Ullmann Fixed-Bed Reactors

 

Catalytic Fixed-Bed Reactors GERHART  E IGENBERGER,  Institut fu¨r Chemische Verfahrenstechnik, Universita¨t

Stuttgart, Stuttgart, Germany WILHELM   RUPPEL,  Retired from BASF SE, Ludwigshafen, Germany

1. 1.1. 1.2.. 1.2

FixedFixe d-Be Bed d Re Reac acto tors rs wi with th GasGa s-Ph Phas asee Reac Reacti tion onss . . . . . . . . . . . . . . Introduction. . . . . . . . . . . . . . . . . . .

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Catalyst Cataly st For Forms ms for Fix Fixed ed-Be -Bed d Reactors. . . . . . . . . . . . . . . . . . . . . .

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1.6. 1. 6.

1.2.1. Fluid Fluid Flow, Mass and Heat Heat Transfer, and Chemical Reaction in Cata Ca taly lyst st-F -Fil ille led d Tub Tubes. es. . . . . . . . . . . . . 1.2. 1. 2.2. 2. Re Regu gula larr Cat Catal alys ystt Str Struc uctu ture ress . . . . . . . . 1.2.3. Compa Compariso rison n and Evalu Evaluatio ation n of  Diff Di ffer eren entt Cat Catal alys ystt For Forms ms . . . . . . . . . .

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1.3. 1. 3.

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Adia Ad iaba bati ticc Fi Fixe xedd-Be Bed d Re Reac acto tors rs . . . . .

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1.3.1. 1.3. 1. Axia Axiall and and Ra Radi dial al Fl Flow ow Re Reac acto tors. rs. . . . . 16 1.3.2. Multistage Reactors Reactors with Interstage Heat Heat Tra rans nsfe ferr . . . . . . . . . . . . . . . . . . . . . . 17 1.4.

Fixed-Bed Reac Fixed-Bed Reactors tors with Inte Integrate grated d Hea eatt Exc Exch han ange. ge. . . . . . . . . . . . . . . . .

1.4.1. 1.4. 1. Heat Heat-E -Exc xcha hang ngee Con Conce cept ptss . . . . . . . . 1.4.2. Heat-Transfer Media for Fixed-Bed Reac Re acto tors rs . . . . . . . . . . . . . . . . . . . . 1.4.3. Coole Cooled d Reactors for Exothermic Exothermic Reac Re acti tion onss . . . . . . . . . . . . . . . . . . . 1.4.4. Heat Heated ed Reactors Reactors for Endot Endothermi hermicc Reac Re acti tion onss . . . . . . . . . . . . . . . . . . . 1.4. 1. 4.5. 5. In Influ fluen enci cing ng th thee Cou Cours rsee of of Rea React ctio ion. n. 1.5. 1. 5.

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Heat He at-I -Int nteg egra rate ted d Re Reac acto torr Co Conc ncep epts. ts. .

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1. Fixed Fixed-Bed -Bed Reactors Reactors with with Gas-Phase Reactions 1.1. Introd Introductio uction n The core part of any fixed-bed reactor is the solid catalyst where the reaction takes place. A larg la rgee va vari riety ety of ca catal talys ystt st stru ructu cture ress ar aree ap appl plie ied d in practice. One class of structures, used in randomly packed beds, consists of cataly catalyst st pelle pellets ts of different shapes. A second class comprises

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1.5.1. Autotherm Autothermal al Reactors with Extern Ext ernal al and Int Intern ernal al Hea Heatt Exc Exchan hange ge . 34 1.5.2. Heat Heat-Inte -Integrate grated d Reac Reactors tors for Coupli Coupling ng of  Endo En do-- and and Ex Exot othe herm rmic ic Re Reac acti tion onss . . . . 35

Wiley-VCH Verlag GmbH & Co. KGaA, Weinheim 10.1002/14356007.b04_199.pub2

Oper Op erat atio iona nall an and d Sa Safe fety ty Is Issu sues es . . . . .

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1.6.1. Para 1.6.1. Paramet metric ric Se Sensi nsitiv tivity ity and Run Runawa away. y. . 1.6.2. Movin Moving g Temperat Temperature ure and Reac Reaction tion Fronts . . . . . . . . . . . . . . . . . . . . . . . . 1.6. 1. 6.3. 3. Ot Othe herr Saf Safet ety y Asp Aspec ects ts . . . . . . . . . . . . .

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1.7.. 1.7

Periodic Period ic Oper Operati ation on of of Fixe Fixed-B d-Bed ed Rea eaccto tors. rs. . . . . . . . . . . . . . . . . . . . . .

1.7.1.. Fixed-Bed 1.7.1 Fixed-Bed Reactors Reactors with Perio Periodic dic Flow Reve Re vers rsal al an and d Exo Exoth ther ermi micc Re Reac acti tion on . . . 1.7.2.. Fixed 1.7.2 Fixed-Bed -Bed Reactors Reactors with Perio Periodic dic Flow Reversal for Coupled Exo- and Endo En doth ther ermi micc Rea React ctio ions ns . . . . . . . . . . . 1.7. 1. 7.3. 3. Pe Peri riod odic ic Fe Feed ed Cy Cycl clin ing g .. . .. .. . .. .. 2. 2.1 .1.. 2.2.. 2.2 2.3.. 2.3 2.4. 2. 4.

FixedFixe d-Be Bed d Re Reac acto tors rs fo forr Liqu Li quid id-P -Pha hase se Rea React ctio ions ns . . . . . . . . . . . Int ntrrod oduc ucti tion. on. . . . . . . . . . . . . . . . . . . FixedFix ed-Bed Bed Cat Cataly alyzed zed Liqu Li quid id-P -Pha hase se Re Reac acti tion onss . . . . . . . . . . Upward Upw ard Liq Liquid uid Flo Flow w throu through gh Fixe Fixed d Beds . . . . . . . . . . . . . . . . . . . . . . . . . Reac Re acto torr La Layo yout ut an and d Op Oper erat atio ion n ....

2.4.1. 2.4. 1. Safe Safety ty Is Issu sues es . . . . . . . . . . . . . . . . . . . 2.4.2. 2.4 .2. Exa Exampl mple: e: Ami Aminat nation ion of Alc Alcoho ohols ls . . . .

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48 52 54 54 55 59 62

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regularly regula rly arr arrang anged ed str struct ucture uress lik likee mon monoli oliths ths with wit h flo flow w ch chan anne nels ls of di diff ffer eren entt sh shap ape. e. Th This is article does not address the topic of catalyst design with respect to its intrinsic catalytic and internal transport properties, but assumes that thee ca th cata talys lystt fo forr a sp spec ecifi ificc re reac actio tion n is gi give ven. n. Section Sec tion 1.2 foc focuse usess ins instea tead d on the dif differe ferent nt catalyst shapes and on their influence on fluid flow and flow distribution in the packed bed, on pressu pre ssure re dro drop, p, on mas masss tra transf nsfer, er, and in par partic ticula ularr on heat transfer within the packed bed.

 

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Cataly lyttic Fix ixed ed--Bed Reactors

Figure 1.  Adiabatic fixed-bed reactor (left) and multitubular fixed-bed reactor (right)

With regard to application and design, it is convenient to differentiate between fixed-bed reac re acto tors rs fo forr ad adia iaba bati ticc an and d no nona nadi diab abat atic ic operation. Since temperature control is one of  thee mo th most st im impo port rtan antt me mean anss of in influ fluen encin cing g a chemical reaction, adiabatic reactors are used prim pr imari arily ly wh wher eree th thee ad adia iaba batic tic te temp mpera eratu ture re change during the reaction is small or where there is only one major reaction pathway. In these cases no adverse effects on selectivity or yield due to the adiabatic temperature developmentt are exp men expect ected. ed. In adi adiaba abatic tic rea reacto ctors rs the catalyst is present in the form of a fixed bed whic wh ich h is su surr rrou ound nded ed by an ou oute terr in insu sula latin ting g

fixed-bed reactor’’. The most common arrangementt for iso men isothe thermal rmal reactor reactor ope operati ration on is the multitu mul titubul bular ar fixe fixed-b d-bed ed rea reacto ctor, r, in whi which ch the catalyst is arranged in the tubes, and the heat carrierr circu carrie circulates lates externally around the tubes (Fig. 1, right). right). Since isotherma isothermall reacti reaction on control control does not necessarily provide optimum selectivity or yield, heat-exchange sections with changchanging in g te temp mper eratu ature ress of th thee he heat at ca carri rrier er ca can n be design des igned ed to est establ ablish ish an opt optima imall tem temper peratu ature re profile along the flow path. Fixed-bed reactors with integrated heat supply or removal are discussed in Section 1.4. Since the reactor feed must be heated to the

 jacket (Fig.1.3.1. 1, left). Such reactors are discussed in Section If the reaction temperature must be maintained taine d within a speci specified fied range, multis multistage tage adiabatic reactors can be used, whereby between each stage the temperature can be influenced by heat he at ex exch chan ange ge or by co cold/ ld/ho hott ga gass in inje ject ction ion.. Su Such ch reactors are discussed in Section 1.3.2. Reactions with a large heat of reaction and reactions that are very temperature sensitive are usually carried out in reactors in which heat of  reaction is provided to or removed from the fixed bed via a circulating heat-transfer medium. Since in most cases the task of the heattransfer cycle is to maintain the temperature in thee fix th fixed ed be bed d wi with thin in a sp spec ecific ific na narr rrow ow ra rang nge, e, th this is concept is frequently described as ‘‘isothermal

ignition temperature reaction before reaction starts, of thethe hot catalytic reactor effluent is often used to heat the cold reactor feed. This causes a thermal feedback which results in socalled autothermal reactor concepts. This type of rea reacti ction on control, control, whi which ch off offers ers cer certai tain n spe specific cific features and development perspectives, is discussed in Section 1.5. Fixed-bed reactors for industrial syntheses aree ge ar gene nera rally lly op oper erat ated ed in a st stat atio iona nary ry mo mode de (i (i.e .e., ., unde un derr co cons nstan tantt op oper erat atin ing g co cond nditi ition ons) s) ov over er prolonged prolo nged produ production ction runs. Design therefore concentrates on achieving optimum stationary operation. Instationary operation, however, is unavoi una voidab dable le dur during ing sta startu rtup p and shu shutdo tdown wn as well we ll as du duri ring ng lo load ad ch chan ange ge or in th thee ca case se of au auto toma matic tic co cont ntro roll ac actio tions. ns. In pa part rticu icula lar, r,

 

Catalytic Fixed ed--Bed Reac acttor ors s

fixed-bed reactors with a strongly exothermic reacti rea ction on exh exhibi ibitt an, at tim times, es, sur surpris prising ing dyn dynami amicc behavior which can affect operational safety. Thiss is dis Thi discus cussed sed tog togeth ether er with oth other er ope operati rationa onall and safety issues in Section 1.6. Contrary to the above-mentioned stationary operation concepts, a deliberately unstationary, mostly periodic operating mode has proved to be of ap advantage a odic number special This Th is appl plie iess to in peri pe riodi c flo flow wof re reve vers rsal al cases. or to periodic feed cycling, which are discussed in Section 1.7. In a pr prod oduc ucti tion on pl plan antt th thee re react actor or ca can n be regard reg arded ed as the cen centra trall app apparat aratus. us. How Howeve ever, r, compared to the remaining parts of the plant for preparing the feed and for separating and workin wor king g up the pro produc ducts, ts, oft often en it is is by no mea means ns the largest and most cost-intensive component. In many cases the achievable conversion in the reactor is limited for thermodynamic (equilibrium) and kinetic reasons (selectivity). It is then necessary to separate the reactor effluent into produc pro ducts ts and unc unconv onvert erted ed fee feed d com compon ponent entss (se (seee Fig. 2), which are recycled to the feed. This recycling procedure involves costs   For product separation .   For recycle compression .

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 For repeated heating and cooling of the circulating reactants to the reaction temperature and back to the temperature of the separating device .   Due Due to lo loss ss of pr prod oduc uct, t, re resu sulti lting ng fr from om th thee ne need ed to re remo move ve pa part rt of th thee re recy cycl clee as a bl blee eed d st stre ream am to limit accumulation of inert substances or harmful byproducts in the recycle loop. .

To mi mini nimiz mizee th thes esee co cost sts, s, im impr prov ovem emen ents ts should aim at increasing the product yield per pass pa ss an and d at de decr crea easi sing ng th thee am amou ount nt of in iner ertt substances in the reaction mixture. A reaction process as depicted in Figure 2 follows the classical unit operation concept of  chemica che micall eng engine ineeri ering, ng, acc accord ording ing to whi which ch a clear division into preparation (mixing and preheatin hea ting) g) of fee feed d com compon ponent ents, s, che chemic mical al rea reacti ction, on, product separation, and cooling is achieved in differe dif ferent nt uni units. ts. A mor moree rec recent ent dev develo elopme pment nt under the heading of ‘‘integrated’’ or ‘‘multifunctional’’ reactor concepts contrasts the unit operation concept. Its aim is to provide optimal reac re actio tion n co cond ndit itio ions ns in th thee re reac acti tion on un unit it by incorporating optimal heat and addition or removal of reaction components at the reaction site.. Hea site Heatt inte integra gratio tion n is is dis discus cussed sed in som somee det detail ail in Se Sect ctio ions ns 1. 1.4, 4, 1. 1.5, 5, an and d 1. 1.7; 7; fo forr ad addi diti tion onal al

Figure 2.  Fixed-bed reactor with product separation and recycle.

a) Fixed-bed reactor, b) Feed preheater, c) Product cooler, d) Recycle compressor, e) Separation column

 

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Cataly lyttic Fix ixed ed--Bed Reactors

infor information mation on reacto reactors rs with integ integrated rated react reactant ant additi add ition on or rem remova oval, l, see  Membrane React Reactors ors and  Chromatographic Reactors. Throug Thr oughou houtt all fol follow lowing ing sec sectio tions ns mod model el simulation results are used to explain and discuss specific points, but the underlying models and an d th thei eirr nu numer meric ical al si simu mulat latio ion n wil willl no nott be addressed. addre ssed. See   Mod Model el Rea Reacto ctors rs and the their ir

designed design ed pro process cess cha chain in of fee feed d pre prepar paratio ation, n, reaction, and product separation. In the case of selectivity-sensitive multistep reactions, react ions, any devia deviation tion from the optimum operation values inevitably leads to yield losses. Such deviations could be the result of a nonuniform residence-time distribution in the packed bed be d du duee to flo flow w di disp sper ersio sion n an and d flow by bypa pass ss

Design Equations as well asr standard textbooks and monographs monograph s for further furthe references refer ences [1–4].

Since many reactions may take place with a considerable heat of reaction, a corresponding heat transport is usually superimposed on the mass transport. Catalyst Catal yst desig design n and deve developmen lopmentt by contr controlol-

phenomena, well as ofindeviations from uniform reactionasconditions the catalyst itself as a re resu sult lt of ma mass ss-- an and d he heat at-t -tra rans nspo port rt re resi sista stanc ncee in the particles and the outer boundary layer [5]. Temper Tem peratu ature re con contro troll pla plays ys a pre predom domina inant nt role in selective reaction control in general and in particular in the case of exothermic multistep reactions. In nonadiabatic reactors the catalyst must therefore be arranged properly in the fixed bed be d in or orde derr to en ensu sure re go good od he heat at tr tran ansfe sferr to a he heat at carrier (see Sections 1.2.3.3 and 1.2.3.4). A further requirement is a low pressure loss of the fixed bed. This applies particularly particularly if the reaction conversion in a single pass is low, so that th at th thee re reac acti tion on mu must st be ca carr rrie ied d ou outt wi with th a la larg rgee recirculation ratio (Fig. 2). A low pressure drop is of prime importance in off-gas purification, where large off-gas streams must be treated at minimall additi minima additional onal cost (Sectio (Section n 1.5.1) 1.5.1).. Finally, the catalyst should be available in sufficiently high concentration to keep the reactorr vo to volu lume me lo low. w. Th Thee ma main in pa param ramet eter erss ar aree th thee th thee void fractio fractions ns e of th thee fix fixed ed be bed d [c [cub ubic ic me mete ters rs of  free gas space per cubic meter of reactor volume] and the specific external catalyst surface ap  [square meters of catalyst surface per cubic meter reactor volume], which is decisive if the

lin ling g the aboveabo ve-men tioned ned microk mic rokine inetic tics s and intern int ernal al tra transp nsport ortmentio proper pro pertie tiess as well wel l as provid pro viding ing the necessary mechanical stability of the catalyst pellets is the specific task of the catalyst designer. For details, see [94]. Once the catalyst is specified, including its kinetic and (internal) transport properties, reaction conditions (feedstock concentrations, pressure, temperature, and residence time) can be found fou nd tha thatt lea lead d to opt optimu imum m yie yields lds.. Rea Reacti ction on engineers must determine these conditions and ensure that they are maintained in an industrial reacto rea ctor. r. Thi Thiss is not a sim simple ple str straigh aightfo tforwa rward rd proc pr oced edure ure bu butt ma may y re requ quire ire a tr tria iall an and d er erro rorr approa app roach ch wit with h mul multip tiple le ite iterat ration ion loo loops ps to ens ensure ure thatt the optimal tha optimal cata catalys lystt is ava availab ilable le for the optimal reactor configuration in an optimally

reaction isecontrolled by Thee ab Th abov ove requ re quire ireme ment ntssexternal aree tomass ar some so metransfer. exte ex tent nt cont co ntra radic dicto tory ry,, wh whic ich h ha hass le led d to a va vari riet ety y of  differ dif ferent ent cat cataly alyst st sha shapes pes and arr arrang angemen ements. ts. Thee mo Th most st co comm mmon on sh shap apes es ar aree co comp mpar ared ed in Figure 10. The simplest kind of fixed catalyst bed is a random packing of catalyst pellets in a tube (Fig. 1). Different pellet shapes are in use, such as spheres, cylinders, rings, flat disk pellets, or crus cr ushe hed d ma mate teri rial al of a ce certa rtain in si siev evee fr frac acti tion on (Fig. (Fi g. 3A 3A). ). Me Mean an pa part rticl iclee di dime mens nsio ions ns ra rang ngee fr from om 1 to 10 mm; the minimum particle or channel size is limited primarily by pressure drop considerations, and the maximum diameter by the specific outer surface area for mass and heat transfer.

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1.2. Catalys Catalystt Forms for Fixed-Bed Fixed-Bed Reactors The essential part of a fixed-bed reactor is the catalyst, where the reaction takes place. Seen from the flowing gas, the following steps of the overall reaction can be distinguished:  Diffusion of the reactants from the flowing gas through the outer gas–particle boundary layer lay er and thr throug ough h mac macrop ropore oress and mic microp ropore oress to the catalytically active sites .  Chemisorption on active sites .   Surface reactions .  Desorption of the products .   Bac Backdi kdiffus ffusion ion of the pro produc ducts ts int into o the flow flowing ing gas .

 

Catalytic Fixed ed--Bed Reac acttor ors s

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Figure 3.  Catalysts for fixed-bed reactors

A) Different pellet types; B) Honeycomb monolith; C) Crossed corrugated plate packing

Th Thee de deta tail ilss of flu fluid id flo flow, w, he heat at an and d ma mass ss transfer, and chemical chemical reaction in catalyst-filled tubes have received considerable attention over the last decades. The state of the art is discussed in Sections 1.2.1–1.2.3. It leads to the conclusion sio n tha thatt a det detail ailed ed con consid sidera eratio tion n of the dif differ ferent ent transport and reaction effects both on the pellet and the tube tube scale is still out of reach reach for curren currentt reacto rea ctorr mod modeli eling ng and des design ign.. Ins Instea tead d mea mean-fi n-field eld models based on local averages are generally used use d in rea reactor ctor des design ign.. The They y inc includ ludee exp experi erimen men-tally tal ly ver verifie ified d cor correl relati ations ons for hea heatt and mas masss transfer and allow for a general comparison of 

with wall openings or in cross-corrugated plate packings (Fig. 3C). Regular catalyst structures are discussed in Sections 1.2.2 and 1.2.3.2.

di diff ffer eren entt1.2.3. cata ca taly lyst st sh shap apes es,, as di disc scus usse sed d in Section Much effort has been devoted to replacing random packings by a regular catalyst arrangement me nt to pr prov ovid idee mo more re pr pred edic icta tabl blee re reac actio tion n conditions, to reduce the pressure drop of the reacting fluid, and/or to increase the heat transferr to th fe thee co cool olin ing g or he heati ating ng wa walls lls.. Pre Press ssur uree dr drop op considerations led to the introduction of monolith catalysts with parallel channels (Fig. 3B) which meanwhile are domin dominating ating catalytic gas purification purific ation applic applications ations like autom automotive otive exhaust purification and flue gas treatment. However, due to the absence of radial flow components, nen ts, lat latera erall heat transf transfer er to a hea heatt transfer transfer wal walll is str strong ongly ly red reduce uced d in mon monolit olith h cat cataly alysts sts.. A radial flow can be induced in catalyst structures

vective transporteliminating in randomly packedgas–catacatalyst beds, essentially external lyst mass transfer resistance as compared to the transport resistance in the catalyst pores. Conversely, the thermal conductivity of the catalyst matrix is generally larger than that of the gas. This Th is me mean anss th that at th thee ex exte tern rnal al he heat at tr tran ansp spor ortt resistance between gas and catalyst surface is higher than the heat transport resistance inside the cataly catalyst st partic particles. les. Typicall tempe Typica temperature rature and conce concentrati ntration on profiles established for a spherical full catalyst and for a shell-type catalyst pellet are illustrated for a partial oxidation reaction in Figure 4. Such profiles can be calculated from the model equations tio ns gi give ven n in   Mod Model el Rea Reacto ctors rs and the their ir Design Equations under the assumptions made

1.2.1. Fluid Flow, Flow, Mass and Heat Heat Transfer, and Chemical Reaction in Catalyst-Filled Tubes

Industrial fixed-bed reactors are usually operated wit with h a cro crossss-sec sectio tional nal loa loadin ding g G z  1 kg ga gass per square meter of reactor cross section per second. This loading produces a sufficient con-

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Cataly lyttic Fix ixed ed--Bed Reactors

there. Whereas part of the desired product B woul wo uld d be ox oxid idiz ized ed fu furth rther er to CO2 in th thee ce cent nter er of  the full-type catalyst, catalyst, leadi leading ng to a higher pellet temp te mper eratu ature re an and d a re redu duce ced d se sele lecti ctivi vity ty,, th this is is no nott the case for the shell-type catalyst, where the reac re actio tion n is co confi nfine ned d to a th thin in ca cata taly lyst st la laye yerr on an inert pellet carrier. However, due to the lower cataly cat alyst st mas mass, s, the she shellll-typ typee cat cataly alyst st wou would ld

Figure 4.   Typical temperature and concentration profiles n for a partial oxidation reaction A O2 B; B  x O2 CO2 m H2O in a spherical full catalyst pellet (right) and a

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shell-type catalyst (left) for similar gas-phase conditions

require a larger Catalyst modificatio ifica tions ns with reactor respec res pectt volume. to imp improv roved ed int intern ernal al masss and hea mas heatt tra transf nsfer er can sub substa stanti ntially ally improve conversion and selectivity in fixed-bed reactors. An essential precondition precondition of Figure 4 is that the catalyst particle must be uniformly exposed over its entire surface to a flow with uniform velocity, temperature, and concentration. This is actually not the case in random packings. Figure 5 shows images of the quite inhomogeneou ne ouss lo loca call ma mass ss-t -tran ransf sfer er di dist strib ributi ution on in a random rando m pack packing ing of rings or cylinders [12]. The visual vis ualizat ization ion tec techniq hnique ue use used d is bas based ed on the

Figure Fig ure 5.   Visu Visualiz alizationof ationof exte external rnalmass mass-tran -transferdistribut sferdistribution ion arou around nd indiv individua iduall ring ring-sha -shaped ped or cylin cylindrica dricall cata catalyst lystpelle pellets ts from

a fixed-bed packing [12]

 

Catalytic Fixed ed--Bed Reac acttor ors s

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re reac acti tion on of a pu puls lsee of am ammo moni niaa in th thee ga gass st stre ream am conta co ntain inin ing g ai airr pl plus us hy hydr drog ogen en pe pero roxi xide de on a catalyst surface impregnated with manganese chloride. The conversion to dark brown manganese dioxide dioxide is so fast under under ambien ambientt conditions conditions that it depends only on the external mass-transfer resistance of the gas boundary layer. The intensity of the dark coloration is thus propor-

a tu tube be of di diam amet eter er d T   5 d P [8 [8]. ]. On th thee ri righ ghtt th thee computed compu ted axial compo component nent of the velocity distribution inside the packed tube is shown at severa sev erall cro cross ss sec section tionss ins inside ide,, bef before ore,, and beh behind ind the packing. The coloration marks the local gas concentration of a component which reacts at the pellet surface in an isothermal first-order reaction like in the flow visualization experi-

tional to the reaction rate of pulse the surface reaction. For local a constant ammonia in the gass flo ga flow w it is al also so pro propo porti rtion onal al to th thee lo loca call external mass transfer, and, if mass transfer and heat transfer are equivalent, also to the local external heat transfer [6]. Figure Fig ure 5 ind indica icates tes tha thatt the loc local al con condit dition ionss in random packings are much more complex than assume ass umed d in cur curren rentt mea mean-fi n-field eld mod models. els. Rec Recent ently ly availab ava ilable le det detaile ailed d CFD sim simula ulation tionss tog togeth ether er with tomographic images of the flow in catalyst packed tubes provide a more detailed picture, which is briefly reviewed in the following. The focus is on catalyst-filled tubes with low aspect ratio, i.e., a small number of catalyst particles across the tube radius, since such arrangements are used in multitubular reactors if a large heat of reaction must be removed or added over the tube wall. Figur Fig uree 6 sh show owss le left ft a sim simul ulat ated ed ra rand ndom om packing of catalyst spheres of diameter   d P   in

ment of Figure Clear Cle arly, ly, thee 5.flo th flow w di dist stri ribu butio tion n is st stro rong ngly ly irregul irre gular, ar, wit with h sta stagna gnant nt zon zones es beh behind ind pel pellet letss and flow channels of different flow velocities between pellets. Only after averaging the flow velocity around the circumference and along some dista distance nce   x   does does a mo more re re regu gula larr flow flow-velocity variation over the radius appear. This is shown in Figure 7 (right) for spherical pellets in a tube of diameter   d T   3   d P. Due to the averagi ave raging ng pro proced cedure, ure, the rem remain aining ing vel veloci ocity ty maxima are now considerably dampened. The radial flow profile closely resembles the mean radial void fraction or porosity profile of the packed tube, which is again averaged over the circumference and a certain height   x   (Fig. 7, left). Due to point contact at the wall the void fraction approaches unity and drops to a minimum at a distance  d P /2 away from the wall. A basica bas ically lly sim similar ilar pic pictur turee can be obt obtain ained ed for other pellet shapes [9, 10].

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Figure Fig ure 6.   Left:Random pack packing ing of sphe spherica ricall cata catalyst lystss of diam diamete eterr d P ina tu tubeof beof di diam amet eter5 er5 d P; Right Right:: Comp Computedaxial utedaxial velo velocity city distri dis tribu butio tion n at sev severa erall hei heigh ghts ts of the pac packin king g (1. (1.8 8 to 5.7 d P) as we wellas llas infr infro ontof ( 0.1 d P) an and d be behin hind d the pa packi cking ng ( 0.1 d P and 2  d P) [8]. The color indicates the change in concentration of a gaseous starting material due to a single reaction

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Cataly lyttic Fix ixed ed--Bed Reactors

Figure 7.   Circumferen Circumferentially tially averaged radial porosity (left) and radial velocity profiles (right) in a tube of diameter 3  d P  filled with spheres of diameter  d P  after several repackings [8]

Circles: values of the individual packings; lines: mean value and standard deviation

The det detail ailed ed exp experim eriment ental al and sim simula ulatio tion n results can be summarized as follows: .

  The local flow pattern in randomly packed tubes, characterized by vortices and the frequent separation and remixing of flow paths, is highly irregular and changes in detail over the flow length and with every repacking. For a given packi packing ng it is laminar/sta laminar/stationar tionary y below a pellet Reynolds number   ReP  of about 100 and tur turbul bulent ent/in /instat station ionary ary for   ReP   >   120 [11]. .   Due to the maximum void fraction at the wall there is a certain flow bypass close to the wall as compared to the bulk flow in the center of  the packing, in particular with tube diameters of a fe few w pa parti rticl clee di diam amet eter erssmixing and an d lo low w flo flow w velocities. However, lateral over the cross section section reduce reducess the influence influence of the wall bypa by pass ss,, wi with th th thee re resu sult lt th that at it ha hass no si sign gnific ifican antt influence on reaction conversion except for very shallow beds. .   As a re resu sult lt of th thee lo loca cally lly va vary ryin ing g flow ve velo loci city ty gas–pe gas –pelle llett mas masss and hea heatt tra transf nsfer er var vary y str strong ong-ly, both around individual pellets (Fig. 5) and over the cross section of the tube. .  At industrial flow rates on the order of 1 kg m2 s1 heat transfer between the tube wall and the catalyst pellets is dominated by convectivee radial trans vectiv transport port throu through gh latera laterally lly mixing flow str stream eamss in the the bulk bulk of the the pack packed ed bed bed and by heat transfer through a boundary layer at the wall. Again, due to the random position

of th thee pe pell llet etss at th thee wa wall ll,, lo loca call wa wall ll he heat at transfer trans fer strongly fluctuates. This can be seen in Figure 8, where the local gas–wall heat transfer has been visualized with the same technique as applied in Figure 5 [12]. In this case a sheet of paper impregnated with manganese chloride solution was placed around the inner tube wall. The white spots mark the area ar eass wh where ere th thee pe pelle llets ts to touc uch h th thee wa wall. ll. In th thes esee areas heat transfer takes place through heat conduction between the pellet–wall contact points poi nts and the sur surrou roundi nding ng sta stagna gnant nt fluid fluid,, while in the darker areas mass and heat transfer is due to lateral convection. The differences enc es in dar darkne kness ss bet betwee ween n pac packed ked bed bedss of  sphe sp here ress (Fi (Fig. g. 8, to top) p) an and d Ra Rasc schi hig g ri ring ngss (Fi (Fig. g. 8, bottom) indicates that wall heat transfer is

Figure 8.   Visualization of convective convective wall mass (and heat) transfer for spheres (d P  8 mm, top) and rings (6 10

 ¼

 

11 mm, bottom) [12]. Downflow of air with 0.4 m/s

 

Catalytic Fixed ed--Bed Reac acttor ors s

higher for spheres than for rings (see also Section 1.2.3.3).

