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ADMIN:

I.W

McGRAW-HILLINTERN ATIO N ALBOOK COM PANYNew Yor kSt Lo uisSan Fra nciscoAuckla ndBeirutBogot aDu sseldor fJoh annesburgLisbonLo ndo nLucern eM adridMexicoMont realNe w Delh iPan amaParisSan Ju anSao PauloSingaporeSydneyTo kyoToronto

YATISH T. SHAHDepartment ofChemica! and Petroleum EnqineerinqUniversity or Piu sburqhPut sburqhUSA

Gas-Liquid-Solid Reactor Design

]O.D.--;- . :

M. E. T. 1.3 . L 1.B R !\ i"~ \

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Th is boo k was set in Times Rom an , Series 327

British Lihrary Cataloging in Puhlication Data

Shah , Y. T.Ga s-l iquid-solid reactor design.I. Chemica l reacto rs - Desig n andco nst ruc tionI. Titl e660.2'83 TP I57 78-40 100

ISBN 0-07-05637G-5

To my parents

191079

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GAS-LIQUID-SOLID REACTOR DESIG NCop yright 1979 by McGraw Hill Inc. All r ight s reserve d.Printed in the United State s of Ame rica . No part of this pub lica tion ma y be reproduced,sto red in a retr ieval system. or t ransmitte d in any form or by any mean s,electron ic, mechanical, pho tocopying, reco rdin g o r ot herwise,with out the prio r permission of the publisher.

I 2 3 4 5 MP 8 0 7 9 8

Printed an d bo un d in the Un ited States of America

1,,

Preface

CONTENTS

ix

1 Practical Systems and Types of Industrial Reactor

,./ I-I Introduction and types of gas-liquid-solid reactionv' 1-2 Important design parameters for a gas-liquid-solid reactorv 1-3 Types of industrial gas-liquid-solid reactor

References

1

169

18

,j 2 Film and Penetration Theory Analyses of Gas-Liquid and Gas-Liquid-Solid Reactions 22

cCIIt:JC't'

2-1 Introduction2-2 Film theory an alysis of gas-liquid-solid reactions2-3 Methods for estimating transport resistances2-4 Heat effects2-5 Recommendations for future study

NomenclatureReferences

./3 Residence-time Distribution and Models for Macromixing in theReactors

3-1 Introduction3-2 Tracers3-3 Methods for obtaining residence-time distribution3-4 Problem areas3-5 Models for macromixing in the reactor3-6 ~TD and scaleup problems3-7 Recommendations for future study


22254749515455

60

60616465699394

100102

l'

vi CONTENTS

4 Mathematical Models for Gas-Liquid-Solid Reactors 105

4-1 Models based on effect iveness of co ntact, with no exte rna l mass-transfer resistan ces (mo dels for tri ckle-bed reactors) 105

4-2 Reactor perfor ma nce based on residence-time distr ibu tion 1124-3 Model when reacta nt presen t in both liquid and vapor phases 1134-4 Models for no niso therma l trickle-bed reactors 1154-5 Models which include external mass-t ran sfer effects 1284-6 Mo de ls for th ree-phase slurry rea ctors 1334-7 Models for the packed-bubble-column gas- liquid reactors 1354-8 Gene ral rem ar ks 1404-9 Recommendati ons for future study 141

Nomenclature 143References 147

5 Laboratory Reactors 149

5-1 Introduction 1495-2 Laboratory gas- liquid-solid reactors 1515-3 Reactors used for gas-solid reactions that ca n be ad apted to

three-phase systems 1605-4 Reactors normally used for gas-liq uid reaction s which ca n be used

for the measuremen t of ab sorption rat es in dilute gas - liquid-sol id slurr ies 171

Referen ces and b ibliog ra phy 175

i ,~ .

6 Dynamics of the Cocurrent-downftow Fix ed-bed Column- - --

6-1 Flow regimes6-2 Pressure dro p6-3 ILiquid holdup~_6-4 Radial and axia l gas and liquid dist ributions6-5 Effect ive ca ta lyst wetti ng6-6 Axial d isper sion6-7 Gas-liquid ma ss transfer6-8 Sol id-l iquid ma ss tran sfer6-9 Heat transfer


7 Dynamics of the Cocurrent-upftow Fixed-bed Column

7-1 Flow regimes7-2 Pressure drop7-3 Gas andliq uid holdups "7-4 Axia l d ispersion in the gas and liquid phases7-5 Gas-liquid interphase mass transfer

180

1801841901992022062 12216220224227

230230232237247251

1

,I

7-6 Liquid- solid mass tran sfer7-7 Hea t transfer

Nome nclatureReferen ces

8 D ynamics of Countercurrent-flow Fi xed-bed Column

8-1 Fl o w reg imes8-2 \ Pressure drop ,8-3 6as a nd liquid holdu ps8-4 Gas- an d liquid-p hase axial d isper sion8-5 Wetted area8-6 G as-liquid mass tra ns fer8-7 Liq uid- solid mass transfer8-8 Heat transfer

Nomencl atureReferences

9 D ynamics of the Gas-Liquid Suspended-solid Colum n

9- 1 Int ro ductio n9-2 H yd rodynamics9-3 G as, liq uid , a nd soli d holdups9-4 Axial dispersion in the gas, liq uid , and soli d phases9-5 G as-liqu id in terface mass transfer9-6 Liqu id- sol id mass tr a nsfer9-7 Wa ll mass transfer in the slurry co lum n9-8 Heat transfer

No mencla tureReferences

Ind ex

CONTENTS vii

261266271273

27527527627728 1292293297297300302

304

304305312328334348356356364367

371

PREFACE

The analysis and design of multiphase reactors is probably the most widelyresearched subject in the area of chemical reaction engineering at the present time.While the subject of two-phase reactor design (i.e., gas-solid and gas-liquid) hasbeen extensively reviewed in numerous texts, no similar treatment of three-phase(i.e., gas-liquid-solid) reactor design is available.

Theonly unified review of three phase operations was published by Ostergaard(Adv. in Chern. Series, vol. 7, p. 71, 1968). Since then, considerable progress hasbeen made on this subject. Numerous reviews (Satterfield, Ci N, AIChE J., vol.21, p. 209, 1975; Go~o, S., J. Levee and J. M. Smith, Catal. Rev . Sci . Eng., vol.15, no. 2, p. 187, 1977; Charpentier, J. C, The Chern. Eng. J ., vol. 11, p. 161,1976; etc.) on various aspects of three-phase reactors have been published. Thismonograph attempts to bring about a more complete and timely review of theentire subject matter.

Three-phase reactors are widely used in hydroprocessingoperations _a!14..f

,iI

x PREFACE

Chapters 2, 3, and 4 review the tools for modeling the performance of three-phasereactors . Chapter 2 evaluates the use of film and penetration theory for thecalculation of absorption rate in three-phase reactors. Chapter 3 describes varioustechniques for characterizing residence time distribution and the models whichtake into account the macromixing in a variety of three-phase reactors. Theconcepts described in these two chapters are vital to the simulation of an entirereactor. Chapter 4 illustrates the development of the mathematical models forsome important pilot scale and commercial reactors. In Chapter 5 some advantagesand disadvantages of three-phase laboratory reactors are outlined.

The modeling and design of a three-phase reactor requires the knowledge ofseveral hydrodynamic (e.g., flow regime, pressure drop, holdups of various phase s,etc.) and transport (e.g., degree of backmixing in each phase, gas-liquid, liquid-solid mass transfer, fluid-reactor wall heat transfer, etc .) parameters. During thepast decade, extensive research efforts have been made in order to improve ourknow-how in these areas. Chapters 6 to 8 present a unified review of the reportedstudies on these aspects for a variety of fixed bed columns (i.e., co-current down-flow, co-current upftow, and counter-current flow). Chapter 9 presents a similarsurvey for three-phase fluidized columns.

The monograph is primarily designed to be used by industrial researchersand graduate students in order to bring them up to date on the state-of-the-artin three-phase reactor design . Th e book can also be used as a reference text forgraduate level courses in reaction engineering.

This kind of an effort could not have been possible without cooperation froma large number of associates. I would first like to thank Mr. H. Taylor of GulfScience and Technology Company for allowing me to present the concepts ofthis monograph as a series of seminars to Gulf scientists and engineers. Onecannot find better critics of the book than the students who read it and hopefullylearn something new from it. My special thanks are to Gary Stiegel andSowmithri Krishnamurthy for proof-reading and finding numerous errors in themonograph. Many helpful discu ssions with John Paraskos, Howard McIlvriedand Norman Carr are gratefully acknowledged.

Acknowledgements are also due to a very efficient group of ladies, Ms.Dolores Persun, Janet Bradley, Quandra Nickols and Susan Mateya of the WordProcessing Center at Gulf Research and Development, whose persistent effortsallowed the preparation of the first and most important rough draft of the mono-graph. Thanks are also due to Mrs . Angela Cheyne for typing the final draft, andher unending effort is gratefully acknowledged. Finally, needless to say, that thisproject would not have been started without the love and understanding of mywife, Mary, for her patience and sacrifice during the many long hours it took toput the manuscript in final form .

Yatish T. ShahJune 1978

I

1

CHAPTER

ONEPRACTICAL SYSTEMS AND TYPES OF

INDUSTRIAL REACTOR

1-1 INTRODUCTION AND TYPES OF GAS-LIQUID-SOLIDREACTION

Reactions involving gas, liquid, and solid are often encountered in the chemicalprocess industry. The most common occurrence of this type of reaction is inhydroprocessing operations, in which a variety of reactions between hydrogen,an oil phase, and a catalyst have been examined. Other common three-phasecatalytic reactions are oxidation and hydration reactions. Some three-phasereactions, such as coal liquefaction, involve a solid reactant. These and numerousother similar gas-liquid-solid reactions, as well as a large number of gas-liquidreactions, are carried out in a vessel or a reactor which contains all three phasessimultaneously. The subject of this monograph is the design of such gas-liquid-solid reactors.

The correct choice of a gas-liquid-solid reactor depends, to a certain extent,on the nature of the reaction. There are three types of gas-liquid-solid reactions.

1. Reactions where gas, liquid, and solid are either reactants or productsII. Gas-liquid-solid reaction with the solid acting as a catalyst

III. Two reacting phases with the third phase inertA. Gas-liquid reaction in packed bed - the solid imparts momentum, better

transfer coefficient, and contactB. Gas-solid reaction with the liquid inert - the liquid acts as a heat-transfer

medium or an agent'for redistributing the concentration of various reactingspecies at the catalyst surface

C. Liquid-solid reaction with inert gas - the gas provides mixing.

2 GAS-LIQUID-SOLID REACTOR DESIGN

Absorption of carbon dioxide in a suspension of lime and thermal coalliquefaction are examples of Type I reactions. In the first example , calciumcarbonate is produced by carbonation of suspensions of lime, whereas, in thesecond example, coal is liquified in the presence of hydrogen and oil to producea host of products. These and several other examples of this type of reaction aresummarized in Table 1-1.

