11.2 Dehydrogenation processes

The dehydrogenation of hydrocarbons involves thebreaking of two carbon-hydrogen bonds with thesimultaneous formation of a hydrogen molecule and amolecule containing a double carbon-carbon bond,which usually represents the desired product. Thedouble bond is a highly reactive point that permits theuse of molecules which contain it as intermediates forthe production of typical petrochemical products suchas polymers.

Dehydrogenation reactions appear very simple; theirthermodynamic and kinetic characteristics have,nevertheless, contributed to make the development oftechnologies that allow for a reliable and efficientindustrial application, rather complex.

The dehydrogenation reactions of greatestindustrial interest are those paraffins containing 2-5carbon atoms which are used to produce thecorresponding olefins; those linear paraffins with 10-15 carbon atoms, needed for the production oflinear-alkyl-benzenes (LAB) and ethyl benzene,dehydrogenated to styrene, the monomer which is thestarting point for the production of polystyreneplastics. The dehydrogenation of alcohols to ketonesand aldehydes represent other dehydrogenationreactions of industrial interest.

11.2.1 Dehydrogenation of lightparaffin (2-5 carbon atoms)

Light olefins continue to serve as a fundamental basisfor the petrochemical industry and refining.Traditionally, steam cracking and Fluid CatalyticCracking (FCC) are the major industrial sources of lightolefins, such as ethylene and propylene. Both processes,in which various hydrocarbon cuts can be used as theraw material, produce various compounds of interest atthe same time. Variations in the operating conditions

make it possible to roughly modify the composition ofthe mixture of products, but this may not be sufficientwhen an increase in market demand for one particularproduct is much higher than that for the otherco-products. For example, the demand for propylene isgrowing faster than that for ethylene in manygeographical areas. In this case, the ability to synthesizea pure product (such as propylene through thedehydrogenation of propane) can prove far moresuccessful in comparison with the classic ‘‘multi-product’’ approach. In the recent past, thedehydrogenation of isobutane to isobutene, anintermediate stage in the production of such high octaneoxygenates as methyl-tert-butyl ether (MTBE), had anextremely important role. However, the decision tochange the composition of gasolines in the UnitedStates has led to a decreased interest in this particulardehydrogenation product (Sanfilippo et al., 2003);nevertheless, existing facilities are continuing their largescale production, and the technology used in thedehydrogenation of isobutane constitutes the basis for the development of similar techniques for thedehydrogenation of propane.

Kinetic and thermodynamic aspectsDehydrogenation removes a molecule of hydrogen

from paraffin forming an olefinic double bond thatwill be the preferred point of attack for successivereactions:

R2CH�CHR2��R2C�CR2�H2

Notwithstanding its apparent simplicity, thedehydrogenation of paraffin is one of the most complexchemical processes to realize industrially, since thethermodynamic equilibrium limits the conversionpossible per pass and the reaction is highlyendothermic, that is to say, a large amount of heat mustbe supplied to the reactant system. In Fig. 1, the

687VOLUME II / REFINING AND PETROCHEMICALS

11.2

Dehydrogenation processes

conversion equilibrium is shown versus thetemperature, for paraffins with 2 to 15 atoms of carbon.It is clear that, in order to make the industrial processmore economic, it is advisable to increase theconversion per pass as much as possible, so that thecost of recycling the unconverted reagent and that ofseparating the reagent and the product is reduced. Asshown in Fig. 1, it is evident that temperatures between500 and above 700°C are needed to reach a conversionrate of 50%. However, high temperatures also enhancean undesired parallel process, i.e., thermal crackingreactions that reduce the selectivity to the desiredproduct and make purification operations downstreamthe reaction zone necessary. A typical scheme thatincludes the most important types of secondary

reactions, such as cracking, oligomerization,isomerization, aromatization and alkylation, is shownin Fig. 2. Oxidative dehydrogenation has been widelystudied, as one route to circumvent the thermodynamiclimitations. This involves supplying oxygen to thereagent mixture so that it combines with the hydrogenproduced by dehydrogenation and forms water. Theremoval of the hydrogen allows the dehydrogenationreaction to proceed with a higher conversion thanwould be suggested by the equilibrium. However, thepractical application of this approach ran into theproblem of making the oxygen react only with thehydrogen, without causing direct combustion of thehydrocarbons and a resulting loss of selectivity. Anadditional advantage of this approach results from theexothermic nature of the reaction producing water,which contributes heat for the dehydrogenationreaction, reducing the need for an external source.

In non-oxidative dehydrogenation, in order to reachhigh reaction rates and minimize the effect of itscomplex secondary reactions, a suitable heterogeneouscatalyst is needed.

The energy required to remove two atoms ofhydrogen from the alkane molecule is almostindependent of the molecular weight of the paraffin andin the range of 113-134 kJ/mol.

The highly endothermic nature of the main reaction(and of the greater part of the secondary reactions) leadsto a strong temperature decrease, if the reaction iscarried out in an adiabatic manner. For example, in thecase of the dehydrogenation of propane with a 25%

688 ENCYCLOPAEDIA OF HYDROCARBONS

SYNTHESIS OF INTERMEDIATES FOR THE PETROCHEMICAL INDUSTRY

equi

libr

ium

con

vers

ion

(%)

01020

4030

5060708090

100

temperature (°C)300 400 500 600 700 800 900

C15

C10

C5

C4

C3

C2

Fig. 1. Equilibrium conversion in the dehydrogenation of some paraffins at atmospheric pressure.