9

(witho (wit hout ut ch chem emic ical al re reac acti tion on)) wi with th a st step ep change of the axial wall temperature. In [11] it is claimed that the respective correlations reported in VDI-Wa¨rmeatlas [13] provide a betterr agree bette agreement ment with heat-u heat-up p expe experiments riments than an alternative model [14] with radially varying heat conductivity and no temperature  jump at the wall.

As a consequence of the above observations it ca can n be co conc nclu lude ded d th that at de deta taile iled d mo mode deli ling ng of th thee complex interactions in case of an industrially relevant multistep reaction is still out of reach. However, it is also unnecessary, since a fixed.

be bed d re reac acto torr ca can n no notdetails t be op oper erat ated edhave prope pro perly rly in regions where these would a strong impact. Mean-field models which average over a ce cert rtai ain n re regi gion on of th thee pa pack cked ed be bed d an and d us usee empirical transport models and correlations are used instead. They have been shown to provide the necessary information for a rational, modelbased bas ed eva evalua luatio tion n of the rea reacto ctorr des design. ign. The modeling results presented and discussed in the foll fo llow owin ing g se sect ctio ions ns ar aree al alll ba base sed d on su such ch mo mode dels ls.. Their main assumptions are listed below:

 The radial transport of matter convective lateral mixing is described by aby radial dispersion coefficient with a similar correlation as for the radial convective transport of heat. .   Fo Forr a mo more re co coar arse se co comp mpar aris ison on of de desig sign n al alte terrnatives or features, radial variations can be neglec neg lected ted and a spa spatial tially ly one one-di -dimen mensio sional nal model with the main flow direction as only spatial coordinate can be used. Such a simplification plifica tion is often appropriate appropriate for tubes with small tube to particle diameter ratio.

  The mod models els ass assume ume axi axial al sym symmet metry, ry, i.e i.e., ., onl only y variations in the axial and the radial coordinate are considered. .   The influence of the radial variation of the axial velocit velocity y including including the flow bypa bypass ss at the wall results in an axial dispersion of a concentration or temperature pulse in the feed. This dispersion can be approximated by an axia ax iall di disp sper ersi sion on te term rm ad adde ded d to th thee me mean an convective flow term (plug flow plus axial dispersion). The value of the dispersion coefficient depends on the flow and packing characteri act eristi stics cs and is giv given en by cor correl relati ations ons der derive ived d from feed pulse experiments [13]. .  Pressure variations across the radius are gen-

Catalyst packings with different pellet shape are compared with respect to the above-mentioned mean-field properties in Section 1.2.3.3. Thiss all Thi allows ows sel select ection ion of app approp ropria riate te cat cataly alyst st shapes for specific fixed-bed processes. processes.

er eral ally ly ne negle glect cted edtubes and an d th the axia ax iall pr pres essu sure re dr drop op in catalyst filled ise described by simple correlations, as listed in Section 1.2.3.1. .  Mostly a quasi-homogeneous energy balance is used, assuming that catalyst and fluid temperature perat ure are equal equal.. .   The Then n fo forr th thee he heat at tr tran ansp spor ortt be betw twee een n th thee packed bed and the cooled or heated wall a two-parameter model is used. It assumes a radiall heat cond radia conductivit uctivity y lR in into to wh whic ich h th thee he heat at conductivity between adjacent pellets and the convec con vective tive hea heatt tra transp nsport ort thr throug ough h lat latera erall mix mix-ing is lumped. For the heat transfer through the close-to-wall region a heat transfer coefficient aw is applied, leading to a temperature step at the wall. The respective correlations are de deri rive ved d fr from om hea eatt-u up ex exp per erim imen ents ts

contrast randominpackings, the external heat and masstotransfer monolith catalysts is much more uniform and reproducible. But since flow in honeycomb monoliths with parallel walls is usually laminar, mass transfer to the wall may become a limiting factor at high reaction rates. This is not the case in such regular structures which induce a vortical flow mixing, like crosscorrugated plate packings (Fig. 3 C). Intensity and local distribution distribution of heat and mass transfe transferr on cross-corrugated structures depends on their geometrical parameters (wave form, amplitude, wavelength, angle of incidence [15]). Higher trans tra nsfe ferr co coef effici ficien ents ts ge gene nera rally lly re resu sult lt in an increased pressure drop. So far, crossed-corrugated plates have obtained industrial relevance in parallel-plate heat

.

1.2.2. Regula Regularr Catalyst Structures Structures

Monolith catalysts are used primarily in environmen ron mental tal cat cataly alysis sis whe where re lar large ge gas str stream eamss must mu st be pr proc oces esse sed d wi with th lo low w pr pres essu sure re dr drop op.. Examples Exampl es inclu include de squar square-chan e-channel nel monol monolith ith catalys al ysts ts fo forr ex exha haus ustt pu puri rifica ficati tion on or fo forr th thee re remo mova vall of nitrogen oxides from flue gases (Fig. 3 B). In

 

10

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

exchangers, as column packings, and as static mixe mi xers rs (e (e.g .g., ., Su Sulz lzer er ty type pe SMV SMV)) bu butt le less ss as catalyst packings. For adiabatic reactors honeycomb monoliths with small channel diameter would be prefer preferable able over crosse crossed-cor d-corrugate rugated d packings due to their lower pressure drop at comparable heat/mass transfer, as explained in Section 1.2.3.2. But since crossed-corrugated

can have advantages whenever lateral mixing and lateral heat and mass transport is an issue (see Sectio Section n 1.2.3. 1.2.3.3). 3). Asymptotic laminar velocity distributions as well as the reactant concentration of a single mass transfer limited reaction in catalyst coated mono mo nolit liths hs wi with th st stra raig ight ht pa para ralle llell wa walls lls ar aree sh show own n in Figu Figure re 9 [17 [17]. ]. Due to lam lamina inarr flow flow,, the the lat lateral eral

plate packings provide good lateral mixing they

reactant concentration profile in the flowing gas

Figure 9.   Axial-flow velocity distribution distribution under fully developed laminar conditions conditions (left) and starting gas concentration for a

single isothermal reaction (right) in monolith channels of different cross section (from [17])

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

11

Figure 10.  Specific outer surface area ap, void fraction e, wall heat-transfe heat-transferr coefficient aW, effective radial heat conductivity  lr, and pre pressu ssure re dro drop p D p, fo forr di diff ffer eren entt pa pack ckin ings gs in a tu tube be of 50 mm in inte tern rnal al di diam amet eter er an and d an ai airr ma mass ss flu flux x G z  1 kg m2 s1)

 ¼  ¼

under ambient conditions (from [12]) a) Glass spheres,  d P  5 mm; b) Glass spheres,  d P  10 mm; c) Raschig rings, ceramic,  d o  10 mm,  d i  6 mm,  l  11 mm; d) Rasc Raschig hig rings, rings, stain stainless less steel steel,, d o  10 mm, d i  9 mm,  l  11 mm; e) Hollow Hollow cera ceramic mic cylin cylinders ders of irregu irregular lar leng length, th, d o  9 mm,   d i  4.5 mm, l  15 mm; f) Full ceramic cylinders of irregular length,   d o  5 mm,   l  11 mm; g) Crossed corrugat corr ugated ed meta metal-pl l-plate ate pack packing ing (Sulz (Sulzer er Kata Katapak) pak),, wide chan channels nels;; h) Cros Crossed sed corr corrugat ugated ed meta metal-pla l-plate te pac packing king,, narr narrow ow chan channels nels;; i) Automotive exhaust monolith, cordierite, square channels,  d   0.9 mm

 ¼

 ¼

 ¼

 ¼

 ¼

 ¼  ¼

 ¼

 ¼

 ¼

 ¼

 ¼

 ¼  ¼

is determ ermined ined by molecu mol ecular larin diffus dif fusion ion. . Thi Thissof leads lea ds to det particularly low values the corners the channels where two reaction surfaces meet and the flow velocity drops to zero. The more acute thee en th encl clos osed ed an angl glee be betw twee een n th thee tw two o su surfa rface ces, s, th thee greater is the depletion in the corner region and thee sm th smal alle lerr is th thee co cont ntri ribu buti tion on of th thee wa wall ll su surf rfac acee to further mass/heat transfer. The efficiency of  channel monoliths of equal cross-sectional area but different shape therefore decreases in the sequence: circle, hexagon, rectangle, triangle. This is illustrated in Table 1, which gives the asympto asy mptotic tic dim dimens ension ionles lesss mas mass-tr s-trans ansfer fer and heat-transfer coefficients for tubes of different cross sections [16]. At the entrance of a tube or monolith the massmas s- and hea heatt tra transf nsfer er coe coeffici fficient entss are max maxima imall

Table 1.   Asymptotic dimensionless laminar-flow heat- or masstransfer coefficients  Nu  a wd h /  lG  or  Sh  k d h /  D   (constant wall conditions) and fanning friction factor  f  for pressure drop   D p  2 f  d h2 )vG  for ducts of different cross section [16] (h Z L / d 

 ¼

Geometry

 

 ¼

 ¼

Nu Nu

f

dh

 

2.47

13.33

 

2a= 3

2.98

14.23

 

a

3.34

15.05

 

2 3a

3.66

16.00

 

a

7.54

24.00

p   ffiffi

 ffiffi ffi p 

2a

 

12

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

under laminar conditions and decrease rapidly to the asymptotic values as the flow profile and boundary layers develop. The asymptotic value is usually established after a flow length of less than ten times the hydraulic channel diameter. The asymptotic values from Table 1 can therefore fo re be us used ed to pe perf rfor orm m es estim timat atee ca calc lcula ulatio tions ns fo forr mono mo noli lith th ca catal talys ysts ts wi with th a la larg rger er le leng ngth th-to -to--

loss coeffi coefficient cient z  an and d th thee pr pres essu sure re dr drop op eq equa uatio tion n

diameter ratio.

where

1.2.3. Compar 1.2.3. Comparison ison and Evaluation Evaluation of  Different Catalyst Forms

 f 1

 Pressure Drop and Uniform Flow Distribu tion   Con Conven ventio tional nal cat cataly alyst st pel pellet let sha shapes pes differ considerably as regards pressure loss. For equa eq uall me mean an di dime mens nsio ions ns an and d th thee sa same me vo void id fraction   e, random packings generally have a considerably higher pressure loss than regular structures, and among these corrugated structures have a higher pressure loss than monoliths with straight parallel channels. External mass and an d he heat at tr tran ansf sfer er ar aree st stro rong ngly ly co corr rrela elate ted d wi with th th thee pressure loss, which means that the benefits of 

low pressure (partly) compensated bye incr in crea ease sed d he heat atdrop and an d are mass ma ss tran tr ansf sfer er resi re sist stan ance. ce. Th The appropriate catalyst geometry for a given processs sho ces should uld the theref refore ore com combin binee the req requir uired ed mass ma ss an and d he heat at tr tran ansp sport ort wi with th th thee lo lowe west st pr pres essu sure re loss. Thee Ha Th Hage gen– n–Po Poise iseui uille lle eq equa uatio tion n (1 (1)) is us used ed to calcula cal culate te the pre pressu ssure re los losss   D p   in lam lamina inarly rly traversed trave rsed mono monolith lith chann channels els where   h   is th thee dynamic viscosity of the fluid,   L   the channel length, vG  the channel velocity and  d h the channel hydrau hydraulic lic diame diameter: ter: 32hvG =d h2

¼

1

 

2 G G =2

¼ zr  v

 

ð2Þ

where  r G  is the fluid density or by the Ergun equation D p= L 

2 2 G

¼  f  v þ f  v 1 G

2

 

3 2 p

ð3Þ

3

 ¼ 150hð1eÞ =ðe d  Þ; f   ¼ 1:75r  ð1eÞ=ðe d  Þ

The choice of a suitable catalyst form is always an optimization problem that can be specified only on ly fo forr a sp spec ecific ific pr proc oces ess. s. Ev Even en th then en,, we weig ight hting ing the target function is intricate, since small pressure losses, uniform flow through the reactor, and goo good d mas masss- and hea heat-t t-trans ransfer fer pro proper pertie tiess genera gen erally lly rep repres resent ent opp opposi osing ng req requir uireme ements nts.. The follow fol lowing ing ass assess essmen mentt sho should uld pro provid videe som somee guidelines.

D p= L 

D p= L 

ðÞ

The pressure loss of randomly packed tubes can be described either by means of a pressure

2

G

p

ð4Þ

The pre pressu ssure re dro drop p the therefo refore re dep depend endss str strong ongly ly on the void fraction  e of the packing and on the pellet diameter  d p. From the standard standard forms for packed-bed catalysts, hollow cylinders of thin wall thickness (e 0.6–0 0.6–0.8) .8) are there therefore fore preferred over spheres (e   0.37–0.4) and solid cylinders cylin ders (e 0.35). The strong dependence of the pressure loss on the void fraction requires careful packing of  catalyst beds to achieve a uniform flow profile over the bed cross section and to avoid bypass flow flo w du duee to lo loca call va varia riatio tions ns in th thee pa pack cking ing dens de nsit ity. y. An ev even en flo flow w di dist stri ribu butio tion n fr from om th thee fe feed ed pipe pi pe th thro roug ugh h th thee di dist stri ribu butio tion n ho hood od ov over er th thee whole fixed-bed or tube bundle is also of importance. In the case of a single packed bed, a uniform flow distribution can usually be obtained if the pressure drop in the fixed-bed is substa sub stantia ntially lly hig higher her tha than n in the ent entran rance ce and exi exitt hoods. This is the case if a random packing of  pelletss is used and the length of the packe pellet packed d bed exceed exc eedss the tub tubee dia diamet meter, er, sin since ce the (lo (local cal))





 

pressure drop increases with the mean me an lo loca cal l flo flow w ve velo loci city ty (E (Eqs qs..square 2, 3) 3)..ofItthe is however not the case with monolith reactors with parallel channels and laminar flow. They need special care in the design of the hoods or must be equipped with static mixers to avoid nonuniform nonun iform flow patter patterns ns [18]. A rel relate ated d pro problem blem occ occurs urs in mul multitu titubul bular ar fixed fix ed-b -bed ed re reac acto tors rs wi with th ra rand ndom om pa pack ckin ings gs if  differ dif ferent ent tub tubes es hav havee dif differe ferent nt pre pressu ssure re dro drops ps due to nonuniform packing or different batches of cataly catalysts. sts. In partic particular ular multit multitubular ubular reactors for highl highly y exoth exothermic ermic,, selecti selectivity-s vity-sensiti ensitive ve reactions require careful filling, often assisted by means of special filling devices and, if necessary, individual compensation of each tube for equal pressure loss [19].

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

Ca Cataly talyst st For Forms ms for Adi Adiaba abatic tic Op Opera eration tion Decisiv Dec isivee par parame ameter terss for adi adiaba abatic tic ope operati ration on include:

 The active catalyst surface available per unit reactor volume .   The mass and heat transfer between between the flowing gas and the active catalyst surface

13

The computation is particularly simple for straigh stra ight, t, lam laminar inarly ly tra traver versed sed cha channe nnels. ls. The asymptotic limiting value given in Table 1 for the square channel

.

¼ k  Dd    3 G

Sh

h

together toget her with

. .

 The pressure drop along bed   The The un unifo iform rmit ity y of th the e flo flow wthe thro th roug ugh h th thee re reac acto torr and thus the degree of utiliz utilization ation of the fixed bed.

aP

 ¼ d   4e h

gives  D    12 d  e

k G aP

The ma The majo jorr po port rtio ion n of th thee ac activ tivee ca cata talys lystt surface is located in the interior of the porous cataly cat alyst st str structu ucture. re. How Howeve ever, r, if it is ass assume umed d tha that, t, in the case of suf sufficie ficientl ntly y fas fastt rea reacti ctions ons,, the reaction is restricted to a thin layer underneath the external surface, surface, the amount of activ activee catalyst surface area can be taken as proportional to the specific external catalyst surface area   ap. Then The n for giv given en kin kinetic etic con condit dition ionss the eva evalua luatio tion n can be restricted to three parameters: specific extern ext ernal al sur surfac facee are areaa ap, gas–so gas–solid lid massmass-transfe transferr coefficient   k G, and the flow pressure loss   D p. The evalu evaluation ation becomes particularly particularly simple if, with a single dominant reaction, the external gas–solid mass transfer limits the reaction rate. In this case neither the reaction kinetics nor the temperature but only the external mass transfer determines the conversion. The relationships for the concentration c of a key ke y co comp mpon onen entt A an and d a si simp mple le re reac actio tion, n, A . . .   produc products, ts, are out outline lined d in [18 [18]. ]. Assu Assumin ming g mass-transfer limitation, the material balance

þ

!

for is the key component along the flow path z e

v  qq zc ¼ k  a c G

G

 

p

ð5Þ

If the reaction has a conversion of 99 %, i.e., an ex exit it conc concen entra trati tion on of 1 % of th thee inlet inlet conc concen en-trat tr atio ion, n, th thee le leng ngth th of th thee fix fixed ed be bed d   L   is, by integration,  L 

¼  ln0k :01aev ¼ 4:k 6eav G

G

p

G

G

p

ð6Þ

Then The n the req requir uired ed rea reacto ctorr vol volume ume V R wi with th th thee cross-sectional area  A  is given by: V R

_

 ¼  Z   A A ¼ 4:6 k  V a L

G

p

ð7Þ

ð8Þ

2 h

and _

2 h

 ¼ 0:38 V  Dd  e

V R

 

ð9Þ

If Equation (1) is used for the pressure loss, then for the laminarly traversed monolith channel, ne l, fr from om Equ Equat atio ions ns (6 (6)) an and d (8 (8)) th thee pr press essur uree dr drop op is given by D p

¼ 12: D28h v e 2 G

2

ð10Þ

According to Equation (9) the required fixed bed volum volumee   V R  with square monolith catalysts _  decreas and a given throughput   V    decreases es with   d h2 , while whi le the geo geometr metric ic arr arrang angeme ement nt (sm (small all bed cross section A  and long bed length  Z L, or large bed be d cr cros osss se sect ctio ion n an and d sh shor ortt be bed) d) ha hass no in influ fluen ence ce   _, on   V R. Th Thee flo flow w ve velo loci city ty   vG   for a given V  _= Ae  and the however, howeve r, depen depends ds on  A  via  v G  V  pressu pre ssure re dro drop p inc increa reases ses wit with h v2G . Th This is me mean anss th that at a mi mini nimu mum m ca cata talys lystt vo volu lume me wit with h mi mini nimu mum m pressure loss is obtained by using a shallow bed

 ¼

  A   and with a large cross-sectional small hydraulic channel diameter orarea particle diameter d h, as sh show own n in Fig Figur uree 12 12B. B. Th This is re resu sult lt ap appl plie iess in general to all catalyst forms. The main difficulties with a flat bed arrangementt res men result ult fro from m the dif difficu ficulty lty to uni unifor formly mly distrib dis tribute ute the flow flow acr across oss the large large bed diam diamete eterr and from mechanical stability problems of the housing in case of operation under vacuum or extended pressure. A possible alternative is to changee from an axial chang axial (Fig. 12 12 B) to a radial radial flow configuration (Fig. 12 C). This is discussed in Section 1.3.1. Comparing pellet packings with regular catalyst structures of equal external surface, it can be stated that the laminar flow prevailing in monoliths with straight parallel walls always

 

14

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

Figure 11.   Metallic open foam monoliths (top) and tubular square monoliths, extruded from copper (bottom left) and

aluminum (bottom right), from [21]

provides a lower pressure drop than a tortuous flow pat path h in eit either her reg regula ularr pac packin kings gs (li (like ke cro crosse ssed d corrugated corru gated packing packingss (Fig. 3 C), foams foams (Fig. 11), or random packings of pellets whatever their shape. Catalyst Forms for Nona Catalyst Nonadiaba diabatic tic Oper Operaa tion   Wit With h non nonadi adiaba abatic tic rea reacti ction on con control trol,, heat must be transported laterally through the fixed-b fixe d-bed ed bet between ween the cat cataly alyst st and the hea heattexchange surfaces. Despite the fact that radial

heat tra heat transp nsport ort tak takes es pla place ce mai mainly nly by con convec vectio tion, n, it is formally described by an effective radial thermal conductivity   lR  transverse to the flow direct dir ection ion and a wal walll hea heat-t t-trans ransfer fer coe coeffici fficient ent aW, as explained in Section 1.2.1. Heat transport in packed tubes has been investigated and is discussed in detail in [8–12, 14]. In [10] it is shown that a plug-flow model with axial dispersion, using adequate  lR and  a W values, is usually able to describe even strongly exothermic reactions in catalyst-filled

Figure 12.  Main design concepts for adiabatic reactors

A) Adiabatic packed-bed reactor; B) Disk reactor; C) Radial-flow reactor

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

15

tubes (with radially nonuniform flow) with reasonable accuracy. Only close to runaway conditions does the more detailed model with an additional radial coordinate predict a somewhat earlier onset of runaway. However, considering the uncertainties in determination of transport and kinetic parameters, these differences can often be neglected.

arrange short packing sections in series, each section being displaced by 90 . The heat-transport parameters in Figure 10 were determined for structures arranged in this way.

A comparison ofwhere different shapes is given in Figure 10, the catalyst two heat-transfer parameters  l r  and  a w  plus the external catalyst surface   ap, th thee be bed d vo void id fr frac acti tion on   e, and the pressure drop   D p/L  are   are given for a selection of  different random and regular catalyst packings in a tube of 50 mm internal internal diameter and an air 2 1 mass flow velocity of  G  G z  1 kg m s . The dimensions of the packings were chosen so that their the ir spe specific cific ext extern ernal al sur surfac facee are areaa   ap   is ca ca.. 2 3 500 m  /m . Un Unde derr th thes esee co cond nditi ition onss ho holl llow ow,, thin-w thi n-walle alled d cyl cylind inders ers hav havee cle clear ar adv advant antage agess over other packing forms, exhibiting the lowest pressure loss and the highest thermal conductivi ti vity ty.. On Only ly as re rega gard rdss wa wall ll he heat at tr tran ansf sfer er ar aree th they ey inferior to full cylinders or to spheres, as was already mentioned in connection with Figure 8. However, good wall heat transfer is apparently less decisive from a reaction engineering viewpoint poi nt tha than n goo good d rad radial ial the therma rmall con conduc ductiv tivity ity,, sinc si ncee th thee fo form rmer er ca can n be co comp mpen ensa sate ted d by an appropriate wall temperature profile, whereas the radial thermal conductivity directly influences the unifor uniformity mity of the reaction conditions conditions over the tube cross section. Monolith structures can have very high specific surface areas combined with a very low

rea reacto ctors. rs. The They y inc includ ludee cer cerami amiccmonoliths and met metalli allicc open foams and square-channel with high hig h wal walll the therma rmall con conduc ductiv tivity ity.. Figu Figure re 11 sho shows ws respective substrates which can be catalytically coated coa ted lik likee com common mon aut automo omotiv tivee mon monoli oliths ths [21 [21]. ]. Open foams combine a high specific surface area ar ea wi with th a hi high gh vo void id fr frac acti tion on (e   >  80 %) %),, i.e i.e., ., a low thermal mass. Due to the irregular, vortical flow gas–solid mass and heat transfer is higher than tha n in squ square are-ch -chann annel el mon monoli oliths ths wit with h the their ir laminar flow structure. Under adiabatic operation this allows for rapid heatup and a smaller catalyst cataly st volum volumee compa compared red to squar square-chan e-channel nel monoliths. Considering a ‘‘merit index’’ which rates rat es ext extern ernal al mas masss tra transf nsfer er ver versus sus pre pressu ssure re drop [21], open foams are, however, slightly inferi inf erior or to squ squareare-cha channe nnell mon monoli oliths ths,, bec becaus ausee of  their higher pressu pressure re drop. Unlike Unl ike squ square are cha channe nnell mon monolit oliths, hs, ope open n foa foam m structur stru ctures es als also o pro provid videe a goo good d lat latera erall con convec vectiv tivee transport and mixing, which makes them suitable as packings in multitubular reactors with heat transfer through the tube walls. But their somewhat delicate structure and their relatively high hi gh co cost stss ha have ve so fa farr pr prev even ented ted ex exte tens nsive ive us usee as catalyst supports. As indicated in Figure 10, standard ceramic

pressure loss, with straight, parallel all el cha channe nnels, ls, but such suc hmonoliths as tho those se use used d for automo aut omobil bilee exhaust control, have only very poor radial heat transport trans port proper properties. ties. Cross Crossed ed corru corrugated gated struc struc-tures are considerably more favorable for isother th erma mall re reac acti tion on co cont ntro rol. l. Du Duee to th thei eirr ra radi dial al flow compon com ponent ent,, the they y off offer er a hig high h rad radial ial the thermal rmal conduc con ductiv tivity ity whi which ch is alm almost ost ind indepe epende ndent nt of  the specific specific surfac surfacee area; the latter latter can be varied varied over ov er a wi wide de ra rang ngee by th thee ge geom omet etry ry of th thee corrugations. In crossed corrugated structures, convective lateral heat transport occurs only in one plane perpendicular to the main flow direction. This mean me anss th that at th thee flo flow w be beha havi vior or in tu tube bess of ci circ rcul ular ar cros cr osss se secti ction on is ra rath ther er no nonu nunif nifor orm m ov over er th thee circ ci rcum umfe fere renc nce. e. It is th ther eref efor oree ad advi visa sable ble to

square-channel dowall-cooled not provide sufficient radial heatmonoliths transferr for transfe wall-coole d fixedbed reactors. This is a result of the missing lateral convection and the low thermal conductivity of the usually applied ceramics. By using monoliths with relatively thick walls and high wall heat conductivity, this deficit can be transformed into an advantage [20, 21]. In this case the catalyst must be coated onto the monolith walls, and the monolith tube replaces a conventional catalyst-filled tube of a multitubular reacto ac tor. r. Th This is me mean anss th that at su such ch a re reac acto torr op oper erate atess as a wall-coated tubular reactor rather than as a fixedfixe d-be bed d re reac acto tor. r. As sh show own n in [2 [21] 1] th thee ef effe fecti ctive ve radial conductivity of an aluminum monolith tube as shown shown in Figure Figure 11 11 (bottom) (bottom) with a void 1 1 frac fr acti tion on of 80 80% % is 30 W m K  , wh whic ich h is on onee

 ¼

 Novel Catalyst Structures and Wall-Coated  Tubes   Re Rece cent ntly, ly, se seve vera rall no nove vell ca cata talys lystt structur stru ctures es hav havee bee been n pro propos posed ed for fixe fixed-b d-bed ed

 

16

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

order of magnitude higher than the effective radial conductivity of any packing (Fig. 10). Ceramics with high heat conductivity like aluminum nitride (AlN) or silicon carbide (SiC) would also provide a high radial conductivity, but monolith tubes of these materials are still considered too crack-sensitive for larger scale operation.

reversal revers al (se (seee Sec Sectio tion n 1.7 1.7), ), spe specia ciall mea measur sures es must be taken to stabilize the packing. If fixed bedss com bed compos posed ed of mon monoli olith th cata catalys lystt sec sectio tions ns are used, any flow direction can be applied. As discussed in Section 1.2.3.2, the requirement for a low pressure loss leads to a fixed bed of large diameter and low height (Fig. 12 B). Such an arrangement (‘‘disk concept’’) is used

It has also proposed sections of  such metal or been ceramic foams to or use monolith tubes as pac packing kingss in con conven ventio tional nal mul multitu titubul bular ar rea reacto ctorr tube tu bess [2 [20, 0, 21 21]. ]. Ho Howe weve ver, r, co cons nsid ider erin ing g th thee manufacturing tolerances of reactor tubes, such packin pac kings gs nee need d to hav havee an ext extern ernal al dia diamet meter er which is 1–2 mm smaller than the mean inner tubee dia tub diamet meter er to pre preven ventt int interl erlock ocking ing dur during ing operat ope ration ion and rep replac lacemen ement. t. How Howeve ever, r, a gas gas-filled gap of about 1 mm at the wall almost eliminates the advantage of the high monolith conductivity. Wall-coated tubular reactors have been considered as an alternative to catalyst-filled tubes for hig highly hly exo exothe thermi rmicc or end endoth otherm ermic ic rea reacti ctions ons,, since such reactions are generally heat transfer limited in catalyst-filled tubes. Due to the direct contact of the thin catalyst layer with the tube wall, the cataly catalyst st temperature will stay close to the wall temperature, and thermal runaway will be prevented even for highly exothermic reactions. Mass transfer to the catalyst-coated tube wall wa ll is di disc scus usse sed d fo forr em empt pty y tu tube bess an and d tu tube bess fill filled ed with inert particles in [17]. From the practical point of operation the main drawback of wallcoat co ated ed tu tube bess is th thee fa fact ct th that at re repl plac acem emen entt of  deactivated deac tivated catalyst requi requires res eithe eitherr chan changing ging

particularly when very short times, follow fol lowed ed by dir direct ect que quench nching ing ofresidence the rea reactio ction, n, are required. Examples include ammonia oxidation in ni nitri tricc ac acid id pr prod oduc ucti tion on an and d ox oxid idati ative ve de dehy hydr droogenation on silver catalysts (e.g., synthesis of  formaldehyd forma ldehydee by dehyd dehydrogen rogenation ation of metha metha-nol). In the first case the ‘‘fixed bed’’ consists of sev severa erall layers layers of pla platin tinum um wire wire gauze, gauze, and and,, in the second case, of a porous silver layer several cent ce ntim imete eters rs in he heig ight ht.. Be Bed d di diam amet eters ers up to several meters are common. Flat-bed reactors with a residence time on the order of milliseconds have recently achieved considerable attention for partial oxidation of hydrocarbons to synthesis gas, olefins, or oxygenates under the heading ‘‘short contact time reactors’’ [22]. On account of the difficulties involved with obtaining uniform flow as well as for structural reasons, the disk concept is limited to relatively small catalyst volumes and low synthesis pressure. For other conditions the radial-flow concept (Fig. 12C) can be an alternative. alternative. The catalyst ly st pa pack ckin ing g is ac acco commo mmoda dated ted in th thee sp spac acee betwee bet ween n two con concen centric tric scr screen een or per perfor forate ated d plate cylinders, and is traversed in radial direction, tio n, ei eith ther er fr from om th thee in insi side de to th thee ou outsi tside de or fr from om the outside to the inside. This design is particu-

the tube walls or in in situ the tube thecleaning reactor. and re-coating of 

larly suitable for largepressure, catalyst volumes and for operation at elevated since at moderatee re at reac acto torr di diam amet eters ers th thee ca cata taly lyst st vo volu lume me ca can n be varied over a wide range by altering the reactor leng le ngth th,, wi with thou outt af affe fect ctin ing g th thee flow flow-th -throu rough gh length of the packing and hence the pressure drop. Compared to axial flow reactors, the short flow length allows small catalyst catalyst particl particles es with reduce red uced d mas masss tra transf nsfer er res resista istance nce to be use used d withou wit houtt the dra drawba wback ck of exc excess essive ive pre pressu ssure re drop. The flow design at the top of the backed bed be d re requ quire iress sp spec ecia iall atte attent ntion ion in or orde derr to av avoi oid da bypass byp ass.. Mor Moree det detaile ailed d inf inform ormati ation on on rad radial ial flow flo w re reac acto tors rs is pr prov ovide ided d in  Radial-Flow Packed-Bed Reactors. Monoli Mon olith th cat cataly alysts sts with str straig aight, ht, par paralle allell channels are particularly well suited for axial

1.3. Adiab Adiabatic atic Fixed-Bed Fixed-Bed Reactors Reactors 1.3.1.. Axial and Radial 1.3.1 Radial Flow Reactors Reactors

Adiabatic fixed-bed reactors constitute the oldest fixed-bed reactor configuration. In the simplest case they consist of a cylindrical jacket in which the catalyst is loosely packed on a screen supp su pport ort an and d is tr trav avers ersed ed in ax axia iall di dire rect ctio ion n (Fig (F ig.. 12 A) A).. To av avoi oid d ca catal talys ystt ab abra rasi sion on by pa parti rtial al fluidization, fluidiz ation, rando random m catal catalyst yst packi packings ngs should alwa al ways ys be tr trav aver erse sed d fr from om to top p to bo bott ttom om.. If th this is is nott po no poss ssib ible le,, e. e.g. g.,, be beca caus usee of pe peri riod odic ic flo flow w

 !