The second type of gas-liquid-solid reaction is the one most often encounteredin the petroleum industry . Hydroprocessing reactions are characterized byreactions between hydrogen, one or more components of the oil phase (such assulfur, nitrogen, vanadium, nickel, etc.), and a catalyst. A large portion of thediscussion of reactor design given in this monograph is most relevant to thesereactions. The reactions may produce a volatile, a nonvolatile, or a mixture ofvolatile and nonvolatile products. Some important examples of this type ofreaction are given in Table 1-2. It should be noted that there could be a reactionsystem such as catalytic liquefaction of coal, where the solid phase could bepresent simultaneously as a reactant (coal) and as a catalyst.

In the third type of gas-liquid-solid reaction, only two of the three phasestake part into the reaction, the third phase being an inert phase. This type ofreaction can be further subdivided into three catagories. Some reactions arestrictly gas-liquid reactions, but they are often carried out in~cked-bedreactQ!soperating under countercurrent-flow conditions. Here , the solid imparts momen-tum transfer and allows better gas-liquid contact and gas-liquid interfacial mass

it

Table 1-1 Examples of gas-liquid-solid-reaction systems where all three phases areeither reactants or products

No. Reaction system Referenc e no. r12

34

5

6

78

91011

Thermal coal liquefactionProduction of calc ium acid sulfite (sulfur dioxide, water, and limestone)

reacting to produce calcium bisulfite and used in the manufacture ofsulfite cellulose

Flotation and special type s of fluidized crystall izat ion processesProduct ion of acetylene by the react ion between water a nd calcium

carbide - desorption of C2H2Production of gas hydrate s in desaturation processes - propane and sea

water produce a solid pha seMelting of gas hydrate or ice crystals - reaction between gas a nd so lid

forming a liquidReact ion of phosphid~ of Ca and Al with wat er (desorpt ion of pho sphine)Manufacture of calcium hypophosphite by the treatment of white phos-

phorus with a boiling slurry of lime (desorption of phosphine, diphos-phine. and hydrogen )

Absorption o f CO 2 in a suspen sion of limeWet oxid ation of active ca rbo n - desorption of ca rbon dioxideBiological and photo-oxidation of suspended or ganic so lids in water

purification

42,79,11312,31,123,124

12960

83

83

6262

5619

4, 100

PRACTICAL SYSTEMSAND TYPES OF INDUSTRIAL REACTOR 3

Table 1-2 Examples of gas-liquid-solid-reaction systems where the gas and liquidare either reactants or products and the solid is a catalyst

No . Reaction system

Oxidation of an aqueous solution of sodium sulfite with copper ionsserving as a catalyst

2 Hydrogenation of sesame seed oil with a nickel-on-silica catalyst3 Hydrogenation of cyclohexanc in an aqueous suspension of 30-lIm

palladium black particles4 Hydrogenation of a-methyl styrene containing a slurry of palladium

black or alumina-supported palladium catalyst5 Hydrogenation of benzene (dilute solution of cyclohexane in benzene)

by 2 percent Pt on alumina6 Hydrogenation of ethylene by Raney nickel particles in a paraffin oil7 Oxidation of SOlon wetted carbon8 Hydrogenation of crotonaldehyde over pelleted palladium-on-alumina

catalyst9 Liquid-phase xylene isomerization on

a. H-Mordcnite (zeolite) catalystb. Silica-alumina catalystc. Dual-function catalyst

JO Oxygen transfer in fermentation11 Carbon dioxide absorption by an aqueous buffer in the presence of

an enzyme (carbonic anhydrase)12 Absorption of oxygen in immobilized enzyme systems13 Absorption of oxygen in an aqueous medium containing activated

carbon14 Catalytic coal liquefaction and upgrading of coal liquids15 Hydrogenation of acetone by Raney nickel catalysts16 Catalytic hydrocracking of petroleum fractions _ri" - Hydrogenation of an aqueous solution of glucose to form sorbitol by a;.

__,s o lid catalyst consisting of nickel on.diatomaceous earth carrier " 'cI X Production of 2-butene-1 Adiol and propargyl alcohol by reaction

between acetylene and formaldehyde in aqueous solution over acopper acetyhide catalyst supported on nickel

19 Hydrodenitrogenation of a lube oil distillate20 Hydrogenation of aromatics in a naphthenic lube oil distillate21 Absorption of SOz in a suspension of magnesium oxide22 Hydrodesulfurization of petroleum fractions

23 Hydrogenation of l-octyne and phenylacetylene in C , to Il-C4 alcoholsand n-C6 to n-Cs alkenes by palladium oxide catalysts

24 Organofunctional group hydrogenation25 Hydrogenation of unsaturated fats using Raney nickel catalysts26 Catalytic hydrogenation of carboxylic acid to form alcohols

u, Reduction of an aqueous solution of adepic acid to producehexane-l.S-diol

h. Reduction of a reaction mixture resulting from cyclohexaneoxidation to produce a mixture of hexane-I.o-diol, pentane-l,5-diol, and butane-Ld-diol

27 Conversion of the oxygen-containing products of propylene oxidationon bismuth molybdate catalyst

Reference no,

24

23107

37,88,96, 107. 110

84,95

7048.7159, 103,104

5447,10922,86117. 118116

7481,87

42.79, 1137520,36.51,53, J02, 1148,10

14.52.99

385112, J2413, 35, 50, 55, 57. 82,92.101,119.120.12115

5.46. 12565

39

4 G AS- LIQU ID -SOLID REACTOR DESIG N

Table 1-2 Con tinued

No . React ion sys tem Referen ce no.

I '

28

29303132

353637

38

39

40

41

42

4344

4546

4748

49

505152

Hydrogenat ion of C 4 hydrocarbons at low temperatures (10--20 ' C ) inthe presence of a noble metal ca talyst - reaction gives high yield(nea rly comp lete hydrogenati on of a cetylene) and high selectivity(only a small loss of butadiene by hydrogenation), a lsocr. Selecti ve hydrogen ation of butad ieneb. Selective hydrogen ati on of methyl acetyl ene a nd propadiene in

propylene feedstocksHydrotreating reactionsDenitrogen ation o f ga s oilsCatalytic hydrogenation of phcnylacetyl ene and sty reneOxidati on o f dilute sol ut ion s (132 pa rts per million I o f formic ac id in

water by a CuO . ZnO cat alystCatalytic ox idation of phenol in aq ueous so luti o n ove r co pper oxideHydrodenitrogenation of var iou s co mpo unds and of a ca ta lyt ically-

crac ked light furnace oilOxidation of acetic acid by co ppe r chro mite ca ta lystsCatalyt ic isome rization of cyclopropaneReaction of ph osphides of Ca an d Al with wat er (desorpt ion of

phosph ine )Hydrogenation of nitr o compounds in the presence of Pt or Pd

catalysts (des orption o f water)Hydrogenat ion of carbo nyl co mpo unds in the presence of nick el

cataly st (deso rption o f water)Reaction betw een CzH z and aqueo us formald ehyde in the presen ce

of copper-bismuth ace tylide cataly st to give butyned iol ,Hydrogen at ion of aqueo us butynediol to butenediol in th e presen ce o f

Ni-Cu-Mn o n silica -based catalystCo nve rs ion of acry lonitrile to aeryla mide using copper chro mite

ca ta lystO xidation of SO, in water co ntai ning M nS0 4 as a catalystP roduct ion of aceta lde hyde fro m oxid at ion of C ZH 4 in a solu tio n of

Cu Cl , co ntaini ng Pd Cl z as a cata lystLiquid-phase esteri ficati on o f terephthal ic acid with methan olHydrog enation o f methyl linolea te in the presence of a palladium

ca ta lys tO xidat ion of sodiu m sulfite with coba lto us sulfate ca ta lys tHydro gen at ion of allyl alcoh ol in the solvents water and ethano l and

in the pr esence of Ran ey nickel ca ta lystHydrogenat ion o f fuma ric acid in the solvent ethan ol a nd in the

presence of Ran ey nickel catalys tHydrogenati on o f a niline to cyclohexyla niline by nickel cata lystsHyd rod eselfuriza tion of narrow-bo iling-ran ge fractions o f gas oi lO xidat ions of sulfide ions (hydroge n su lfide) ( 0 thi osul fate ion s a nd

methyl merca ptan to dim eth yl d isul fide in th e pr esence of ac t iva tedcarbo n

Oxidati o n of aqueou s solut ions of so dium sulfide in th e presence ofactiva ted ca rbon

- ---------------------------

72,73

7777787, 40

9333

769862

30

39

11

II

80

27,4317,49

2,6125

11291

91

4112718

1,11 5, 122

PRACTICAL SYSHMS AND TYPES OF INDUSTRIAL REACTOR 5

transfer. A number of such reactions are listed by Danckwerts ;2 9 they are alsosummarized in Table 1-3 and Table 2-4. Some reactions, such as Fischer- Tropschreact ions, are strictly gas-solid reactions. In these reactions, the liquid does nottake an active part in the reaction but is either used as a heat-transfer med iumor as an agent for redistributing the concentration of the reacting species at thecatalyst surface. Since liquids are bett er heat-co nducting mediums than gases,excessive heat ing of the catalyst in some gas-solid reactions (e.g., Fischer- Tropsch

Table 1-3 Examples of gas-liquid-solid-reaction systems where only two phasestake active parts in the reaction. The third phase is inert.

N o.

I2

3

4

5

6

78(}

10

Rea ct ion systemt

Pol ymer izat ion o f ethylene or propylene in cyc lohexa neCata lytic hydration of olefins H ydration of light o lefins suc h as ethylene,

propylene, and butenes (a t high pre ssure and h igh wa ter-to-olefin ratio Iin the pr esence of ca ta lyst

H ydrogenat ion of ethylene using a large co ncen trat ion of Raney nickelca ta lys t suspended in so lut io n

C 2H 4(gj + H , (g) -> C 2 H,(gjTh e Fischer-Tropsch process Reaction o f carbon monoxide with hydro-

gen in the pre sen ce of a solid ca ta lys t to produce a mixture ofhydro carbo ns, a lco ho ls, aldehydes, ketones, and acids dep ending uponoperating conditio ns and the nature of t he ca ta lyst

Catalytic ox idation of olefin Production o f epoxides such as e thy leneo xide and higher o lefin ox ides by t he oxi da t ion o f olefins in th epresen ce of silver ox ide o n silica -gel car rier - process ap plica ble too ther organic ox idation processes

Ca ta lyt ic hydr ogenation of dio lefins to for m mon o- olefins a nd sa turatesin the presence of a "wash oi l"

Diolefins --+ M ono-olefins --+ Saturates

Isomerizati on of cyclo pro paneH ydrogen at ion o f cro tona ldehydeClea ning o fsa nd filters in water-treat ing plan ts - gas is ine rt a nd provid es

stirring and m ixingGas-liquid reacti on s in packed tower s - so lid is inert, e.g.,

a. Rem oval o f lean H 2S fro rrl a variety o f s treamsb. Rem oval of lean 50 2 fro m a var iety o f s treamsc. Absorption of lea n 503 in aqueous H2S0 4 as well as aroma tic

su bstances for sulfonatio nd. Absorption of nit rous gases NO x in water and aq ueo us alka line

so lutionse. Abso rpt ion o f lean CO C I2 in aqueo us a lka line so lutio nI Removal of ph osphin e from C2 H 2 by a bsorption in aqueous NaOCI

or H 2S0 4

Referen ce no .