CH3 CH CH3

CH3

CH3 CH CH2 CH CH2

CH3 CH2 CH3

CH3

CH3 C CH2CH3 CH CH CH3

CH2 CH2

CH3

CH2 CH CH2 CH3

CH2 CH CH CH2

�H2

�CH4

�CH4

CnH2n�2

CnH2n CnH(n�y)

CnH(n/x)coke

CH3 CH CH2

�H2

R

dimerization

polymerization

oligomerization

coking

dimerization

cracking

dehydro

dehydro

aromatization

alkylation

side chainaromatization

dehydro

side reactions increase withtemperature and conversion

isomerization

Fig. 2. Reactionnetwork ofdehydrogenationof isobutane.

conversion, there is an adiabatic temperature decrease of200°C. Since no catalyst is able to operate in such a widetemperature range, it is clear that a single adiabaticreactor cannot be used, and a heat source must also besupplied in the course of the reaction. We will discusslater how different ways of supplying heat for thereaction characterize the different technologies available.

Examining Fig. 2 once again, one can note how allthe reaction paths lead to the depositing of carbon(coke) on the surface of the catalyst. Even with a highlyselective catalyst, the depositing of coke graduallymakes it less active, so that a periodic regeneration ofthe catalyst is necessary, normally by burning the cokein air streams. The catalysts for dehydrogenation musttherefore be able to accept repeated cycles whichalternate between reducing and oxidizing atmospheres.The maximum level of coke accumulation andconsequent time needed for each phase in the reactioncycle and regeneration depend on the nature of theparticular catalyst used.

From the thermodynamic point of view it isnecessary to also consider the negative role played bypressure, due to the increase in the gaseous volumeproduced by the reaction. In addition, using a dilutingagent or operating in a vacuum have the advantage ofincreasing the driving force that acts in the process.

CatalystsThe principal catalytic systems that have a

dehydrogenation activity reported in scientific andpatent literature are: • Group VIII metals (mainly platinum with tin)

supported on alumina and with promoters.• Chromium oxides on alumina or zirconium, with

promoters.• Iron oxides supported, with promoters.• Gallium, as a supported oxide or included in zeolitic

structures: gallium in/on mordenite (Pogue et al.,1993), on SAPO-11 (Machado et al., 2002), onMCM-41 (Fajula et al., 2003), on TiO2 (Nakagawaet al., 2001), on Al2O3 (Iezzi et al., 1994).

• Copper, for the dehydrogenation of alcohols toaldehydes.

Including the most recent literature, the situation fordehydrogenation catalysts and their applications isoutlined in Table 1.

Commercial application has restricted potentialcatalysts to those indicated below:• For the dehydrogenation of ethyl benzene to styrene,

oxides of Fe as the only choice.• For long chain paraffins for LAB, Pt/Sn promoted

on Al2O3.• For light paraffins and olefins, Pt/Sn promoted on

Al2O3 and Cr2O3 on Al2O3.The two large families of catalysts for the

dehydrogenation of paraffins (based on Pt and Cr) weredeveloped at the same time. The two families do notdiffer substantially in terms of activity and selectivity,but rather in the quality of some by-products and in thetreatment needed to complete the regeneration after thecombustion of the coke. In addition, the phenomena thatlead to the irreversible deactivation (sintering,volatilization of the active components andtransformations of the morphology or the state of thesupport) are typically connected to the chemical typesthat characterize the various catalysts.

Platinum catalystsPlatinum catalysts are less active in the formation of

heavy by-products, compared to chromium catalysts,but cause skeletal isomerization in the case ofhydrocarbons with 4-5 atoms of carbon, resulting inpossible complications in the process stages.

Patent literature shows a very high number offormulations based on the noble metals of Group VIIIand elements of Group IV. However, only twoformulations are applied at industrial level:• Pt/Sn with alkaline metals (and possibility of further

promoters) on an alumina support, used in theOleflex (UOP) (Pujado e Vora, 1990).

• Pt/Sn, preferably with alkaline promoters on zinc ormagnesium aluminates, used in the STARtechnology.Both of these applications use tin, believed to be a

promoter which improves the activity, selectivity andstability of the catalyst. Various hypotheses have been


DEHYDROGENATION PROCESSES

Active component Dehydrogenation Dehydrogenation Dehydrogenation ofof catalyst of light paraffin of paraffins C10-C14 ethyl benzene to styrene

Pt/Sn excellent excellent poor

Cr oxides excellent moderate good

Fe oxides absent/poor not available excellent

Ga system excellent not available excellent

Table 1. Performance of dehydrogenation catalysts

formulated to explain the resistance to ageing found incatalysts promoted with tin (Barbier et al., 1980). Suchhypotheses can be summarized as models of threetypes (Fig. 3), involving: uniform dispersion ofplatinum on a surface constituted by alumina and tinaluminate; formation of a solid solution with atoms orclusters of tin within the platinum lattice; or formation of platinum-tin alloys.

The type of model that best represents thecharacteristics of the catalyst depends on many factors. For example, it is possible to move fromsituations best-described by the first model tosituations that fall in the second and third category inthe following ways:• Increasing the ratio Sn/Pt.• Increasing the calcination temperature of the

catalyst. • Changing from acid supports to neutral supports. • Modifying the preparation procedure, from an

impregnation of the platinum compound on oxidesof co-precipitated Sn/Al to a preparation wherecomplexes of Sn/Pt are impregnated on alumina.

• Increasing the concentration of the two metals onthe support.

• Decreasing the surface area of the alumina.The mechanism by which tin modifies the catalytic

behaviour of platinum can be explained on the basis oftheories that are not closely connected to the types of tinpresent on the surfaces.