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

17

Figure 13.  Left: Reaction chamber for conversion of nitrogen oxides from power station flue gases by selective catalytic

reduction (SCR); Right: detail of the monolith catalyst used [23]

flow adiabatic reactors, since they allow for a low pressure drop and provide a high specific outer surface area. Since monolith catalysts are usually produced with a rectangular cross section, the fixed bed is constructed by arranging thes th esee in indi divid vidua uall mo mono noli lith thss in ro rows ws in th thee fo form rm of  a large ‘‘box’’. Conventional DENOX reactors forr re fo remo movi ving ng NO x fr from om po powe werr st stati ation on flu fluee ga gase sess aree th ar thus us de desig signe ned d as re rect ctan angul gular ar ch cham ambe bers rs (Fig. 13, left). The monolith catalyst is often arranged arran ged in the form of severa severall layers in series series,, and the spaces between the individual layers permit cross-mixing, so that the influence of 

outlet pipes of which often induce a swirl component in the flow, which leads to a nonuniform flow through the monolith [18]. Purel Pur ely y ad adia iaba bati ticc fix fixed ed-b -bed ed re reac acto tors rs ar aree used mainly for reactions reactions with a small heat of  reaction. A typical example is gas purification, in which small amounts of noxious components nen ts are con conver verted ted.. The cha chambe mbers rs use used d to remove rem ove NO x   from from pow power er sta statio tion n flue gas gases es (Fig. 13), with a catal catalyst yst volume of more than 3 1000 m , are the largest industrial adiabatic reactors, and the exhaust catalyst for internal combustion engines, with a catalyst volume of 

nonu nonuniform flow and any possible possib lempen local blockagee ofniform ag mono mo noli lith th ch chan anne nels ls ca can n be co comp ensat sated ed to somee ext som extent ent.. Squ Square are-ch -chann annel el mon monoli olith th cat cataalysts, lys ts, coa coated ted par paralle allel-p l-plate late cat cataly alysts sts or cor corrug rugatated cer cerami amicc mat matrix rix com compos positeite-typ typee cat catalys alysts ts (Fig. 13, right) are in use [23]. Refe Re fere renc ncee was ma made de in Se Sect ctio ion n 1. 1.2. 2.3. 3.1 1 to th thee importance of uniform flow into and through adiabatic fixed-bed reactors. This is not easy to achiev ach ieve, e, par particu ticular larly ly wit with h low pre pressu ssure re los losss monolith reactors, and requires careful design of th thee in inflo flow w ho hood od.. On ac acco coun untt of th thee lo low w pressure loss, unfavorable flow conditions in the outflow hood may also affect the flow behavio ha viorr th thro roug ugh h th thee ca catal talys ystt be bed. d. Th This is is of  particular importance for the design and operation ti on of au auto tomo moti tive ve ex exha haust ust ca cata taly lysts sts,, th thee in inle lett or

ca ca. .emic 1 L, thee in th smal sm alle lest st.. Ty Typi pica calel ap appl plic icat atio ions ns in the e chem ch ical al indu dust stry ry incl in clud ude meth me than anat atio ion n th of  traces of CO and CO2  in NH3  synthesis gas, as wel welll as hyd hydrog rogena enation tion of sma small ll amo amount untss of uns unsatu aturat rated ed com compou pounds nds in hyd hydroc rocarb arbon on streams.

1.3.2. Multis Multistage tage Reactors Reactors with Interstage Heat Transfer

In the majority of fixed-bed reactors for industrial synthesis reactions, direct or indirect supply or removal of heat in the catalyst bed is necessary to adapt the temperature profile over the rea reacto ctorr len length gth to the requirem requirement entss of an optimal reaction pathway.

 

18

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

Figure Fig ure 14.   A) Adia Adiabati baticc reac reactor tor withinterstag withinterstagee gas feed feed;; B) Multi Multibed bed adia adiabaticfixed-be baticfixed-bed d reac reactor tor with inter interstag stagee heat heatexch exchange ange

A par particu ticular larly ly sim simple ple for form m inv involv olves es inj inject ecting ing hot or cold gas between the stages of an adiabati ba ticc mu mult ltis ista tage ge re reac acto torr (F (Fig ig.. 14 A) A).. It is ap appli plied ed if the composition change by injection has a posi po siti tive ve ef effe fect ct on th thee de desir sired ed re reac acti tion on.. To achieve uniform composition and temperature in the subsequent catalyst stage, the injection is usually situated in an inert section of the bed which whi ch pro provid vides es for the nec necess essary ary lat latera erall mix mixing ing.. Alternatively Altern atively,, inters interstage tage heat exch exchangers angers are used, through which the required or released hea eatt of re reaact ctio ion n is sup upp pli lied ed or re remo mov ved (Fig (F ig.. 14 14B) B).. A mo more re st strin ringe gent nt co cont ntro roll of th thee

reaction reacti on tem temper peratur aturee is pos possib sible le if the hea heattexchange surface is integrated in the fixed bed. This Th is ty type pe of re reac acto torr is co cons nsid idere ered d in Se Sect ctio ion n 1. 1.4. 4. Adiabatic multistage fixed-bed reactors with interstage cooling or heating are used primarily if the reaction proceeds selectively to a single product but is limited by the equilibrium conditions. Intermediate cooling or heating is used to change the gas temperature towards higher equilib equ ilibrium rium con conver versio sion. n. In exo exothe thermic rmic rea reacctions the equilibrium conversion to the target product produ ct decre decreases ases with increa increasing sing tempe temperature rature (Fig. 15, left), while the opposite is true for an

Figure 15.   Three-stage Three-stage fixed-bed reactor for equi equilibri libriumum-limite limited d reac reactions tions with inte interstag rstagee heat remo removal/ val/supp supply ly in the

temperature conversion plane Left: Exothermic reaction with equilibrium and maximum reaction rate curve; Right: Endothermic reaction with equilibrium and maximum temperature limit. In both cases the feasible temperature–conversion range is shaded

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

19

endothermic reaction (Fig. 15, right); the reaction rate always increases with temperature. In thee ca th case se of an ex exot othe herm rmic ic re reac actio tion n th thee ma maxi ximu mum m reaction temperature is therefore bound by the equilibrium conversion curve and the minimum temperature by the slowest acceptable kinetics. For endothermic reactions the thermal stability of the reactor material or of the catalyst will provide a maximum temperature limit, whereas the minimum temperature is either limited by the equilibrium curve or also by the slowest accept acc eptabl ablee kin kineti etics. cs. Thi Thiss defi defines nes the sha shaded ded feasibility regions of Figure 15 in the temperature–conversion plane, in which any equilibrium limited reactor should operate. The optimal operation line in the case of an endothermic reaction is the maximum temperature tu re lim limit. it. In th thee ca case se of an ex exot othe herm rmic ic re reac actio tion na trajec tra jector tory y of max maximu imum m rea reactio ction n rat ratee exi exists sts in the conversion conv ersion/tempe /temperature rature plane, i.e., the maximum reaction rate curve of Figure 15, left. In the case of an adiabatic reaction the temperatur per aturee cha change ngess lin linear early ly wit with h the ach achiev ieved ed conversion   D X  according   according to the equation R

G PG

feed and product of stage 1 [24]

0

 ¼ ðr Dhc Þc D X 

DT ad ad

Figure 16.  Layout of aintegrated two-stage heat ammonia converter with radial flow stages and exchanger between

 

ð11Þ

The adi adiaba abatic tic rea reacti ction on pat path h the theref refore ore lies on a straight line of gradient   DT  / D X , which is positive ti ve fo forr an ex exoth other ermi micc re reac actio tion n (Fi (Fig. g. 15 15,, le left ft)) an and d negative for an endothermic reaction (Fig. 15, right) rig ht).. A rea reason sonabl ablee rea reacti ction on pat pathwa hway y for a multistage adiabatic reaction can therefore be composed of the straight lines of an adiabatic reaction section, followed by the vertical lines

layout of an Uhde ammonia synthesis reactor, designed on this basis [24]. For structural reasons the heat exchanger for the inflow is often incorporated in the pressure casing. In the specific design of Figure 16 the feed gas is preheated in a multitubular heat exchanger in the upper center of the reactor, from where it enters the first reaction stage radially from the outside. This flow scheme protects the pressure-bearing

of indi in dire rect ctThe inte in term rmed edia iate te cool co olin ing g or he heat atin ing g (Fig. 15). optimal reaction pathway with thee sm th small alles estt re requ quir ired ed ca cata taly lyst st vo volu lume me re resu sults lts fo forr an exothermic reaction when the trajectory follows lo ws,, in a la larg rgee nu numb mber er of sm smal alll st step eps, s, th thee li line ne of  maximu max imum m rea reacti ction on rat rate. e. For an end endoth otherm ermic ic reaction the steps must stay clear of the maximum temperature limit. In practice, the apparatus and equipment expenditure involved in a larg la rgee nu numb mber er of sta stage gess mu must st be ba bala lanc nced ed ag again ainst st the sav saving ingss in cata catalys lyst. t. Con Conven ventio tional nal mul multist tistage age reacto rea ctors rs for this this clas classs of reacti reaction on theref therefore ore oft often en have three to five stages, leading to reaction paths as shown in Figure 15. If large reactor volumes and/or high synthesis pressure are required, also radial-flow multistage reactors are used. Figure 16 shows the

reactor wall against highisoutlet temperature of the first stage. Thethe outlet then cooled down by the feed in the heat exchanger and is subsequently fed from the outside into the second radial-flow stage, from where it leaves the reactor.. A thi tor third rd rad radial ial flow sta stage ge can be con contai tained ned in a separate radial flow reactor. This design allows for larger catalyst volumes of the second and third stages, which is usually required because cau se of the dec decrea reasing sing tem tempera peratur turee and the approach to equilibrium. The capacity of such a reactor can be increased by expanding the height of the radial-flow radial-flow stage stagess withou withoutt chan changging in g th thee re reac acto torr di diam amete eterr an and d he henc ncee th thee wa wall ll thickness. Dehydr Deh ydroge ogenati nation on rea reacti ctions ons lik likee eth ethylb ylbenenzene dehydrogenation for styrene synthesis are

 

20

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

Summarizi Summar izing, ng, mul multist tistage age adi adiaba abatic tic fixe fixeddbed be d re reac acto tors rs ca can n be ap appl plie ied d wh wher eree a ce cert rtain ain varia va riatio tion n of th thee re reac actio tion n te temp mpera eratu ture re ov over er the reactor length has no adve adverse rse effects on the desir de sired ed pr prod oduc uctt yi yiel eld. d. Th This is is in pa part rticu icula larr the case for equilibrium-limite equilibrium-limited d reacti reactions ons with one mai main n pro produc duct. t. Alt Althou hough gh a mul multis tistag tagee rea reactor ctor with a small number of stages is not able to follow maximum reaction rate path closely (see (s ee th theethe exam ex ample pless in Fi Fig. g. 15 15), ), th the e co comp mpar arat ative ively ly simp si mple le set etup up an and d se sev ver eral al de desi sign gn opt ptio ion ns (axial and radial flow) provide supporting arguments. In particular in the following cases a multistage arrangement should be considered for struc structural, tural, kinetic, or econ economic omic reasons:

Figure 17.   Two-stage radial flow reactor scheme of the

SMART process [25] for ethylbenzene dehydrogenation. Catalyst (1) is the dehydrogenation catalyst and catalyst (2) the hydrogen combustion catalyst

pre presen sently tly oft often en car carried ried out in a seq sequen uence ce of  radial rad ial-flo -flow w rea reacto ctors rs with int inters erstag tagee hea heatin ting g thro th roug ugh h th thee ad addi ditio tion n of su supe perh rhea eated ted st stea eam. m. Since minimal synthesis pressure is beneficial to overcome equilibrium limitations, a radialflow design can be favorable because of its low pressure drop. In addition, heat input between the sta stages ges by add additio ition n of sup superh erheat eated ed ste steam am reduces reduc es the product partia partiall press pressure ure furth further. er. In the UOP/Lummus design of Figure 17, called the SMART process [25], superheated stea st eam m an and d ai airr ar aree ad adde ded d to th thee fe feed ed of th thee se seco cond nd stage. A first layer of noble metal oxidation catalyst (2)geoxidizes the hydrogen produced in the first sta stage prefer pre ferent entiall ially, y, bef before ore the reactan rea ctants ts enter the second dehydrogenation stage (catalyst 1). The advantages of this design are twofold: hydrogen combustion provides additional heat for the second dehydrogenation stage and thee co th cons nsump umptio tion n of hy hydr drog ogen en in incr crea ease sess th thee equilibrium equil ibrium conve conversion rsion for the dehy dehydrogen drogenation ation step. Care must be taken that the oxygen concent ce ntra ratio tion n to th thee se seco cond nd st stag agee al alwa ways ys st stay ayss below the flammability limit and that the oxygen ge n is to tota tall lly y co cons nsum umed ed in th thee ox oxid idat atio ion n ca cata taly lyst st layer,, becau layer because se most dehy dehydroge drogenation nation catalysts aree de ar deac activ tivat ated ed by ox oxyg ygen en.. As in indi dica cate ted d in Figure 17, the two catalysts (1) and (2) of the second stage must be separated by concentric screens.

1. If, in the case of lar large ge single-tr single-train ain plants, plants, subdivision into several individual units is nec eceess ssar ary y fo forr re reaaso son ns of tr tran ansp spor ortt or construction. 2. If a cat cataly alyst st mus mustt be replaced replaced in ind indivi ividua duall stag st ages es at di diffe ffere rent nt ti time mess on ac acco coun untt of lo loca cally lly different catalyst deactivation. 3. If ste stepwi pwise se add additi ition on of rea reacta ctant nt has kinetic kinetic advantages compared to total addition to the feed. In a combination of interstage gas feed and an d in inte terst rstag agee he heat at ex exch chan ange ge a su suit itab ably ly design des igned ed hea heatt exc exchan hanger ger ens ensures ures uni unifor form m distribution and mixing of the side feed with the reaction gas stream. 4. If, in the case of equ equilib ilibriu rium-li m-limit mited ed rea reacctions, tio ns, a limi limitin ting g pro produc ductt is ext extract racted ed bet betwee ween n the stages. In addition to the example of  Figure 17, intermediate absorption of SO3 before the last stage of SO3  synthesis is a example. 5. classic With rea reactio ction n tem temper peratu atures res abo above ve 300 C, intermediate cooling of the reactant in an exte ex tern rnal al he heat at ex exch chan ange gerr ca can n st stil illl be pe perrformed directly with boiling water, whereas in a multitubular fixed-bed reactor a hightemperature heat-transfer medium must be used as coolant (see Section 1.4).

1.4. Fixed Fixed-Bed -Bed Reactors Reactors with Integrated Heat Exchange The development of reactors in which the heatexchan exc hange ge sur surface facess are int integr egrate ated d in the fixe fixed d bed occurred in parallel with the development of  multistage adiabatic reactors with intermediate

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

21

heating or cooling. The main aim is to supply or remove the heat of reaction as close as possible to the reaction site. Earlier concepts tried to provide reaction conditions which are as isothermall as possib therma possible le (‘ (‘‘isoth ‘isothermal’ ermal’’’ fixed-b fixed-bed ed reactors). A more recent goal is to provide and adjust an optimal temperature profile for the given reaction. In case of an exothermic equi-

alternatives. Design (A) can be used for exothermic the rmic and end endoth otherm ermic ic rea reacti ctions ons,, whe where re a liqui liq uid d he heat at ca carr rrie ierr is ci circ rcul ulate ated, d, us usua uall lly y in cross-cocur crosscocurrent rent or crosscross-count countercurr ercurrent ent flow around the tube bundle. In the case of endothermic reactions hot gases could also be used as heat carrie carriers. rs. The spec specific ific adva advantages ntages/disad /disadvanvantages tag es of the res respec pective tive heat car carrie riers rs are dis-

librium-limited reaction this implies following thee ma th maxi ximu mum m re reac actio tion n ra rate te curv cu rve e of Fi Figu gure re 15 as closely close ly as possib possible. le. The design of fixed-bed reactors with integrated heat exchange can be divided into two groups: Either the heat exchange surfaces are immersed in the fixed bed of catalyst, or the catalyst is placed inside of channels, usually in tubes, tub es, whi which ch are sur surrou rounde nded d by the hea heat-t t-trans ransfer fer mediu me dium. m. A se seco cond nd di dist stin inct ctio ion n ca can n be ma made de wit with h respect to the heat-transfer medium, the main option opt ionss bei being ng non nonboi boilin ling g liq liquid uids, s, boi boiling ling liquids, and gases (see Section 1.4.2).

cussed inongly Section 1.4.2. For str strong ly exo exother thermic mic rea reacti ctions ons at ele elevat vated ed temperatures usually a molten salt is used as heat carrier. A standard salt-bath reactor design is shown in Figure 18 B. A circulation pump prov pr ovid ides es a hi high gh ci circ rcul ulati ation on ra rati tio o of th thee liq liqui uid d sa salt lt through the multitubular reactor to ensure conditions dit ions as iso isothe therma rmall as pos possibl sible. e. A sma smalle llerr portion of the flow is directed over a steam generator, where the heat of reaction is used to generate steam while this part of the flow is cooled coo led dow down. n. The tem temper peratu ature re of the cir circul culati ating ng salt bath is thus controlled by the flow over the steam generator. Typical large-scale multitubularr rea ula reacto ctors rs con contai tain n sev severa erall ten tenss of tho thousa usands nds of  tubes of inner diameter between 15 and 50 mm, with wit h tu tube be le leng ngth thss be betw twee een n 0. 0.5 5 an and d 6 m an and d reac re acto torr di diam amet eter erss up to 8 m. Th Thee ma maxi ximu mum m dimensions are usually limited by the mode of  transportation from the construction site to the plan pl ant. t. Ty Typi pica call ex exam ampl ples es ar aree th thee sy synt nthe hesi siss reactors for acrolein/acrylic acid, for phthalic

1.4.1.. Heat-E 1.4.1 Heat-Exchan xchange ge Concepts

The most widely used design concept is the multitu mul titubul bular ar rea reacto ctorr in whi which ch the cat cataly alyst st is filled fill ed in a bu bund ndle le of st stra raigh ightt pa para ralle llell tu tube bess that are surrounded by a heat-transfer medium. Fig igu ure 18 shows thre reee common design

Figure 18.  Design options for multitubular fixed-bed reactors

A) Multi Multitubu tubular lar fixed fixed-bedreactorfor -bedreactorfor exot exotherm hermic ic or end endothe othermicreaction rmicreactionss with liqu liquid id cool cooling ing or heat heating; ing;B) B) Multi Multitubu tubular lar fixed fixed-bed reac reactor tor for exot exotherm hermic ic high high-temp -temperat erature ure reac reactionswith tionswith molte molten n salt saltcool cooling ing and stea steam m gene generatio ration; n; C) Multi Multitubu tubular larfixed fixed-bed -bed reactor for exothermic reactions with evaporation cooling

 

22

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

anhy anhydride, dride, and for maleic anhyd anhydride ride by partia partiall oxidation of the respective hydrocarbons. For exothermic reactions with coolant temperatures up to 300 C a boiling liquid, preferably water, can be used as coolant medium. A standard design is shown in Figure 18 C, where thee bo th boili iling ng wa wate terr su surro rroun unds ds th thee tu tube bes, s, an and d st steam eam and the coolant circulate by natural convection.

main advantage of all such concepts over the multitubular reactor with boiling-water cooling (Fig. 18 C) is the fact that the reactor casing no longerr need longe needss to withstand the boilin boiling g pressu pressure. re. This allows boiling-water cooling to be applied up to ab abou outt 300 300 C (c (ca. a. 100 ba barr bo boili iling ng pr pres essu sure re)) and the reactor volume to be extended considerably.

The is separated separat in drum and led to theliquid bottom bott om of the reac red eactor tora by bsteam y a dow down n pip pipe. e. The Th e adva ad vant ntag ages es ov over er Fig Figur uree 18 B ar aree th thee mu much ch simpler design concept and fail-safe operation in case of an emergen emergency cy shutdown shutdown,, as discus discussed sed in Sec Sectio tion n 2.4 2.4.. A pos possib sible le dis disadv advant antage age can ari arise se from the fact that the reactor shell must withstand the steam pressure, which limits the maximum im um re react actor or di diam amet eter. er. Th Thes esee re reac acto tors rs ar aree therefore often built longer (up to 15 m) with smalle sma llerr tot total al dia diamet meter. er. Typ Typica icall exa exampl mples es are  ethylene oxide synthesis (at 220–280 C coo coolan lantt temper tem peratu ature) re) and met methan hanol ol syn synthes thesis is (ca (ca.. 250 C coolant temperature). Compared to the standard multitubular reactor concepts of Figure 18, more recent design concep con cepts ts emb embed ed the hea heat-t t-trans ransfer fer sur surfac faces es in the packed bed. These concepts have so far been applie app lied d exc exclus lusive ively ly to exo exothe thermic rmic rea reacti ctions ons,, preferably with water-evaporation cooling. A

One of the first design concepts wastubes Linde’s isothermal reactor with coil-wound for methanol synthesis (Fig. 19) [26]. In this case cool co olin ing g is pr prov ovid ided ed by a se sequ quen ence ce of co conc ncen entri tric, c, spiral-wound cooling tubes which are embedded in the catalyst bed. Figure 19 (left) shows a schematic of a possible design [27]. Inert packing and active catalyst are placed between the cool co olin ing g co coils ils an and d ca can n be ea easi sily ly lo load aded ed an and d emptied through manholes. The coolant circulates lat es by nat natura urall con convec vectio tion n dow down n the cen centra trall down do wn pi pipe pe an and d th then en th thro roug ugh h th thee co cool olin ing g co coil ils, s, in which partial evaporation takes place, back to the steam drum. Care must be taken that all coiled tubes have about the same total length in orde or derr to ac achie hieve ve si simi mila larr flo flow w ve velo loci city ty an and d st stea eam m content. This means that the outer coils must have a steeper slope than the inner ones. Besides methanol synthesis, the Linde isothermal reactor has been applied to a number of 

Figure 19.  Linde’s isothermal reactor with coil-wound tubes

Left: reactor schematic; Right: coil-wound cooling tube internals (bottom) and complete reactor (top) [27]

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

23

exothermic, synth other, preferably exothermic, synthesis esis reactions inclu including ding hydro hydrogenat genation, ion, partia partiall oxida oxida-tion, and Fischer–Tropsch synthesis. Compared to multitubular reactors of Figure 18, increased heat transfer on the catalyst side, resulting in a moree iso mor isothe therma rmall beh behavi avior, or, is cla claime imed d bec becaus ausee of  the cross-flow of the reacting gas over the tube bundle. Due to the flexible arrangement of the

the flow direction here is reversed, which requires qui res cir circul culati ation on pum pumps ps with spe specifi cificc flow control. As a third concept, the so-called Thermoplate design of ITS Reaktortechnik, is shown in Figur Fig uree 21 [2 [29] 9].. The re resp spec ecti tive ve re reac acto tors rs ar aree based upon so-called thermoplate heat exchangers ge rs.. A th ther ermo mopl plat atee co cons nsis ists ts of tw two o me meta tall sh shee eets ts

coiled tubes, large and temporal temperature gradients canlocal be tolerated, and thus fast startu sta rtup p and shu shutdow tdown n pro proced cedure uress are fac facilit ilitate ated. d. The cooling intensity over the height of the packed bed can be adjusted according to the heat-re hea t-relea lease se pro profile, file,by by var varyin ying g the axi axial al spa spacin cing g of the coiled tubes [27]. As sec second ond con concep ceptt wit with h hea heat-t t-tran ransfe sferr sur sur-face fa cess em embe bedd dded ed in th thee pa pack ckin ing g is TO TOYO YO’s ’s MRF-Z MRFZ rad radialial-flow flow rea reacto ctorr con concep ceptt (Fig (Fig.. 20) [28]. It extends the radial-flow concept by incorporating heat-exchange tubes in the annular pack pa cked ed be bed. d. In th thee sp spec ecific ific de desi sign gn do doub uble le-walled so-called Field tubes are used in which water fed into the central tube partly evaporates in the outer annulus. Contrary to conventional Field tubes, where the water in the center tube flows down downwards wards while the boilin boiling g steam/ steam/water water mixture flows upwards by natural convection,

welded together along theirAfter edgeswelding, and pointwelded across their surface. the space between the plates is hydraulically expanded by injecting a liquid between the metal sheets. This opens coolant channels, as shown schematically in Figure 21 A. A thermoplate fixedfixe d-be bed d re reac acto torr is co comp mpos osed ed of a st stac ack k of  vertica ver ticall the thermo rmopla plates tes whi which ch for form m the hea heat-tr t-trans ans-fer surfaces while the catalyst is filled in between the plates. Coolant distribution and collecting tubes are welded to the bottom and the top to p of th thee th ther ermo mopl plat ates es,, as sh show own n in Fi Figu gure re 21 B and C. Wat Waterer-eva evapor poratio ation n coo cooling lingis is use used d for the reac re acto tors rs sh show own n in Fig Figure ure 21 21.. The de desig sign n is claimed to be less expensive than multitubular reac re acto tors, rs, si sinc ncee on only ly a co comp mpar arat ativ ivel ely y sm small all nu nummberr of th be ther ermo mopl plate atess is re requ quir ired ed,, an and d th thei eirr cu cutt ttin ing g and welding is fully automated.