16

9, 16, 211, 32, 44,58, 64, 65, 66,67, 68, 69, 85,10597, 108

I II

126103, 10445

3, 29, 90, 106

Man y ot he r reac t ions ar e list ed in Refs. 3.29, 90, and 106 and in Table 2-4.t In reactions 1- 6 th e liquid phase act s as a heat-tran sfer med ium .


reactions) is avoided by using an inert, nonvolatile liquid (oil) as a heat-removalmedium. If a reaction is reversible and approaohes equilibrium, introduction ofan inert liquid phase may markedly increase conversion if the product is relativelymore soluble in the liquid than the reactant. A good example is the hydrationof olefins to the corresponding alcohols in the presence of tungsten oxidecatalyst. 128 These reactions can be carried out either in single-phase flow or inmixed flow. A higher conversion at equilibrium of propylene to isopropanol canbe obtained in a two-phase system compared to a completely vapor-phaseoperation. The improved performance can be attributed to the difference betweenthe solubility of propylene and isopropanol in the excess water present. Theconcentration of the product at the catalyst surface is reduced by the presenceof liquid and, thus, alters the reactant-product ratio in a favorable direction. Insome other reactions, the presence of liquid may redistribute the concentrationsof the reacting species at the catalyst surface.126.128 Some oxidation reactions arealso marked by three-phase operations. Conversion of primary alcohols to thecorr esponding sodium salt of the acid (e.g., air oxidation of ethanol to acetic acid)in the presence of palladium- and platinum-based catalysts is one such class ofreactions. These and some other examples of gas-liquid-solid catalytic reactionsare summarized in Table 1-3.

1-2 IMPORTANT DESIGN PARAMETERS FOR AGAS-LIQUID-SOLID REACTORAn appropriate design and model of a gas-liquid-solid reactor requires theestimation of various transport (momentum, mass, and heat), kinetic, and mixingparameters. Specifically, the following parameters are needed.

1-2-1 Knowledge of Flow Regime and Flow UniformityThe mixing characteristics and the transport processes within a reactor dependstrongly on the prevailing flow regime . The flow regime largely depends on theflow rates of gas and liquid phases (and solids in the case of a three-phasefluidized bed) and their relative orientation (cocurrent upwards, cocurrent down-wards, or countercurrent), the nature, size, and status of the packing material, thefluid properties, and the nature of gas and liquid distributors. In a three-phasefluidized -bed reactor, the flow regime would also be dependent upon the concen-tration of the solids and the length and diameter of the reactor.

The flow r~ime plays a very important role in reactor scaleup. If the dataobtained in the pilot-scale reactor are to be useful for a larger-scale reactor, theflow regime in these two reactors must be the same. The flow regimes in a varietyof fixed-bed operations are described in Chaps. 6 to 8.

Flow uniformities are important for the proper reproduction of the data. Inlarge-scale reactors, uniform distribution of gas can be difficult. Nonuniformitiescan cause channel ing or bypassing, which can be harmful to the reactor per-

IJ. '

b---....------------

~..

PRACTICAL SYSTEMS AND TYPES OF INDUSTRIAL REACTOR 7

formance. Furthermore, the effects of these nonuniformities on the transport andmixing processes would be difficult to estimate. In small-scale reactors, flownonuniformities can be encountered at low flow rates. Flow uniformities can beachieved by adding a calming section before the reactor. The effects of flownonuniformities on the residence-time distribution in a reactor and the reactorperformance are discussed in Chaps. 3 and 4.

1-2-2 Pressure DropThe pressure drop across the reactor constitutes an important parameter becausepumping costs could be a significant portion of the total operating cost. As shownin Chaps . 6 through 9, various transport variables such as gas-liquid and liquid-solid mass-transfer coefficients can be correlated to the pressure drop using theanalogy between mass- and momentum-transfer processes. S!gnific~!)l_ p.ress!!!edrop can also cause large undesired changes in the_p'!!.!i~l.p.r.~_s.!!r.~ ofthereacting

ga~~WilQhiiJie~t~~fo.r: .. ,~-_ . , ...... -- ----.----.--- .

1-2-3 Holdups of Various PhasesThe holdups can play an important role in the reactor performance. For example,in a pilot-scale trickle-bed reactor, the liquid holdup can play an important rolein changing the nature of the apparent kinetics of the reaction. When homogeneousand catalytic reactions occur simultaneously, the liquid holdup plays an importantrole in determining the relative rates of homogeneous and catalytic reactions. Ina three-phase fluidized-bed reactor , the holdup of the solid phase plays an im-portant role in the reaction rate, particularly when the solid phase is a reactant.The gas holdup, of course, always plays an important role in reactor performancewhen a gaseous reactant takes part in the reaction.

The holdup of a phase is usually defined as the volume of the phase per unitreactor volume. However, for a fixed-bed reactor, the gas and liquid holdups areoften defined on the basis of void volume of the reactor. In a fixed-bed reactor,the liquid and sometimes gas holdups are divided into two parts : dynamicholdup, which depends largely on the gas and liquid flow rates and the propertiesof the fluids and the packing material, and static holdup, which depends to amajor extent on the nature of the packing (e.g., porosity of the packing) and thefluids' properties. The relationships between the holdups of various phases andthe system variables for a variety of three-phase reactors are discussed in Chaps.6 through 9.

1-2-4 Residence-time Distribution and Axial MixingWhen the fluid elements pass through the reactor, the exchange of mass betweenthe fluid elements occurs both on a microscale as well as on a macroscale. Themixing process on a macroscale is characterized by the residence-time distributionof the fluid elements. Usually, only the macromixing is considered to have a


significant effect on the reactor performance. In a three-phase reactor, theresidence-time distribution for each flowing phase is measured separately. Thereactor performance must take into account the role of residence-time distribution,which is normally measured by tracer techniques. Various macromixing modelsused to correlate the residence-time distributions of various phases in a three-phasereactor ar..t discussed in Chap. 3.

1-2-5 Gas-Liquid Mass and Heat TransferThe importance of gas-liquid mass transfer on the reactor performance dependsupon the nature of the reaction system and the flow conditions in the reactor.Two important parameters characterizing the gas-liquid mass transfer are thegas-liquid mass-transfer coefficient and the gas-liquid interfacial area. Both ofthese parameters depend on the flow conditions and the.nature and status of thesolid packing. The relationships between gas-liquid mass-transfer coefficients,gas-liquid interfacial area, and the system conditions for various types of reactorsare described in Chaps. 6 through 9.

Estimation of gas-liquid mass-transfer rates also requires the knowledge ofsolubilities of absorbing and/or desorbing species and their variations withtemperature (i.e., knowledge of heats of solution) . In some reactions, such ashydrocracking, significant evaporation of the liquid occurs. The heat balance ina hydrocracker would thus require an estimation of the heat of vaporization ofthe oil as a function of temperature and pressure. The data for the solubility,heat 01 solution, and heat of vaporization for a given reaction system should beobtained experimentally if not available in the literature.

1-2-6 Liquid-Solid Mass and Heat TransferJust as in the case of gas-liquid mass transfer, two important parameterscharacterizing the liquid-solid mass transfer are the liquid-solid mass-transfercoefficient and the liquid-solid interfacial area . Various correlations for theestimation of these parameters under a variety of system conditions are discussedin Chaps. 6 through 9. The importance of liquid-solid mass transfer on the reactorperformance depends, once again, on the nature of the reaction and the flowconditions.

Very little work has been done on the liquid-solid heat-transfer coefficient ina three-phase reactor. Under many industrial conditions, the temperatures ofthe liquid and the adjoining solid particles are assumed to be equal.

1-2-7 Intraparticle Mass and Heat TransferThe methods outlined by Satterfield't" for taking into account the effects of intra-particle mass- and heat-transfer resistances on the effective reaction rate areapplicable to three-phase reactors and, therefore, they will not be repeated here.The importance of these resistances depends upon the nature of the reaction and


the catalyst properties. It should be noted that, unlike those in gas -solid orliquid-solid reactions, the complications in gas-liquid-solid reactors arise whenthe pores of the catalyst particles are filled with both the gas and liquid phasessimultaneously. In this situation, a knowledge of the static holdup along withtotal available pore volume of the cata lyst will allow the estimation of the totalmass- and heat-transfer resistances of the gas and liquid phases within thecatalyst.

1-2-8 Wall Heat TransferLittle is known about the fluid-wall heat transfer in the case of gas -liquid flowin a fixed-bed reactor. Some research on this subject, however, has been carriedout for the specific case of cocurrent downflow over a fixed-bed reactor. This issummarized in Chap. 6. Some work on the slurry-wall heat-transfer rate for athree-phase fluidized bed has also been reported. The heat-transfer rate ischaracterized by the convective heat-transfer coefficient between the slurry andthe reactor wall. Some correlations for the heat-transfer coefficient in a three-phaseslurry reactor are discussed in Chap. 9.

1-2-9 Intrinsic KineticsFor most reaction systems, the intrinsic kinetic rate can be expressed either bya power-law expres sion or by the Langmuir-Hinshelwood model. The intrinsickinetics should include both the deta iled mechanism of the reaction and the kineticexpression and heat of reaction associated with each step of the mechanism. Forcatalytic reactions, a knowledge of catalyst deactivation is essential. Film andpenetration models for describing the mechanism of gas-liquid and gas-liquid-solid reactions are discussed in Chap. 2. A few models for catalyst deactivationduring the hydrodesulfurization process are briefly discussed in Chap. 4.

1-3 TYPES OF INDUSTRIAL GAS-LIQUID-SOLID REACTOR

The types of industrial gas-liquid-solid reactor used in industry can be largelydivided into two categorie s, i.e., one where the solids are fixed and the other wheresolids are in a suspended state (fluidized bed). Although the choice of the statusof the solid depends mainly on the nature of the reaction system, often bothfixed- and fluidized-bed systems are examined for the same reaction system (e.g.,coal liquefaction).