The theory, based on an electronic effect,hypothesizes that the tin in metallic form (both in solidsolution and in small bimetallic clusters or as an alloywith platinum or, finally, as an ion Sn2� in close contact with the platinum atoms) gives up its electronsto the vacancies of the band 5d of the platinum atoms.In this case, small quantities of tin (below the limit ofanalytical measurement) should influence the platinumproperties. A quantity of tin which corresponds to 15%of the platinum’s atomic weight is sufficient to fill theband 5d. In addition, solid solutions or alloys can leadto the formation of small crystallites of platinum,characterized by different reactivity (Burch, 1981;

Burch and Garla, 1981). On the basis of the electroniceffect, it is possible to explain the promoting effect ofthe tin, by assuming that the coke precursors cannot fixthemselves to the surfaces of the ‘‘promoted’’ crystals ofplatinum, but instead migrate towards the support,avoiding the deactivation of the active sites (Lieske etal., 1987). The smaller amount of coke formation can also beconnected to an inhibition of the condensation reactionsthat lead to the synthesis of polymers. For the samereason, the hydrogenolysis of the �C�C� bond isalso inhibited, because the hydrocarbon cannot adsorbitself strongly on the catalytic surface.

The formation of coke and the hydrogenolysisrequire large ‘ensembles’ of platinum atoms, whiledehydrogenation requires small agglomerates ofplatinum, or well-dispersed active centres (Barbier et al., 1980). The formation of an alloy leads to adilution of the platinum atoms, with the formation ofsmaller agglomerates. However, where the formation ofalloys is not observed, small quantities of tin in thecrystallites of platinum can produce a dilution effect.

A high level of activity of the platinum in theinteraction with surface ions Sn2� (but not with metallictin) can be explained by the exchange of electronsbetween tin ions and metallic platinum (electroniceffect) or by the stabilization of small crystallites ofplatinum, caused by a strong interaction with analuminate of surface tin.

It has been suggested (Beltramini and Trimm, 1987)that tin added in small quantities tends to segregate onthe surface of the crystallites of platinum in sites withlow coordination (angles and extremities), inhibiting thereactivity of these sites.

The tin also produces a reduction of the surfaceacidity of the support, responsible for both the crackingreactions and skeletal isomerization, and for theformation of the coke precursors.

The main properties of the support which determinethe characteristics of platinum catalysts are as follows: a) surface acidity; b) intrinsic stability and stability of thedispersion of platinum during reaction and regeneration;



support tin aluminate

increasing Pt�Sn, Sn/Pt, reduction T

increasing carrier surface area

Pt Pt-Sn alloytin cluster

Fig. 3. Models for Pt/Sn catalysts.

c) chemical interaction with the promoters; d) distribution of the dimensions of the pores.

Zinc and magnesium aluminates are neutral orslightly basic and therefore do not need alkalinepromoters; alumina supports, on the other hand, needalkaline promoters to minimize the acidic properties ofthe system responsible for a decrease in the selectivityand life of the catalyst. The main role of the supportconsists in stabilizing the dispersion of the platinum,especially during the regeneration of the catalyst, bymeans of the combustion of the coke. While platinumdeposits on silica sinters following oxidative treatments,the sintering is very limited for platinum on alumina andpractically absent for platinum supported on magnesiumaluminate, even after many regeneration cycles.

The alumina support is optimal for platinum and isbased on transition alumina (d e ÿ), characterized by ahigh level of thermal stability and low acidity. The poresshould be large in diameter, to avoid rapid plugging bythe coke, thus prolonging the life of the catalyst(Bricker et al., 1990).

The addition of alkaline metals or rare earths isnecessary to mitigate the acidity of the alumina; cesium,lithium and potassium are considered optimal alkalinepromoters, both individually and in combination. Thismitigation of the acidity is fundamental to reaching highselectivity towards olefins, in that the acidic centres areresponsible for cracking reactions, at the expense ofboth the reagents and the products, as well as theoligomerization-polymerization reactions of the olefins,which lead to the formation of coke.

Platinum catalysts are subject, through continuousand discontinuous operations, to reactivation treatmentsin situ, using substances such as oxygen, chlorine,steam, hydrogen or small quantities of sulphur.

The oxygen is supplied to the catalyst during the firststage of regeneration to burn the coke. The eliminationof the coke can be conducted at a relatively lowtemperature (about 720 K) and with a low concentrationof oxygen, to minimize overheating of the catalyst andsinterization of the platinum crystallites. It seems toproceed in two stages: first, the coke in contact with theplatinum crystallites is burnt; then, by increasing thetemperature to about 820 K, combustion of the coke,adsorbed on the support by spill-over from the platinumcrystallites activated by the oxygen, takes place (Handyet al., 1990). The treatment with oxygen has otherbeneficial effects, such as the passivation of platinum,thanks to the formation of a mono-molecular layer ofoxide, and the re-dispersion of the platinum itselfthrough the formation of platinum dioxide (PtO2), whichis more mobile than metallic platinum. That happens ifthe temperature is maintained below 870 K; at highertemperatures, the dioxide becomes unstable and tends tobreak up, forming metallic platinum which sinterizes

easily. The oxidation of metallic tin to Sn2� andsuccessively to Sn4� takes place by spill-over of oxygenfrom the platinum. In the case of alloys Pt/Sn, tin isoxidized preferentially, while platinum remains in themetallic state. Not all the oxidized tin returns to the alloyafter reduction, due to interactions with the support.

The treatment with chlorine is aimed attransforming the surface layer of tin and platinumoxides into chlorides and oxychlorides, which cansublime or move on the surface of the catalytic systemthus aiding dispersion. However, this treatment alsocreates acid sites that must be eliminated or minimizedto avoid isomerisation of the paraffins.