1.4.2. Heat-T Heat-Transfe ransferr Media for Fixed-Bed Reactors

Fixed-bed Fixed-b ed rea reacto ctors rs wit with h int integr egrate ated d (in (indir direct ect)) heat exchange require an assortment of heattransfer media that covers the whole temperature range of interest. It is convenient convenient to distin distin--

Figure 20.  TOYO’s MRF-Z radial flow reactor for metha-

nol synthesis with water-evaporation cooling through Field tubes, embedded in the packed bed [28]

guish between gaseous, liquid, and vaporizing heat-transfer media. Gaseous heat-transfer media in the form of hot flue gases are used in the temperature range above 500 C to exclusively supply heat for endo endothermic thermic reactions. reactions. Conversely, Conve rsely, vapo vaporizing rizing heat-t heat-transf ransfer er media are used exclusively to remove heat from exothermic reactions. Whereas formerly petroleum fracti fra ctions ons suc such h as ker kerose osene ne (e. (e.g., g., in eth ethylen ylenee oxide synthesis) were more widely used, they have ha ve no now w be been en la larg rgel ely y re repl plac aced ed by bo boil iling ing wa wate terr on account of their flammability, lower heat of  vapo va poriz rizat atio ion, n, an and d th thee ne need ed to pr prod oduc ucee st stea eam m in a downst dow nstrea ream m con conden denser ser/he /heat at exc exchan hanger. ger. Depending on the saturated vapor pressure, the temperature range from 100 to 310 C (100 bar) can be covered with boiling water. In this range

 

24

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

Figure 21.  Thermoplate fixed-bed reactors from ITS Reaktortechnik [29]

A) Schematic flow part and heat exchange betweenreactor two thermoplates; Innertubes. part ofOuter a smaller thermoplate reactor with header tube; C)ofInner of a larger thermoplate with severalB) header reactor wall and hoods are notone yet installed

it is th thee pr pref efer erre red d he heat at-tr -tran ansf sfer er me mediu dium m fo forr evapor eva porativ ativee coo coolin ling g if an iso isothe thermal rmal coo coolin ling g temp te mper eratu ature re is re requ quir ired ed.. Fig Figur ures es 18 C to 21 sh show ow respective reactor concepts. Locally Local ly varia variable ble coola coolant nt tempe temperature rature profile profiless can be established most easily with liquid heattransfer media that do not vaporize in the intended ten ded ope operat rating ing ran range. ge. To avo avoid id cav cavitat itation ion pitting, pressurized water should be used only up to ca. 220 C; hea heat-tr t-trans ansfer fer oil oilss cov cover er the

temperature salt melts are used almost exclusively [30]. Compared to heat-transfer oils they have the advantage that they are incombustible and have no extended vapor pressure, although they have the disadvantage that they solidify at about abo ut 200 C (n (nit itra rate te me melts lts)) or 40 400 0 C (ca (carbo rbonat natee melts). Nitrate melts decompose above about 500 C, while carbonate melts have a high corrosion ros ion pot potent ential. ial. The tem tempera peratur turee ran ranges ges of  poss po ssib ible le he heatat-tr tran ansf sfer er me medi diaa ar aree sh show own n in

temp te mper eratu ature re ra rang ngee up to 30 300 0 C, wh while ile ab abov ovee th this is

Figure 22.

Figure 22.  Application range of common heat-transfer media

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

25

 ¼ h : =ðc :   r  Þ, whi  ¼ which ch de deter termin mines es the req requir uired ed pum pumpin ping g ene energy rgy of a hea heat-t t-tran ransfe sferr

Figure Fig ure 23.   The mate materialrial-spec specific ific fact factor or F 

medium, as a function of temperature [31]

0 25

2 75 2  p

In addition to the thermal stability, the energy required to circulate a heat-transfer medium thro th roug ugh h th thee co cool olan antt sy syst stem em is an im impo porta rtant nt select sel ection ion cri criter terion ion.. The fol follow lowing ing equ equati ation on   _ per unit yields the required pumping power  N  _ for heat-transfer amount of transported heat  Q media without phase change if the temperature of the heat transfer medium between reactor entrance and exit changes by   DT c  [31]: _=DT  Q c 0:36 _  N 

0:73

  / r h : c

 p 0 09

 

ð12Þ

where  r ,  c p, and  h  are the density, specific heat capacity, viscosity of the heat-transfer medium. Forand liquid heat-transfer media, the equation ti on is de deri rive ved d fr from om th thee pr pres essu sure re dr drop op in a turbul tur bulentl ently y tra traver versed sed,, hyd hydrau raulic lically ally smo smooth oth pipe. It is instructive to resolve Equation (12) with respect to the required pumping power: _  N 

/

  _ Q

DT c

2:75

0:25

  ch:   r 

2 75 2  p

 

ð13Þ

Thus the required power varies as almost the _=DT  . This means that a proper, third power of  Q c energy ene rgy-ef -effici ficient ent des design ign of the coo coolin ling g or hea heatin ting g syst sy stem em sh shou ould ld av avoi oid d st stro rong ngly ly is isot othe herm rmal al conditions with very small values of   DT c. On the contrary, a certain change of the cooling/  heat he atin ing g te temp mpera eratu ture re al alon ong g th thee re reac acto torr flo flow w path pa th ca can n be be bene nefic ficia iall fo forr th thee yi yiel eld d of th thee

reaction, as is explained in Section 1.4.5.2. The factor  F  h0:25 = c p2:75   r 2 , which depends only on the physical properties of the heat-transfer medium, is plotted as a function of temperature in Fig Figur uree 23 23.. Th This is sh show owss th thee ex exce cept ptio iona nally lly go good od heat-tr hea t-trans ansfer fer pro proper perty ty of liq liquid uid wat water, er, whi which ch results from its density and, in particular, from its hig high h spe specific cific hea heat. t. Com Common mon hea heat-tr t-trans ansfer fer oil oilss and sal saltt mel melts ts (so (sodiu dium m nit nitrat rate) e) lie clo close se tog togeth ether, er, but about one order of magn magnitude itude above water. The poor performance of molten sodium compare pa red d to he heatat-tr tran ansf sfer er oi oils ls an and d sa salt lt me melts lts is noteworthy.

 ¼  ¼

ð

Þ

Gases Gas esentire are thetemperature only onl y hea heat-tr t-trans ansfer fer med media usable usa ble over the range, but ia because of their low density they would require excessive circulation energies with an   F  factor   factor five orders of magnitude above that of liquids. They are therefore used exclusively as flue gases to supp su pply ly he heat at at hi high gh te temp mper erat atur ures es,, wh wher ere, e, in addition to convection, radiation contributes to a large extent to the heat transfer. Even then, large temperature differences between the heattrans tra nsfe ferr me mediu dium m an and d th thee re reac acto torr wa wall, ll, wi with th possible adverse effects on the uniformity of  the heat supply, must be tolerated. For higher temperatures, molten salts are the preferred liquid heat-transfer media of choice. Salt melts cover a larger temperature range and have the particul particular ar advan advantage tage over over oils that they

 

26

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

incomb ombust ustible ible.. The pot potent ential ial dan danger ger of a are inc relatively large amount of hot salt melt obviously exists, but is reliably dealt with by experienced rienc ed react reactor or const construction ruction compa companies. nies. Specia Speciall nitrate melts (HITEC) can be used in the temperature perat ure range 200– 200–500 500 C [3 [32] 2].. Gr Grad adua uall de de-compos com positio ition n beg begins ins abo above ve thi thiss tem temper peratu ature, re, and  can acc accele elerat ratee vio violen lently tly abo above ve 600 C. Ac Acce cess ss of  organic to (steam the meltexplosion) (nitrate decomposition)components and of water must be exc exclud luded ed [49 [49]. ]. Sal Saltt mel melts ts bas based ed on car carbon bonate atess are used for the temperature range 400–800  C. In th this is ca case se it is no nott so mu much ch th thee th ther erma mall st stab abil ilit ity y of the molten salt but rather the corrosion of the reactor react or materia materials ls that causes probl problems. ems.

good heat transport from the packing to the tube walls. The fixedfixed-bed bed reactor conc concepts epts discussed in connection with Figures 18 to 21 are typically used for exothermic reactions. Whereas waterevaporation cooling is the preferred concept up to ab abou outt 30 300 0 C coo coolin ling g tem temper peratu ature, re, mol molten ten sal saltt

  La Larg rgee he heat at tra trans nsfe ferr ar area eass mu must st be av avail ailab able le pe perr catalyst volume, which determines the tube diameter diame ter in multit multitubular ubular reactors. .  The temperature of the heat-transfer medium must be close to the desired catalyst temperature to ensure effective catalyst temperature control. .  A sufficiently high mass flow velocity of the

is thee co th cool olin ing g me medi dium um of ch choic oicee fo forr hi high gher er temperatures. Although often referred to as ‘‘isothermal’’, the temperature profiles of a liquid-cooled or liquid-heated packed-bed reactor with a strongly exothermic or endothermic reaction are far from fro m iso isothe therma rmal. l. Figu Figure re 24 sho shows ws sim simula ulated ted temp te mper eratu ature re an and d co conv nvers ersio ion n pr profil ofiles es of a st stro rong ng-ly exo exothe thermi rmicc par partia tiall oxi oxidat dation ion rea reactio ction n in a wall wa ll-c -coo oole led d tu tube be of 25 mm di diam amet eter er un unde derr typical synthesis conditions [10]. It compares thee 2D mo th mode dell re resu sult ltss co cons nsid ideri ering ng pl plug ug flow (dotted lines) or using a radially varying void fracti fra ction on and flow vel veloci ocity ty pro profile file,, com compar parabl ablee to Figure 7 (solid lines). The comparison shows that the plug-flow assumption is well justified, since the conversion and yield profiles are almost identical and only the temperature in the tube center is a little higher in the more detailed model. Even if the packed-bed temperature profiles may differ substantially from the coolant temperature, it is of decisive importance that the reaction conditions are as uniform as possible for all of the many thousand tubes of a largescale multitubular reactor. This refers both to the conditions inside the tubes (flow velocity

reacti rea ction on gas gases es is gen genera erally lly nec necess essary ary to ens ensure ure

and pressure drop, catalyst activity) and to the

1.4.3. Coole 1.4.3. Cooled d Reactors for Exothermic Exothermic Reactions

Strongly exothe Strongly exothermic rmic synthe synthesis sis react reactions ions such as partial oxidations can only be carried out in fixed-bed reactors if the catalyst temperature is contro con trolled lled wit within hin a nar narrow row opt optima imall win window dow.. This can be achieved if the following requirements are fulfille fulfilled: d: .

 o -xylene to Figure 24.  Fixed-bed temperature  T , conversion  X , and yield profiles  F , computed for the partial oxidation of  o

phthalic anhydride in a tube of 25 mm ID under typical operating conditions [10] Solid lines: calculated with a quasi-homogeneous 2D model assuming the detailed radial flow and void fraction profile (comparable to Fig. 7); Dotted lines: calculated with a quasi-homogeneous, 2D plug-flow model

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

27

coolant outside the tubes. For evaporation cooling, unifor uniform m coolan coolantt conditio conditions ns can can more easily be guaranteed than for liquid circulation systems. Here both parallel and crossed cocurrent or countercurrent flow of the heat-transfer medium diu m are commonly commonly employed, employed, as sho shown wn in Figure 18 A and B. The main portion of the heat-tr hea t-trans ansfer fer med medium ium is gen general erally ly cir circul culate ated d

baffles, as illustrated in Figure 18 A. Since the heat-removal conditions are poorly defined in thee re th regi gion on of th thee flo flow w de defle flect ctio ion n wi with th th thee ch chan ange ge from a transverse flow to a parallel flow and back, this region should be free of tubes. A chan ch ange ge of th thee ra radia diall flo flow w ve velo loci city ty ov over er th thee reactorr radius reacto radius can be compe compensate nsated d by provi providing ding pressu pre ssurere-rel relief ief bor bores es in the baf baffle fle pla plates tes,, in-

with a high-capacity pump in orderand to achieve uniform heat-exchange conditions a small temperature tempe rature difference difference betwe between en coola coolant nt feed and an d ex exit it,, wh while ile a pa part rtia iall st stre ream am is pa pass ssed ed th thro roug ugh h a he heat at ex exch chan ange gerr to su supp pply ly or re remo move ve th thee he heat at of  reaction. The steam generator usually produces saturated steam at a pressure corresponding to a boiling point of 30–80 K below the minimum cooling temperature. Apparatus construction companies specializing izi ng in mol molten ten sal saltt coo cooled led mul multitu titubul bular ar rea reacto ctors rs havee dev hav develo eloped ped det detail ailed ed and com compre prehen hensiv sivee know-how regarding flow control of the heattran tr ansfe sferr me mediu dium. m. Th This is co conc ncer erns ns th thee flo flow w co contr ntrol ol within the reactor as well as the uniform feed and discharge of the heat-transfer medium to/  from fro m the tub tubee bun bundle dle,, whi which ch gen genera erally lly tak takes es place via external annular channels [34]. A distinction can be made between parallelflow control and cross-flow control of the coolant. an t. Pa Para ralle llel-fl l-flow ow co cont ntro roll (Fi (Fig. g. 18 B) is ac achi hiev eved ed by two rectifier plates with narrow bores. On account of the pressure loss through the bores, suita su itabl bly y ar arra rang nged ed ov over er th thee re reac acto torr ra radi dius us,, a uniform flow profile can be achieved over the cross section with uniform heat-removal conditions between the distributor plates. The ad-

creas creasing ing in number numbe r towards the reacto reactor r axis. In this way more constant external heat-transfer conditi con ditions ons can be ach achiev ieved ed ove overr the rea reacto ctorr radius. In summary, ensuring uniform heat-transfer conditions as regards the heat-transfer medium requires requir es consi considerab derable le flow-te flow-techno chnologica logicall knowhow. Some publications illustrate the effect of  an inadequate design and layout of the heattransfer medium circuit on reactor performance and runaway behavior [35–37]. Truly isothermal conditions could only be achiev ach ieved ed if the cat cataly alyst st were coa coated ted dir direct ectly ly onto on to a re reac acto torr tu tube be wa wall ll of co cons nstan tantt te temp mper erat ature ure,, as mentioned in Section 1.2.3.4. Such a ‘‘wallcoated’’’ catalytic reactor is a concept which has coated’ been bee n dis discus cussed sed fre freque quently ntly but har hardly dly tra transfe nsferre rred d into int o ind indust ustrial rial app applic licatio ation. n. The mai main n rea reason sonss are problems proble ms with in situ uniform wall coatin coating g and catalyst replacement in large-scale reactors. Recently, microreactors using the design of  parallel-plate heat exchangers with a channel wid wi dth on th thee or ord der of 1 mm or le less ss hav avee be beccom omee available (see  Micro Process Technology, 2. Process Proc essing ing). ). In the these se rea reacto ctors rs the cat cataly alyst st is generally coated on the walls, which allows for exce ex celle llent nt co contr ntrol ol of it itss te temp mper eratu ature re vi viaa th thee he heat at--

vantage of this arrangement that Due the whole reactor can be equipped with is tubes. to the nonuniform flow conditions in the inflow and outflow regions of the liquid coolant, only the region between the distributor plates should be filled fill ed wi with th ac acti tive ve ca cata taly lyst st.. Fo Forr pa paral ralle lell flo flow, w, he heat at transfer between coolant and tube wall is lower than in case of cross-flow, but the difference is not too important, since the main heat-transfer resistance is inside the tubes between wall and catalyst. Because of the increased external heat transfer between coolant and tube wall, the crossflow design design of Figure Figure 18 A is more widely widely used, used, in particular for highly exothermic partial oxidation reactions. It comprises a radially symmetrical arrangement with disk and doughnut

transfe transfer r medium and with makes conducti conducting ng highl highly y exothermic reactions undiluted feed possible. bl e. Ho Howe wever ver,, si simi mila larr to tu tubu bular lar re reac acto tors rs,, a unifor uni form m cat cataly alyst st dep deposi osition tion on the cha channe nnell wal walls ls and the possibility to replace the catalyst without dismantling the microreactor are among the unso un solve lved d pr prob oble lems ms.. Th This is is th thee re reas ason on wh why y microreactors have so far mainly been applied for homogeneous liquid-phase reactions at elevated pressure.

!

1.4.4. Heated Reactors Reactors for Endothermic Endothermic Reactions

Since en Since endo doth ther ermi micc re reac acti tion onss ar aree no nott in da dang nger er of  thermal runaway, close coupling of the reaction

 

28

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

temper peratu ature re to the tem tempera peratur turee of a hea heat-t t-tran ransfe sferr tem medi me dium um is le less ss es esse sent ntia ial. l. Th Thee li limit mitin ing g po poin ints ts ar aree usually transfer of the necessary amount of heat into the catalyst bed and control of the hightemp te mper eratu ature re re resi side denc ncee tim timee in or orde derr to li limi mitt si side de reacti rea ctions ons lik likee cra cracki cking ng and cok cokee for format mation ion.. Although molten-salt reactors have also been used for high-temperature dehydrogenation re-

Figure 25 only about 50% of the heat of combustion is transferred into the endothermic reforming reaction. The rest leaves the reformer box as hot flue gas, requiring a complex heat exchanger network for heat recovery. Such a design is therefore only economical in combination with a larger chemical complex where the excess heat can be utilized.

actio actions ns like ethylbenze ethyl benzene dehydrogen dehyd rogenation ationusufor styrene synthesis, suchne reactions are now ally carried out in multistage adiabatic reactors with wi th in inte ters rsta tag ge he heaati tin ng, as di disc scus usse sed d in Section 1.3.2. The alt altern ernativ ativee for str strong ongly ly end endoth otherm ermic ic high-t hig h-temp empera eratur turee rea reactio ctions ns lik likee (ste (steam) am) ref reform orm-ing of hydrocarbons are multitubular reactors heated directly or indirectly by burners. A typical examp example le is metha methane ne steam reform reforming, ing, where reformer tubes of about 50 mm ID and 10 m leng le ngth th ar aree ar arran range ged d al alon ong g th thee wa walls lls of bi big g bo boxe xes, s, fired fir ed by bu burn rner erss at th thee ce ceil ilin ing, g, th thee bo bott ttom, om,or or/a /and nd the sidewalls (Fig. 25). At the required catalyst temperature tempe ratures, s, the reaction reaction is fast and essentiall essentially y heatt tra hea transf nsfer er con control trolled led.. Thi Thiss mea means ns tha thatt the whole tube length is needed to transfer the heat from the hot burne burnerr gases into the packed catalysst be ly bed. d. In th thee ge gene nera rall de desi sign gn sh sho own in

Forwith decentralized or remote locations, concepts improved heat integration have been developed. An example is the Haldor Topsoe heat he at-i -int nteg egra rate ted d re refo form rmer er HT HTCR CR sh show own n in Figu Fi gure re 26 26,, le left ft.. As sh show own n in th thee in inse set, t, th thee reform ref ormer er tub tubes es con consis sistt of a tub tube-i e-in-tu n-tube be arr arrang angeement with the catalyst in the annulus in good heat he at co cont ntac actt to th thee ho hott flu fluee ga gase ses. s. Th Thee he heat at of th thee hot products leaving through the (empty) inner tube tu be is tr tran ansf sfer erred red to th thee re react actan ants ts en ente terin ring g through the annulus. The highest catalyst temperature resulting in maximum conversion is therefore reached at the lower end of the tube. No ca cata taly lyst st is pl plac aced ed in th thee in inne nerr tu tube be in or orde derr to prevent back-reaction of the equilibrium-limited rea reactio ction n wit with h dec decrea reasin sing g tem temper peratu ature. re. A similarr tubesimila tube-in-tub in-tubee arran arrangemen gementt is used in the ICII ga IC gass-he heat ated ed re refo form rmer er (G (GHR HR)) co conc ncep eptt (Fig. 26, right). Here the hot gas leaving the

Figure 25.  General flow sheet of a top-fired reformer for synthesis gas production. In spite of a flame temperature of about

2000 C, the catalyst temperature at the exit of the reformer tubes reaches only about 850  C

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

29

Figure 26.  Heat-integrated reformer concepts for methane steam reforming

Left: Haldor Topsoe HTCR concept; the inset shows the flow in the concentric reformer tubes with catalyst in the annulus and countercurrent heat exchange between the hot product in the central tube and the feed in the annulus [38]; Right: Similar gasheatedrefo hea tedreforme rmerr (GH (GHR) R) con concep ceptt of ICI[39],wher ICI[39],wheree theconc theconcen entri tricc ref reform ormer er tub tubes es areheat areheated ed by thehot pro produ duct ct of theauto theautothe therma rmall reformer (ATR)

secondary autothermal reformer (ATR) is used to heat the primary steam reformer tubes. The limiting factor for the above concepts is the temperature stability of the reformer tubes used. This limits the maxim maximum um reforming tem perature to about 850 C in the catalyst tubes, leading to about 90% equilibrium conversion for methane steam reforming at 20 bar. Higher temper tem peratu atures res wou would ld req requir uiree ver very y exp expens ensive ive high hi gh-t -tem empe pera ratu ture re al allo loys ys or ce cera rami micc tu tube be materials. To our knowledge the only large-scale synthesis using a ceramic tube multitubular reactor

directi dire ction on,, al alon ong g wh whic ich h sh shor ortt co cont ntac actt ti time me,, high-t hig h-temp empera eratur turee syn synthe thesis sis in wal wall-c l-coate oated d tub tubuular reactors is going to be developed.

  Evonik BMA Process is theProcess)   (former  BMA forr HC fo HCN N sy synt nthe hesi siss fr from om Degussa meth me than anee and ammonia [40]. A sketch of a single ceramic tube is given in Figure 27 together with reactor reactor temperature profiles. This reactor also seems to be the onl only y lar largege-sca scale le mul multitu titubula bularr rea reactor ctor today, where the catalyst is coated at the inner tube wall. The tubes are freely suspended from the top where they are connected to the gas quenching heat exchanger. There the products are cooled down from above 1000 C to about 200 C to prevent ammonia decomposition to nitr ni trog ogen en,, th thee ma main in si side de re reac acti tion on.. Th Thee de desi sign gn an and d the safe operation of such a reactor and the manufacturing of the ceramic tubes cannot yet be considered a standard technology due to the risk ri sk of tu tube be fa fail ilur ure. e. Bu Butt it po poin ints ts in into to th thee

with a pro pronou nounce nced d tem temper peratu ature re the maximu max imum m (‘ (‘‘ho ‘hott spot’’) close to the entrance of reactor. This may lead to increased formation of undesired byproducts and more rapid deactivation of the cataly cat alyst st in the hot hot-sp -spot ot reg region ion.. Mea Measur sures es to influence influen ce the course of react reaction, ion, particularly particularly to damp da mpen en ex excces essi siv ve ho hott sp spo ots ts,, ar aree br brie iefly fly reviewed in the following.

1.4.5. Influe Influencing ncing the Course Course of Reaction

In a standard multitubular fixed-bed reactor one type of catalyst is filled in all tubes and the coolant conditions are kept as uniform as possible over all tubes of the bundle. In case of a strongly exothermic reaction the resulting temperature profile may look like that in Figure 24,

 Reaction Control through Catalyst Activity  Profiles   An ea earl rly y ap appl plie ied d me meth thod od to da damp mpen en hot spots consists of diluting the active catalyst with wit h in iner ertt ma mate teria rial. l. Th This is ca can n ei eith ther er be do done ne at th thee cata ca taly lyst st pr prep epar arati ation on st stag age, e, wh wher eree di diff ffer eren entt batc ba tche hess of th thee sa same me ca cata taly lyst st wi with th di diff ffere erent nt amounts of inert material or with different shell thickness (in case of active shell type catalysts)

 

30

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

Figure 27.   Sketch of one tube of the Evonik BMA multitubular reactor for HCN synthesis with axial and radial temperature

profiles [40]

are produced. Alternatively, the catalyst pellets aree mi ar mixe xed d wi with th di diff ffere erent nt am amou ount ntss of in iner ertt pe pell llet etss duri du ring ng fill fillin ing g of th thee re reac acto torr tu tube bes. s. The fir first st option is generally superior to the second from the vie viewpo wpoint int of sel select ectivi ivity, ty, as dis discus cussed sed in connection with Figure 4, but the second option is much easier to apply. Thee mo Th most st ob obvi viou ouss so solu lutio tion n wo woul uld d be to fil filll th thee catalyst tubes sectionwise with catalysts or catalyst mixtures of different activity. For partial

to a relative activity of 1 as the reaction rate drops due to depletion of the reactants. The result res ultss are com compar pared ed to the das dashed hed pro profile files, s, where activity rises steadily from an entrance value to maximum activity. Although the maximum temperature is substantially dampened in the latter case compared to a case with constant rela re lati tive ve ac acti tivi vity ty of 1, a ho hott sp spot ot is st stil illl vi visi sibl ble. e. An expe ex peri rimen menta tall ve verifi rifica cati tion on of th thee tw two o ca case sess wi with th a slightly simplified activity profile is shown in

oxi oxidat dation ion rea reactio ctions, ns, a front lesss act les active ive catalys cat alyst t is occasionally used in the part of the reactor to av avoi oid d to too o hi high gh ma maxi ximu mum m te temp mper erat atur ures es.. Figure 28 A illustrates the use of two catalysts of dif differ fering ing act activi ivity ty in ser series. ies. The res result ulting ing temper tem peratu ature re pro profiles files hav havee a typ typica icall dou double ble-hump hu mp sh shap apee [4 [41, 1, 42 42]. ]. Gi Give ven n an ap appr prop opria riate te reac re acto torr mo mode del, l, it is ea easy sy to de dete term rmine ine an ac acti tivi vity ty profile which limits the maximum temperature temperature to a pre-specified value. The resulting activity and temperature temperature profiles are given in Figure 28 B [43]. In this case the length of a fully active (relative activity  1) entry region is such that the temperature rises to the pre-specified maximum value. To maintain the temperature at this level, lev el, the activity activity in the followin following g reg region ion is sharply decreased and is then gradually raised

Figure 28 C. In this casemix theing activity has bee been n ach achieve ieved d by mixing active act ivevariation cataly cat alyst st pelle pe llets ts wi with th in iner ertt pe pell llet etss [4 [43]. 3]. For a mu multi ltitu tubu bula larr reactor this requires a fully automated filling procedure proce dure with compu computer-co ter-controll ntrolled ed belt weigh weigh-ers for active and for inert pellets to ensure uniform filling of all reactor tubes. The control of the maximum temperature by cataly cat alyst st act activi ivity ty pro profile filess pre presen sents ts pro proble blems ms if the main reaction zone moves into the region of  high hi gh ca cata taly lyst st ac activ tivity ity du duee to ch chan anges ges in th thee operating conditions. For example, in the case of Figure 28 B and C, a decrease in throughput may ma y al alrea ready dy re resu sult lt in re reac acti tion on ru runa nawa way y in the short, fully active front region. This should be av avoid oided ed by re redu duci cing ng th thee ac activ tivity ity of th this is zone. Catalyst deact deactivatio ivation n occu occurring rring during

 ¼

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

31

Figure 28.  Influence of activity profiles on the temperature profile of a strongly exothermic reaction

A) Catalyst with 66 % activity in the front section of the tube and 100 % activity in the rear section of the tube [41]; B) Linear (broken line) and optimal catalyst activity distribution (full line) for limiting the maximum temperature to 370  C (simulation result); C) Experimental verification of B [43]

operation may have equally severe effects. If, for example, the catalyst is poisoned in a front migr mi grat atin ing g fr from om th thee en entra trance nce to th thee re rear ar,, th thee ma main in reaction zone finally reaches the highly active rear catalyst region, which may also lead to temperature runaway. Rapid catalyst deactivation may also lead to a potentially dangerous tran tr ansi sien entt ru runa nawa way, y, wh whic ich h is ad addr dres esse sed d in Section1.6.2. Section 1.6.2.1. 1. Nevert Nevertheles heless, s, filling subseq subsequent uent tubee sec tub sectio tions ns wit with h cat cataly alysts sts of dif differ ferent ent act activi ivity/  ty/ 

shows a respective reactor together with measured temperature profiles. The great advantage over the somewhat similar temperature profiles in Figure 28 A is the possibility to adjust the peak temperatures of the two sections independently den tly by cha changi nging ng the tem temper peratu atures res of coo coolan lants ts 1 and 2. This is of particular importance in case of gradual catalyst deactivation. A second option of influencing the temperaturee pro tur profiles files is thr throug ough h the pur purpos posefu efull uti utiliz lizatio ation n

selectivity a method of improving the performancecan of be existing multitubular fixed-bed reactors.

of the temperature change the of the heat-transfer medium flowing through reactor: This is technically simpler than a second cooling section and may result in a substantial saving of  pumping energy for the coolant. It is first discussed for the example of cocurrent or countercurrent cooling of a fixed-bed reactor with a strongly exothermic partial oxidation reaction. Figure 30 (left) shows temperature profiles for three different cooling options. If the coolant is circulated so fast that its temperature in the reactor hardly changes, then its flow flow direc direction tion is irrelev irrelevant ant and a temper temperature ature profile with a pronounced temperature maximum mu m ma may y be es esta tabl blis ishe hed, d, si simil milar ar to th thee ex examp ample le disc di scu uss ssed ed in Fi Figu gure re 24. Th This is is sh sho own in Figur Fig uree 30 A. If th thee co cool olan antt is ci circ rcul ulate ated d in

 Reaction Control through Axial Control of  general, influencing influencing the  Heat Transfer   In general, reaction via control of the axial heat-transfer conditi con ditions ons is mor moree flex flexibl iblee and eff effect ective ive but als also o more elaborate than incorporating catalysts of  different activities. An obvious solution is to arra ar rang ngee di diff ffer eren entt he heat at-tr -tran ansf sfer er ci circ rcui uits ts to achieve a stepwise approximation of an optimum tem temper peratu ature re pro profile. file. Liq Liquid uid sal saltt coo cooled led multitubular reactors with two cooling sections meanwhile belong to the state of the art, in particular for highly exothermic reactions like partial oxidations or hydrogenations. Figure 29

 