1-3-1 Fixed-bed Gas-Liquid-Solid ReactorsIn principle, the gas-liquid-solid reactor with the fixed bed of solids can beoperated in three ways, depending upon the relative orientation of the gas andliquid flow (see Fig. 1-1).The gas and liquid can each flow cocurrently downwards,


Gas Liquid(a)

Liqu id Gas(b )

(c )

Figure 1-1 Types of gas-l iquid-fixed-bed-solid reactor, (a) Fixed-bed cocurrent downflow, (b) fixed-bed countercurrent flow, (e) fixed-bed cocurrent upflow.

cocurrentIy upwards, or countercurrently (normally, the flow directions are liquiddownwards and gas upwards) .The hydrodynamics and the mass- and heat-transferconditions are different in each of these flow conditions. Some practical examplesusing these types of reactors are shown in Table 1-4A.

One of the most widely-used three-phase reactors is the trickle-bed reactor.This type of reactor is particularly favored by the hydroprocessing industry. In atrickle-bed reactor, the liquid flows down over the packings in the form of a thinliquid film ,.and the continuous gas phase flows in between the packing eithercocurrently or countercurrently. In the normal mode of operation, gas and liquidflow cocurrently downward . In commercial hydroprocessing reactors, the liquidvelocity ranges from a minimum of 3 m h- 1 (corresponding gas velocity rangesfrom 542.5 to 2,712.7 m h - 1 at STP) to a maximum of 91.4 m h- 1 (correspondinggas velocity ranges from 16,300 to 81,075mh - 1 at STP):94 These velocities areabout an order of magnitude higher than those used in pilot-scale hydroprocessingtrickle-bed reactors.

-b


Table 1-4 Practical examples

A. Fixed-bed reactorI. Tr ickle-b ed reactor

a. Catalytic bydrodesulfurizatjonb. Catalytic hydrocrackingc. Catalytic hydr otreatingd. Ca talytic hydrogenation such as diolefin hydrogenation, hydrogenat ion of va rious petro leum

fractions , hydr ogenation of lubricating oils, hydrogenation of nitro-compounds, carbonylcompounds, carbo xylic acid, benzene, a-methyl styrene

e. Production of calcium acid sulfite - Jenssen tower opera tionf . Synthe sis of but ynediolg. Production of sor bitolh. Oxidation of formic ac id in wate ri.~urries of act ivated carbonj . Hydrogenation of aniline to cyclohcxylaniline

2. Cocurrent-upfiow reactora. Coal liquefaction (SY NT HO IL react or )b. The Fischer-Tropsch processc. Selective hydrogenation of phenylacetylene and styrened. Cat alyt ic hydrodesulfur izat ion

3. Segment ed-bed reactors .a. Coal liquefaction (Gulf process)b. Fermentation reactions

B. Gas-Iiquid-ssuspended-sotid reactors1. S~ or ftui dized-bed reactors

a. Product ion of ca lcium acid sulfite - fluidized-bed react orb. Cata lytic hyd rogenation of car boxy lic acid - slurry reactorc. Th e F ischer- Trops ch process - slurry reacto rd. Catalytic oxida tio n of olefins - slurry reactore. Cat alytic hydration of o lefins - sl~rY__reacto rf. Polymerizat ion of ethylene - sll.!!2'. reactor

q ' -9~~d ~!]..i.(I,~~ter - tre~!i!ls..J)lants - fluidized-bed react orh. Flu idized crysta lliza tio n processi. Coal liquefact ion (H-CO AL process, SRC process) - fluidized-bed reacto rsj . Absorpt ion of S02 in a suspension of limestone part icles - slu rry reacto rk. Manufa cture of calcium hyd rophosphite by treati ng white phosph o rus with a boiling slurry

of limeI. Liquid-phase xylene isomerization - slur ry reactor

m. Cat alytic hydrogenation of x-methyl sty renen. Catalytic oxida tion of sodi um sulfite

2. Agita ted-slurry reactora. Cat alyti c hydrogenat ion of unsaturated fats and fatt y oi lsb. React ion between HCl and CH 30H in the presence of ZnCl 2 ca ta lystc. Hydrogenation of ace tone

Several advantages and disadvantages of a tr ickle-bed reactor are listed inTable 1-5. The commercial trickle-b ed reactors are operated under plug-flowconditions.~aJ}'S1sare effectively.wetted. These factors allow high conversionto be achieved in a single reactor. The liquid-to-solid rati o (or liquid holdup) ina trickle-bed reactor is small, thus minimi zing the importance of homogeneous


reactions, This may be important in th(h'#rooeSUIfiiEi7,ation reacti011...y.'here lowliquid holdup minimizes the importance of thermal or hydrocracking of the'oil.Normally, in a gas-liquid-solid operation, both the gas-liquid and' the liquid-solidinterfacial mass- (or heat-) transfer resistances are important and are treatedseparately. Because of the thinness of the liquid film in a trickle-bed reactor, thesetwo resistances can be combined and the overall resistance of the liquid filmwould be smaller than those obtained in other types of gas-liquid-solid operation .The trickle-bed operation is normally carried out under cocurrent-downflowconditions; hence, flooding is not a problem in such a reactor. The trickle-flow I,operation gives a lower pressure drop than bubble-flow (cocurrent-upflow orcountercurrent-flow) operation. The low pressure.drop.also .allows for a uniformpartial pressure of the gaseous reactant (i.e., hydrogen in hydroprocessingoperations) in the reactor. This would be important for ensuring hydrogen-richconditions at the catalyst surface along the entire length of the reactor. Hydrogenstarvation at the catalyst surface is known to cause rapid decay of the catalyst.

The trickle-bed reactor can be operated as a partially or completely vapor-phase reactor. It minimizes the energy costs associated with reactant vaporization.Mixed flow conditions at the' catalyst surface exist in hydrocracking reactions,hydrogenation of crotonaldehyde and isomerization of cyclopropane. Whenthe temperature rise in a trickle-bed reactor is significant (e.g., hydrodesulfurizationand hydrocracking reactions), it can be conveniently controlled by the addition

Table 1-5 Advantages and disadvantages of trickle-bed reactors

AdountaqesI. Flow is close to plug 110w, allowing high conversion to be achieved in a single reactor.2. Liquid-to -solid ratio is small, minimizing the homogeneous side reactions if possible.3. Liquid flo'ws as a film, thus offering very small resistance to the diffusion of gaseous reactant to

the catalyst surface.4. Flooding is not a problem. Pressure dr op is lower than in cocu rrent-upflow and countercurrent-

flow reactors .5. If temperature rise is sign ifica nt, it may be controlled by recycling the liquid product or liy the

addition of "quenches" from the side of the reactor. The recycling of liquid would cau se thereactor to behave more like a CSTR; hence . recycling will not be possible when high conversionsare desired.

6. Can beoperatcd as a partially or completely vapor-phase reactor. A trickle-bed reactor minimizedr. the energy costs assoc iated with reactant vaporizat io n.(y Lower pressure drop will allow an essentially uniform partial pre ssur e of reactant across the

length of the reactor.8. In the commercial reactor, uniform distribution of gas and liquid are achieved. The catalyst is

uniformly and effectively welted by the liquid.

DlsaduantaqesI . Poor radial mixing of heat .2. At low liquid flow rates , flow rnaldistributions such as channeling, bypassing, and incomplete

Gcatalyst wetting may occur. These adversely affect the reactor performance.r. 3. The cat alyst particles cannot be very small. The intraparticle-diffusion effects can be significant., The catalyst pore-mouth plugging can cause rapid deactivation.

~ \

I.~


of one or more streams of "quench fluids" (normally hydrogen) along the lengthof the reactor.

A major disadvantage ol...the.trickle-bedreactor is the poor radial mixing ofheat in large-scale.reactors. The poor mixing can cause excessive localized heatingof the catalyst. The excessive heating has harmful effects, in that it causes rapiddeactivation of the catalyst and excessive vaporization of the liquid film. In atrickle-bed reactor, the catalyst particles cannot be very small, because they cangive a large initial pressure drop and a faster increase in pressure drop as thecatalyst ages. The large particles will give pronounced intraparticle diffusion effectsand, in processes such as hydrodesulfurization, the catalyst pore-mouth pluggingwould casue rapid deactivation of the catalyst. Finally, in trickle-bed reactorsoperating at low liquid flow rates, flow maldistributions such as channeling, by-passing, and incomplete catalyst wetting may occur. This can adversely affectthe reactor performance:/.>

For some reactions listed in Table 1-4A, the fixed-bed reactor is operatedunder cocurrent-upfiow conditions. Unlike the trickle-flow condition, this type ofoperation is normally characterized by bubble-flow (at low liquid and gas rates)and pulsating-flow (at high gas flow rates) conditions. Normally, the bubble-flowconditions are used. In the SYNTHOIL coal-liquefaction process, both pulsating-and spray-flow conditions are used , so that the solid reactant (coal) does not plugthe reactor. In bubble flow, the gas is the dispersed phase and the liquid is acontinuous phase. In pulsating flow, pulses of gas and liquid pass through thereactor. In the spray-flow regime, the gas is a continuous phase and the liquidis a dispersed phase.

A comparison between the cocurrent-upflow and the cocurrent-downflowreactor is shown in Table 1-6. This comparison shows that the upflow reactorgives better mixing (both radial and axial) , higher gas-liquid mass-transfercoefficients, higher liquid holdup, better liquid distribution, better heat transferbetween liquid and solid, lower concentration of solid particles, less solids plugging(e.g., coal liquefaction), and , sometimes, better aging of the catalyst than thedownflow reactor under equivalent flow conditions. However, it also gives higherpressure drops (total and partial pressure of the reactant) , poorer conversion (dueto axial mixing), more homogeneous reactions, and more intraparticle diffusionaleffects than downflow operation. Flooding is not a problem in cocurrent-downflowoperation, but it could, however, be a problem in an upflow operation,

The countercurrent-flow fixed-bed operation is often used for gas-liquidreactions rather than gas-liquid-solid reactions. Examples of reactions using thistype of reactor are given by Danckwerts.l? A comparison between a gas-liquid-solid (catalytic) fixed-bed reactor and a gas-liquid-solid (inert) fixed-bed reactoris shown in Table 1-7. The major difference between packed-bed gas-liquidreactors and gas-liquid-solid catalytic reactors is in the nature and size of thepacking used and the conditions of gas and liquid flow rates. The packed-bedgas-liquid reactors use nonporous, large-size packing, so that they can be operatedat high gas and liquid flow rates without excessive pressure drop. The shape of

,14 GAS-LIQUID-SOLID REACTOR DESIGN

Table 1-6 Upftow versus downflow cocurrent fixed-bed reactors

I. Larger pressure drop in an upflow reactor .2. Better mixing in an upflow reactor. This may give better heat transfer. but larger a xial mixing

would give poorer conver sion in an upflow reactor.3. At low flow rates upflow behaves like a bubble column. i.e.. gas as a dispersed phase. liquid as a

continuous phase. In downflow trickle-bed operation. gas is a continuous phase a nd liquid flowsas a film.