Successive treatment with hydrogen reduces theoxychlorides of platinum and tin to the metallic state.The reduction of tin is favoured by the spill-over ofhydrogen from the platinum. The same phenomenonhelps prevent the deactivation of the catalyst, reducingthe coke precursors present on the surface of thesupport. The chlorine is then eliminated by means of asteam treatment, which also contributes to theelimination of coke precursors through the water/gasshift reaction. Moreover, sulfur compounds can beadded continuously to inhibit the coking andhydrogenolysis reactions. This last treatment, though,can only be used with supports based on spinels, while,with aluminates, sulphates are formed during theregeneration which contribute to the deactivation of thecatalyst and worsen its mechanical properties.

Pt/Sn catalysts are thus subject to two types ofdeactivation: slow irreversible sinterization, whichrequires the replacement of the catalyst after a relativelylong period of time (years), and the reversible pluggingwith coke. The catalyst Pt/Sn maintains a good catalyticactivity even in the presence of a small percentage ofcarbon deposits and therefore allows a relatively longperiod (hours or days) between regenerations.

Daniel E. Resasco (Resasco and Haller, 1994)proposes that the dissociative adsorption of paraffin,which involves two metal atoms, constitutes the slowstage of the reaction with these catalysts. Thedissociative adsorption is practically irreversible, whilethe elimination of the second atom of hydrogen is veryfast and almost in equilibrium. In the case of thedehydrogenation of isobutane, the global mechanismcan be represented by the following reactions.

i-C4H10�2 Pt��Pt�i-C4H9�PtH (slow stage)Pt�i-C4H9�Pt��

��PtH�Pt�i-C4H8Pt�i-C4H8�

� i-C4H8�Pt2PtH → H2�2Pt

Chromium catalystsIn scientific and patent literature, two types of

chromium catalysts are mentioned for the



dehydrogenation of light paraffins:• Chromium oxides supported on d/ÿ-alumina

promoted with alkali, used in the Catofintechnologies (ABB Lummus) and FBD(Snamprogetti-Yarsintez).

• Chromium oxides supported on zirconium; thissupport was developed because of its lower acidityand greater thermal stability with respect toalumina.The nature of the active sites in supported

chromium catalysts has been amply discussed. Cr2O3 isthe most stable of the chromium oxides and is formedby heating CrO3 to 770 K, even in oxidizingatmospheres. The difference between the standard freeenergy of formation in CrO3 and Cr2O3 (Fig. 4) showsthat, even at room temperature, Cr2O3 is the morestable form. However, when chromium oxides aresupported on alumina, even after calcination at 970 K,a certain percentage of chromium remains in theoxidized state 6�, due to strong interaction with thesupport (Cavani et al., 1996). Only after reduction withhydrogen or a hydrocarbon, all the Cr6� is transformedinto Cr3�. The reduction takes place in two stages: avery rapid first stage leads to complete transformationinto Cr3� and then a slow stage of transition towardslower oxidation follows. The majority of researchershold that Cr3� constitutes the active sites fordehydrogenation, but there is no lack of reference inscientific literature to a catalytic activity carried out inaddition, or principally, by Cr2� (Grünert et al., 1986;Ashmawy, 1980). There is agreement, though, on theconsideration that the formation of solid solutions ofchromium and aluminium oxides is responsible for theirreversible deactivation of the catalysts. In Fig. 5 thedifferent types of Cr3� present on the reduced chromia-alumina system are illustrated (for simplicity the atomsof oxygen between the Cr3� ions are omitted; Puurunenand Wekhuysen, 2002) along with their evolution

towards species of a-chromia-alumina (a-4Al2O3-Cr2O3) with reduction both of the surfacearea and catalytic activity.

The entry of Cr3� into the structure of the alumina ishelped by localized overheating, caused by combustionof the coke during regeneration of the catalyst. Thespeed of interaction between Cr and Al can be slowedby an opportune modification of the surfaces(Sanfilippo et al., 2001; Iezzi et al., 2002). Fig. 6 showshow, in the presence of various heteroatoms, such assilica and tin, a reduction in the speed of formation ofa-4Al2O3-Cr2O3 is observed, both in terms of a smallerconcentration of chromia-alumina operating in reactionconditions, and in terms of a reduction in the loss ofsurface area due to ageing simulated with thermaltreatment based on increasing temperatures. It ishypothesized that the heteroatom makes the surface ofthe alumina less reactive for the Cr3�.

The alkali metals have been indicated as promotersof activity and selectivity. Their presence increases thenumber of active chromium sites and reduces thesurface acidity. Not all alkaline metals have beenidentified as promoters of chromium catalysts, however.Only cesium, potassium and rubidium, due to astabilizing effect deriving from the large dimensions oftheir respective cations, have had this characteristicconfirmed. All research has, in any case, shown how theactivity of dehydrogenation is directly proportional tothe amount of chromium present on the surface,independent of any other chemical groups.

Chromium catalysts are reactivated by treatmentwith air, which has the following effects: partialoxidation of Cr3� to Cr6�, which contributes to thedispersion of the crystallites of Cr2O3; and combustionof coke, very probably catalysed by Cr6�, whichincreases the temperature of the catalyst and can beused to supply heat for the endothermicdehydrogenation reaction.

A.G. Zwahlen (Zwahlen and Agnew, 1992), whoreports a kinetic study on the dehydrogenation ofisobutane, carried out in a complete mixing reactor on aCr2O3/Al2O3 catalyst, interprets the experimental databy means of a Langmuir-Hinshelwood type of kinetic



DG

° (kc

al/m

ol)

0

5

10

15

20

25

30

35

40

45

50

temperature (°C)0 200 400 600 800 1,000

Cr2O3�1.5 O2(g)�2 CrO3

Fig. 4. Free energy for the reaction of the formation of CrO3from Cr2O3.

�

isolated Cr3�

unexposed Cr3� a-chromia-alumina

alumina

clusters Cr3�

Fig. 5. Models for Cr/Al catalysts.

equation, assuming that the adsorption of isobutanerepresents the slow stage of the reaction and thatcompetition with adsorption of hydrogen does not exist.