32

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

Figure 29.  Multitubular reactor with two cooling sections for a partial oxidation reaction (right) and measured reactor

temperature profile (left)

cocurrent and its velocity is chosen so that it beco be come mess no noti tice ceab ably ly ho hott tter er ov over er it itss pa path th,, an almost isothermal temperature behavior can be achi ac hiev eved ed (F (Fig ig.. 30 B). B).Th Then en th thee re reac acti tive ve ga gass at th thee inlet is in contact with the coldest coolant and the cooling temperature rises in step with the cons co nsum umpti ption on of th thee re reac acta tant nts, s, su such ch th that at th thee reaction rate remains virtually constant over a fairly long sectio section n [44–4 [44–46]. 6]. The stabilizing effect of cocurrent cooling with wit h sub substa stantia ntiall tem temper peratu ature re cha change nge of the coo cooll-

reactor tubes should always be surrounded by liquid coolant to exclude melting of the reactor tubes. For cocurrent flow of reactants and coolant, the gas flow should then also be directed upwards, with the danger of increased catalyst attrition due to (partly) fluidizing the packed bed. A third concern is related to proper control of a substantial coolant temperature rise (30 K  in Fig. 30 B), which will require a more refined control contr ol and safety conc concept. ept. On the other hand, the benefit of cocurrent

ant along theexploited flow pathupastoinnow Figure 30 B has hardly been in industrial reactors for different reasons. First of all, at the required low flow velocity (in the example of  Fig. 30 B,   vS  0.01 m/s), natural convection would strongly distort the coolant temperature profile. However  vS describes the mean coolant velocity parallel to the tube axis. With crosscocu co curr rrent ent flo flow w of th thee co cool olan antt lik likee in Fig Figur uree 18 18A, A, the actual flow velocity will in fact be substantially larger, depending on the number of deflections, and the aforementioned problems do nott ar no aris ise. e. A se seco cond nd co conc ncer ern n is re rela late ted d to th thee flo flow w direction of the coolant. Normally coolant flow is directed upwards as in Figure 29, in order to avoid formation of gas pockets below the upper tube plate. As mentioned in Section 1.6.3, the

co cool olin ing g to as coolant in Fig Figur ureecirculation 30 B wi will ll be subs su bstan tantia l wit with respect power. Astial can beh seen from Equation (13), the pumping power  N  decrea dec reases ses wit with h (1/ DT c)2.75. If th thee co cool olan antt is  heated up by   DT c   25 C as in Figure 30 B inst in stea ead d of 2. 2.5 5 C unde underr ‘‘isothe ‘isothermal’ rmal’’’ condit conditions, ions, thee pu th pump mpin ing g po powe werr de decr crea ease sess by a fa fact ctor or of 50 500. 0. This will easily overcompensate any additional pressure drop by an increased number of disk  and doughnut baffles. Compared to cocurrent flow, countercurrent flow has a markedly destabilizing effect for an irreve irr eversi rsible ble exo exothe thermic rmic rea reacti ction on at low flow velocities (Fig. 30 C). Since the incoming reaction mixture in this case is in contact with the warm coolant outflow, the maximum temperature rises to much higher values. Through the

 ¼

 ¼

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

33

Figure 30.  Influence of coolant flow direction and flow velocity  v S  on reaction temperature profile

Left: Isothermal (A), cocurrent flow (B), and countercurrent flow (C) for a partial oxidation reaction; Right: Adiabatic (D), isoth isotherma ermall (E), and coun counterc tercurre urrent nt coo cooling ling (F) for an end endothe othermic rmic reac reaction tion (styr (styrene ene (ST) synt synthesi hesiss from ethy ethylben lbenzen zene, e, EB) a) Reactor temperature; b) Coolant temperature; c) Reference temperature

positive feedback of heat, countercurrent cool-

Counte Cou ntercu rcurre rrent nt flow of hea heat-t t-trans ransfer fer med medium ium

ing canstates, even lead to general the occur occurrence rencethe of runaway multiple steady and in favors of a str strong ongly ly exo exothe thermic rmic,, irre irrever versib sible le rea reacction [47, 48]. In contrast to pure heat exchange without a reaction, react ion, count countercur ercurrent rent heat trans transfer fer in reactorss inv tor involv olving ing exo exothe thermi rmicc rea reacti ctions ons sho should uld therefore be chosen only in special applications. The temperature control of an exothermicc eq mi equi uili libr briu iumm-co cont ntro rolle lled d re reac actio tion n ca can n cons co nsti titu tute te su such ch a ca case se.. As il illu lust stra rate ted d in Figure 15 (left), the optimum temperature profile sh shou ould ld in th this is ca case se de decr crea ease se wi with th in incr crea easi sing ng conv co nver ersio sion, n, i.e i.e., ., al alon ong g th thee tu tube be le leng ngth th.. On account of the equilibrium inhibition of the reac re actio tion, n, it is no nott po poss ssib ible le fo forr th thee rea react ctio ion n to ru run n away in the front region.

can also be of advantage endothermic equilibrium-limited reactions,for since the increasing temperature allows for almost constant reaction rate along the flow path and favors high conversion. Figures 30 D–F show the calculated temperature and concentration profiles with different heating conditions for the case of styrene synthesis synthe sis (dehy (dehydrogen drogenation ation of ethyl ethylbenze benzene). ne). Here the temperatures for the three different heating regimes have been adjusted to give the same styrene conversion. With adiabatic and isothermal isothe rmal react reaction ion contr control, ol, the styrene forma forma-tion rate decreases along the tube, whereas it remains roughly constant with an appropriately adjust adj usted ed cou counte ntercur rcurren rentt flow of the hea heatin ting g medium.. Howev medium However, er, consi considering dering coke forma formation, tion, an increasing temperature profile may be less

 

34

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

favorable than the decreasing adiabatic temperature profile. Generally, the combination of several heattransfer medium circuits (Fig. 29) and the purposeful utilization of the temperature change of  the heat-transfer medium in the reactor offer a wide range of possi possibilitie bilitiess to estab establish lish optim optimum um temperature profiles for a given reaction. This

to th thee ig ignit nitio ion n te temp mper erat atur uree of th thee ca catal talyt ytic ic reacti rea ction. on. Two exa exampl mples es of hea heatt int integr egrati ation on forr en fo endo doth ther ermi micc re reac acti tion onss ar aree di disc scus usse sed d in Figure 26. If the reaction system is all together modera mod erately tely exo exothe thermic rmic,, an ‘‘au ‘autot tother hermal mal’’’ operation results in which no other addition or remo re mova vall of he heat at is ne nece cess ssar ary. y. Au Auto toth ther erma mall operation always requires a special startup pro-

is part pa rtic ularl ytalys impo im port rtan ant t tyfo forrdu comp co mpen ensat satin ing g chan ch ange ges sicul inarly cata ca lyst t ac acti tivi vity duri ring ng time ti me on stream via the coolant temperature profiles. In any case, single-tube experiments with adjustable coolant temperature profile should be performed in order to verify respective advantages and/or disadvantages.

cedure to raise the catalyst temperature the ignition temperature of the respectiveabove reactions. This means that an autothermal reactor always operates in a region of multiple steady states where the ignited state must be established and maintained. 1.5.1. Autoth Autothermal ermal Reactors Reactors with External and Internal Heat Exchange

1.5. Heat-In Heat-Integrate tegrated d Reactor Concepts In heat-integrated reactors the hot effluent of a fixed fix ed-b -bed ed re reac acto torr is us used ed to he heat at up th thee co cold ld fe feed ed

The conve conventiona ntionall autot autotherma hermall react reactor or design consis con sists ts of an adi adiaba abatic tic pac packed ked-be -bed d rea reactor ctor coupled with a countercurrent heat exchanger (Fig. 31 A). As mentioned above, the reaction

Figure 31.   Autotherma Autothermall fixed-bed reactors with recuperativ recuperativee heat exchange. Two basic designs with typical temperature and concentration profiles of a weakly exothermic reaction for two different feed concentrations  c 0(1) and  c 0(2) [50, 53]

A) Conventional design with separate heat exchanger and adiabatic fixed bed; B) Countercurrent fixed-bed reactor

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

mu must st be st start arted ed wi with th he help lp of a se sepa parat ratee pr preh ehea eater ter through which the catalyst bed temperature is raised rai sed abo above ve the ign igniti ition on tem temper peratur aturee of the reaction. react ion. During operation, operation, contro controll measur measures es must mu st be ta take ken n to pr prev even entt th thee re reac actio tion n fr from om extinction if, e.g. the feed to the reactor is too lean le an.. Sin Since ce th thee he heat at of re reac acti tion on is lib liber erat ated ed in th thee cataly cat alyst st sec sectio tion n aft after er the the fee feed d has bee been n heate heated d in

35

the slope doubles if the feed conc concentra entration tion and hence DT ad doub uble les. s. In th thee st stan anda dard rd de desi sign gn (A (A)) a ad do doubling doubli ng of DT ad resu sult ltss in a do doub ublin ling g of th thee to tota tall ad re temperature tempe rature rise DT , si sinc ncee th thee le leng ngth th of th thee he heat at-exchange section is fixed. In the countercurrent reac re acto torr co conc ncep eptt (B (B)) th thee re reac acti tion on st start artss as so soon on as the igniti ignition on tempe temperature rature  T ign ign  of the reaction is exceed exc eeded. ed. If the act activa ivatio tion n ene energy rgy is suf sufficie ficientl ntly y

D

the heat hea t exc exchan hanger ger,, operation the tot total al tem temper peratur ature e ris risee T  under autothermal can be a multiple of the adiabatic temperature rise  D T ad ad (Eq. 11). To prevent overheating, autothermal operation is usually limited to reactions with a maximal adiabatic temperature rise on the order of 300– 400 40 0 K, de depe pend ndin ing g on th thee te temp mper erat atur uree st stab abili ility ty of  the catalyst and the reactor construction. It is best be st su suit ited ed fo forr re reac acti tion onss wi with th a lo low w ov over eral alll exothermicity with   DT ad ad   on the order of 20– 100 K, but requires very efficient gas–gas heat exchange if DT ad small.. Catalytic combustio combustion n ad is small of traces of organic compounds and catalytic hydrogenations are typical examples. Figure 31 B shows the concept of a countercurr cu rren entt fix fixed ed-b -bed ed re reac acto torr in wh whic ich h th thee ca cata taly lyst st is placed inside and outside of a tube bundle that forms a countercurrent heat exchanger for the reacting react ing gas. Altern Alternatively atively,, a paralle parallel-plate l-plate design can be used with the catalyst deposited at the plate surface or between the plates. One advantage of this design is the improved heat transfer caused by the presence of the catalyst packing. Another advantage is a self-adaptivity of this concept with respect to feed concentration, resulting in considerably lower maximum temperatures if the feed concentration varies.

high, the temperature tempe rature increase incre rapidly and thee re th reac actio tion n wi will ll be will comp co mple lete ted daseaf afte terr a sh shor ortt distance. In a limiting case the heat exchange over the reaction front can be neglected and the maximum temperature reached can be estimated fr from om th thee ig igni niti tion on te temp mper erat atur uree pl plus us th thee adiabatic adiab atic tempe temperature rature rise   DT ad show own n in ad   as sh Figure 31 B:

This explained asreactor follows. Incan the be countercurrent of Figure 31 B the reaction starts as soon as the ignition temperature of the reaction   T ign ign  is exceeded. This means that the reaction will start earlier if the slop sl opee of th thee te temp mpera eratu ture re pr profi ofile le in th thee he heat at exch ex chan ange gerr se sect ctio ion n in incr crea ease sess as sh show own n fo forr the dotted profiles. The temperature slope in the heat-exchange section can be estimated by the following equation [51, 52]:

Section isns, obviously impossible for endother th ermic mic 1.5.1 reac re actio tions , bu butt it is po poss ssib ible le to co coup uple le th thee endothermic reaction with an exothermic reaction tio n whi which ch pro provid vides es the req requir uired ed hea heatt of rea reacti ction on in such a way that the combined reaction is sligh sli ghtly tly ex exot othe herm rmic ic an and d no ad addi ditio tiona nall he heat at in inpu putt is requi required. red. Autothermal reforming (Fig. 26, right) and oxidative dehydrogenations (Fig. 17), whereby an exothermic combustion provides the heat for the following endothermic reaction in the same adiaba adi abatic tic pac packed ked bed bed,, are exa example mpless of a sub subcla class ss of reactors for simultaneous coupling of exoand endothermic reactions. In th thee fo follo llowin wing, g, ca case sess ar aree di disc scus ussed sed in wh whic ich h the exo exothe thermi rmicc rea reacti ction, on, usu usuall ally y a cat cataly alytic tic comb co mbus ustio tion, n, is se sepa para rated ted fr from om bu butt in cl clos osee

 ¼ DT 2 lG c

dT  d z

ad ad

 z  pG

eff 

ð14Þ

where   G z   [kg m2 s1] is the gas mass flow velocity and leff  an  an effective axial conductivity combin com bining ing the con conduc ductiv tivitie itiess of the cat cataly alyst, st, the gas, and the separating walls. This means that

 ¼ T  þDT 

T max max

ign ign

ad ad

 

ð15Þ

This is This is of co cour urse se onl only y a ro roug ugh h es estim timat ate, e, si sinc ncee neither is the ignition temperature a properly defined define d quantity nor is the influence influence of the heat exchange over the length of the reaction front really neg really neglig ligibl ible. e. The esti estimat matee is, how howeve ever, r, sufficient to explain the main differences between the two types of autothermal reactors. A mod modifie ified d con concep ceptt of aut autoth otherm ermal al ope operat ration ion is di disc scus ussed sed in Se Sect ctio ion n 1. 1.7. 7.1, 1, in wh whic ich h th thee recu re cupe pera rativ tivee he heat at ex exch chan ange ge of Fi Figu gure re 31 is replaced by regenerative heat exchange.

1.5.2. Heat-I Heat-Integr ntegrated ated Reactors Reactors for Coupling of Endo- and Exothermic Reactions

Auto Au toth ther erma mall op oper erat atio ion n as di disc scus usse sed d in th thee

 

36

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

thermal contact with the endothermic reaction in the same reactor. Then the heat of the exothermic reaction must be taken up by the endothermic one as soon as it is set free. In addition, all hot product gases should be used to heat up alll co al cold ld fe feed eds. s. Su Such ch co conc ncep epts ts ha have ve re rece cent ntly ly be been en developed in particular in the context of decentralized traliz ed hydro hydrogen gen production for fuel cells.

strong separation of the reforming and the combustio bus tion n zon zones, es, and the tem temper peratu ature re pro profile filess show sh ow an ex extre treme me ho hott sp spot ot wi with th a ma maxi ximu mum m temperature tempe rature substantially substantially exce exceeding eding the temperature for total conversion of the reforming reaction. A detailed analysis reveals that both thee re th repu puls lsio ion n of th thee re reac acti tion on zo zone ness an and d th thee resulting hot spot result from an inherent insta-

An parallel obvi ob viou oussflow desi de sign gn is an whereby arra ar rang ngem emen enttenof  many channels, the dothermic reaction and the exothermic reaction each take place every other channel. To incorporate efficient heat exchange between the hot produc pro ducts ts and col cold d fee feeds, ds, cou counte ntercu rcurren rrentt flow with approximately equal heat capacity of both streams is necessary. Figure 32 (left) shows the schema sch ematic tic des design ign and the sim simula ulated ted tem tempera peratur turee and conversion profiles for countercurrent coupling of methane steam reforming and methane combus com bustion tion.. The con conver version sion pro profiles files exh exhibit ibit

bility, changeinduced [54, 55].by the countercurrent heat exUnder countercurrent operation, the desired overlapping of the exo- and endothermic reaction ti on zo zone ness ca can n on only ly be ac achi hiev eved ed if th thee fu fuel el of th thee combustion reaction is added over the length of  thee re th reac acti tion on zo zone ne,, as in indi dica cate ted d in Fig Figur uree 32 (right (ri ght). ). Her Heree uni unifor form m fue fuell add additio ition n ove overr the leng le ngth th of th thee co comb mbus ustio tion n ch chan anne nels ls wa wass as as-sumed. sum ed. The lar largege-sca scale le rea realiza lizatio tion n of suc such h a uniformly unifo rmly distributed distributed side feed in a multich multichanannel arrangement at temperatures exceeding the

Figure 32.  Countercurrent coupling of methane steam reforming with methane combustion in a parallel-plate reactor with

catalysts deposited on the separating walls Left: Simulation results with premixed methane/air at the combustion side [54]; Right: Simulation results with equally distributed side feed of methane on the combustion side [55, 56]

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

37

Figure 33.  Cocurrent/countercurrent reactor for coupling of methane steam reforming with methane combustion [57, 58]

A) Flow scheme of the total setup with two countercurrent heat exchangers attached to the cocurrent reaction section; B) Temp Temperat erature ure and conv conversio ersion n profil profiles es in the cocu cocurren rrentt reac reaction tion sect section; ion; C) Layo Layout ut of the reac reaction tion sect section ion with two comb combustio ustion n catalysts in each combustion channel before and behind the methane side feed; D) Photo of the full pilot plant setup

met methan hanee ign ignitio ition n tem temper peratu ature, re, how howeve ever, r, is a thus-far unsolved task. Under cocurrent flow of the exo- and endothermic reactions, overlapping of the two reac-

reforming and methane combustion. Feed temperatures and flow rates must be chosen such thatt nei tha neithe therr met methan hanee com combus bustio tion n is que quench nched ed by the strongly endothermic steam reforming, nor

tion zonesoperation will automatically but cocurrent requires a take moreplace, complex flow flo w de dessig ign, n, as sh show own n sc sch hem emaati tica call lly y in Figure Fig ure 33 (to (top). p). Now the coc cocurr urrent ent rea reacti ction on section sec tion is sep separa arated ted fro from m the two hea heat-e t-exch xchang angee sect se ction ions, s, in wh whic ich h fe feed ed an and d ef efflu fluen entt of th thee reactants from the reforming and from the combustion side are in heat exchange. This design hass th ha thee ad addi ditio tiona nall ad adva vant ntag agee th that at th thee he heat at exchange in both heat exchangers always takes place at about equal heat capacity flux. This allo al lows ws di diff ffer eren entt flo flow w ra rate tess in th thee re refo form rming ing and in the combustion channels to be chosen [56–58]. Figure 33 shows the simulated temperature and conversion profiles for the reaction section, opti op timi mize zed d fo forr co coccur urre ren nt me meth than anee st stea eam m

does methane combustion lead to m a runaway. In additio add ition, n, the exi exit t tem tempera peratur tures es fro from the rea reacti ction on section should be high enough for total conversion of the reforming mixture and for subsequen qu entt he heat at ex exch chan ange ge in th thee he heat at ex exch chan ange gerr sections. To fulfil these requirements, methane forr th fo thee co comb mbus ustio tion n si side de is sp split lit be betw twee een n th thee fe feed ed and a side feed in the middle of the reaction section. This leads to double-hump temperature profiles. The simulation results of Figure 33 B have been verified in a pilot reactor for decentralized hydr hy drog ogen en pr prod oduc uctio tion n (Fi (Fig. g. 33 D) [5 [58]. 8]. Th Thee reacti rea ction on sec section tion con consist sistss of nin ninee par parall allel el ref reform orm-ing and ten combustion channels of 2.3 mm width in a folded-sheet design [56] with a total volumee of 0.7 L for production volum production of 5 m3(STP)/h

 

38

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

hy hydr drog ogen en (c (cor orre resp spon ondi ding ng to 15 kW lo lowe werr heating value). The same commercial Pd catalyst, coated on corrugated spacers, was used in the ref reform orming ing and the com combus bustion tion cha channe nnels; ls; >98% reforming conversion and full combustion conversion was reached with ca. 90% thermal efficiency in an operating window between 30 and 100% load.

1.6.1. Param Parametric etric Sensitivity Sensitivity and Runaway

Similar cocurrent coated catalysts haveflow beenconcepts proposedwith andwallanalyzed, among others in [59, 60]. If the catalysts for the exo exothe thermic rmic and the end endoth otherm ermic ic rea reacti ction on are coated on opposite sites of the same separating wall, excellent heat transfer between the endothermic and exothermic reactions can be achieved. In addition, homogeneous ignition of  the combustion reaction can be prevented if the channe cha nnell dim dimens ension ionss are in the sub submil millime limeter ter range [61, 62]. Wall heat conduction assures an al almo most st un unif ifor orm m te temp mper eratu ature re pr profil ofilee as lo long ng as thee re th reac acto torr di dime mens nsio ions ns ar aree in th thee lo low w ce cent ntim imet eter er range. However, wall heat conduction prevents efficient heat recovery from the hot effluents to heat he at up th thee co cold ld fe feed edss if th thee he heat at ex exch chan ange ge sect se ction ionss ar aree no nott ex exte tend nded ed to se seve vera rall te tens ns of  centimeters centi meters in length [56].

Gas-phase Gas-ph ase fixe fixed-b d-bed ed rea reactor ctorss are amo among ng the biggest and most widely used reactors in the chemical industry. Since in these reactors large amounts of combustible or potentially decomposable gases are processed, often at high tem-

for this s beh behavi avior or hav have e bee been n ide identi ntified. fied. Colloq loquiuiallythi they are all considered under theCol heading ‘‘runaway’’. In fixed-bed reactors, runaway usually occurs cu rs un und der ope pera rati tin ng co cond ndit itio ions ns of hi high gh parametric sensitivity, where small changes in thee op th oper erat atin ing g pa para rame mete ters rs ca can n le lead ad to la larg rgee change cha ngess in max maximu imum m tem temper peratu ature re and yie yield ld [63]. [63 ]. One mai main n reason reason is the expon exponent ential ial dep depenendenc de ncee of th thee re reac acti tion on ra rate te on te temp mper erat atur uree (Arrhe (Ar rheniu niuss law law). ). Fig Figure ure 34 sho shows ws cal calcul culate ated d temperature profiles for a partial oxidation reacti ac tion on in a wa wall ll-c -coo oole led d fix fixed ed-be -bed d re reac acto torr tu tube be of  typical dimensions. In Figure 34 A only the main reaction is considered, while in Figure 34 B the total combustion to CO2   and water is additionally taken into account. Both cases lead to almost identical results up to coolant temperatures of 330 C. Since considerably more heat is liberated in total combustion combustion than in the desired main reaction, the sensitivity increases substa sub stantia ntially lly as soo soon n as the ign igniti ition on tem temper peratu ature re of the second reaction is exceeded. As a measure of the parametric sensitivity, Figure 34 C shows the change in maximum temperature versus the cooling temperature for case B. The sensitivity is only moderate at low

peratures and safety issues are imp import ortant ant elevated for the their irpressures, design des ign and operat ope ration ion.. Since the early 1970s, such issues have been treated mainly under the heading ‘‘thermal runaway’’ in a large number of publications. This could give the impression that fixed-bed reactors with gas-phase reactions reactions are a reactor type with wi th a hi high gh po pote tent ntia iall ri risk sk.. In fa fact ct,, th thee op oppo posi site te is true. tru e. Com Compar pared ed to liq liquid uid-ph -phase ase rea reacti ctions ons the mass of reactants in gas-phase reactions is several orders of magnitude smaller. This implies that the danger of decomposition of reactants accumu acc umulate lated d in the rea reacto ctorr is con conside siderab rably ly sma smalller than in liquid-phase reactors. The heat capaci pa city ty of th thee ca cata taly lyst st pa pack ckin ing g ad addit ditio iona nall lly y damp da mpss th thee un unco cont ntro roll lled ed te temp mpera eratu ture re ris risee in fixed-bed reactors.

cooli cooling ng tempera techanges mperatures, tures, whereas wherea s abov abovee T C   343 C small in   T  C   — and also in other parameters such as throughput, feed concentration, and pressure — lead to large changes in reactor behavior. In fact under the conditions considered a new steady state is approached if  thee co th cool olan antt te temp mper eratu ature re is ra raise ised d ab abov ovee   T C  344 C (Fig. 34 D). This means that the maximum temperature remains at an upper value, even eve n if the coo coolan lantt tem temper peratu ature re is dec decrea reased sed  below   T C   340 C. Due to the unavoidable differences between individual tubes, multitubular reactors cannot be op oper erat ated ed in th thee ra rang ngee of hi high gh pa param ramet etric ric sensitivity. In the case considered, the cooling temperature must be kept below ca. 340 C, and the tubes made longer to achieve a sufficient

1.6. Operat Operational ional and Safety Safety Issues

Nevertheless, Nevertheles s, instab instabilities ilities can arise in fixedfixed-bed bed reactors, react ors, partic particularly ularly with strong strongly ly exoth exothermic ermic reactions, and can lead to temperature excursions which can damage the catalyst and the reactor construction materials. Several causes

 ¼

 ¼

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

39

Figure Fig ure 34.   Par Parame ametri tricc sen sensit sitivi ivity ty of a par partia tiall oxi oxida datio tion n rea reacti ction on in a fixe fixed-b d-bed ed rea reacto ctorr of typ typica icall dim dimens ension ionss as a fun functi ction on of the  T   T   z   T  coolant temperature C  with (  0) C

 ¼  ¼

A) Temperature profile over reaction length (main reaction only); B) temperature profile including total oxidation as side reaction reac tion;; C) maxi maximum mum temp temperatu erature re T max functi ction on of coo coolan lantt tem temper peratu ature re T C incas incasee B;D) T max func ncti tionof  onof T C max andyieldas a fun max as a fu for both cases

conversion. This example emphasizes the requireme qui rement nt dis discus cussed sed in Sec Sectio tion n 1.4 1.4.3 .3 for mak making ing thee co th cond ndit itio ions ns in th thee tu tube bess of th thee tu tube be bu bund ndle le an and d in the cooling circuit as uniform as possible to avoid premature runaway reaction in individual tubes. Under runaway conditions, initially only a few particularly sensitive tubes of the bundle will be affected. In a multitubular reactor with thou th ousa sand ndss of tu tube bess no nott ev ever ery y tu tube be ca can n be equipped equip ped with tempe temperature rature-profil -profilee measu measurerements; it is therefore likely that this runaway will wi ll re rema main in un und det etec ecte ted d by te tem mpe pera ratu ture re

measu meas ure reme men nts ts.. Alth Al tho oug ugh h temp te mper erat atu ure ress  >1000 C can often be reached in the catalyst during such runaways, there is no increased safe sa fety ty ris risk, k, pr prov ovid ided ed th thee tu tube be is su surr rrou ound nded ed by a liquid heat-transfer medium. Because of the good heat transfer to the fluid, the tube temperature will remain close to that of the heattransfer medium, and melting of the tube does nott oc no occu curr ev even en if th thee ca cata taly lyst st pe pell llet etss ar aree destroyed by the high temperatures. The most certain method of detect detecting ing a runaway is on-line analysis of a produ product ct formed in the runaway reaction. For example, CO 2 can be

 

40

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

monitored in the off-gas during the runawaysens se nsiti itive ve sy synt nthe hesi siss of et ethy hyle lene ne ox oxid ide. e. If it itss concentration increases above a specified limit, thee re th reac acto torr mu must st be sh shut ut do down, wn, pu purg rged ed wi with th nitrog nit rogen, en, and coo cooled led to a low lower er tem tempera peratur turee forr a cer fo erta tain in pe peri rio od be befo fore re op oper erat atio ion n is recommenced. Numerous runaway criteria can be found in the with which ranges of  high hig hliterature parame par ametri tric c sen sensiti sitivit vity y operating can be pre precal calcul culated ated for known reaction kinetics. In practice these criteria are of only limited importance, because they rarely take into account the peculiarities of  individ ind ividual ual cas cases. es. Sen Sensiti sitive ve rea reacti ctions ons suc such h as partial parti al oxida oxidation tion and partial hydro hydrogenat genation ion are therefo the refore re gen genera erally lly tes tested ted in sin single gle-tub -tubee rea reacto ctors rs of the same dimensions as those in the subsequen qu entt mu mult ltitu itubu bula larr re reac acto tor. r. Th This is al allo lows ws th thee rang ra ngee of pa param ramet etri ricc se sens nsiti itivi vity ty to be de dete termi rmine ned d directly. Recalculation of the results for other tube diameters is possible, but the uncertainties in the quantification of the heat-transfer characteristics (see Section 1.2.1) should be kept in mind. A fixed-bed reactor can enter the region of  high parametric sensitivity through changes in thee op th oper erat ating ing co cond nditi ition onss or in th thee ca cata taly lyst st properties. In addition, rapid changes in feed temperature, feed concentration, or throughput may induce moving temperature and reaction fronts which can lead to transient excess temperatures, perat ures, a pheno phenomenon menon known as ‘‘wron ‘wronggway behavior’ behavior’’. ’.

1.6.2 1.6.2. . Movin Moving g Temp Temperatur eraturee and Reacti Reaction on Fronts

In addition to parametric sensitivity, a second cause of thermal runaway of a reaction can be attributed to dynamic phenomena resulting in moving temperature and reaction fronts. This is expl ex plai aine ned d wi with th Fig Figur uree 35 wh wher eree an in initi itiall ally y co cold ld and suf sufficie ficientl ntly y lon long g adi adiaba abatic tic fixe fixed-b d-bed ed rea reacto ctorr is considered [57]. It is fed from one side at low temperature with the reactants of an irreversible, exothermic reaction. If a certain part of the bed is hea heated ted abo above ve the ign igniti ition on tem temper peratu ature re T ign ign of th thee re reac acti tion on,, a tra trave veli ling ng ho hott zo zone ne is ge gene nera rate ted. d. The hot hot zone is framed by two two fronts. fronts. At At the left side a reaction front with total conversion of the reaction is established, which travels with the

Figure 35.  Traveling thermal front (velocity  wT) and reaction fron frontt (velo (velocity city   wR) for an irr irreve eversi rsible ble ex exoth otherm ermic ic

reaction in a locally preheated, sufficiently long adiabatic fixed-bed reactor with feed temperature  T 0  well below the ignition temperature  T ign ign

reaction front velocity   wR. At the right side a puree thermal pur thermal fro front nt tra travels vels wit with h the therma thermall fro front nt velocity   wT. Thee ac Th actu tual al re react actio ion n fr fron ontt ve velo loci city ty is th thee re resu sult lt of two co coun unte terac ractin ting g ef effe fect cts: s: Th Thee co cold ld fe feed ed tend te ndss to co cool ol do down wn th thee ca catal talys ystt in th thee fr fron ontt region while the heat of reaction, released in the front, results in a backward conduction of  heat. This reduces the reaction front velocity compared to the thermal front velocity (wR   < wT) wh whic ich h ma may y ev even en le lead ad to an up upst stre ream am mo movi ving ng front (wR   <   0). If the feed temperature   T 0   is well below the ignition temperature T ign ign  of the reaction and the adiabatic temperature rise is limited, both fronts will travel downstream as shown in Figure 35. Typical for the exothermic reaction front is its sha shapepe-pre preser servin ving g cha charac racter ter.. This dis distin tin-guishes it from thermal fronts, which disperse (flatten) due to the combined influence of convective gas transport, heat transfer between gas and and axial dispersion of amass and heat.packing, The mean transport velocity of thermal front   wT   can be described by the well-known Equation (16).