4. High pressure drop in an upflow reactor would cause significant drop in the partial pressure ofthe reactant across the length of the reactor.

5. Under similar flow conditions. a higher gas-liquid mass-transfer coefficient is obtained in anupflow operat ion than in a downflow operation. .

6. High liquid holdup and liquid-to-sol id rat io in an upflow reactor. High liquid holdup will offermore liquid-phase resistance to the mass transfer of the gaseous reactant to the catalyst surface .High liquid-to-solid ratio will give more importance to the role of possible homogeneousreactions.

7. At low liquid flow rates , upflow will pro vide better distribution of liquid and, thus . in many cases,better performance of the reactor than the downflow reactor under similar operating conditions.

8. If reaction is rapid and highly exothermic. heat transfer between liquid and solid is more effectivein an upflow reactor.

9. In an upflow reactor, the catalyst must be kept in place by suitable mechanical methpds, otherwisethe bed will be fluidized. In a downflow reactor, the catalyst is held in place tightly by the flow.This may cause undesired cementation of the soft cataly st particles.

10. In an upflow reactor , the catalyst pores are more likely to fill completely with liquid than in adownflow reactor. The catalyst effectiveness factor is lower when the catalyst pores are completelyfilled with liquid compared to the case when they are o nly partially filled with liquid.

11. Better sweeping of the catalyst by liquid in an upflow reactor may sometimes give better agingof the catalyst. If a solid reactant ' is used (e.g., coal liquefaction) then an upflow would causeless solids plugging problems than the downflow oper ation.

12. In an upflow reactor, flooding may be a problem.

Table 1-7 Gas-liquid-solid (catalytic) fixed-bed reactor versus gas-liquid-solid(inert) fixed-bed reactor

1. Gas-liquid- solid catalytic (GLSC) reactors are usually run with cocurrent flow of both gas andliquid . Gas-liqu id reaction in an absorption tower (GSLI) is often operated under countercurrent-flow conditions.

2. The function of solids in a GSLI reactor is to impart momentum tran sfer and better contact betweengas and liquid .

3. The GSLI are often operated under very higb gas and liquid flow rate s (near flooding ) comparedto the ones used in GSLC (in particul ar tr ickle-bed) reactors . An exception is the SYNTHOILreactor for coal liquefaction.

4. In order to obtain high flow rates, the pack ings for GLSI reactors are large compared to the onesused in GSLC reactors. Catalysts in GSLC reactors are porous and as small as possible to avoidintraparticle diffusion effects. The packing in GSLI reactors are nonporous. The liquid holdup inthe catalyst pore s (usually called static holdup) may contribute significantly to the overall liquidholdup in the reactor and it would also affect the liquid -phase residence-time distribution andbackmixing in GSLC reactors.

5. The wetting characteristics of the catalyst particl es in a GSLC reactor may be substantiallydifferent from the wett ing characteristics of packing s in a GSLl reactor.


the packing used is also designed to give larger gas-liquid interfacial areas. Thepacked-bed gas-liquid reactors are often operated near flooding conditions.

In recent years, segmented fixed-bed reactors have also been used for avariety of applications. A few specific types of segmented fixed-bed reactor areshown in Fig. 1-2. Although, in this figure, both gas and liquid are shown to flowcocurrently upward, they could, in principle, flow cocurrently downward or in acountercurrent fashion. The segmented bed ofthe type shown in Fig. 1-2(a)wouldgive a reactor with different mixing zones. It has been examined in connectionwith gas-liquid reactions. The segmented-bed reactors shown in Fig. 1-2(b) and(c) are used for the catalytic, liquefaction of coal.i ' These configurations permit

t tGas Liquid

(a)

Gas Liquid

t t

...

Liquid

t

tLiquid

(b)

t tGas Liquid

(d)

Gas

t

tGas

Catalyst

Catalyst

tLiquid

Liquid

t

(e)

t

tGas

Gas

Catalyst-,-

Catalyst

.1

Figure 1-2 Various types of segmented fixed-bed gas-liquid-solid reactor. (a) Horizontal segmentsof bed, (h) vertical segments of bed, (e) annular segments of bed, (d) catalyst impregnated at the wall.

,16 GAS-LIQUID-SOLID REACTOR DESIGN

Table 1-8 Advantages and disadvantages of segmented-bed reactors

1. Allows more flexibility of mixing characteristics in the reactor.2. In a vertically-segmented bed, the three phases can be transported without plugging the react or.3. Belter flexibility of liquid-to-solid (catalyst) ratio, thus allo wing better variation in homogeneous

and heterogeneous reaction rate s when both are possible .4. High liquid-to-solid ratio. thus allowing more homogeneou s reactions. This may not be desirable.

High liquid holdup will also offer more resistance to the transfer of gas \0 the ca talyst surface.5. Poor liquid distribution and mixing in the catalyst baskets if they are large and suspended

vert ically.

the three-phase reactors to be operated without plugging of the catalyst bed . Theopen section of the reactor behaves like a well-mixed column. In order to usethe catalyst (which is packed in vertically-suspended baskets) effectively, goodradial mixing within the catalyst baskets is desirable. This is achieved by theintense agitation in theopen sections of the column. It is clear that in order touse the catalyst surface effectively, the size (diameter) of the screen baskets shouldnot be very large.

The segmented-bed reactor allows better flexibility of liquid-to-solid (catalyst)ratio, thus allowin g better variations in homogeneous and heterogeneous reactionrates when both are possible. High liquid holdup will, however, offer moreresistance to the transfer of gaseous reactant to the catalyst surface.

The segmented-bed reactor of the type shown in Fig. 1-2(d) is useful whenthe reaction requires only a small amount of catalyst activity and a high degreeof mixing.This reactor would have good heat-transfer characteristics, which wouldmake it useful for highly exothermic reactions. This type of reactor will allowmore homogeneous reactions to occur if they are possible. Some advantages anddisadvantages of the segmented-bed reactors are summarized in Table 1-8.

1-3-2 Gas-Liquid-Suspended-Solid ReactorsSome practical examples using this type of reactor are illustrated in Table 1-4B.This second major type of gas-liquid-solid reactor can be further subdivided intofive categories :

1. Agitated gas-liquid-suspended-solid reactors2. Nonagitated three-phase slurry reactors3. Nonagitated three-phase cocurrent-upflow fluidized-bed reactors4. Nonagitated three-phase countercurrent-flow reactors (spouted-bed reactors)5. Pulsating three-phase reactors.

The pulsating three-phase reactor has been examined only at the laboratorylevel. The pulsation gives good mixing and qeat- and mass-transfer characteristicsin the column. The first three types of gas-liquid-suspended-solid reactor are themost commonly used in practice. Schematic diagrams for these reactors are shownin Fig . 1-3(a), (b), and (c), respectively. The agit ated and nonagitated slurry

r

,.

I

r. PRA CTICAL SYSTEMS AN D TYPES OF INDUSTRIAL REACTOR 17

..-Gas

Solid G~s

bubble

,0~ 0

(a)

I

Gas LiquidGas

t tGas tt 0 0 00 0

0 0 0 SolidSolid 0 0. Q. 0 QGas 0

'V 0 o 0bubble 0 0 0 0

'\)0

1 1~ Gas Liquid

(b ) (c)Figure 1-3 Schematic diagrams of some gas- liquid-suspended-solid reactors. (a) Agitated slurryreactor. (b) nona gitated slurry reactor, (e) fluidized-bed reactor,

-..

reacto rs are batch reactors in which the liquid does not flow through the reacto r.These reactors are used when a small quantity of product is required. The con-tinu ou s three-phase fluidized-bed react or is used for Fischer-Tropsch and catalyt iccoal-liquefaction (H-COAL) processes. Several advantages and disadvantages ofslurry and fluidized-bed reactors are listed in Table 1-9. The major advantages .

~f the gas-liquid-suspended-solid reactors are. that they giVe' b etter flexibility ofmixing , heat recovery, and temperature control.They allow the use of fine.cataJystparticles.whichmiaimize the intraparticle diffusion-effects. Such reactors can beeffectively used for a reaction which involvesa rapidly decaying catalyst and~ph.aseJ:EidiD:ns::in.v_alYitigbo~h sol id reactant ~D_d solid catalyst (e.g.,catalyticcoal liquefaction) . These reactors, however~- give poor conversion due to axialmixing. The separation of cataly.st from the product mixture may cause problems .


Table 1-9 Advantages and disadvantages of slurry or fluidized-bed reactors

AdvalltagesI . High heat capacity providing good temperature control.2. Potentially_Q!g~ r~action.J2l~~r unit volu~~.()r reactor i~ the catalyst is highly active,3. Ea se of heat recovery,4. Can be easily used as a bat ch (slurry) reactor or continuous-flow (fluidized-bed) reactor.5. The ca talyst ca n be easily removedand replaced if it decay s rapidly.: Steady-state operation can

....!!~ .i!chieved even in a rapidly decaying system.6, l~_!!llow~~_~e._~.'5rY.W1e.calPJ.x~.P~s. which cangive an effectiveness fact~r:lppr()~ c.~!!.1g

unity. This is especially important if d,ifll,lsipn limitations cause rapid cat alytic deactivation or' ''po orer s~ I~~.i i\,l!.Y. ' -

7:1i'aliows three-phase gas-liquid-solid (reactant) reactions to operate in the presence of a solidcatalyst Y.:iUiouti!~!i.!g.of tJi~.r.c::.ac~~ e.g., the H-COA L process for coal liquefaction.

8. It allows more flexibility for mix ing, e.g., agitated slurry reactor.

DisadvalltagesI. High degree of axial mixing reduces conversion, High degree of conversion is obtained only by

staging several reactors in series.2..Ca~~t_~eP:l,r~.t i?n fro~.0:_prodll~ ..!!:il'ture by filtrat~~_I~.l~yy~s~.pr_o.~J'l1). .9LQ\!JU i.!!&.!..h....

filters . The cost of filtr at ion maJl be-'experiS~ - -- .3, 'Thchigh -ra tio cifllq-;;i'd"t~'~~i id ;;'ay-aTIo;-homogeneous side reactions to become important,

if they are possible,4. High liquid holdup may cause the liquid-phase diffusional resistance to the gaseous reactant to

be an important facto r affect ing the global rate of reaction.

e,.~ .... :..~:.. \-~ .:,,,,-;..~~)rt:t- {~,cc'-I.C: ~" . re""l~\ .C."''' '\u e: 'y..l:':' ,.~'-e~.C!' ~ .

HIgh llquid holdup may also give significant homogeneous reactions and signifi-cant mass-transfer resistances for the gaseous reactant.