Reactor design and commercial applicationsPropane, n-butane, isobutane and isopentane are

presently dehydrogenated to the correspondingmonolefin in various industrial facilities around theworld. Catofin, Oleflex, STAR and FBD (Fluidized BedDehydrogenation) are technologies used industrially,while the PDH (Propane DeHydrogenation) has beenused only on a demonstration unit.

An ideal reaction system for the industrial processof dehydrogenation of light paraffins must supply alarge amount of reaction heat at a temperature above770 K, while maintaining strict control of thetemperature, to minimize the formation of by-products.In addition, it must permit a periodic regeneration of thecatalyst to remove carbon compounds that accumulateon its surface. These conditions, together with thethermodynamic and kinetic limitations described above,have pushed researchers to develop optimal reactordesign solutions for the industrial exploitation ofdehydrogenation reactions. The available commercialtechnologies offer different choices as regards thereaction system, tending to simultaneously optimize the reaction conditions and the supply of energy to thereaction.

Catofin TechnologyCatofin technology (Arora, 2004), commercialized

by ABB Lummus, derives directly from the Catadieneprocess, originally developed by Eugène Houdry for thedehydrogenation of n-butane to butadiene. Thetechnology uses fixed-bed adiabatic reactors in which

there is an alternation between reaction andregeneration cycles. Many reactors are used in parallelin a plant, to permit continuous production. Thereaction heat is stored up in the catalyst duringregeneration and released during the reaction. The swiftpassage between an oxidizing environment and areducing environment, along with a short reactionperiod, is compatible with the choice of chromiumcatalysts. The diagram of the Catofin process relative tothe reaction zone is shown in Fig. 7 A.

The reaction takes place on catalysts based onchromium and alumina oxides, at sub-atmosphericpressure, and is conducted using a certain number offixed-bed (at least three) adiabatic reactors, in whichthe reaction and the regeneration of the catalyst withair (as described above) are cyclically alternated. Thecatalyst, in the form of cylindrical pellets, consists ofbetween 18-20% in weight of chromium oxide(Cr2O3) and 1-2% in weight of alkali metals,supported on alumina with a surface area of about120 m2/g. During the ageing of the catalyst, which iscompletely replaced every 1-2 years, the averagetemperature of reaction is progressively increased inorder to maintain a constant production output,causing a gradual loss of selectivity which cannot beoverlooked. Typical operating conditions require atemperature ranging from 860 to 920 K, a pressure ofbetween 33 and 50 kPa and a space velocity (based onliquid feed) of between 0.4 to 2 normal-litre/h/litre ofthe catalyst.

Other interesting characteristics of this technologyare:• Low consumption of reagent, due to the high

selectivity ensured by the catalyst. • Heat of reaction, supplied both by the combustion of



Me

O

�Me (II/IV)(Me � Si, Sn)

Al

O O

OK

Al

O

Al

Al

O O

OOO

Al

O

Al

K0

10

20

30

40

50

800700 900 1,000

a-C

r-A

l (%

wt)

surf

ace

lost

(%

)

16

12

8

4

00 1,000 2,000 3,000 4,000

hours on stream temperature (°C)

phase evolutioncalcination time 1 hour

a-chromia alumina formation (%) thermal stability (simulated aging)

carrier d-q Al2O3 as such carrier Al2O3 “Si modified” carrier Al2O3 “Si�Sn modified”

Fig. 6. Interaction of Cr with the carrier.

the coke, and by the surface reduction of the catalystwith natural gas and possibly by co-feed of a fuel

during the regeneration phase. • High yield per pass, made possible by carrying out

the reaction below atmospheric pressure. • Catalyst endowed with high level of thermal

stability, resistance to friction and tolerance topotential poisons such as water vapour and heavymetals.

Oleflex technologyOleflex technology, developed and commercialized

by UOP, utilizes mobile adiabatic-bed reactors in series.The catalyst flows slowly, pulled by gravity through thereaction system, while the reagents flow radially withinthe various reactors. The heat for the reaction issupplied directly to the reagents by intermediate heatingin fired furnaces, placed between one reactor andanother. The catalyst is collected on the bottom of thelast reactor, transported pneumatically to theregenerator, and then sent back into the reaction zone.For further details regarding the Oleflex technology, seeSection 10.2.1

STAR technologySTAR technology (STeam Active Reforming),

originally developed by Phillips, is now commercializedby Krupp-Uhde. It uses an idea similar to that normallyused in the steam reforming of hydrocarbons to producesynthesis gas. The catalyst is contained in tubes placedinside a fired furnace, in which a fuel is burnt to supplythe heat of reaction. In this case, too, the need forperiodic regeneration of the catalyst imposes the use ofseveral parallel reactors. An innovative characteristic of this technology is represented by the insertion,downstream of the main reactor, of a secondary reactor(fixed-bed adiabatic) for oxydehydrogenation, a processin which oxygen is fed and reacts with part of thehydrogen formed, permitting further progress of the main reaction. This technology is suitable for the useof platinum catalysts, which allow long reaction times,on the order of hours, before regeneration. The schemefor the STAR process relative to the reaction zone isshown in Fig. 7 B.