 ¼ evr r  cc

wT

G  pG

S S

ð16Þ

where   ev  is the superficial velocity,   r Gc pG   the gass he ga heat at ca capa pacit city, y, an and d   r ScS   the the me mean an he heat at capacity of gas and packed bed. The eff effect ective ive tem temper peratu ature re rise of the rea reacti ction on front   DT R  obviously depends on the difference between wT and wR. If th thee di diff ffer eren ence ce is la larg rge, e, th thee heat liberated by reaction is distributed over an incr in crea easin sing g po porti rtion on of th thee be bed d (th (thee sh shad aded ed ar area ea in Fig.. 35) Fig 35),, lea leadin ding g to a mod moderat eratee tem temper peratu ature re increase DT R. Conversely, if  w  wR approaches  wT

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

th thee he heat at li libe bera rate ted d ac accu cumu mulat lates es in a sm smal alll zo zone ne of  thee be th bed, d, le lead adin ing g to hi high gh va value luess of   DT R. An integral heat balance around the traveling hot zone zo ne,, ba bala lanc ncing ing th thee he heat at of re reac acti tion on an and d th thee he heat at accumulated in the hot portion of the bed, leads to

 ¼ ðw w Þr  c DT 

evr G c pG DT ad ad

T

R

S S

R

 

ð17Þ

and, if Equation (16) is inserted, to DT R

 ¼ DT 

ad ad

wT

ð18Þ



wT wR

or wR

 ¼ w ð1DT  T

ad =DT R ad

Þ

ð19Þ

Equation (19) provides a simple explanation for the shape-preserving or self-sharpening behavio ha viorr of ex exot othe herm rmic ic re reac acti tion on fr fron onts ts as sh show own n in Figure 35: Below the igniti ignition on temperature T ign ign of the rea reacti ction on con consid sidered ered,, no rea reacti ction on tak takes es place and hence the thermal front velocity   wT should sho uld pre prevail vail.. Abo Above ve   T ign front nt sho should uld ign   the fro move with   wR   <   wT. The local difference in the front velocities tends to drive the reaction front into a shock, which is counteracted by the unav un avoid oidab able le di disp sper ersiv sivee ef effe fects cts me ment ntio ione ned d above. A moving reaction front can also occur in a fixed fix ed-be -bed d re react actor or if th thee flo flow w ve velo loci city ty is su sudd dden enly ly increased above the blow-out velocity of the reac re acti tion on an and d if th thee fe feed ed co conc ncen entr trati ation on or th thee fe feed ed temperature is suddenly decreased [48, 64, 65]. Forr th Fo thee re resu sulti lting ng in incr crea ease se of th thee ma maxi ximu mum m temperature tempe rature the term ‘‘wrong ‘wrong-way -way behav behavior’ ior’’’ has been established established [65].  Rapid that Catalyst Deactivation   Equation (19) confirms in a moving exothermic reaction front   wR   <   wT. This means that the heat of  reaction is distributed over an extended portion of the bed, as already discussed for Figure 35. Hence the maximum temperature rise   DT R   remains bounded. A notable exception occurs if  the rea reactio ction n fro front nt vel veloci ocity ty   wR   is no nott fr free eely ly established but is forced to increase. This may happen hap pen if the cat cataly alyst st is rap rapidly idly dea deactiv ctivate ated, d, e.g., by a catalyst poison in the feed [66]. Then the deactivation deactivation front pushe pushess the reaction front ahead. A critical situation obviously occurs if  the velocity of the deactivation front equals the ther th erma mall fr fron ontt ve veloc locity ity.. Th Then en th thee he heat at of th thee reaction can no longer be distributed over an exte ex tend nded ed be bed d po port rtio ion n as in Fi Figu gure re 35 bu butt

41

accumulates in a narrow region, resulting in a steadily increasing maximum temperature. In this case (wR   wT) Equation (18) predicts a constant increase of the maximum temperature to infinity. Althoug Alth ough h it is rat rather her unl unlike ikely ly tha thatt the vel veloci ocity ty of a poison front coincides with that of a temperature front, the situation is different if deac-

 ¼

tivatio tivation n resul results ts fromcircle thermal catalyst st deact deactivaivation. Then a vicious cancataly develop in which the hig high h cat cataly alyst st tem temper peratu ature re in the rea reacti ction on front deactivates the catalyst, which induces a reacti rea ction on fro front nt mov moveme ement nt thr throug ough h whi which ch the catalyst temperature is further increased. Such situati situ ations ons hav havee bee been n obs observ erved ed and rep reporte orted d in [67, 68]. Figure 36 (left) shows experimental results from the startup of a laboratory reactor for vin vinyl yl ace acetate tate syn synthe thesis sis usi using ng a the therma rmally lly inst in stab able le zi zinc nc ac acet etate ate ca cata talys lystt [6 [68] 8].. On th thee right-hand side simulation results following an incident during startup of a pilot plant reactor are shown [67]. The specific example was the oxid ox idat atio ion n of tr trac aces es of CO in of offf-ga gass in an industrial indust rial adiaba adiabatic tic fixedfixed-bed bed reacto reactor, r, using an Ni catalyst with a too low thermal stability. In spit sp itee of th thee lo low w ad adia iaba batic tic te temp mper erat atur uree ri rise se (<   100 K) of th thee re reac actio tion n (s (see ee te temp mper erat atur uree profile pro file for   t    500s 500s), ), a sub substa stanti ntial al tra transi nsient ent temperature rise developed. Since, due to the low feed concentration, concentration, no speci special al preca precautions utions where taken, the gradual development of the high-temperature front remained unnoticed as long as the hot front was inside the packed bed. The more dramatic was the sudden substantial exit ex it te temp mper erat atur uree ri rise se,, as so soon on as th thee fr fron ontt

 ¼  ¼

approa approached the cataly catalyst st Theched example exa mples s sho show w bed. thatt dyn tha dynami amicc eff effect ectss associated with the movement of exothermic reaction fronts may lead to unexpected temperature excursions. This is in particular the case if  instead of a gas-catalyzed reaction also a solid takes part in the reaction. The most prominent exam ex ample ple is th thee bu burn rn-o -off ff of co coke ke in co coke ked d ca cata talys lystt beds.  Decoking of Fixed-Bed Reactors   So far only gas–gas reactions catalyzed by a fixed-bed cataly lysst have been consid ideere red d. Under certain circumstances, however, also gas–solid reac re acti tion onss ta take ke pla place ce in ca cata talyt lytic ic fix fixed ed-be -bed d reac re acto tors. rs. A ty typi pica call ex exam ample ple is th thee pe peri riod odic ic burn-off of coke formed during reforming or

 

42

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

Figure 36.  Evolution of temperature profiles during the start-up of fixed bed reactors with thermally instable catalysts

Left: vin Left: vinyl yl ace acetat tatee syn synthe thesis sis in a fixe fixed-b d-bed ed rea reacto ctorr wit with h an ins instab table le zin zincc ace acetat tatee cat cataly alyst st [68 [68]; ]; Rig Right: ht: sim simula ulatio tion n res result ultss sho showin wing g the startup of the oxidation of off-gas with a low CO concentration in a pilot-plant reactor [67]

de dehy hydro droge natio tion n surface. reac re actio tions ns, , wh which ichexample tend te ndss of  to block thegena catalyst Another partic par ticular ular imp import ortanc ancee is the reg regene enerat ration ion of  diesel soot filters. In both cases, the fixed-bed entran ent rance ce is per period iodica ically lly hea heated ted to abo above ve the cok cokee or soo soott ign ignitio ition n tem temper peratu ature, re, and a gas flow containing oxygen is introduced. The developing temperature fronts can be treated similarly to Fig Figur uree 35 [5 [57, 7, 69 69]. ]. Th Thee ma main in di diff ffer eren ence ce is th that at G now a gas component A (oxygen) is reacting with a solid reactant BS (coke) which has been deposited at the fixed-bed surface. In the reaction front the initial coke loading q0B  is assumed to be co comp mple lete tely ly bu burn rned ed of off, f, le lead adin ing g to th thee simple stoichiometric relation wR n q0B

0 A v z

 ¼ c

  and   wR

 ¼ cn qv

A Z 0 B

ð20Þ

where n is th thee st stoi oich chio iome metr tric ic co coef effic ficien ientt of co coke ke combustion comb ustion.. Hence Hence,, the reacti reaction on front velocity wR  is now completely specified by the gas feed velocity v z, the coke loading  q0B , and the oxygen feed concentration   c0A . If the coke loading q0B  is low and the oxygen concentration   c0A  is high,   wR  will be large and exceed exc eed the the therma rmall fro front nt vel veloci ocity ty wT. In th this is ca case se the reaction front will precede the thermal front (Fig (F ig.. 37 37,, le left ft). ). If th thee co coke ke lo load adin ing g is hi high gh an and d th thee oxygen concentration is low (Fig. 37, right), the thermal front will precede the reaction front. In both cases the heat of coke combustion will be

dispersed overin a larger themaximum fixed-bed (shaded areas Figureportion 37) andofthe temperature rise   DT R  is limited. Again, a simple heat balance between the heat released and stored in the fixed-bed gives an app approx roxima imatio tion n for   DT R. In the case of  Figure 37 (left) the solid component is obviously limiting and will be completely consumed in the rea reactio ction n fro front, nt, whi while le c0A is in ex exce cess ss.. Th Then en th thee heat balance with   r    denoting the mean heat r  c capacity of catalyst and soot loading yields 0 R B

ðw w Þr r cDT   ¼ w q ðDh Þ T 

R

R

R

leading leadi ng to 0 B

 ¼ w  ww q ðr r cDh Þ R

DT R

R

R

  p

T

ð21Þ

 |{z}  D T S

ad

Alternatively, Alternativel y, if in Figure 37 (right) (right) the solid comp co mpon onen entt is in ex exce cess ss an and d th thee ga gass co comp mpon onen entt is limiting, the respective heat balance leads to 0 A

Dh Þ  ¼ w  ww c eð r  c R

DT R

R

T

R

G  pG

ð22Þ

 |{z}  D T G

ad

A sin singul gulari arity ty wit with h (th (theor eoreti eticall cally) y) unb unboun ounded ded temp te mper eratu ature re ri rise se is pr pred edic icted ted if th thee re reac acti tion on fr fron ontt velocity approaches the thermal front velocity (Fig.. 38) (Fig 38).. Saf Safee ope operati rating ng con condit dition ions, s, on the oth other er

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

43

Figure Fig ure 37.   Reaction front velocity wR and ther thermal mal front front velo velocity city wT during a gas–solid gas–solid reac reaction tion (cok (cokee comb combustio ustion) n) in a fixed-

bed reactor Top: Simplified model of coke (BS) combustion with an oxygen-containing gas (AG); Left: Low solid loading and high gas concentration; Right: High solid loading and low gas concentration

hand, can be expected in the two limiting cases of Figure 37. If the coke loading is low and the oxygen concentration is high, the reaction front velocity exceeds the thermal front velocity considerably and the maximum temperature rise S approaches   DT ad , which is proportional to the coke loading (see Eq. 21). More realistic for the decoking of fixed-bed reactors is the opposite case ca se wi with th hi high gh co coke ke lo load adin ing. g. The Then n th thee ma maxi ximu mum m

Figure Fig ure 38.   Temperatu Temperature re rise DT R depe dependin nding g on the reac reaction tion front velocity w R (see Eq. 11) for the gas (oxygen) limiting

case, fol case, follow lowing ing Equ Equati ation on (21 (21)) an and d for the sol solid id (co (coke ke)) limiting case, following Equation (20). If the reaction is ignited at the reactor exit (‘‘rear-ignition mode’’), a backward traveling reaction front (wR  <  0) may occur [57]

temper temp erat atur uree ri rise se ca can n on only ly be li limi mite ted d by re redu duci cing ng 0 the oxygen concentration   cA  to very small values. This reduces the reaction front velocity to almost zero and the maximum temperature rise G to DT ad , which in turn is proportional to  c0A  (see Eq. 22). The decoking of industrial fixed-bed reactors reacto rs is therefore started with very low oxygen concentrations, and only gradually is the oxygen concentration increased to achieve total regeneration. This allows for safe regeneration, irrespective of the initial coke loading. Diesel soot filters, on diesel the other hand,conare usually regenerated with exhaust, tain ta inin ing g ab abou outt 10 vo vol% l% of ox oxyg ygen en.. Wi With th th this is hi high gh oxygen concentration the regeneration temperature can can only be controll controlled ed by limiting limiting the soot 0 loading  q B . Since on-line monitoring of actual soot loading is far from trivial, soot filter regeneration is usually initiated in shorter intervals to sa safel fely y av avoi oid d ex exce cess ss te temp mper erat atur ures es du durin ring g regeneration. Interestingly, a third inherently safe regeneration option exists. In the two options discussed so far, regeneration is started by heating the reactor inlet above the coke ignition temperature (front-ignition mode of Fig. 37). Then the reaction front moves downstream from left to right. However, if coke combustion is started

 

44

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

with limited gas flow by igniting the reaction at the end of the reactor, the reaction front may move upstream. This leads to a maximum temG perature rise   DT R   <   DT ad  (see Fig. 38, left for wR   <  0), since in this case the reaction front must mu st he heat at up th thee ga gass pl plus us th thee pa pack cked ed be bed d upstream.

stoich stoi chio iome metr tric ic op oper erat atio ion n in th thee ig igni nita tabl blee range [70]. Referring to multitubular fixed-bed reactors react ors of conv convention entional al dimen dimensions, sions, a prere prereqquisite uis ite for thi thiss is a pre pressu ssure-r re-resi esista stant nt con constr struct uction ion with check valves and flame barriers so a possible ignition is confined to the interior of the reactor. Meltin Mel ting g of rea reacto ctorr tub tubes es dur during ing run runawa away y

1.6.3.. Other Safety 1.6.3 Safety Aspects Aspects

rea reacti ction on ifisthe only onl y to be fea feared redis in multitu titubul bular ar reactors respective tube notmul completely surrou sur rounde nded d by liq liquid uid hea heat-tr t-trans ansfer fer med medium ium.. Thus, appropriate appropriate design must ensur ensuree that running dry of reactor tubes cannot occur. In the case of corrosive reaction gases, provision for the detection of leaks caused by corrosion must be made, particularly when pressurized or boiling water is used as coolant. A ne new w op opti tion on to co cope pe wi with th so some me of th thee ab abov ovee mentioned mentio ned safet safety y conce concerns rns are micror microreacto eactors rs (see  Micro Process Technology, 2. Processing and Section 1.2.3.4). It has been shown [61,, 62] tha [61 thatt ext extreme remely ly exo exothe thermic rmic rea reacti ctions ons with pot potent ential ially ly exp explos losive ive fee feed d com compos positio ition n lik likee stoichiometri stoich iometricc hydro hydrogen–a gen–air ir mixtu mixtures res can be carried out without thermal runaway in catalyst-coated microreactors if the cross-sectional dimensions of the reaction channels are sufficiently small [71, 72]. In addition, micromixing of reactants prior to reaction or controlled addition tio n of rea reacta ctants nts wit with h imm immedia ediate te mix mixing ing becomes possible. This opens the possibility to carr ca rry y ou outt re reac acti tion onss wh whic ich h so fa farr ha have ve be been en impo im possi ssibl blee be beca caus usee of la lack ck of su suffic fficie ient nt te temp mper er-ature and residence-time control. Due to optima optimall react reaction ion conditions, conditions, the vol-

Becausee of the sma Becaus small ll mas masss sto storag ragee cap capaci acity ty compared to liquid-phase reactors, the danger of a su sudd dden en re reac actio tion n of ac accu cumu mulat lated ed re reac acta tant ntss in gas-phase fixed-bed reactors is comparatively low. Leaving out the peculiarities of individual cases, the following safety risks can generally be assumed for fixed-bed reactors: 1. Leaks Leaks which result result in the rel releas easee of large amou am ount ntss of ga gass or va vapo porr an and d th thee fo form rmat atio ion n of  explosive clouds. 2. Le Leak akss re resul sulti ting ng in re rele leas asee of la larg rgee am amou ount ntss of  liquid heat-transfer media (oils, salt melts). 3. Formati Formation on of ignitable ignitable or decomposable decomposable gas mixtures in the reactor. 4. Mel Melti ting ng of th thee re reac acto torr du duee to a ru runa nawa way y reaction. Safety aspects of liquid heat-transfer media were mentioned in Section 1.4.2. Ignitable gas mixtures mixtur es can arise particu particularly larly during partia partiall oxidation reactions. They are especially critical where large, packing-free volumes are present. This is the case in the inflow and outflow hoods of the reactor, in the tubes of thea catalyst packingwhile dampens thereactor propagation flame front due to its heat capacity. Complete avoidance of ignitable mixture is generally not possible in partial oxidations because at least during mixing of the gas streams prior to the reactor the ignition limit is exceeded locally. Operat Ope ration ion of fixe fixed-b d-bed ed rea reacto ctors rs wit with h gas mixt mi xtur ures es wh whic ich h af afte terr mi mixi xing ng re rema main in in th thee ignitable range has in the past generally been avoided, either by dilution with inert gas or by operating in the nonstoichiometric range. The former for mer req requir uires es add additio itional nal cos costs ts for hea heatin ting, g, cooling, and separation of the inert gas, while thee la th latte tterr gi give vess on only ly lo low w co conv nver ersio sions ns of th thee excess reactants in a single pass. New developments men ts in par partia tiall oxi oxidat dation ion the theref refore ore aim for

 !

umetric productivity oftivity a ity single can exc exceed eed the pro produc ductiv of microreactor conven con ventio tional nal largelar ge-sca scale le rea reacto ctors rs by ord orders ers of mag magnit nitude ude.. However, the challenges to be solved are the equal equ al dis distrib tributi ution on and col collect lection ion of the fee feed d stream str eams/ef s/efflue fluents nts to/ to/fro from m a lar large ge num number ber of  microreactors and the replacement of the catalyst after activity degra degradation dation..

1.7. Period Periodic ic Operation Operation of Fixed-Bed Reactors Industr Indu stria iall fix fixed ed-be -bed d re reac acto tors rs ar aree ge gene neral rally ly operated under operating operating conditions as constant as pos possib sible, le, clo close se to the their ir ste steady ady-sta -state te yie yield ld maximum max imum.. The There re are are,, how howeve ever, r, two not notabl ablee

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

45

exceptions to this rule. One exception refers to heat-integrated reactor concepts, where instead of rec recupe uperati rative ve hea heatt exc exchan hange ge (as in Sec Section tion 1.5 1.5)) regenerative heat exchange is applied and part of the fixed bed is used as a regenerative heat exchanger. This requires periodic flow reversal. Such Suc h hea heat-in t-integ tegrat rated ed con concep cepts ts wit with h per period iodic ic flow reversal are treated in this Section both

reactions. The main goal in all cases is autothermal operation, i.e., to recover the heat of a high-temperature reaction by using part of the fixed bed for regenerative heat exchange between twe en th thee co cold ld fe feed ed an and d th thee ho hott pr prod oduc uctt st stre ream ams. s. Alth Al thou ough gh in init itia iall lly y pa pate tent nted ed in th thee la late te 1930s [73], the breakthrough of periodic flow reve re vers rsal al ca can n be at attr trib ibute uted d to th thee pi pion onee eerin ring g wo work  rk 

for andoffor endothermic reactions. Theexothermic second class periodically operated reactors uses periodic feed cycling to influence or period per iodica ically lly reg regene enerat ratee the cat cataly alyst st bed bed.. The periodic decoking of fixed-bed reactors treated in Sec Sectio tion n 1. 1.6. 6.2. 2.2 2 is on onee ex exam ample ple.. Fur Furth ther er examples are discussed in Section 1.7.3.

of B ORESKOV  and MATROS 39an A gives a sketch of the basic [74–76]. concept,Figure showing adiabatic fixed-bed reactor with feed/exit tubes and two valves for periodic flow reversal [51]. Before start of operation the fixed bed must be heated to above the ignition temperature of  the catalytic reaction. If the reactor is then fed with wi th co cold ld fe feed ed fr from om on onee si side de,, th thee co cold ld fe feed ed ga gass is heated up by the hot catalyst on one side and cooled down by the cold catalyst on the other side. As in Figure 35, two temperature fronts develop and move through the packed bed. In thee fir th first st fr fron ontt th thee ga gass is he heat ated ed up an and d th thee reactants are converted. After a certain period the direction of flow is reversed via the valves and the temperature fronts are pushed back. In the cyclic steady state a hot plateau in the middle of the packed bed thus moves up and down, while the exit and inlet sections of the

1.7.1. Fixed 1.7.1. Fixed-Bed -Bed Reactors Reactors with Periodic Flow Reversal and Exothermic Reaction

The most prominent example of periodically operate ope rated d fixe fixed-b d-bed ed rea reacto ctors rs is the fixe fixed-b d-bed ed reactor with periodic flow reversal for weakly exothermic reactions. Fixed-bed reactors with periodic flow reversal have also been proposed for reversible exothermic reactions and more recent rec ently ly for str strong ongly ly end endoth otherm ermic ic syn synthe thesis sis

Figure 39.  Reverse-flow reactor with direct (regenerative) heat exchange for an irreversible reaction [18, 51]

A) Bas Basic ic arr arrang angem emen ent; t; B) Loc Local al con concen centra tratio tion n andtempe andtemperat ratur uree pro profile filess pri prior or to flowrever flowreversalin salin theperio theperiodic dicall ally y ste stead ady y sta state; te; C) Variation of outlet temperature with time in the periodically steady state

 

46

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

packed bed serve as regenerative heat exchangers. Since the reactor is adiabatic, the heat of  reaction can only be removed with the leaving gas, and the gas exit temperature shows a sawtoot to othh-lik likee va vari riat atio ion n in ti time me (F (Fig. ig. 39 C) C).. Fro From m an overall energy balance it is obvious that the integral of the exit temperature over time must exceed the feed temperature just by the adiabat D ad. ic temperature tempe raturely,rise Interesting Inter estingly, in theT limit of rapid flow rever rever-sal the reverse-flow reactor and the countercurrent re nt fix fixed ed-be -bed d re react actor or of Fi Figu gure re 31 B sh show ow iden id enti tica call te temp mpera eratu ture re an and d co conv nver ersio sion n pr proofiles [51, 52]. This can be explained with the help of Figure 40: With rapid flow reversal the catalys cat alystt tem tempera peratur turee wil willl rem remain ain alm almost ost con con-stan st antt du duee to th thee la larg rgee he heat at ca capa paci city ty of th thee pack pa ckin ing, g, wh whil ilee th thee ga gass te temp mper erat atur uree wi will ll be below the catalyst temperature in the respective feed section and above it in the exit section (Fig (F ig.. 40 B) B).. Th This is be beha havi vior or is eq equi uiva vale lent nt to th that at of  a countercurrent fixed-bed reactor in which the catalyst is located at the separating walls betwee tw een n th thee ba back ck an and d fo fort rth h flo flowin wing g ga gass (F (Fig ig.. 40 C) C).. It need only be considered that instead of pushing in g th thee rea react ctin ing g gas gas fo forr a sh shor ortt per perio iod d in on onee an and d for another period in the other direction, half of  thee ma th mass ss flo flow w wi will ll go pe perm rman anen entl tly y in on onee an and d th thee second half in the other direction. Since it is much easier to solve the steadystate sta te mod model el of the cou counte ntercu rcurre rrent nt fixe fixed-b d-bed ed reac re acto torr th than an th thee dy dyna nami micc mo mode dell of th thee re re-verse-flow reactor, the equivalence can be used for shortcut design calculations for the reverseflow flo w re reac acto torr [1 [18, 8, 51 51]. ]. Th This is is sh show own n in Fi Figu gure re 40

D, temperature profiles for the the limit of  fastwhere flow reversal are constructed from tempera pe ratu ture re sl slop opee as gi give ven n by Equ Equat atio ion n (1 (14) 4) an and d th thee maximum temperature estimated from the ignition temperature   T I  and the adiabatic temperature rise (Eq. 15). Figure 41 A shows the influence influence of the ignition temperature  T I on the temperature profiles. Different ignition temperatures may result from different catalysts or from gradual catalyst deactivation. The reverse-flow reactor (as well as the countercurrent reactor of Fig. 31 B) is selfadaptive to catalyst decay in that a higher maximum mu m te temp mpera eratu ture re is es esta tabli blish shed ed if th thee ca cata taly lyst st is deactivated deac tivated.. Only if the ignit ignition ion tempe temperature rature approaches the temperature  T S  of the intersecting slopes will the autothermal operation break 

Figure Fig ure 40.   Equiv Equivalen alence ce of reac reactor tor oper operatio ation n with peri periodic odical al

flow rev revers ersal al an and d cou counte ntercu rcurre rrent nt he heat at exc exchan hange ge for a we weakl akly y exothermic irreversible reaction A) Fixed-bed reactor with periodic flow reversal; B) Temperature pera ture profil profiles es with rapid flow reve reversal; rsal; C) Coun Counterc tercurre urrent nt reacto rea ctorr wit with h cat cataly alyst st at thewall; D) Sch Schem emati aticc con concen centra tratio tion n and temperature profiles in both reactors [18]

down and the reactor be extinguished. The selfadaptivity concerning different feed concentrations tio ns (Fi (Fig. g. 41 B) wa wass al alre read ady y di disc scus usse sed d in connection with Figure 31 B. The obvious advantage of the reverse-flow reactor over the countercurrent fixed-bed reactor is the excellent efficiency and inherent simplicity of the regenerative heat exchange. Reactions with an adiabatic temperature rise as low as 10 K can be run autothermally at maximum temperature tempe raturess excee exceeding ding 500 C in a properly designed fixed bed. It is therefore presently a favorite design for the catalytic oxidation of 

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

47

Figure Fig ure 42.   Alter Alternati native ve desi designs gns for auto autother thermal mal reac reactor tor with

periodic flow reversal A) Rad Radialflow ialflow con concep ceptt [18 [18]; ]; B) Thr Threeee-be bed d arr arran angem gementwith entwith bed purge prior to flow reversal [51]

Figure 41.  Influence of operating parameters on reactor

behavior A) Temperature and concentration profiles for gases with different ignition temperatures  T I1 I1,  T I2 I2; B) Influence of the 0 0 feed concentration   c1 ,   c2   resulting in a different adiabatic temperature rise  D T ad ad  [18, 51]

traces of combustible components in exhaust air. ai r. It Itss pr prim imee di disa sadv dvan antag tagee is th thee un unst stea eady dy mo mode de of operation and the necessity to switch large gas streams periodically. periodically. In the basic design of  Figure Fig ure 39 39,, un unco conv nver erte ted d ga gass in th thee en entr tran ance ce hood and the preheating section of the packed bed be d is flus flushe hed d in into to th thee ex exit it wi with th ev every ery flow reversal. To avoid this breakthrough most commercial mer cial air pur purifica ification tion uni units ts use a thre three-b e-bed ed design (Fig. 42 B) in which one bed is purged with clean air prior to flow reversal. In addition to the standard design, a number of mo modi dific ficat atio ions ns ha have ve be been en pr prop opos osed ed an and d applie app lied. d. As exp explain lained ed in Sec Sectio tion n 1.3 1.3,, lar large ge adiabatic packed beds are preferably designed

as radial-flow reactors to avoid excessive pressure drop with increasing bed height. In the radial flow design depicted in Figure 42 A, the hot reactor part in the middle insulates itself  against heat losses. This design can easily be extended to a fixed three-bed with rotating beds,arrangement. similar to theDesigns Ljungstro¨ m heat-exchanger design, have been proposed and tested as well [18]. The rotating fixed-b fixe d-bed ed des design ign off offers ers the pos possib sibilit ility y of a continuous, valveless operation at the expense of rotating seals at the cold end of the rotating fixed bed. Typical applications for exhaust air purification are char characte acteriz rized ed by rapi rapid d conc concentr entratio ation n changes. An efficient control strategy must preventt th ven thee re react action ion fro from m ext extinc inctio tion n dur during ing tim times es of  low concentrations and the catalyst from overheating and sintering during times of rich feed. Severa Sev erall pos possib sibili ilitie tiess to ach achiev ievee thi thiss goa goall ar aree discussed in [50]. Also heat losses from the hot partt of th par thee re rever verse se flow rea reacto ctorr sho should uld be lim limite ited, d,

 

48

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

since nonadiabatic operation with reactions of  low lo w exo exoth therm ermic icity ity may lea lead d to sp spuri uriou ouss and eve even n chaotic quasiperiodic oscillations [51, 57]. Aside from air purification, autothermal operatio era tion n wit with h per period iodic ic flow rev revers ersal al has bee been n propos pro posed ed and dem demons onstrat trated ed for a num number ber of  exothermic exoth ermic synth synthesis esis react reactions ions with equil equilibrium ibrium limitat lim itation ionss suc such h as met methan hanol, ol, amm ammoni onia, a, and

bed reactors of Figure 31. Alternatively it could be fed to the reverse-flow reactor of Figure 39. Pionee Pio neerin ring g wor work k of Amo Amoco co on sim simulta ultaneo neous us operat ope ration ion [78 [78]] stu studie died d met methan hanee ste steam am ref reform orming ing together with methane combustion in a reverseflow flo w re reac acto torr co cons nsist istin ing g of tw two o in iner ertt en end d po porti rtion onss and an active center part. Successful operation of several laboratory and pilot plant reactors up

sul sulfur fur trio trioxid synthes Andecreasing advanta adv antage getemperahas bee been n claimed inxide thee syn factthesis. thatis.the ture tu re pr profi ofile le in th thee ex exit it se sect ction ion of th thee re reac acto torr le lead adss to an inc increa reased sed equ equilib ilibriu rium m con convers version ion ove overr tha thatt obtained obtai ned adiab adiabatical atically. ly. Howev However, er, a two-st two-stage age adiabatic reactor with interstage cooling will still lead to highe higherr conv conversion ersion [51, 77] with the additional advantage that the heat of reaction after each stage can be used for the additional production of steam as in Figure 15. Purely autothermal concepts for the above-mentioned bulk chemicals will therefore probably be restricted strict ed to sites where simple reacto reactorr operat operation ion withoutt wastewithou waste-heat heat utilization is requi required. red.