The three-phase continuous countercurrent fluidized-bed reactor and thespouted-bed reactor have been used on the laboratory scale . Pruden and Weber8 8have shown that the countercurrent mode of operation for hydrogenation ofa-methyl styrene performs better than the cocurrent fixed-bed operation undersimilar reaction conditions .

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\'\


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1970.31. Ericsson, E. 0., J . L. McCarthy, and D. A. Pearson , in Pulp and Paper Science aHd Technology

Ie.E. Libby, ed .), McGraw-Hili, New York, 1962, vol . 1, p. 240 .32. Farley, R., and D. J. Ray, J. lnst, Petrol., vol. 50, p. 27, 1964.33. Flinn, R. A., O . A. Larson, and H. Beuther, Hydrocarbon Proc. Petro Refiner; vol. 42, no . 9,

p. 129, 1963.34. Freifelder, M ., Practical Catalyt ic Hydrogenation- Technjque.~ and Application, Wiley Interscience,

New York, 1971.35. Frye, e. G. , and J. F . Moshy, Chern . Eng. Proqr.; vol. 63, no. 9, p. 66, 1967.36. Gerdes, K. F ., B. E. Strangeland, G . T. S. Chern, a nd R. J. Gurnerrnan, "Simulation of Hydro-

cracking Reactor System s" (pe rsonal communica tion).37. Germain, A. H ., A. G . LeBabvre, and G. A, L'H omme, Chern. React. Enq.. 11 Adv. Chern. Ser., no .

133, p. 164, 1974.38. Gilbert, J . B., and R. Kartzrnark, Proc. Amer. Inst . Petr.; vol. 45, no . 3, 1965.39. Gorshkov, A. P., I. K. Kolchin, A. M. Gribov, and L. Ya. Margolis, Killet. Katal ., vol. 9, p. 1086,

1968.40. Goto, S., and J . M. Smith, AIChE J., vol. 21, no. 4, p. 714,1 975.41. Govindarao, V. M. H., and K. V~ Ramana Murthy, J . App!. Chern. Biotechnol., vol. 25, p, 169,

1975.42. Gui ll, J ., A. Tarrer , L. Taylor, Jr.. J . Prather, and S. Green, Jr ., 1&EC Process Design Dev., vol. IS,

no . 4, p. 490, 1976.43. Gunn, D. J ., and A. Saleem. Trans. lnst . Chern . Eng., vol. 48, p. 146, 1970.44. Hall , e. C, D. G all , and S. L. Smith, J . lnst . Petrol ., vol. 38, p. 845, 1952.45. Halvorson, H. 0 ., in Encyclopedia of Chemical Technology (R. E. Kirk and D. F. Othmer, eds.),

Wiley Interscience, New York, 1955, vol. 14, p. 946.46. Hanika, J., K. Sporka, a nd V. Ruzicka, Coil. Czech. Chern. Comm., vol . 36, p. 2903 ,1971.47. Hanson , L. L., and A. J. Engel , AIChE J., vol.B, p. 260, 1967.48. Hartman , M., and R. W. Coughlin, Chern. Eng. Sci., vol. 27, p. 867,1972.

20 GAS-LIQUID-SOUD REACTOR DESIGN

49. Hatch, 1. F., Hydrocarbon Processinq, vol. 49, no . 3. p. 101,1970.50. Hellwig, 1. R., R. P. Van Dri csen, S. C. Schuman, and C. E. Slyngstad, Oil Gas J ., vol. 60, no. 21,

p. l19, 1962.51. Henry, H. C , and J. B. Gilbert, I &EC Process Design Dev., vol. 12, p. 328, 1973.52. Hofmann, H., Thesis, Technische Hochschule, Darmstadt, 1954.53. Haag, H., in Proceedinqs of th 3rd Internati onal Conqress on Catalysis, North Holhind Publishing

co, Amsterdam, 1965, p . 7.54. Hooper, J. R., and D . S. Sh igemura, AIChE J ., vol. 19. no. 15, p. 1025, J973.55. Horne, W. A., and J . M cAfee, Adv. Petrol. Chern. Refinin g, vol. 3, p. 193, 196{).56. Juvekar, V. A., and M. M. Sharma, Chern. Eng. s, vol . 28, p. 825, 1973.57. Karolyi, J ., A. Zal a i, R. Birthler, and H. Spitzner, Tilt . Chem. Enq.; vol. 3, p. 597,1 963.58. Kastens, M. 1. , 1. 1. Hirst, and R. G . Dressler, Ind. Eng. Chern.. vol. 44, p. 450, 1952.59. Kenney, C. N., and W . Sedricks, Chern. En.

.J

,


90. Ramchandran, P. A., and M. M. Sharma, Trans. lnst. Chern. En.q., vol. 49, p. 253, 1971.91. Reuther, J. A" and P. S. Puri, Can. J. Chern. Eng., vol. 5 t, p. 345, 1973.92. Ross, L. D., Chern. Eng. Prngr., vol. 61, no . 10, p. 77, 1965.93. Sadana, A., and J. R. Katzer, 1&EC Fundamentals, vol. 13, p. 127, 1974.94. Satterfield, C. N., AIChE J ., vol. 21, no . 2, p. 212,1975.95. Satterfield, C. N., and F. Ozel, AIChE J., vol. 19, no. 6, p. 1259,1973.96. Satterfield, C. N., A. A. Pelossof, and T. K. Sherwood, AIChE J., vol. 15, p. 226, 1969.97. Satterfield, C. N., and T K. Sherwood, The Role (~f Diffusion in Catalysis, Addison-Wesley,

Reading, Mass., 1963.98. Satterfield, C. N., and P. F . Way, AIChE J., vol. 18,p . 305, 1972.99. Schoenemann, K., Dechema Monograph, vol. 21, p. 203,1952.

100. Schorr, V., V. Boval, V. Hancil, and J. M. Smith, I&EC Process Desiqn De; vol. 10, p. 509,1971.

101. Schuit, G . C. A., and B. C. Gates, AlChE J., vol. 19, no. 3, p. 417, 1974.102. Scott, 1. W., and A. G . Bridge, Adv. Chern., no. 103, 197 t.103. Sedricks, W., PhD Thesis, University of Cambridge, 1969.104. Sedricks, W., and C. N. Kenney, Chern. Eng. Sci., vol. 28, p. 559, 1973.105. Shah, Y. T, and A. J. Perrota,l&EC Product Res. Dco.; vol. 15, no. 2, p. 123, June, 1976.106. Shah, Y. T ., a nd M. M. Sharma, Trans. l nst. Chem. Eng., vol. 54, p. 1,1976.107. Sherwood, T . K ., and E. J. Farkas, Chern. Eng. Sci ., vol. 21, p. 573,1966.108. Shingu, H., US Patent 2,985,668 (1961).109. Silvestri, A. J., and C. D. Prater, J. Phys. Ch em.; vol. 68, p. 3268, 1964.110. Snider, J . W., and J. J. Perona, AI ChE 1., vol. 20, no. 6, p. 1172, 1974.111. Somers, A., Y. T . Shah, and J. Paraskos, Chern. Eng, Sci., vol. 31, p. 759, 1976.112. Srivastava, R. D., A. F . McMillan, and I. J. Ha rris, Can. J. CIJem. Eng., vol. 46, p. 181, 1968.113. Stanulonis, J . J ., B. C. Gates, and J. M. Olson, AIChE J., vol, 22, no . 3, p. 576, 1976.114. Strangeland, B. E., I &EC Process Design Dev., vol. 13, p. 71, 1974.115. Titova, G . A., Chern. Abstr., vol. 70, p. 1314, 60621h, 1968.116. Tsao, G . T, Chern. Eng. Sci., vol , 27, p. 1593, 1972.117. Tsao, G . T., and D. D. Lee, AfChE J., vol. 21, no . 5, p. 979,1975.118. Tsao, G. T., A. Mukerjee, and Y. Y, Lee, in Proceedings o{ the 4th lnrernational Fermentation

S:rmposium, Fermentation Technology Today, Kyoto, 1972, American Society for Mic robiology,Washington D.C., 1972, p. 65.

119. Urban, W.. Erdoei Kohle, vol . 8, p. 780,1955 ,120. Van Deemter, 1. J.. in Proceedings of the 3rd European Symposium on Chemical Reaction

Engineering, Pergamon Press, Oxford, 1965, p. 215.121. Van Driesen, R. P., and N. C. Stewart, Oil Gas J ., vol, 62, no. 20, p. 100, 1964.122. Verno, Y. J., Water Pol. COil!. Fed., vol. 46, no. 12, p. 2778, 1974.123. Volpicelli, G., and L. Ma ssimilla, Chern. Eng. Sci; vol. 25, p, 1361, 1970.124. Volpicelli, G., and L. Massimilla, Pulp Paper Mag. Can., vol. 66, p. T512, 1965.125. Ware, C. H., Jr., PhD Thesis, University of Pennsylvania, University Park, 1959.126. Way , P. F ., PhD Thesis, MIT, Cambridge, Mass, 1971 ; also Satterfield , C. N ., and P. F. Way,

AIChE J ., vol. 18, p. 305, 1972.127. Yitzhaki, D., and C. Aharoni, ATChE J., vol. 23, no. 3, p. 342,1977.128. Zabor, R. C., R. C. Odiozo, B. K. Schmidt, and J. R. Kaiser, in Actes du Deuxieme Conqresse

Internationale de Catalyse. Paris, 1960, Editions Technip, Paris, 1961, p. 2601.129. Zenz, F. A., and D. F .Othmer, Fluidization and Fluid-Particle Systems, Reinhold, New York,

1960.

,CHAPTER

TWOFILM AND PENETRATION THEORY ANALYSESOF GAS-LIQUID AND GAS-LIQUID-SOLIDREACTIONS

2-1 INTRODUCTION

The gas-liquid and gas-solid reaction processes can be analyzed by severaldifferent physical models, namely film, penetration, surface renewal , Danckwerts,film-penetration, etc.These models are described by Danckwerts.:'? Although eachof these models gives a somewhat different physical picture of the reaction process,in many instances the final desired answer for the rate of absorption of gas in thepresence of a liquid- or a solid-phase reaction is similar. Since film and penetrationtheories are most widely used, we review their applications here.