STAR technology is used for the dehydrogenation oflight paraffins and the dehydrocyclation of paraffins C6-C7 in tubular, fixed-bed reactors, placed inside afired furnace, as described above. The process uses acatalytic system which is very stable in the presence ofwater vapour as a diluting agent. The active phase of thecatalyst is made of platinum (0.01-5% in weight),promoted with tin (0.1-5% in weight), supported onzinc or magnesium aluminate and bound with calciumaluminate. The use of alkaline metals with the functionof promoting the selectivity is optional. The steamadded to the reagent mixture plays an important role,since it increases the driving force of the reaction, while



A

B

C

feedair

products exhaust air

fuel flue

feed

feed

steam

reactorin dehydrogenation

boiler feedwater

products

products

steam

reactor indehydrogenation

boiler feedwater

reactorin regeneration

air

flue

A

B

C

Fig. 7. Configuration of commercial reactors for paraffindehydrogenation (A, Catofin technology; B, STARtechnology; C, PDH technology).

allowing a relatively high total pressure to bemaintained (around 5 bar); that, in turn, allows anoptimization of investment and operating costs of thereaction section and compression of the effluent, whileat the same time avoiding the possible infiltration of airfrom the outside. Steam also acts as a thermalmoderator, thus preventing an excessive decrease intemperature that could block the progress of thereaction. In addition, steam inhibits the side reactionsthat lead to the deposition of coke on the catalyst. Allthis enables maintaining a cycle of about seven hours ofreaction, followed by one hour of regeneration. In thesuccessive adiabatic oxydehydrogenation reactor, theoxygen (about 90% pure) is fed in, which, reacting withpart of the hydrogen formed in the dehydrogenationreaction, allows a further advancement of thedehydrogenation reaction itself, because of both theelimination of the hydrogen and the increase intemperature. In the oxydehydrogenation reactor, oxygenis supplied in a molar ratio of between 0.08 and 0.16with respect to the hydrocarbon supply. Steam, on theother hand, is fed into the main reactor in a molar ratiobetween 3.5 and 4.2 with respect to the hydrocarbonsupply. The starting temperature to the main reactor isabout 880 K, with an exit temperature ranging between920 and 940 K, while the exit temperature on theoxydehydrogenation reactor is about 950 K. The spacevelocity, based on the supply in liquid state, is typicallyequal to 4 normal-litre/h/litre of catalyst in thetraditional version and increased to 6 normal-litre/h/litreof catalyst in the version with oxydehydrogenation,which allows a reduction in the dimensions of the costlymain reactor, while adding a secondary reactor ofsimple construction.

PDH technologyA similar technology but without the oxidative

dehydrogenation stage and using a chromium catalyst,was developed originally by Linde, Borealis and Statoil(Zimmermann and Versluis, 1995).

This technology has elements in common with theSTAR technology, in that it uses fixed-bed reactors inwhich the heat of reaction is supplied by burning a fuelexternally (Fig. 7 C). The process requires three parallelreactors, two of which are normally working in thereaction, while the third is the regeneration phase. Thedistinctive characteristic of PDH technology is theabsence of reagent dilution, which permits a reductionof the reactor dimensions and simplifies the purification of the product. This is made possible thanksto strict control of the reaction temperature, whichmakes the reactors practically isothermal. The catalyst isbased on chromium oxide supported on alumina.Recently, Pt catalysts on various oxides have also beenindicated (Schindler et al., 2004).

FBD technologyFBD technology (Sanfilippo, 2000; Miracca and

Piovesan, 1999), commercialized by Snamprogetti, usesa fluidized-bed reactor with a catalyst that circulatescontinuously from the bottom of the reactor to the topof the regenerator plant and viceversa. The heat neededfor the reaction process is supplied by a fuel burneddirectly in the regenerator and transported to the reactorthrough the heat capacity of the regenerated catalyst. Inthe reactor, suitable baffles are fitted internally (forexample, horizontal grids or tubes) with the end oflimiting the phenomenon of back-mixing, typical ofbubbling fluidized-bed systems. The FBD processscheme relative to the reaction zone is shown in Fig. 8.

FBD technology derives conceptually from FCCtechnology, widely used in refineries to increase theyield of gasoline from crude, converting heavy cuts oflesser value by means of controlled breaking of thebonds between carbon atoms.

Cracking reactions, even if not limited byequilibrium, have characteristics that are very similar todehydrogenation reactions. In fact, they are alsoendothermic and lead to the formation of carbon by-products that inhibit catalytic activity. In fluidized-bedreactors a catalyst in microspheroidal form is used(average diameter of particles less than 0.1mm) withparticularly suitable fluid-dynamic characteristics andresistance to attrition. The catalyst contains chromium



air flue

stack

hydrogenand light

ends

isobutyleneand isobutane

feedisobutane

aircompressor

Fig. 8. Fluidized-Bed Dehydrogenation (FBD) processscheme.

oxide (Cr2O3) in quantities ranging between 12 and20% by weight, potassium oxide (1-2% by weight) andsilica (1-2% by weight) supported on alumina. Typicaloperating conditions include a reaction temperaturebetween 820 and 870 K, an operating pressure slightlyabove atmospheric pressure (1.1-1.5 bar) and a spacevelocity (calculated on liquid feed) of between 0.4 and 2 normal-litre/h/litre of catalyst. In this technology, the flow rate of the circulating catalyst is dictated by theneed to transport heat to the reaction rather than that ofregenerating the catalyst itself. It is necessary, in fact, tohave from 5 up to 15 kg/h per kg of feed of hotregenerated catalyst circulating, to satisfy the thermalrequirements of the reaction. The use of fluidized-bedtechnology allows continuous operation under stableoperating conditions. The catalyst can be replaced orfilled-up while the plant is operating, eliminating theneed to halt operations. In this case, too, the physicalseparation between the reaction zone and theregeneration zone, together with an operating pressurewhich is higher than the atmospheric pressure, ensuresthe safety of the operation. The thermal profile of thegas is characterized by reaching the highest temperatureat the exit of the reactor; therefore, the conversionobtainable per pass is higher, due to the approach to amore favourable equilibrium.