If endo endothermic thermic synthesis synthesis reactio reactions ns are coup coupled led with wit h an exo exothe thermic rmic aux auxili iliary ary rea reacti ction on lik likee a catalytic cataly tic combu combustion stion the combi combined ned react reaction ion can be weakly exothermic. This would allow the combined reactions to be run autothermally if  efficient heat exchange between the hot effluent (s) and the cold feed( d(ss) is provid ideed. In

to 60 cm and an d 4ed. m le leng ngth thper with wi th CO yie ield lds s up to 90% wasIDrep report orted. The period iodic ic stea steadydy-sta state te profiles obtained are shown in Figure 44. They are symmetric in the sense that they are mirror images of each other in successive semicycles, but the they y dif differ fer fro from m the ste steady ady-sta -state te pro profile filess discussed in Figures 39–41. The rapid methane combus com bustion tion res results ults in a dis distin tinct ct tem temper peratu ature re peak, pea k, whi which ch is fol follow lowed ed by the tem temper peratu ature re decrease caused by the endothermic steam reforming. The length of the active catalyst zone and the cycle time must be properly adjusted to preven pre ventt exc excess ess tem temper peratu atures res,, bac back k rea reacti ction on with decreasing temperatures, or coke formation at temperatures below 750 C. This sensitivity obviousl vio usly y led Amo Amoco co to sto stop p fur furthe therr dev develo elopme pment. nt. In the meantime, a similar process has been patented paten ted by Exxon ExxonMobil Mobil [79], where wherein in instea instead d of pre premix mixed ed oxy oxygen gen one or two oxy oxygen gen-co -contai ntainning in g si side de fe feed edss ar aree ad adde ded d to th thee ce cent ntra rall pa part rt of th thee fixed bed. The case of asymmetric operation with recuperative heat exchange (Fig. 43b) has already been discussed in connection with Figure 32, left. Here it was shown that the exothermic and endoth end othermi ermicc rea reactio ction n zon zones es ten tend d to sep separa arate, te, leading to excessive maximum temperatures in

Sectio Section n 1.5.2 options for a respec respective tive coupling using recuperative (indirect) heat exchange are discussed. In this section the analogy between fast flow reversal and countercurrent heat exchange, introduced in Section 1.7.1, is used to discus dis cusss alt altern ernativ atives es for the aut autoth otherm ermal al cou cou-plin pl ing g of en endo do-- an and d ex exot othe herm rmic ic re reac actio tions ns in recuperative or regenerative mode. Figure 43 is a schematic representation of the corresponding alt altern ernati atives ves for the cou coupli pling ng of met methan hanee stea st eam m re refo form rming ing wit with h me meth than anee co comb mbus ustio tion n [51]. In th thee si simu mult ltan aneo eous us op oper erat atio ion n mo mode de of  Figure 43a oxygen or air is ‘‘simultaneously’’ mixed with the steam reforming feed such that the total reaction is weakly exothermic. This mixture could be fed to the autothermal fixed-

case methane steam Similar results sul ts of can be fou found nd for reforming. regene reg enerat rative ive heatt exhea change with periodic flow reversal [51]. The symmetric operation mode of Figure 43c leads to more favorable conditions, since the addition of he heat at th thro roug ugh h a ho hott si side de st stre ream am or an ox oxid idiz izin ing g compou com pound nd trig trigger gerss the end endoth otherm ermic ic rea reactio ction n immediately [51]. Multiple addition points as in Fi Figu gure re 32 (r (rig ight ht)) wi will ll fu furt rthe herr sm smoo ooth then en th thee ho hott temperature zone. More recently, a further alternative of conductin duc ting g hig high-t h-temp empera eratur turee end endoth otherm ermic ic rea reacctions tio ns un unde derr th thee co cond nditi ition onss of pe peri riod odic ic flow reve re vers rsal al ha hass be been en de deve velop loped ed [8 [80– 0–82 82]] (s (see ee Section 1.7.2.2). This process is closely connect ne cted ed wi with th th thee ph pheeno nome men non of mo movi ving ng endothermic reaction fronts.

1.7.2. Fixed-B 1.7.2. Fixed-Bed ed Reactors Reactors with Periodic Flow Reversal for Coupled Exo- and Endothermic Reactions

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

49

Figure Fig ure 43.   Recu Recupera perativeand tiveand corre correspon spondingregener dingregenerativ ativee alter alternati natives ves for the auto autother thermal mal coup coupling ling of meth methanesteam anesteam refo reforming rming

with methane combustion (from [51])

 Moving Endothermic Reaction Fronts   The simple front velocity Equation (19) also applies for the case of endothermic reactions, if we consider that   DT ad ad   is now negative and hence thee fr th fron ontt ve veloc locity ity is al alwa ways ys fa fast ster er th than an th thee thermal front. This is obvious, since the heat of re reac acti tion on mu must st be ta take ken n fr from om th thee he heat at st stor ored ed in the packing. In the following, again methane steam reforming is considered under the additional assumption that reaction equilibrium is quickly qui ckly est establ ablishe ished d at eac each h tem temper peratur ature. e. Hen Hence, ce, conversion is strictly coupled to the local temperature. Figure 45 (top) shows resulting temperature profiles for a catalyst bed preheated

uniformly to 800 K and fed with a cold reaction mixture from the left [80]. At this temperature the equilibrium conversion is only about 45%. As int intuit uitive ively ly exp expect ected ed a str strong ongly ly dis disper persiv sivee front develops, since below about 600 K equilibrium conversion is too low and the thermal front velocity (short arrow) prevails, while with increasing increa sing tempe temperature rature and conve conversion rsion more heat is taken from the packing and the reaction front veloc velocity ity incre increases ases (long (longer er arrow arrow). ). Surprisingly, after the bed temperature has been raised to 1000 K (Fig. 45, middle) and to 1500 K (Fig. 45, bottom), the dispersive part of  the reaction front turns into a pure shock at a

 

50

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

Figure 44.  Methane steam reforming with simultaneous methane combustion in a reverse-flow reactor [78]

The fact thattemperature this shock  certain temperature. travels slower at higher preheating is obvious, since more heat for the reaction is stor st ored ed in th thee pa pack cked ed be bed. d. Fo Form rmat atio ion n of th thee sh shoc ock  k  is first of all a consequence of the equilibrium assumption made and is explained through the thre th reee ar arro rows ws in Fig Figur uree 45 (m (mid iddle dle): ): A pu pure re thermal front prevails at low temperatures since equili equ ilibriu brium m con conver version sion is low (sh (short ort arr arrow) ow),, but also at high temperatures and complete conversion ( X   >  98%) no heat will be required for the reac re acti tion on an and d he henc ncee al also so he here re a pu pure re th ther erma mall fr fron ontt should prevail. Only at intermediate temperatures will most of the heat stored in the bed be used for reaction, causing the reaction front to trav tr avel el fa fast st (l (lon ong g ar arro row). w). Fr From om th thee re resu sult ltin ing g arrows, shock formation becomes evident. This

has first been shown together withitssimplified equations for the shock height and traveling velocity in [80]. The que questio stion n whe whethe therr a suf suffici ficient ently ly sha sharp rp temper tem peratur aturee fro front nt dev develo elops ps for end endoth othermi ermicc reactions is of great practical importance, since, if so, it allows a large part of the heat stored in a preheated catalyst bed to be used for the endothermi the rmicc rea reacti ction. on. Figure Figure 46 sho shows ws exp experi erimen mental tal and simulation results whereby at time zero the reforming feed starts flowing into a sufficiently preheated bed, composed of an inert entrance section followed by an active catalyst bed [57]. The points connected by thin lines are results of  measure mea suremen ments, ts, and the thi thick ck lin lines es sho show w the respective simulations. In addition, the temperature atu re rec record ording ingss at fou fourr dif differ ferent ent loc locati ations ons in the

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

51

Figure Fig ure 45.   Simula Simulated ted tem tempera perature ture profi profiles les duri during ng meth methane ane stea steam m refo reformin rming g in unif uniforml ormly y preh preheate eated d cata catalyst lyst bed beds, s, assu assuming ming

an equilibrium controlled reaction [80]. The arrows mark the thermal front (short) and the reaction front velocities (long)

active catalyst bed are shown in Figure 46 B. They The y cle clearl arly y sho show w a sha sharp rp fro front nt mov moveme ement. nt. Over the whole period of 250 s required by the

profile in Figure 47 (top) should be established. The cold reforming feed, entering from the left, will be heated up in a thermal front. As soon as

front to reach the outlet, methane conver-a sion was achieved (Fig. full 46 C). This allows reve re vers rsee flo flow w pr proc oces esss to be se sett up in wh whic ich, h, during one half-cycle, the reforming reaction take ta kess pl plac acee in an in initi itial ally ly pr preh ehea eate ted d be bed d un unti till th thee reaction front approaches the end of the active catalyst. In the second half cycle the flow direction is reversed and a hot gas stream is used to reheat the catalyst bed.

the sufficiently high, the traveling temperature shock of theisreforming reaction develops. When it reaches the end of the catalytic bed, the production period ends and the flow must be reversed to restore the initial temperature profile. A mere flow reversal with equal heat capacity flux would (in the simplified ideal case cons co nsid idere ered) d) le lead ad to th thee so soli lid d profile profile in Fig Figur uree 47 B, where the shaded area of the profile must be filled up by additional heat input. This can be achieved as shown in Figure 47 C if air or fuel for a combustion reaction is added at discrete positions along the reactor. Each injection initiates itia tes the therma rmall fro fronts nts whi which ch (un (under der the ide ideal al conditions considered) would fill the temperature gaps by the end of the regeneration semicycl cy cle. e. To ac achi hiev evee th this is,, th thee di dist stan ances ces be betw tween een th thee

Coupling of Exo- and Endothermic Reacperi riod odic ic pr proc oces esss in wh whic ich h th thee mo move ve- tions   A pe ment me nt of an en endo dothe therm rmic ic re reac acti tion on fr fron ontt th thro roug ugh ha preheated catalyst bed can be exploited is proposed in [80]. Figure 47 provides a simplified pict pi ctur uree of th thee pr proc oces esss sc sche heme me.. At th thee start start of th thee endothermic production period the solid line

 

52

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

Figure Fig ure 46.  Tempera  Temperature ture profi profiles les duri during ng meth methane ane stea steam m refor reforming ming in a unif uniformly ormly preh preheate eated d cata catalyst lyst bed, cons consistin isting g of an iner inertt

pre-section followed by an active catalyst section A) Measured (points) and simulated temperature profiles (solid lines) at specified times; B) Temperature measurements over time at the location 1–4 specified in A); C) Methane conversion over time [57]

must be to adjacent injection disp di spla lace ceme ment nt of a ports ther th erma mal l fr fron onttequal durin du ring g the thee th regeneration semicycle. Based on this simplified picture, cyclic processes have been set up and studied both experimentally and through detail det ailed ed mod modeli eling. ng. Exp Experi erimen mental tal res result ultss are presented in [81, 82]. Related processes have the potential of carrying out endothermic reactions at high temperatures and with normal or elevated pressure with high energy efficiency and with moderate equipment costs. The critical step of high-temperature heat exchange takes place regeneratively in the inert packings at both ends of the adiaba adi abatic tic pac packed ked-be -bed d rea reacto ctor. r. The hea heatt inp input ut can be ach achiev ieved ed via app approp ropriat riately ely des design igned ed burner sections [80, 82] and the flow-reversal

va valv lves es ar aree op opeera rate ted d at th thee lo low w fe feeed/ d/ex exit it temperatures.

1.7.3. Period Periodic ic Feed Cycling Cycling

The option to increase conversion and/or selectivity by periodic cycling of the feed concentration was studied in the 1990s in a number of  publications for different reactor types including in g fix fixed ed-b -bed ed re reac acto tors rs [8 [83, 3, 84 84]. ]. An id idea ea wa wass th that at through a periodically changing feed composition with relatively high frequency (ca. 1 Hz), a metastable state of the catalyst can be stabilized which would be unattainable under steady-state operation. In catalyst fixed beds with industrial dimensions, however, it is hard to achieve the

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

53

slightly red slightly reduci ucing ng (st (stoic oichio hiomet metric ric oxy oxygen gen exc excess ess ratio   l   <  1) and oxidizing conditions ( l   >   1) partt of the mon par monoli olith th cat cataly alyst st wit with h suf sufficie ficient nt sto torred oxygen can oxid idiize the off-gas compon com ponent ents, s, whi while le par partt of the cat catalys alystt wit with h oxygen deficiency can reduce NO to molecular nitrogen. Several studies have shown that an oxygen oxy gen con conten tentt whi which ch cyc cycles les aro around und   l   1

 ¼

gives a higher conversion of NO x  and neously of oxidizable components thansimultasteadystate operation at  l  1 [96]. A second class of periodic processes uses cycl cy cles es wi with th lo long nger er cy cycl clee pe peri riod ods, s, du duri ring ng wh whic ich ha subs su bsta tanti ntial al ch chan ange ge of th thee ca cata taly lyst st vi viaa ad adso sorp rptio tion n or chemisorption or via conversion of catalyst compon com ponent entss tak takes es pla place. ce. An ear early ly exa exampl mplee is the selective catalytic NO x  reduction of flue gases from power stations, wherein ammonia is periodically fed and stored in the catalyst bed in a reverse-flow process [33]. WellWe ll-es estab tabli lish shed ed,, ag agai ain n fr from om th thee ar area ea of  automo aut omotiv tivee exh exhaus austt pur purifica ificatio tion, n, is the soso-cal called led NOx storage catalyst (NSC) for lean-burn combustion engines. In this catalyst NO is oxidized overr Pt und ove under er fue fuel-le l-lean an (ox (oxyge ygen-r n-rich ich)) con condiditions to NO2. This converts the barium compound of of the catalyst catalyst to Ba(NO3)2 in the storage reac re acti tion on sh show own n in Fig Figur uree 48 48,, wh whil ilee CO2   is released. With increasing time under fuel lean conditions most of the accessible Ba has been converted and some of the NO x breaks through, as shown at the top of Figure 48. Subsequently, engine control is manipulated so that for a short peri pe riod od (c (ca. a. 3– 3–5 5 s) fu fuel el-ri -rich ch ex exha haus ustt wi with th a la large rgerr percentage of unburned CO is produced. Under

 ¼

Figure 47.   A simplified simplified pictu picture re of reverse reverse-flow -flow coup coupling ling of 

an endothermic reaction with an exothermic reaction [80] A) Temperature profiles during the endothermic production period per iod,, sta starti rting ng wit with h thesolidline pro profileand fileand end endingwith ingwith the broke bro ken n lin linee pro profile file;; B) Tem Temper peratu ature re pro profile filess aft after er flow reversal with inert feed of equal heat capacity flux as in the endoth end otherm ermic ic per period iod;; thetemp thetempera eratur turee in theshade theshaded d par partt of the profile cannot be recovered; C) Temperature profiles after flow reversal with four side feed points for a combustion reaction fuel

same metastable state over the whole length of  the cat cataly alyst st bed bed.. Ins Instea tead, d, a cat cataly alyst st pro proper perty ty profile will established whereby the catalyst prop pr oper erti ties es be (e.g (e .g., ., it itss ox oxid idat atio ion n st stat ate) e) vary va ry periodically with time and along the length of  the catalyst bed. One example in which this periodic change in catalyst states is exploited in practice is the so-called  l -controlled operation of automotive three-way catalysts. Here the oxidation of CO and hydrocarbons is achieved simultaneously with reduction of nitrogen oxides. It requires a catalyst with both oxidizing and reducing properties. This is achieved by a catalyst matrix of  ceria/alumin ceria /aluminaa with oxyge oxygen-stora n-storage ge capab capability ility and, dispersed over it, noble metal (nano)particles on which either the oxidation or NO reduction ti on re reac actio tion n ca can n ta take ke pl plac ace. e. By pe peri riod odic ic cy cycl clin ing g the automo automotive tive exha exhaust ust compo composition sition between

these conditions barium nitrate compos com posed ed (se (seee the regene reg enerat ration ionis rapidly reacti rea ction ondein Fig. 48) and is converted back to carbonate while NO is released, causing the NO x  peak at the top of Figure 48. Most of the NO however is reduced by CO on the Rh sites of the catalyst to nitrogen and CO2 is formed. The overall regeneration reaction is given in Figure 48, while the main steps on the different catalyst sites are shown below:

ð

Ba NO3

Þ þ3 CO 2

Ba  BaCO 3

þ2 NOþ2 CO

2

Rh 1 NO CO  N2 CO2

þ NOþH

Rh 2

þ 1 N þH O 2 2

2

2

 

ð23Þ 24

   

ð Þ ð25Þ

 

54

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

Figure 48.   NO x   and NH3  Diesel exhaust emissions behind a NO x  storage catalyst (top) and behind Daimler’s BlueTec

combination [85] of an NSC and an SCR (bottom) during two periodic cycles with a long fuel-lean and a short fuel-rich exhaust

ð26Þ

2. Fixed Fixed-Bed -Bed Reactors Reactors for Liquid-Phase Reactions

ð27Þ

2.1. Introd Introduction uction

Pt=Ba

þ

þ

CO H2 O  CO2 H2

þ

Pt

þ

2NO 5H2  2NH3 2H2 O

   

As a side reaction, CO and water vapor form hydr hy drog ogen en in th thee wa wate terr ga gass sh shift ift re reac acti tion on (E (Eq. q. 26 26), ), which in turn produces ammonia (Eq. 27) and

Usually fixed-bed reactors are either operated with gas gas-ph -phase ase rea reactio ctions ns or in a tric trickle kle-be -bed d mode, mod e, whe whereb reby y a liq liquid uid rea reactan ctantt tri trickl ckles es thr throug ough h

ca caus uses an ammo moni nia peak pe akduced inced theeby th exha ex haus ustteni t (Fi (Fig. g. 48 48, top) to p). .es This Th is am peak pe ak can ca nabe redu re shor sh orte ning ng thee, th regeneration time and/or by an additional oxidation catalyst after the NSC. In an interesting development, Daimler [85] added an SCR catalyst immediately after the NSC (Fig. 48, bottom). During the short fuelrich period almost all of the ammonia produced is adsorbed at the SCR catalyst, and the ammonia pea peak k alm almost ost van vanishe ishes. s. Dur During ing the lon longer ger fuel-lean period NO x  leaving the NSC is convert ve rted ed wi with th th thee st stor ored ed am ammo moni niaa at th thee SC SCR. R. Th This is is a nice example of process integration, combining bin ing the use of dif differe ferent nt cat cataly alysts sts tog togeth ether er with a periodic operation mode. It may point to the directi direction on in which further improvements improvements of fixed-bed reactors can be expected.

the bed froms upw top ard to bottom, while ae gaseous reacta rea ctant nt flow flows upward (  Three-Phas  Three-Phase TrickleTrick leBed Reactors). In this chapter, fixed-bed reactors with reactants in the liquid phase will be consid con sidere ered. d. In con contra trast st to tric trickle kle bed beds, s, suc such h reacto rea ctors rs are gen genera erally lly ope operat rated ed in an upfl upflow ow mode to ensure that the catalyst is completely soaked with liquid and to avoid gas-filled portions of the packed bed during startup or load change. In some cases a gaseous reactant is added to the bottom feed and is consumed in thee re th reac actio tion n or a ga gase seou ouss pr prod oduc uctt is fo form rmed ed during the reaction and is carried out with the liquid liq uid flow flow.. The lim limits its of hyd hydrod rodyna ynamica mically lly stable operation of a liquid-filled fixed-bed reactor with upflow of liquid (and gas) is considered in Section 2.3.

!

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

55

The main advantages over the traditionally applied applie d stirred stirred-tank -tank reactors or slurry bubble columns with suspended catalysts are the wellspecified residence time with minimum backmixing and the fixed catalyst bed, which avoids separa sep aration tionof of the slu slurry rry cat catalys alystt fro from m the effl effluen uentt and its recycling. Since little information on the design and operation of fixed-bed reactors with

layout and operation of respective reactors (see Section 2.4). The reactor layout is primarily dictated by safety considerations which are closely related to the large hold-up of potentially decomposable ab le li liqu quid ids. s. Th Thes esee po poin ints ts ar aree di disc scus usse sed d in Section 2.4.1. Specifically, an alcohol amination tio n re reac actio tion n wil willl be co cons nside idere red d as an ex exam ample ple in

li liqu quid idcharacteristics upflo up flow w ca can n be fo foun und ddifferences in th thee op open enfrom lite li tera ratur ture, e, some and conventional packed-bed reactors with gas flow are reported in the following. Liquid-phase fixedbed reactions are usually operated at elevated pressure (>   100 bar) to keep volatile components ne nts in so solu lutio tion. n. On Only ly in ca case se of we weak akly ly so solu lubl blee gase ga sess as re reac acta tant ntss is an ad addi diti tion onal al ga gass fe feed ed at th thee bottom of the reactor applied. Compared to gas-phase reactions in packed beds, liquid-phase reactions are characterized by diffusivities in the fluid phase being three orders of magnitude lower than in the gas phase and densities densities being two to three orders of magnitude higher. This implies that transport resistances in the liquid-filled catalyst pores have a substantially larger impact on conversion and selectivity than in gas-phase reactions and that flow velocities through the packing are usually two tw o to th thre reee or orde ders rs of ma magn gnit itud udee lo lowe wer, r, al al-though the convective heat and mass flux are in the same range (order of several kilograms per square meter of reactor cross section and second sec ond). ). Som Somee imp implic licatio ations ns for liq liquid uid-ph -phase ase reactions in fixed-bed catalyst are considered in Section 2.2. Because of the substantial increase of the

Section 2.4.2. As shut-down a potentially dangerous tion, emergency due to powersitualoss will be discussed.

por pore e tra transp ort res resista istance nce,prefer , a het hetero erogen geneou eously sly catalyzed cataly zednsport reaction would preferably ably be carried carrie d outt in th ou thee ga gass ph phas asee at el elev evat ated ed te temp mpera eratu ture re an and/  d/  or reduced pressure unless thermal stability of  the reactants or other considerations excludes this option. Typical examples of liquid-phase reactions in packed beds are selective hydrogenations of carbon triple or double bonds or aminations of alcohols. As example of the influence of the catalyst pore transport resistance, the hyd hydrog rogena enation tion of but butadi adiene ene to but butene ene is discussed in the next section. Compar Com pared ed to fixed fixed-be -bed d rea reacto ctors rs with gas gas-phase reactions, liquid-phase reactors contain a sub substa stantia ntially lly lar larger ger amo amount unt of pot potent ential ially ly decomposable material and are often operated under high pressure. This must be considered in

ample containswhich simulation results of a selective hydrogenation should elucidate the case. To increase the reaction rate and the selectivity, the decrease of the diffusive resistance is even more important than in the gas reaction case. Shell-type catalysts with a thin, highly poro po rous us la laye yerr of ac activ tivee ca cata taly lyst st on an in iner ert, t, nonporous carrier are in use both for gas- and liquid-phase reactions. If fully active catalysts are applied the trend goes towards strands of  small diameter with a high void fraction. To provide these strands with the necessary mechanical stability is a permanent challenge for catalyst manufacturers. Genera Gen erally lly,, the con convec vectiv tivee tra transp nsport ort of fluid is directed around the catalyst pellets because of  the comparatively dense inner structure. In [87]

2.2. FixedFixed-Bed Bed Catalyzed Catalyzed Liquid-Phase Reactions Transport throug Transport through h the exter external nal cataly catalyst st boun bounddary layer and through the liquid-filled catalyst (macro)pores usually takes place through molecular liquid diffusion, which is about three orderss of magnitude slower than gas diffusion. order diffusion. Inside the catalyst pellets the transport is confined to the void space (void fraction   eP 0.5) and the transport path is extended by the tortuosity osi ty of the por pores es (to (tortu rtuosi osity ty fac factor tor t  4), whi which ch reduces the effective diffusivity by a further factor of   eP / t  t    8 [86]. This means that the concentrations at the active catalyst sites inside thee pe th pell llet et ma may y di diff ffer er su subs bsta tant ntia ially lly fr from om th thee measur mea surabl ablee liq liquid uid-ph -phase ase con concen centra tratio tions. ns. Although the pellet concentration profiles in the case of a conse consecutive cutive reaction reaction are qualitatively qualitatively similar to the gas reaction case, discussed in Figu Fi gure re 4, th thee sl slop opes es ca can n be mu much ch st stee eepe perr an and d th thee consequences more severe. The following ex-





 ¼  ¼

 

56

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

macroporous pellets with high permeability for the flowing fluid are proposed. Then part of the diffus dif fusion ional al res resist istanc ancee cou could ld be ove overco rcome me by convection through the pellets. Another alternative would be metallic open-foam monoliths or rin rings gs with a thi thin-w n-walle alled d hon honeyc eycomb omb str struct ucture ure in the inside as catalyst support, as shown in Figure 11.

basis of the selective hydrogenation. If no more butadiene is present and hydrogen is still available, butene would be hydrogenated to butane. Forr th Fo this is re reas ason on bu buta tadi dien enee is al alwa ways ys ke kept pt in excess. In the following, simulation results are presented which point at the delicate interplay of transport and reaction in this liquid fixed-bed reactor.

 Example: Liquid-Phase Hydrogenation of   Butadiene.   The selective hydrogenation hydrogenation of  butadiene on alumina-supported Pd can be carried out at elevated pressure in the liquid phase. Figu Fi gure re 49 sh show owss a si simp mpli lifie fied d re reac acto torr flo flow w scheme. The feedstock is a crude C 4  fraction. The contained butadiene should be selectively hydrogenated to butene while preventing subsequent sequ ent hydro hydrogenat genation ion to butan butane. e. The react reaction ion is carried out in an adiabatic fixed-bed reactor, in this case with downflow of liquid with small hydrog hyd rogen en bub bubbles bles dispersed dispersed in the liq liquid uid.. A major portion of the reactor effluent is recirculate la ted d to es esta tabl blish ish a su suffi fficie cient ntly ly hi high gh liq liqui uid d downflow velocity and to further convert the recycled recyc led butadi butadiene. ene. The heat of hydro hydrogenat genation ion is extracted through a heat exchanger in the recirc rec irculat ulation ion lin line. e. Hyd Hydrog rogen en is add added ed to the reactor feed and is completely converted in a sing si ngle le pa pass ss wh while ile bu buta tadi dien enee in th thee fe feed ed is al alwa ways ys in exce excess. ss. At the Pd sur surfac facee but butadi adiene ene is pre prefer ferabl ably y chemis che misorb orbed, ed, dis displa placin cing g but butene ene.. Thi Thiss is the

a schematic of the distributionFigure of Pd50 onshows the alumina support and of the catalyst cycle. Butene and butadiene compete forr th fo thee fr free ee Pd si site tes. s. Th Thee am amou ount nt of ch chem emis isor orbe bed d butadi but adiene ene is den denoted oted q1, an and d th that at of ch chem emis isor orbe bed d butene   q2. The chemisorbed species react with dissolved hydrogen to form butene or butane, respectively respe ctively.. In the simulation model the reaction rates are proportional to  q 1 or  q2 and to the local concentration of dissolved hydrogen. It is assumed that the catalyst pores are completely soaked soa ked with liq liquid uid,, whil whilee gas gaseou eouss hyd hydrog rogen en dissolves in the pore liquid and diffuses to the cata ca talys lystt si site tess wh wher eree it is co cons nsum umed ed.. Du Duee to hydrogen dissolution and consumption, the gas bubb bu bble less in th thee do down wnflo flowi wing ng li liqu quid id be beco come me smalle sma llerr unt until il fina finally lly onl only y tin tiny y ine inert rt bub bubble bless lea leave ve the reactor and are separated from the liquid. Butadiene can be adsorbed at the alumina surface and move by surface diffusion to the Pd sites, where it is chemisorbed and converted to butene. If the butene formed is not displaced from the Pd site by nearby butadiene butadiene and hydrogen is present, it reacts further to form butane.

