraction < 0.044 mm) or by the mean grain size d p (via umf , Eq. 18). If the content of fines increases, bubbles collapse sooner and the equilibrium bubble size becomes smaller, with a resultant greater bubble–suspension mass-transfer area. This effect generally is fully developed only in the plantscale reactor, where bubbles can grow without the hindrance of vessel walls. Thus, in principle, the performance of catalytic fluidized-bed reactors can be controlled by modifying the catalyst particle-size distribution [112, 264]. The recommended content by weight of fines (< 0.044 mm) for ‘‘good fluidization’’ is 30–40 % [265], but maintaining this high a fines content in the system over a long span of time requires a very efficient solids recovery system.
Bed Diameter. According to Equation (22), the mean upward bubble velocity increases as the bed diameter d t increases. As a result, the bubbles have a shorter residence time in the bed; hence the exchange area between the bubble and suspension phases is smaller, so conversion is reduced [263]. In case of circulating fluidizedbed combustors, measurements have shown that the downwards velocity of solids in the wall zone increases drastically with increasing size of the combustor [260].
Lateral Mixing of Reactants. On a laboratory scale, reactants experience compulsory uniform distribution over the bed cross section. In plant equipment, on the other hand, reactants often arrive in the reactor via individual feed points. The resulting uneven distribution of reactants can have a marked effect on reactor performance, which has been shown for the effect of coal feeding on the emission properties of fluidized-bed furnaces [248].
Grid Design. In the laboratory, porous plates are the preferred type of gas distributor because of the ease of working with them. Gas distribution becomes worse when these are replaced by industrial distributor designs; thus the exchange area between the bubble and suspension phases is reduced, again with consequently lower conversion [43].
Secondary Reactions in the Freeboard. In a bench-scale apparatus, the fluidized gas is rapidly cooled by the vessel wall in the freeboard space after leaving the bed, so secondary reactions in the freeboard are often negligible. Such is not the case in the plant-scale reactor. The action of wall cooling is not significant here, and the entrainment rate is high because of the higher
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fluidization velocities common in full-scale equipment. Both effects – lack of cooling and high solids concentration in the freeboard – may lead to marked secondary reactions in the freeboard of industrial fluidized-bed reactors. In the case of a system of consecutive reactions where the desired product is formed as an intermediate, the freeboard reactions will generally lower the selectivity. The effect of freeboard reactions has been demonstrated for the example of NO and CO emissions from a fluidized-bed furnace [232].
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5 W.-C. Yang (ed.): Fluidization, Solids Handling and Processing, Noyes Publications, Westwood 1999. 6 D. Gidaspow: Multiphase Flow and Fluidization, Academic Press, London 1994. 7 J. R. Grace, A. A. Avidan, T. M. Knowlton (eds.): Circulating Fluidized Beds, Chapman & Hall, London 1997. 8 W. -C Yang (eds.): Handbook of Fluidization and FluidParticle Systems, Marcel Dekker, New York 2003. 9 M. Pell: Gas Fluidization, Elsevier, Amsterdam 1990. 10 J. G. Yates: Fundamentals of Fluidized-Bed Chemical Processes, Butterworths, London 1983. 11 K.-E. Wirth: Zirkulierende Wirbelschichten – Stromungsmechanische Grundlagen, Anwendungen in der Feuerungstechnik, Springer Verlag, Heidelberg 1990. 12 P. Basu: Combustion and Gasification in Fluidized Bed , Taylor & Francis, Boca Raton 2006. 13 K. B. Mathur, N. Epstein: Spouted Beds, Academic Press, New York 1974. 14 K.-E. Wirth in E.-U. Schl under (ed.): VDI W armeatlas, 8th ed., VDI Verlag, D usseldorf 1997, pp. Lf 1–9. 15 K.-E. Wirth in E.-U. Schlunder (ed.): VDI Heat Atlas, 1st English ed., VDI Verlag, D usseldorf 1993, pp. Lf 1–9. 16 H. Martin inE.-U. Schlunder (ed.): VDIW armeatlas, 8th ed., VDI Verlag, D usseldorf 1997, pp. Mf 1–9. 17 H. Martin in E.-U. Schlunder (ed.): VDI Heat Atlas, 1st English ed., VDI Verlag, D usseldorf 1993, pp. Mf 1–9. 18 Ullmann, 4th ed. 3, 433–460. 19 Ullmann, 4th ed. 3, 480–493. 20 U. Arena, R. Chirone, M. Micchio, P. Salatino (eds.): Fluidization XI , Engineering Conferences International, Brooklyn 2004. 21 M. Kwauk, J. Li, W.-C. Yang (eds.): Fluidization X , Engineering Foundation, New York 2001. 22 L.-S. Fan, T. M. Knowlton (eds.): Fluidization IX , Engineering Foundation, New York 1998. 23 K. Cen (eds.): Circulating Fluidized Bed Technology VIII , International Academic Publishers, Beijing 2005. 24 J. R Grace, J.-X. Zhu H. de Lasa (eds.): Circulating Fluidized Bed Technology VII , Can. Soc. Chem. Engng., Ottawa 2002. 25 J. Werther (ed.): Circulating Fluidized Bed Technology VI , DECHEMA, Frankfurt 1999. 26 L. Jia (ed.): Proc. 18th Int. Conf. Fluidized Bed Combustion, ASME, New York 2005. CD-ROM, ISBN 0-7918-3755-6. 27 S. Pisupati (ed.): Proc. 17th Int. Conf. Fluidized Bed Combustion, ASME, New York 2003. CD-ROM, ISBN 0-7918-3755-6. 28 F. Preto (ed.): Proc. 16th Int. Conf. Fluidized Bed Combustion., ASME, New York 2001, CD-ROM, ISBN 07918-3523-5. 29 S. Ergun, Chem. Eng. Prog. 48 (1952) 89–97. 30 J. Werther, Chem.-Ing.-Tech. 54 (1982) no. 10, 876– 883. 31 C. Y. Wen, Y. H. Yu, AIChE J. 12 (1966) 610–612. €
Catalyst Attrition. Catalyst attrition is minimal in laboratory apparatus, because of the use of porous plates as gas distributors, as well as the low gas velocities and bed depths. Attrition is necessarily greater in industrial reactors. To reduce this risk in scale-up, the attrition tests described in Section 2.11 should be carried out and the results converted to the full-scale conditions with the aid of Equations (34), (35) and (36).
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Other Factors. In addition to the factors just listed, many other effects become apparent when a fluidized-bed reactor is scaled up that are difficult to calculate. Examples are the risk of nonuniform gas distribution over very large cross sections in shallow fluidized beds; the formation of deposits in the bed; the fouling of heatexchange surfaces; and catalyst aging and poisoning. On the whole, accordingly, the scale-up of fluidized-bed reactors is a complex process, commonly requiring a large amount of pilotscale experimentation. Current knowledge about the fluid mechanics in the fluidized bed, however, enables simulation calculations of many of the scale-up effects, so the amount of testing during process development may be decreased and the risk can be at least limited.
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