In Fig. 9, the qualitative thermal profiles inside thecatalytic bed are shown for the types of reactormentioned above, based on the hypothesis ofobtaining the same conversion for every reactor. Thecurve E represents the conversion of equilibriumrelative to temperature. In the case of tubular reactorsinserted into a fired furnace, the reaction can takeplace in almost isothermal conditions (curve A). Inadiabatic reactors, in which the heat of reaction issupplied to the reagents in a gaseous state before theyenter the reactor, the decrease in temperature of thegas caused by the reaction itself is so large, as not topermit the use of a single catalytic bed. The reaction is therefore conducted in a certain number ofreactors in series with intermediate heating of thereacting gas in external fired furnaces (curve B).When, instead, the release of heat by the catalyst (that as a result cools) meets the needs of theendothermic reaction, the angular coefficient of thestraight line that represents the thermal profile (curveC) can be modified, varying the quantity of catalyst inthe reactor and making it possible to carry out thereaction in a single reactor. In this case, it becomesnecessary to heat the solid periodically when it hasexhausted the stored heat; the reactor works thus in acyclical mode, and more reactors in parallel areneeded to maintain continuous production. In afluidized- bed reactor (by definition well-mixed andisothermal), the thermal profile can be modified

towards a course roughly represented by curve D, ifsuitable internal baffles are inserted to limit theinternal mixing. From a kinetic point of view, a higheraverage temperature makes possible a reduction in thequantity of catalyst required and diminishes thereaction volume, while a lower average temperaturemakes it possible to obtain greater selectivity withregard to the desired product, reducing the influenceof secondary reactions. To determine the optimal typeof reactor, additional considerations are necessary,relating to the efficiency with which the heat ofreaction is supplied in the various cases, to theassociated cost of investment and to the possibility ofdistributing the heat in a homogeneous manner, thusavoiding the formation of hot spots that leadinevitably to a reduction in selectivity.

11.2.2 Dehydrogenation of heavyparaffins

The catalytic dehydrogenation of heavy paraffins (C10-C14) leads to a family of linear monolefins with thesame number of carbon atoms and the double bond inrandom position along the chain. This family of olefinsis used to synthesize alkyl benzenes, used as



AC

BE

D

conv

ersi

on (

%)

0

10

40

20

50

30

60

80

70

90

100

temperature (K)700 750 800 850 900 950 1,000

equilibrium

fired tubular reactor, almost isothermal -heat to gas through wall

adiabatic multibed reactor - reaction heatsupplied to gas in interstage heat exchangers

adiabatic fixed bed - reaction heatfrom solid heat capacity

staged fluidized bed - reaction heat fromcountercurrently circulating catalyst

Fig. 9. Qualitative profiles of reactortemperature-conversion in isobutane dehydrogenation.

intermediates in the production of biodegradabledetergents, through reaction with benzene in conditionsof acid catalysis. Around 1960, the growing demand forlinear monolefins, suitable for the production ofdetergents, created excellent prospects for thedevelopment of dehydrogenation technologies. Since1969, UOP has commercialized the combined processPacol-Olex for the dehydrogenation of linear paraffinsand the selective extraction of the olefins produced; aprocess, which even today remains the most successfultechnology in the field.

From a thermodynamic point of view, thedehydrogenation of heavy paraffins must overcome theconstraint resulting from the fact that the energy requiredto break a bond between two carbon atoms (245 kJ/mol)is less than the energy necessary to break the bondbetween an atom of carbon and an atom of hydrogen(365 kJ/mol; Griesbaum et al., 1989). Therefore, thermalreaction is not advisable, since it would lead to apredominance of cracking reactions overdehydrogenation reactions. On the other hand, a catalyticprocess allows the reaction to proceed at a relatively lowtemperature, minimizing the formation of by-productsand obtaining conversion and selectivity such as to makeexploitation of the reaction financially attractive.

In Pacol technology, the reaction takes place in afixed-bed reactor, which works in a wide-range ofoperating conditions, depending on the type of feed andthe conversion per pass desired. The temperature can bebetween 570 and 820 K and the pressure between 100and 300 MPa. The conversion per pass is 10-15% inweight, with selectivity towards useful olefins 90-94%in weight. The catalyst is based on platinum supportedon alumina, with the addition of promoters (tin, indium,lithium). The low conversion per pass implies that theseparation and the purification of the products and therecycling of the non-converted reagents are afundamental factor in the economics of the process. Tothis end, UOP developed a process of extraction basedon molecular sieves with high capacity to absorb olefins(Broughton and Berg, 1970). Even if the catalyst is veryselective, the formation of by-products reduces itsactivity over time and its useful life is limited to aboutone year.

11.2.3 Dehydrogenation of ethylbenzene

Ethyl benzene is dehydrogenated to styrene andhydrogen on iron-based catalysts in the presence ofsteam (James and Castor, 1994):

C6H5�CH2�CH3��C6H5�CH�CH2�H2

Cracking gives rise to unwanted by-products, such as

benzene and ethylene (that are in effect the reagentsfrom which we obtain ethyl benzene industrially),toluene and methane; by- products originating fromconsecutive reactions are phenyl-acetylene and a-methyl-styrene. Since the industrial technologies donot provide for catalyst regeneration cycles, it isabsolutely fundamental to eliminate the formation ofcoke. Carrying out the dilution with steam andoperating under vacuum are essential features; inparticular, the dilution with steam reduces the potentialfor the formation of coke, acts as a diluting agent,maintains the correct state of oxidation of the catalystand supplies sufficient heat capacity to lessen theadiabatic reduction of temperature associated with theconversion of ethyl benzene.