2-1-1 Film TheoryThe film theory was originally proposed by Whitman.l " who obtained his ideafrom the Nernst"!? concept of the diffusion layer. It was first applied to theanalysis of gas absorption accompanied by a chemical reaction by Hatta. 8 5 8 6It is a steady-state theory and assumes that mass-transfer resistances across theinterface are restricted to thin films in each phase near the interface. If more thanone species is involved in a multiphase reaction process, this theory assumes thatthe thickness of the film near any interface (gas-liquid or liquid-solid) is the samefor all reactants and products. Although the theory gives a rather simplifieddescription of the mult iphase reaction process, it gives a good answer for theglobal reaction rates , in many instances, particularly when the diffusivities of allreactants and products are identical. It is simple to use, particularly when the

22

"

FILM AND PENETRATION THEORY ANALYSES OF REACTIONS 23

reaction process is complex, involving several volatile and nonvolatile reactantsand/or products. The governing equations for this theory can be written in agc'neral form, as :

d2WD - -'-R*-OI dx 2 I - , i = 1,2, ... ,n , (2-1)

where Db Wi, and Rt are the diffusivities, concentrations, and rate of generation(or depletion) of the reacting species i, respectively. The number of equations oftype (2-1) involved is generally equal to the number of reacting species, n. Not allthe equations may be relevant for the purpose of calculating global reaction rates.It should be noted that no convective mass transfer, as such , is taken into accountin Eq. (2-1). The relevant boundary conditions for all species, i = 1,2, . .. , etc.,are generally written at the interface (gas-liquid or liquid-solid) and at the filmboundary. The absorption rate is obtained as

dW!R;= -Di - ' .dx x=o (2-2)If there is no reaction in the film, the theory implies that the mass flux across

the film is given by

(2-3)where ~Wi is the difference in concentration at the two ends of the film and K,the phenomenological mass-transfer coefficient for the ith species. K I can beexpressed as equal to D;/b, where b is the film thickness. The values of thephenomenological coefficients K, depend upon the properties of the diffusingspecies and the fluid and the prevailing hydrodynamic conditions during thereaction process.

(2-4)i = 1,2, . . . ,n.

2-1-2 Penetration TheoryClassical penetration theory, often referred to as Higbie's penetration theory, ismore realistic than film theory and takes into account the transient mass -transferconditions. In this theory, no assumption is made regarding the depths ofpenetration of the various reacting species . It is more complex to use than thefilm theory, since it involves partial differential equations for the mass balancesof each reacting species. It gives a more accurate answer for the absorption anddesorption rates when the diffusivities of all reacting species are unequal. Thegoverning equations for this theory can be expressed in a general form :

02 Wi * _ aWlex 2 - R j - at'

The right-hand side of Eq . (2-4) takes into account the time variation of theconcentration of the reacting species in the liquid phase. Just as in the case offilm theory, classical penetration theory neglects the convective mass transfer at

---

24 GAS-LIQUID- SOLID REACTOR DESIG N

the gas-liquid interface. The number of equations of type (2-4) generally requiredis equal to the number of the reacting species, n. Equation (24) requires initialconditions along with the boundary conditions. The boundary conditions arcusually written at x = 0 and at x = 00 (ifliquid film is thick and can be consideredto be a semi-infinite medium for mass penetration) or at x = d, where d is thethickness of the liquid film. Unlike film theory, penetration theory takes intoaccount the effect of time of contact between gas and liquid during a gas-liquidreaction process. The latter theory is particularly useful when the gas-liquidreaction is non isothermal, because in this situation an assumption of the samefilm thickness for mass and heat penetration is obviously not true and the filmtheory is not applicable. When the reaction process is complex, an analyticalevaluation of the absorption and/or desorption rates by the penetration theorymay become too complex for its practical use. Penetration theory has been shownto be equivalent to the boundary layer theory.? At any given time, the absorptionrate is given as

OWlRj = -Di - ' .ax x=o (2-5)

, ~

I '.,

fi\

"II' I

I ,

2-1-3 Mechanism of Absorption with ReactionGas-liquid reactions may conform to various mechanisms. Under certain con-ditions, the absorption and reaction may conform to a "slow" reaction mechanism.By this term, we mean that a gaseous species A is absorbed, diffuses through thefilm.and.then J.:~!l~J!> in the bulk liquid.Thus, according to film theory, the processesof chemical react ion and diffusion become two steps in series for a slow reaction.The absorption rate in this case is almost unaffected by a chemical reaction. Inthe limiting case, where the concentration profile of the absorbing species in theliquid filmis flat, the reaction is often called a "very slow" reaction, and the processof absorption is said to be in the kinetically-controlled regime,

Absorption accompanied by a


in the liquid. A "reaction plane" is formed in the liquid, where the instantaneousreaction occurs, and both absorbed species and the nonvolatile reactant diffusetowards this "reaction plane," where they react. TM..J:e.ac.!io~~~_sol~lymass-transfer-controlled. When the concentration of the nonvolatile reactant is;m;ch larger than that Ofthe absorbed species, the : '.!:~actiQD--p-la~" is almost nextto the gas-liquid interface , and the rate of reaction is solely governed by the rateof diffusion of the nonvolatile reactant from the bulk liquid to the interface. Itshould be noted, however, that even a reaction which may not be intrinsicallyinstantaneous can become "instantaneous" under certain conditions. By the term"instantaneous," it is meant that the rate of reaction is controlled by the rate ofdiffusion oftlu:.abso-r.-'b~?speCies andjhe_n~~~ojit.il~re~-,

(2-8)

(2-7)

(2-9)

Gas film resistant notimportant:~.- Absorption rate, R A '" kLa~'(A-i'L - A L)

Bulk liquid..- -- -_._--..--:~

AL K'-'-'-'-'-'-'-A + C _ B (Q)

AL~~----A+C~ B(Q)

Liquid film

Gas-l iquidint erfa ce

Gas film

26 GAS-LIQUID-SOLID REAClDR DESIGN

2. Fast reaction,

3. Instantaneous reaction,

where K = rate constant, CL = bulk concentration of the liquid reactant, A il =gas-l-iq-trid"interface concentration of the gaseous .reactant, D A = diffusivity ofgaseous reactant in the liquid phase, and kL = .gas:::nquid mass-transfer coefficientin the absence of a chemical reaction. "" . . . ..

TIi.~'~oncentration distributi~ns for species A, 8, and C in these regimesbased on the film theory are described in Figs. 2-1 through 2-3.

If the product 8 in reaction (2-6) is volatile, it will simultaneously desorbin the gas phase. The typical concentration distributions of species A, B, and C

1. Slow reaction,

This reaction can occur in slow, fast, or instantaneous reaction regimes. In eachregction re.giJ11s .1~e .ga~~phase resistance can be important, dependingon "t heconcentration of the ' reactant -in "the gas phase and the solubility of the gas inthe liquid phase. The criteria for these reaction regimes are 2 4 . 3 9

Figure 2-1 The concentration di stribution for species A, B, and C in the slow reaction regime, basedon the film theory. Reaction occurs ma inly in the bulk liquid; mass tran sfer and reaction areprocesses in series. -

FILM AN D PEN ETRATION THEORY ANAL YSES OF REACTI ONS 27

Gas film Liqu id film Bulk liqu id

c

A + C -.!i-- B (~)

Gas-liqu idinterface

OL- +- --..::~ _DC

IJ

I

IIti

-- Gas film resista nt import ant

_.- Gas film resista nt not impo rtan t

Figure 2-2 The concentra tion distribu tion for species A, B. and C in the fast reaction regime, basedon the tilm theory. Reaction occur s mainly in the Iiquid film ; mass transfer and reac tion areparallel processes ; the ab sorption ra!e of the gas i~ increased due to the.chemical reaction,

React ion plane

~------Cl

Gas-li qu idinterface

---Gas film resist ant important- . - Gas film resistan t not import ant

Figure 2-3 The concentra tion distribution for species A, B, and C in the instan taneou s reactionregime, based on the film theory. Reaction occurs at a plan e in theiiq~jd; i~ gaS and the l iq iilacannot coexist and the increase in abso rpt ion rate due to the chemical reaction is a maximum .

28 GAS- LIQUID - SOLID REACTOR DESIGN

(a)

A;~------""Bl

x=o

(b)

(e)

(e) (f) .

(g)

eFigure 2-4 Typical concent ratio n profiles of instan taneou s reaction between the gas A and the reactantC, based on film theory. (a) Diffusion contro lled -slow reac tion, (b) kinetically controlled -slowreaction, (e) gas-film-controlled desorpt ion - fast reaction, @l liquid-film-co.ntrolled desorption - fastreaction , (e) liqu id-film-con trolled absorption - instantaneous reaction between A and C, (f) gas-film-controlled abso rption-i nsta nta neous reaction between A and C, (g) concentra tion pro files for A, B, andC for instanta neous reaction between A and C - both gas- and liquid-phase resistances arecomparable.P?


for various cases of slow, fast, and instantaneous reactions based on the filmtheory are illustrated in Fig . 2-4.

An extensive literature on the film and penetration theory analyses of gas-liquid reactions has been published. A summary of the types of reaction analyzedby the film and penetration theories is given in Table 2-1 for the case of finite

Table 2-1 Summary of theoretical analysis of gas-liquid reactions with finitereaction rate by film and penetration models]

Type of reaction Model Reference

Zero-order reaction (e.g. oxidation reactions of Film and penetration 186hydrocarbon)

Zero-order reaction Penetration 11,177

A(g) ..... B Penetration 7.39.

- --- - - - -- ---- - -- --- - --- ---------- - --

F ilm 122Penet ra tion 125, 156Penetration 125

F ilm 131Film 23

Penetration 174

Film 187, 188

Film 95

Penet rati on 160

F ilm 123

Penetrat ion 125

Penetra tion 125

Penetr a tion 125

Penetrat ion 74

Film 31

Penetration 20,21

30 GAS-LIQUID-SOLID REACTOR D ESIG N

Table 2-1 Cont inued

Type of reaction

YA A (g) -+- YoB -> Product.Y~ E (g) -+ YuB- ProductYAA (g) + YuB~ r EE + FFYAA(g) + Y. B~ YeE -l- YFFA( g) + YoB YeC + Prod uct,YEE(g) + YcC -> \'oRA (g) + B~EA(g) ..... B ..... CA(g) -+ B -+ CA(g) + B-+C,A(g) + C -+ PA (g) + YoB -+ Product,A( g) +- }cC -> Pr oduct.Alg) + YoB -+ YeC(l) + Yo 0 ,YEE(g) .:. Yc C -> YuBY.A(g) + YuB YeC(l) + YoD,YEE(g) + }c C Yu-BA(g) + Yu B -+ YeC,A(g) + Yc C -+ ProductA(g) + YoB -+ Product,A(g) + YcC ProductA(g) + uB Pro duct,YEE(g) + Yo -B - ProductA (g) + YeC ..... Pr oduct,B(g) + c C -> Product

K,CI, + H10 H ' - cr + HOCI

Rate = K F([Cl lJ _ [H_~ ]JCl~~[HOC ll).K* = [H + ] [C.~_- J [HO

FILM AND PENETR ATIO N THEOR Y AN ALYSt'S OF REACTIO NS 31

Table2-1 Continued

Type of reacti on

Condensation polymerization (e.g., synthes is ofpol yamidcs, syn thesis of pol ymers such aspolyethylen e terephthalat e)2A (g) ~A' -I- A.A* + B ~P