Figure 49.  Simplified flow sheet for butadiene hydrogenation

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

57

Figure 50.  Sketch of the catalyst surface and of the catalytic hydrogenation cycle on Pd

Thee ab abov ovee sc scen enari ario o ha hass be been en si simu mula late ted d by an Th isothe iso thermal rmal mul multisc tiscale ale rea reacto ctorr mod model, el, usi using ng mas masss balances for all adsorbed or chemisorbed species ci es at th thee ca cata taly lyst st su surf rfac ace, e, an and d by ma mass ss ba bala lanc nces es forr th fo thee tr tran ansp spor ortt in th thee ca cata talys lystt po pore ress as we well ll as in the flowing liquid. All model parameters have a physic phy sical al mea meanin ning; g; ads adsorp orptio tion n equ equili ilibri briaa and diff di ffus usio ion n co coeffi effici cien ents ts we were re ta take ken n fr from om th thee literature.

Figure 51.  Dissolved hydrogen concentration

Steady-state simulation results are given in Figures 51–53, which display the two-dimensiona sio nall pr profil ofiles es of th thee re react actan ants ts al alon ong g th thee ca cata talys lystt pores and along the height of the packed bed.  Dissolved hydrogen concentration  (Fig. 51): As long as hydrogen bubbles are present in the liquid phase the concentration of dissolved hydrogen in the flowing liquid is constant. Since hydr hy drog ogen enat atio ion n of th thee bu buta tadi dien enee do doub uble le bo bond nd is a

 

58

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

Figure 52.  Butadiene concentration in the liquid

very fast reaction and liquid diffusion of dissolved hydrogen is slow, hydrogen concentration drops rapidly along the pellet coordinate  x and an d is co comp mplet letel ely y co cons nsum umed ed in a th thin in ou oute terr la laye yerr of the catalyst pellet.  Butadiene concentration in the liquid  (Fig. 52): Butadiene in the reactor feed is present in surplus. As also seen in the dissolved hydrog hyd rogen en pro profile file (Fi (Fig. g. 51) it is con conver verted ted to bute bu tene ne in a th thin in ou outer ter la laye yerr of th thee ca cata taly lyst st.. Butene But ene for formed med is qui quickl ckly y rep replac laced ed fro from m the catalyst surface by additional butadiene as long as the but butadi adiene ene con concen centrat tration ion is suf suffici ficient ently ly

high in th high thee wh whol olee ca cata talys lystt pe pell llet. et. Al Alon ong g th thee reactor react or length butadiene conc concentrat entration ion drop dropss both in the flowing liquid and in the catalyst pellets, but due to its stoichiometric surplus its conc co ncen entr trat atio ion n in th thee flo flowi wing ng li liqu quid id is we well ll ab abov ovee zero over the react reactor or length. Inside the cataly catalyst st pellet pel lets, s, however, however, its con concen centrat tration ion dro drops ps to zer zero o at about   z  0.4 m. This can be more clearly seen in Figure 53, which shows the amount of  butadiene chemisorbed at the Pd sites.  Butadiene chemisorbed at the Pd sites (Fig. 53): Since butadiene is preferentially chemisorbed misorb ed on Pd, its loading on Pd is remarkably remarkably

Figure 53.  Concentration of chemisorbed butadiene at the Pd sites

 ¼

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

59

high even at low liquid concentrations. Only when the liquid concentration approaches zero does butadiene loading drop substantially and the Pd sites are now covered with the sufficiently available butene. This is of no importance in the center of the pellets, since no hydrogen is present there, but in a thin outer layer at around  z  0.4 m dissolved hydrogen is still available

prevent abrasi prevent abrasion on throu through gh (part (partial) ial) fluidiz fluidization. ation. Masss flo Mas flow w ve velo loci citie tiess in li liqu quid id fixe fixedd-be bed d re reac acto tors rs 2 1 are on the order of kg m s , res result ulting ing in smal sm alll li liqu quid id flo flow w ve veloc locit itie ies, s, us usua uall lly y on th thee or orde derr of several mm/s. Also in packed-bed reactors with liquid upflow, abrasion by fluidization must be reliably avoided. avoide d. Even with the low flow velocities velocities this

(see Fig. can react with chemisorbed butene to 51) giveand undesired butane. Although the above simulation results can not be considered quantitative, they are able to elucidate the complex interplay of slow diffusion and fast surface reaction in solid-catalyzed liquid liq uid-ph -phase ase rea reactio ctions. ns. In the exa exampl mplee dis discus cussed sed they the y poi point nt at the source source of inc increa reased sed but butane ane formation inside the catalyst, although the measurablee butadi surabl butadiene ene conc concentrati entration on in the liqui liquid d phase always stays well above zero (Fig. 52). This detailed insight could thus be helpful for designing desig ning better cataly catalysts sts and reactor conc concepts epts for liquid fixed-bed reactors.

2.3. Upwar Upward d Liquid Flow Flow through through Fixed Beds

is not no t ea easy sy ifeter the th afor ore e me ment ione ned high hi ghly lyd,po poro rous us, smallsma ll-diam diamete re af cataly cat alysts stsntio shall sha ll dbe used, use since sin ce, they have an appreciable pressure loss and low weig we ight. ht. On Onee op opti tion on wo woul uld d be to lim limit it th thee up upflo flow w velocity by using a shallow disk-type reactor bed as in Figure 12 B, but the resulting large diameter is in conflict with the fact that most of  the consi considered dered reactions require high press pressure. ure. In addition, due to the strong transport resistances tan ces,, a suf suffici ficientl ently y hig high h flow vel veloci ocity ty and longer residence time, resulting in slim and tall reactor tubes, is required. One proven option to prevent fluidization of  the catalyst consists of superimposing an inert packing of larger and heavy pellets — a ballast overlay — over the active catalyst bed. In this case ca se wa wall ll fr fric icti tion on of th thee pe pelle llets ts ha hass an ad addi ditio tiona nall stabilizing effect against lifting the packed bed.

Liquid flow thr Liquid throug ough h pac packed ked-be -bed d rea reacto ctors rs is usually directed upwards to prevent formation of gas pockets and to allow gas bubbles to leave thro th roug ugh h th thee to top. p. Th This is is di diff ffere erent nt fr from om ga gass packed-bed reactors, in which the gas generally flows downward to stabilize the packing and

Stabilization of Layered Packed-Beds un der Upflow Conditions.  To understand the basic behavior, liquid upflow experiments with commercial catalyst packings in glass tubes of  sufficient diameter are helpful. In Figure 54 the cataly cat alyst st pac packin king g con consis sistin ting g of sma small ll cat cataly alyst st

 ¼

Figure Fig ure 54.   Liq Liquidupflo uidupflow w exp experi erimen mentt in a cat cataly alyst st bedsupp bedsupport orted ed on a lay layer er of lar largerinertparti gerinertparticle cless andover andoverlai laid d wit with h a ba balla llast st of 

heavier inert particles Left: Lef t: ons onset et of lif liftin ting g the cat cataly alyst st par partic ticlesfrom lesfrom thesuppo thesupport rt lay layer; er; Rig Right: ht: at hig highe herr flow vel veloci ocitie tiess the gapmove gapmovess int into o the cat cataly alyst st layer with partial fluidization of catalyst particles and subsequent collapse of the layered bed

 

60

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

particles has been placed above a layer of large and heavier inert particles which serve as a flow dist di stri ribu butor tor.. At th thee to top p of th thee ca catal talys ystt pa pack cking ing th thee aboveabo ve-men mentio tioned ned bal ballas lastt of hea heavy vy ine inert rt par particl ticles es has been placed. In the experiments the liquid upflow is gradually increased and the change of  thee fe th feed ed pr pres essu sure re is me meas asur ured ed fo forr di diff ffer eren entt heights (weights) of the ballast overlay.

Then th Then thee pr press essur uree in th thee li liqu quid id ph phas asee p( z) and in the pac packin king g (wit (withou houtt sur surrou roundi nding ng liq liquid uid)) s   z can be calculated by

Abovethe a certain flow velocity small gap between inert pre-packing and athe catalyst pack pa ckin ing g oc occu curs rs (Fi (Fig. g. 54 54,, le left) ft).. Wit With h fu furt rthe herr increa inc reasin sing g flow vel veloci ocity ty thi thiss gap mov moves es upwards into the catalyst packing (Fig. 54, right) and trickling of catalyst particles through the gap can be observed. The gap width expands as the flow velocity is further increased until the layered packing collapses and the heavier ballast particles trickle down through the fluidized catalyst packing. The first occurrence of a small gap between the support packing and the catalyst bed (Fig. 54, left) marks the velocity limit which should not be exceeded in order to safely avoid avo id col collap lapse se of the lay layere ered d bed str struct ucture ure,, resulting in increased catalyst abrasion due to fluidization. Force Balance for a Packed Bed under Liquid Upflo Upflow. w.   For the design and operation of  fixed fix ed-b -bed ed re reac acto tors rs un unde derr di diff ffer eren entt up upflo flow w conditions a force balance of the packing can be set up. In the following, following, a simpli simplified fied version of such a balance is discussed which has been proven to be well applicable to a number of  cases. The balance is based upon the following assumptions:

where  r F  is the fluid density,  r K  the density of  the liquid soaked particle packing,   e   the bed void fraction, and   f SF SF   and   f WS WS  are the specific friction forces between fluid and packing and betwee bet ween n wal walll and pac packin king. g. As men mentio tioned ned abo above, ve, the friction friction term in the liquid liquid for force ce bal balanc ancee equals the pressure drop   D p  through the packing, depending on fluid velocity  v F:

ðÞ

¼ r  g þ f  ðev Þ

d p d z

SF SF

F

F

ð28Þ

 z

 ¼ r  g  1  e e þ f  ðev  Þ  f 

ds  d z

SF SF

F

K   z

ð  Þ ¼ D pðv Þ e h

 f SF SF vF

 

WS WS

F

ð29Þ

 

ð30Þ

The actual vertical pressure force  s V  of the pack pa ckin ing g is th thee di diffe ffere renc ncee of Eq Equa uati tion onss (2 (29) 9) and an d (2 (28), 8), th thee pr pres essu sure re of th thee pe pell llet et pa pack ckin ing g reduced by the buoyancy force of the pellets, leading to ds v d z

¼ dðs d z pÞ ¼ ðr  r  Þg  1  1 e  f  ðev Þ  f  SF SF

K 

F

F

 z

WS WS

 

ð31Þ

The res resulti ulting ng wei weight ght for force ce F of th thee pa pack ckin ing g at height   z  of cross section  A  is

ð Þ ¼ s  ð zÞð1eÞ A

F   z

v

 

ð32Þ

The friction force between reactor wall and packing  f WS WS  can be approximated by a correlation used for the wall flow of pellets in silos [88]  4

.

  The flowing liqui phase and phase of pa part rtic icle less liquid aree dco ar cons nsid ider ered ed the to solid be ho homo mo-geneously geneou sly dis disper persed sed in eac each h oth other, er, inc incomompressible, and of constant density. Plug flow is assumed for the liquid. .  Phase fractions of both phases are considered constant over the height of the column; this means that the liquid flow velocity does not changee with height. chang .   Du Duee to li liqu quid id up upflo flow w a li lift ftin ing g fo forc rcee ac acts ts on th thee particle phase which equals the pressure drop through the packing and depends on the flow velocity. .   Betwe Between en pa pack ckin ing g an and d tu tube be wa wall ll a fr fric icti tion on force may hinder the movement of the packing, in particular if the catalyst packing is pressed down by a ballast overlay.

 ¼  l D

WS  f WS

R

ð Þ

tan w s v

 

ð33Þ

where   l   is the ratio of horizontal to vertical stress,  w  the angle of wall friction, and   DR  t  the he tube diameter. Since wall correlations for  l and w   are lacking for liquid upflow systems, the contrib con tributi ution on of wal walll fri fricti ction on is app approx roxima imated ted heuristicall heuri stically y in the follow following. ing. The force balance (Eq. 31) can be solved separately for each packed-bed section (active catalyst and ballast) with respect to the vertical pressure force   F   or   s V   at each height. Initial condition is the respective pressure  s V. Results of th this is ca calc lcul ulat atio ion n ar aree sh show own n in Fi Figu gure re 55 fo forr th thee case ca se wi with thou outt ba balla llast st la laye yerr (l (lef eft) t) an and d wit with h a ballast layer of 39 wt% of the catalyst packing (right).

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

61



Figure Fig ure 55.   Resu Resulting lting pack packing ing pres pressure sure (1 e) r V over the the hei heigh ghtt of the pack packing ing wit withou houtt (to (top p lef left) t) an and d with a ballas ballastt lay layer er of 39

wt% of the catalyst packing (top right) Bottom: schematic of the situation in the packed bed at zero bottom pressure without and with ballast

The graphs show how the resulting pressure s V  increases linearly with increasing packing height  z  at zero flow velocity. With increasing upflow velocity v 0 the lifting force (identical to the pressure drop of the packing) reduces   s V until  s V approaches zero at every height of the packing (green line in Fig. 55, left at  v0  0.64

increa incr eased sed ab abov ovee 0. 0.64 64 cm/ cm/ss up to 0. 0.83 83 cm/ cm/s, s, when the lifting force of the catalyst packing equals the weight of the ballast. This can be considered a conservative limit for the upflow velocity above which instability of the layered packing may develop. Experiments as in Figure 54 have shown

cm/s). [89], isethe point atcurs, which thee on onse sett ofAfter fluidi flui diza zatio tion nthis of th the pack pa ckin ing g oc occu rs, si sinc nce now no w fo forr ea each ch su subs bseq eque uent nt pa pack ckin ing g la laye yerr th thee weight force and the upflow lifting force are equal. equ al. This is sho shown wn sch schema ematica tically lly at the bot bottom tom left of Figure 55. Stabili Sta bilizat zation ion of Pac Packed ked-Be -Beds ds by a Bal Ballas last  t   Layer. The above situation changes if a ballast is placed on top of the catalyst packing, since now sub subseq sequen uentt cat cataly alyst st lay layers ers are hin hinder dered ed from free floating but are pressed together by thee we th weig ight ht of th thee ba ball llas astt an and d th thee pr pres essu sure re dr drop op of  the upflowing fluid. In Figure 55, right, it is assumed that the weight of the ballast pellets is so high and their pressure drop so low that their resulting weight is not affected by the upflow velocit vel ocity. y. The upfl upflow ow vel veloci ocity ty cou could ld now be

that insreality upflow velocitie loc ities can be substantially applied app lied bef before orehigher the lay layere ered d pac packking in g be beco come mess un unst stab able le.. Th This is ca can n be ex expl plai aine ned d by the above-mentioned fact that a ballast weight on top of the catalyst packing compresses the packing and induces radial forces which press the pellets against the wall, as shown schematically at the bottom right of Figure 55. Correlatio la tion n (3 (33) 3) co coul uld d be us usef eful ul to es estim timat atee th thee resulting wall friction force. Unfortunately, no experimentally validated correlations for the  l and   w   parameters are presently available for liquid upflow. Single Single-tube -tube exper experiment imentss under reactor operating conditions should therefore be pe perf rfor orme med d to ve verif rify y th thee st stab abil ility ity of th thee layer lay ered ed be bed d an and d to mo moni nito torr ev evid iden ence ce of ca cata taly lyst st abrasion.

 ¼

 

62

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

Reactorr Layout and Operation Operation 2.4. Reacto Liquid-phase reactions in fixed-bed reactors are usuall usu ally y con conduc ducted ted und under er upfl upflow ow of liq liquid uid to prev pr even entt fo form rmat ation ion of ga gass po pock cket ets. s. Th They ey ar aree mo most st often operated at elevated or high pressure to ensure a sufficiently high liquid concentration of gaseous reactants. As long as the heat of 

enough to prevent hot-spot formation due to decomposition reactions. Such a hot spot could eventu eve ntually ally lea lead d to a rea reactio ction n run runawa away. y. Mod Modelin eling g results with different reactor diameters could lead to a reactor diameter below which such a runaway can be excluded. Since heat-transfer correlations for fixed beds filled with stagnant liquids are not too well founded, it is recom-

react reaction small to moderate mode rateK) (adi (adiabatic abatic tempertemperatureion riseisbelow about 100 and decomposition tio n rea reacti ctions ons of the reactants reactants can be saf safely ely excluded, larger diameter adiabatic packed-bed reactors are generally applied. In ca case se of an in inci cide dent nt th thee li liqu quid id co cont nten entt of th thee reactor can either remain in the reactor or it is dump du mped ed in a re relie lieff ta tank nk.. Sin Since ce hi high gh-b -boil oilin ing g organic components will remain in the pores of the catalyst pellets, washing of the catalyst bed with a suitable solvent is necessary after dumpin dum ping g to pre preven ventt dec decomp omposi ositio tion n and por poree blocking. This may require major investments for the permanent provision of relief tanks and washing liquid. If the liquid hold-up remains in the reactor after an emergency shutdown, the reactor tempera pe ratu ture re mu must st be ke kept pt lo low w en enou ough gh to pr prev even entt an any y decomposition reactions leading to a temperature and/or pressure increase. This implies that thee he th heat at of re reac acti tion on mu must st be sm smal alle lerr th than an th thee he heat at dispersed in the packing and through the walls, whic wh ich h mu must st be mo moni nito tore red d by a su suffic fficie ient nt nu numb mber er of temperature measurements. In case of partially gas-filled fixed beds, stopping the liquid feed will lead to a drop of the liquid level in the fixed bed with the result that part of the catalyst

mended to perform respective runaway experiments men ts with single sin gle tub tubes es und under er rea realis listic tic ope operati rating ng conditions to verify an intrinsically safe shutdown procedure (see also Section 2.4.1).

falls runaway reactionsliquid may be induced in thedry dryand part. If the reaction remains in the reactor, the reactor can be restarted quickly afte af terr th thee ca caus usee of th thee in inci cide dent nt ha hass be been en re reso solv lved ed.. As a drawback, limiting the reaction rate by limiting the synthesis temperature reduces the reactor productivity. For mor moree exo exothe thermi rmicc rea reacti ctions ons and and/or /or hig higher her prod pr oduc ucti tivi vity ty,, co cool olin ing g of th thee re reac acto torr wi will ll be necessary. A multitubular reactor with evaporative cooling like in Figure 18C could be a favorable solution. If in case of an incident with loss of electrical power the flow through the reactor stops unintentionally, evaporative cooling in g wi will ll al allo low w fo forr a co cons nsta tant nt wa wall ll te temp mpera eratu ture re.. It must be checked whether heat transfer through the catalyst packing with stagnant liquid is high

by differential scanning (DSC), by reaction calorimetry or incalorimetry adiabatic decomposition experiments experiments at eleva elevated ted pressure and temperatu per ature. re. Res Respec pectiv tivee pro proced cedure uress are usu usuall ally y specifie spe cified d by nat nation ional al saf safety ety sta standa ndards rds [90 [90]. ]. Suc Such h expe ex peri rimen ments ts pr prov ovide ide va valua luabl blee in info form rmat atio ion n abou ab outt th thee po pote tent ntia iall de deco comp mposi ositi tion on ri risk sk of a specific reaction. But they provide no sufficient information for the reactor design and its operation strategy. A mathematical simulation of the pre-runaway conditions would contain many uncertain assumptions assum ptions conc concerning erning high-p high-pressu ressure re phas phasee equilib equ ilibria, ria, whi which ch det determ ermine ine the liq liquid uid-ph -phase ase comp co mpos ositi ition on,, an and d th thee in inte terp rpla lay y of re reac acti tion on kinetics and heat transport. Heat transport by ther th erma mall co cond nduc ucti tion on th thro roug ugh h th thee li liqu quid id an and d

2.4.1. Safety Issue Issuess

Compared to fixed-bed reactors with gas-phase reactions, a substantially larger mass of reactants is present in liquid-phase reactors both in the liquid filled space between the cataly catalyst st pellets and in the liquid-filled catalyst pores. Since in addition high pressures are quite common, safety issues can be of major concern, in particular if decomposition reactions of the reactants or pro produc ducts ts can not be exc exclud luded ed at ele elevat vated ed temperatures. Special attention must be paid to the above-mentioned case that the flow through the reactor stops unintentionally due to pump failure or loss of electrical power. Then reactants or products of the reaction may react and decomp dec ompose ose fur furthe ther, r, cau causin sing g hot spo spots ts whi which ch could initiate a runaway of the whole reactor holdup. It is th ther eref efor oree es essen sentia tiall to de dete term rmine ine th thee decomposition behavior of the reaction mixture

 

Cata Ca taly lyti tic c Fix ixed ed-B -Bed ed Re Reac acttor ors s

63

cataly catalyst st phase will be superimposed superimposed by natural convection, which in turn depends on the flow prop pr oper ertie tiess of th thee li liqu quid id in th thee pa pack cked ed be bed. d. Furthermore, experiments from pilot plants are not easy to scale up to large plant dimensions. Wall heat capacity and wall heat conduction conduction in small-diamet smalldiameter er but thick thick-walle -walled d highhigh-pressu pressure re pilot pil ot tub tubes es pro provid videe str strong ongly ly dif differ ferent ent con condit dition ionss

increase increa se wil willl bec become ome ava availa ilable. ble. Thi Thiss is val valuab uable le informa inf ormatio tion n for ass assess essing ing the ris risk k of rea reacto ctorr runaway under different load conditions. Depend Dep ending ing on the out outcom comee of suc such h test tests, s, adiabatic reactor design could be used in full scale sca le if exo exother thermic mic dec decomp omposi ositio tion n can saf safely ely be excluded. exclu ded. Otherw Otherwise ise evapo evaporation ration-coole -cooled d reactor concepts and/or emergency shutdown strat-

compared to large-diameter adiabatic production reactors. Thee re Th relia liabi bilit lity y of pi pilo lott pl plan antt re resu sult ltss ca can, n, however, be strongly improved if specially designed pilot rigs with small wall heat capacity and excellent insulation are used. Such reactors can ca n be bu buil iltt fr from om a th thin in-w -wall alled ed tu tube be of mo mode dera rate te diameter, made of highly stress-resistant steel. Thiss tub Thi tubee sho should uld be sup super-i er-insu nsulat lated ed and ins insert erted ed in a high-pressure vessel. Although height and diam di amet eter er of su such ch a pi pilo lott tu tube be wi will ll st still ill di diff ffer er fr from om the large-scale reactor, the reaction can be run with the same catalyst under comparable, fairly adiabatic reaction conditions. Runaway experiments can be safely conducted and recorded, since if the inner tube breaks, the outer vessel will wi ll pr prev even entt an any y re rele lease ase of re reac actan tants ts to th thee environment. Ofte Of ten n th thee re requ quir irem emen entt of eq equa uall pa pack ckin ing g heightt as in the final reacto heigh reactorr can not be fulfilled in th thee pi pilo lott tu tube be.. In sh shor orte terr pa pack cked ed be beds ds wi with th ga gas/  s/  liquid upflow, the liquid flow velocities must be redu re duce ced d to at atta tain in th thee sa same me re resi side denc ncee tim time. e. Si Sinc ncee the rise velocity of gas bubbles will be hardly affected, liquid holdup in the pilot tube will be substantially larger than in the full-scale reactor [95]. This means that the pilot reactor will

egi egies es must mus bell-sc implem imp lement ented. should sho be veri ve rifie fied d int fu fullscal ale e te test stssed. of They thee fina th final l uld plan pl ant. t. Aspects of such a design will be discussed in thee fo th foll llow owin ing g fo forr th thee ex exam ampl plee of al alco coho holl amination.

re rema main in floo floode ded dofaf afte ter r ce cess ssat atio ion n of th thee liq liqui uid doffe feed ed,, whereas part the upper packed bed the larg la rgee sc scal alee re reac acto torr wo woul uld d fa fall ll dr dry. y. Th Ther eree a proceeding reaction can cause a hot spot with subsequent runaway as long as the gas phase contains reactants. Liquid reactants will still be present to a substantial amount in the catalyst pores, but convective cooling by the outside liquid is no longer effective. Runaway experiments with partially filled packed beds should therefore be included in the pilot tests. As result of such experiments the dynamic behavior of the runaway reaction can be quantified, depending on the reaction condition ti onss at sh shut utof offf of th thee fe feed ed.. In Info form rmat atio ion n on induction time and the formation of hot spots or mo movi ving ng re reac acti tion on fr fron onts ts as we well ll as th thee pr pres essu sure re

bed ofthe 10–20 m height. Under state sta tereactors operat ope ration ion temper tem peratur aturee inc increa reases sessteadysteadi ste adi-ly over the reactor height as diethyl glycol is converted. As a si side de re reac actio tion, n, di diet ethy hyll gl glyc ycol ol sl slow owly ly decomposes to carbon monoxide, which is further th er co conv nver erted ted to me meth than anee if hy hydr drog ogen en is pr prese esent nt or to CO2 with water. Both decompositions are exothermic. Under normal operation the side reacti rea ctions ons are neg neglig ligibl iblee sin since ce the ami aminat nation ion reaction prevails. If electrical power fails, the ammonia feed stops, and the remaining diethyl glycol will decompose, resulting in a temperature increase and, due to gas formation, in a pressu pre ssure re inc increa rease. se. The con conseq sequen uences ces of thi thiss incident incide nt depen depend d on the amoun amountt of dieth diethyl yl glycol glycol present in the reactor, which is proportional to

2.4.2. Examp Example: le: Amination Amination of Alcohols

A patent discloses details of the reactor design and operation for the liquid-phase synthesis of  aminod ami nodigly iglycol col and morp morphol holine ine fro from m die diethy thyll glycol and ammonia in the presence of hydrogen [91 [91]. ]. Maj Major or find finding ingss con concern cerning ing saf safee rea reacto ctorr desi de sign gn an and d op oper erat atio ion n ar aree di disc scus usse sed d in th thee following. The ami aminat nation ion rea reacti ction on is exo exothe thermic rmic wit with h an adiabatic temperature rise of 30–50 K. Diethyl glyc gl ycol ol,, am ammo moni nia, a, an and d hy hydr drog ogen en ar aree fe fed d fr from om th thee bottom as a two-phase mixture. To ensure a sufficie suf ficient nt amo amount unt of dis dissol solved ved amm ammonia onia and hydrogen in the reaction mixture, sufficiently high pressures should be applied. At temperatures around 200 C the reaction rate is moderate, allowing a feed of about 2 kg of reactants per lit liter er rea reacto ctorr vol volume ume and hou hour. r. Tra Tradit dition ionally ally,, the reaction is carried out in adiabatic packed-

 

64

Cata Ca taly lyttic Fix ixed ed-B -Bed ed Re Reac acto tors rs

the previous reactor load. At low load the reactor temperature is low and the heat of decompositi pos ition on of the rem remain aining ing diet diethyl hyl gly glycol col wil willl lead to a moderate temperature increase along the packed bed which gradually disperses over the height of the reactor. The reaction dies out. At high load, however, both the reactor temperature and the glycol reservoir will be increased.

ensure uniform plug flow velocities through all tubes of the bundle, and to prevent flow maldistribution through natural convection. Due to the homogeneous reaction, special care is required to pre preven ventt pre pre-- or pos post-re t-react action ion with unc unconontrolled temperature rise in the in- and outflow hood ho odss of th thee re reac acto tor. r. Th Thee di disc sclo lose sed d sl slit it de desi sign gn of  the hood results in a substantially reduced hood

After period hot spot will form fo rmed edan ininduction an ad adiab iabat atic ic reac re acto torar wi with th th the e da dang nger erbe of  a reactor runaway in which both diethyl glycol and morpholine would be decarbonylated and decompose. deco mpose. To exclud excludee such an event, the maxmaximum reactor load and the maximum temperature must be limited to ensure safe dying out of  the reaction if the ammonia feed stops. In the specific case of diethyl glycol amination the risk assessment led to the design of a multitu mul titubul bular ar rea reacto ctorr wit with h wat water-e er-evap vaporat oration ion cooling, cooli ng, princ principally ipally similar to the multitubular multitubular reactor shown in Figure 18 C. The choice of the tube dimensions was based upon mechanical stab st abil ility ity co cons nsid idera erati tion onss an and d ve verifi rified ed in th thee above-ment abov e-mentioned ioned pilot-p pilot-plant lant expe experiments riments.. On loss of electrical energy, a valve in the steam drum of the cooling-water circuit opens to ambient. bie nt. With dec decrea reasin sing g pre pressu ssure re the coo coolin ling g water wat er sta starts rts to boi boill unt until il its tem tempera peratur turee has  dropped to 100 C. The reactor tubes are thus cooled down at an appropriate rate to a safe temperature at which any decomposition can be excluded. It is necessary to provide a sufficient liqu li quid id wa wate terr co cont nten entt of th thee st stea eam m dr drum um su such ch th that at at th thee en end d of th thee st stea eam m re reli lief ef al alll re reac acto torr tu tube bess ar aree still under cooling water. The cited patent [91]

volum volume e with flow-distribu flow-d istribution tion elements tss for form fo rm flo flow w di dist strib ribut utio ion n ov over er elemen alll tu al tube bes of unithee th bundle. Water-evaporation cooling is used to control the temperature in the tube bundle.

conta contains ins a descr description iption of thewith reactor, its startup by activating the catalyst hydrogen, the reacti rea ction on pro proced cedure ure,, and the saf safety ety pro proced cedure uress on loss of energy. Similar reactor concepts can be applied for other liquid reactions with temperature-sensitivee rea tiv reacta ctants nts.. Exa Example mpless inc includ ludee the syn synthe thesis sis of  cyclododeca cyclo dodecadienon dienonee from cyclod cyclododec odecatrien atrienee and N2O, which is preferably carried out between 200 and 250 250 C. Hig High h pressure pressure is req requir uired ed for suf sufficie ficient nt dis dissol solutio ution n of the gaseou gaseouss rea reacta ctant nt N2O and product N 2  in the reaction liquid. An appropriate inherently safe multitubular reactor design is disclosed in [93]. In spite of the fact that th at th thee re reac acti tion on re requ quire iress no ca cata talys lyst, t, th thee re react actor or tubes are filled with an inert metal-ring packing to increase the heat transfer to tube wall, to

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