The temperature of reaction is between 550 and630°C. At lower temperatures the kinetics are too slowand both ethyl benzene and styrene rapidly undergocracking above 610-620°C. A rapid quenching istherefore necessary to cool the effluent products fromthe catalytic bed and minimize the thermal formation ofby-products. The conversion per pass is generally 60-70%, with selectivity to styrene over 93-95%.

The catalyst is essentially based on iron oxide, withpotassium as a promoter. Other metals are used assupplementary promoters, such as chromium oxides,cerium, tin, calcium, molybdenum, magnesium,titanium and others (Muhler et al., 92). The most recentformulations do not use chromium and are able tooperate at low ratios of steam/carbon, limiting theformation of phenyl-acetylene to 20-40 ppm. Theprincipal technologies are commercialized byLummus/UOP and Atofina/Badger.

Typically, fixed-bed and radial flow adiabaticreactors are used to reduce the load losses subdividingthe reaction between multiple reactors in series withintermediate heating. Recently, the use of moreinnovative reactor concepts has been proposed(Sanfilippo et al., 2004), in particular those involvingfluidized-bed reactors and catalysts which are not ironbased, but use gallium with promoters (Pelati andGulotty, 2004).

11.2.4 Future developments

A high number of R&D initiatives are attempting todevelop new production processes for olefins, but onlya few have real prospect of eventual success on anindustrial level, due to both intrinsic difficulties with thetechnology itself and the question of their economicviability.

In the last decade, research in the field ofdehydrogenation of paraffins has concentrated onoxydehydrogenation, that is to say, the coupling of the



dehydrogenation reaction with a controlled combustionof the hydrogen co-produced through the co-feed ofoxygen to the dehydrogenation reactor.

This would make it possible to get around theconstraints imposed by thermodynamics and conductthe reaction at low temperature, in conditions whichcompletely inhibit the side reactions of thermal crackingand even eliminate the need for periodic regeneration ofthe dehydrogenation catalysts.

Although there are encouraging results, therealization of industrial oxydehydrogenation applied tothe principal reaction still seems far off. In reality, thisapplication requires the development of a catalyst withvery particular characteristics, one that must beextremely selective in the oxidation of hydrogen, avoidcombustion of the hydrocarbons, maintain its activity inthe presence of steam and be sufficiently active toconsume the supplied oxygen completely.

It has been proposed that different catalysts be usedfor the oxidation process and for that ofdehydrogenation, to be installed in different parts of thereactor. The possibility of burning the hydrogenbetween the two stages of dehydrogenation, to move thethermodynamic equilibrium to the right, has also beensuggested in scientific literature (Grasselli et al., 1996).Indeed, it has been proposed in various technologies, forexample the STAR technology, limited only to thecompletion reactor, operating with a very reducedquantity of oxygen with respect to the hydrocarbonspresent. In the dehydrogenation of ethyl benzene, thistechnique is also applied: in the Lummus/UOP/SMARTtechnology an intermediate palladium catalyst bedburns the hydrogen that is formed in the precedingdehydrogenation bed with added oxygen, supplyingheat for the endothermic dehydrogenation reaction. Inthis way, by reducing the hydrogen content, thethermodynamic constraint is removed, and a betterconversion rate reached.

It would appear more difficult to develop a solematerial able to catalyse both reactions simultaneously.Attempts carried out so far have focused on the additionof more promoters to platinum catalysts, to make itpossible to direct the combustion preferentially towardshydrogen. Recent studies have concentrated onmolybdates of cobalt, manganese or nickel. Thepossibility is also being evaluated of using bothmolecular oxygen and structural oxygen, of a catalystable to accumulate or release oxygen with a redoxmechanism (Iglesia et al., 2002). Although there areencouraging results, there is no prospect of an industrialapplication in the short term.

A certain interest has been sparked in theproduction of ethylene from ethane by means ofoxidising dehydrogenation, using extremely lowcontact time (milliseconds) reactors on platinum

catalysts (Schmidt Bodke et al., 1999; Beretta et al.,2001; Font Freide et al., 1990; Bharadwaj et al., 2000;Donsì et al., 2005).

An alternative method involving the separation ofhydrogen as soon as it is formed consists in theutilization of hydrogen permeable membranes placeddirectly inside the dehydrogenation reactor. In this case,the challenge consists in developing membranes whichare able to reliably endure a high temperatureenvironment and potential abrasion from the catalyst.The use of reactors with permselective membranes forH2 that can carry, with supports, the active phase fordehydrogenation, or, be limited to the separation of hydrogen, is being studied. Catalytic combustion ofhydrogen is also being studied (Iglesia et al., 2003).

The extension to ethane of traditional catalyticdehydrogenation, does not appear to be suitable as such:in fact, to reach economically attractive levels ofconversion, it would be necessary to operate at veryhigh temperatures (greater than 700°C) and in this casethermal reactions are already prevalent, as they takeplace in classic steam cracking. Nevertheless, catalyticdehydrogenation at lower temperatures (and obviouslywith lower conversion rates) appears attractive, if it iscoupled with a direct use of ‘‘diluted’’ ethylenecontained in the effluent flow from the reactor, as, forexample, in alkylating benzene, and producing ethylbenzene that is then dehydrogenated to styrene (Pogueet al., 1993; Buonomo et al., 1998) or that of propylenediluted by dehydrogenation of propane, to alkylatebenzene and obtain cumene (Paggini et al., 2002).

Obviously, future developments will depend on thepossibility of obtaining high selectivity (more efficientuse of raw materials), on the energy efficiency of thenew process, on its technological reliability andeconomics with respect to existing technologies.

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Domenico SanfilippoIvano Miracca

SnamprogettiSan Donato Milanese, Milano, Italy

Ferruccio TrifiròDipartimento di Chimica Industriale e dei Materiali

Università degli Studi di BolognaBologna, Italy



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11.2 Dehydrogenation processes