GeneralA(g)~E-+P

b~~~~.A(g) + C - Products.B(g) + C - Products

(e.g., CO a nd C2 H 4 in ammoniacal CuCbsolution, 0 , and CO in haem oglo bin, H ,Sand ca, in amine under ca ustic co nd itions)

Model

Fil m

Film

Film

Film

Reference

149

62105

136. 137, 13R

Film

Fi lmA(g) + C ~ B(g) + S,s

B(g) -+ Pr oducts (S act s as cat alyst )(e.g., Cl2 and CO 2 in Na 2C03 so lut io n, S02and CO 2 in Na2C0 3 so lution)

A(g) + C -+ S,B(g) + S ~ P roduct

(e.g., S02 a nd O 2 in ca ustic so da )A (g) + C -+ S, Fi lmB(g) + C ~ Product,

sB -+ Product

(e.g.. CI2 and O 2 in NaOH . S0 2 and CO, inNaOH)

A (g) + C -+ B(g) + E, F ilmB(g) + C - Pr oducts

(e.g., phosgen e and CO2 in NaOH or Na, CO lso lutio ns)

141

141

141

140

j '

t Not e that not ations A, B, C. D. E. F, etc., in th is ta ble refer to either reac tant or productdepending up on the particular rea ction sys tem und er co nsidera t ion.

reaction rate and in Table 2-2 for the case of instantaneous react ion rates. Someadditi onal react ion systems involving simultaneous absorption of two gases maybe found in the review by Ramchandran and Sharma.l "! Additional reactionsinvolving one or more volatile products may also be found in the review by Shahand Sharma.P?

Experimental systems A large numbe r of gas- liquid reaction systems have beenexamined experimentally. Some of these studies are outlined in Table 2-3. This


Table 2-2 A summary of studieson instantaneous reactions

Reaction Model Reference

A~E Film 120A2E Film l20

A+B~E Film 120A+B~E+F Film 120A(g)~LZ"C, Penetration 136. 142B(g)IZ.,C

(e.g-s simultaneous absorption of C02 and H2S intoamine solution)

YAA + aB = YcC + oD Penetration 156A(g) + Z IBt + ... YiP I + Y2P2 + ... 42

(e.g.S02 +HzO~ H + + HSO "CO 2 + 2R2NH~R2NCOO - + R2NHtl

(S02),oln + SOj + H20~ 2HSO;- Film 126(H2S),oln + OH- HS- + H2O. Film 126HS - + OH-S2 - + H2O(S02),oln + OH - ~ HSO ), Film 126HSO J + OH- sot + H2OA + B -> Product Penetration 6, 18,39,166A + YDB -> Product Penetration 39, 43, 98, 132

(e.g., oxygen into sodium dithionite solution)A(g) + YDB - Product in packed tower Film 155, 163 \A(g) + oB -+ P I. Penetrat ion 39, 73, 148Etg) + YD,B -> P2

(e.g., simultaneous absorption of CO 2 and HzSinto alkaline liquid)

A(g) + YIIB -+ Product, Film 95A(g) + fcC -+ Pr oduct

N N

A(g) + L z.n, -> L E;, Penetrat ion 3l = t i= I

where E"

E2 . . , E. are n products

A(g) + C -+ S + Pr oduct (instantaneous), Film 142B(g) + S -+ C (finite rate )A(g) -> B(g) + D, Film 140A(g) + C(g) -+ B(g) + D

(e.g., absorption of mixture of CO 2 and pho sgeneinto water)

A(g) + C -+ E (gas or liquid) '(instantaneous reaction), Film 137YEE + fcC -+ Products (finite reaction rate )

H2S + MOH -> MHS + H2O, Penetration 127H2S + 2MOH - M2S + H2O,H2S + CO ) -+ HS - + HCO "H2S + HCO, -> HS - + H 20 + CO 2CO] + H20 -HCO, + OH -

FILM AND PENETRATION THEORY ANALYSfS OF REACTIONS 33

Table 2-2 Continued

------ - - - -

Reaction

A(g) ..... B(g) + EA(g) + C ..... S,B(g) + C -+ Products,8(g)~ ProductsA(g) + C ..... B(g) + E,8(g) ~ Products

(S acts as catalyst)

Model

Film

Film

Film

Reference

141

141

141

table indicates that the absorption of carbon dioxide in a variety of solvents isthe most extensively examined process. Other absorbent gases studied are carbonmonoxide, ammonia, hydrogen chloride, hydrogen cyanide, carbonyl sulfide,phosphorous trichloride, hydrogen bromide, mercaptans, oxygen , chlorine, sulfurdioxide , hydrogen sulfide, phosgene, and nitrous gas. Absorption of some hydro-carbons such as acetylene, benzene, ethylene, propylene, dimethyl ether, vinylchloride, heptene, octene , and isobutylene have also been considered. The studiesare made for both single gas absorption as well as simultaneous absorption oftwo gases. Reactions producing both nonvolatile and volatile products areexamined. The solvents used are of varied chemical nature.

Several experimental systems, such as absorption of C1 2 , NH 3, and phosgenein water, hydrogen sulfide in aqueous buffer solutions, and oxygen in glucosesolution have been extensively examined with the help of penetration and/or filmmodels . Others, such as absorption of CO 2 in amine solution and hydro-chlorination of octyl and dodecyl alcohols have been only either partially examinedor not subjected to a significant theoretical evaluation. The absorptions of Cl2and NH 3 in water are accompanied by large heat effects. These effects are alsoexamined with the help of penetration theory. Several other practical systemsinvolving simultaneous absorption of two gases to and /or the one involving oneor more volatile products! 1 have been described, and reviewed by Ramchandranand Sharmal t! and Shah and Sharrna.'P",

r: / .J ( . /~I UJ)2-2-2 Gas-Liquid-Solid Catalytic Reaction t J' I JNo analysis ofthis type of the reaction by penetration theory has been reported.We shall analyze a few examples of this type of reaction by the film theory. Fora gaseous reactant, the following process steps are involvedIn this type of areaction. . . ...

1. Mass transfer from the bulk concentration in the gas phase to the gas-liquidinterface.

2. Mass transfer from the gas-liquid interface to the bulk-liquid phase .

34 GAS-LIQUID- SOLID REACTOR DESIGN

3. Mixing and diffusion in the bulk liquid.4. Mass transfer to the external surface of the catalyst particles.5. Reaction at the catalyst surface, including intraparticle diffusion effects.

Table 2-3 A summary of experimental gas-liquid reaction systems

System

CO, in monoethanolaminc solutionCO 2 in aqueous amine-potash solutionCO 2 in aqueous amineCO 2 in solutions of Na,CO) and triethanolamineCO 2 in diethanolarnine solutionCO 2 in carbonate-bicarbonate , alkaline solutions

CO 2 in water in presence of carbonic anhydraseCO, in weak ion-base ion-exchange resinCO 2 and H 2S into calcium cyan amide solutionCO and CO, into ammoniacal cuprous chlorideca, and sa, in NaOHCO, and carbonyl sulfide in NaOHCO 2 and NH , in waterCO 2 and phosgene in waterH,S and CO , in amine solutionH,S a nd CO 2 in aqueous hydrox ide solutionO 2 in sodium sulfite

O 2 in glucose solutionO 2 in cuprous chloridea, in aqueous alkaline so lution o f sod ium dithionitea, in aqueous solutions of phenol, methanol, and formaldehydeCI2 in waterCI2 in HCI , NaCI , and NaOHChlorination of acetyleneChlorinization of benzene with stannic chloride cat alystFerric-chloride-catalyzed hydrochlorination of l-hexadeceneHydrochl orination of cetyl and dod ecyl alcoholsNH , in H2S04NH ~ in ethyl malonate in alcoholic solutionNH) in waterNH ., in aceti c acidSO , and NH, in water50 2 from K20-S0 2 C.HsCOOH-H,OSO, and CO 2 into aqueous carbonate solution or aqueous caustic soda502 oxidation by V205 catalyst in K,S,0 7 meltSO, in water502 in alkaline solution

'I Oxidation of SO , by aerosols MnSO.H,S in alk aline solutionH,S in water

"i 0y:~ \.i.,:.,' )~, : . I ~ . . . -'.

Reference

12, 22, 32, 42, 46, 61. 18145,48,167191,1921851185,49,80,129,146,147,

164,193,19918219479190747484, 1481401014857,81,107,108,109,128,

19710fl969717220,2117555[69lSI501686329,77,9358153827291,9242,12612636,78, 11470, 1269

FILM AN D PENETRATION THEORY ANA LYSES OF REACTIONS 35

Table 2-3 Continued

System

Nitrous gas in waterCarbonyl sulfide in amines a nd alkalinesEthy lene in aqueous chlo rine solu tio nlso hut ylene in aqueous solu tion of H 2S0 4C2H 4 an d O 2 into so lut ion ofCuCI, conta ining Pd Cl2 catalystNO and H 2 into sulfuric acid so lutionCondensation of po lymer ization - synthesis of polyamides and

polyeste rBenzoic acid in wate r Io-Sa licyclic ac id in waterBenzoic acid in 22.5 percent aqueous glycero l fBenzo ic ac id in 41 percen t aqueo us glycero lIodine dissolution in pot assium iod ine - a fast inter facial reactionHydro gen chlo ride in ethylene glycolDim ethyl sulfid e an d N0 2 into dimethyl sulfoxideC 3 H. CO, a nd H2 into but an olHeptane or octene, CO , an d H, into octa no ls or nonanolsC, H HCl, and O 2 into aqueous CuC l2C2H4 a nd O 2 into wate rHCl a nd O 2 into waterEthylene an d pro pylene into non-Newton ian so lution o f copolyme rS02 and H2S into molten su lfurSO , and CI2 into sulfuryl chlor ideSO , a nd HCl into ehlorosulfonic ac idDimeth yl ether a nd SO J into dimet hyl sulfatePCI3 and O 2 into phosphorous oxychlorideAcetyle ne and Cl 2 into tet rachl or oethaneC,H4 and HCI into et hyl chlorideMercaptan s and CO2 into alk al isNH , an d CO 2 into aq ueo us urea nitrat eVinyl chlor ide a nd CI, in tr ichloroetha neC 2H4 a nd HBr int o eth yl bromideCl H. CO an d H 2 into butanolHC N and CO , in a lkalisIsobut ylen e and bu tenes into aqueo us H, S0 4CO" H ,S and CO" COS, H 2S into aq ueo us potass ium carbonate

so lutionCO" H, S and CO" COS. H,S into aq ueo us tr ieth an olamine and into

aq ueo us ammo niaCO,. H,S and CO" CO S, H IS int

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