ORIGINAL PAPER

Development of Dehydrogenation Catalysts and Processes

Bipin V. Vora

Published online: 13 November 2012

� Springer Science+Business Media New York 2012

Abstract Catalytic dehydrogenation plays an important

role in production of light (C3–C4 carbon range), detergent

range (C10–C13 carbon range) olefins and for ethylbenzene

dehydrogenation to styrene. During the World War II,

catalytic dehydrogenation of butane over a chromia–alu-

mina catalyst was practiced for the production of butenes

that were dimerized to octenes and hydrogenated to octanes

to yield high-octane aviation fuels. The earlier catalyst

development employed chromia–alumina catalyst and

more recent catalytic developments use platinum or mod-

ified platinum catalysts. Dehydrogenation is a highly

endothermic process and as such is an equilibrium limited

reaction. Thus important aspects in dehydrogenation entail

approaching equilibrium or near-equilibrium conversion

while minimizing side reactions and coke formation.

Keywords Paraffin dehydrogena � Olefins production �Platinum catalysis

Significant development work has been done on paraffin

oxydehydrogenation that would allow selective combustion

of hydrogen and thus shift the equilibrium more in favor of

higher conversion. This approach was practiced during

1960s for the production of butadiene by Petro-Tex

Chemical Corporation and by Phillips Petroleum Com-

pany. However, as significant quantity of butadiene

became available as byproduct of naphtha cracker for

ethylene, this route is not practiced today.

Commercial processes for catalytic dehydrogenation of

propane and butanes attain 30–60 % per-pass conversion

while C10–C13 range paraffin dehydrogenation operate at

per-pass conversion range of 10–15 %. Ethylene and pro-

pylene are No.1 and No.2 largest volume petrochemical

intermediates. Figure 1 shows worldwide demand of eth-

ylene and propylene and growth from 1990 to 2010 [1]. For

ethylene thermal cracking of ethane, LPG and heavier

feedstocks continues to be the primary route. Thermal

cracking of LPG or heavier feedstocks also provide sig-

nificant quantity of propylene byproduct. Fluid catalytic

cracking in refineries also produce significant quantity of

propylene byproduct. However, over the past two decades

propylene growth rate has out-paced these conventional

supply routes leading to construction of number of com-

mercial units for selective catalytic dehydrogenation of

propane to propylene. Also during 1990s several isobutane

dehydrogenation units were built producing isobutylene

that were converted to MTBE or dimerized and hydroge-

nated to produce isooctane, a high octane gasoline blending

components.

1 Historical Developments

Chromia–alumina catalyst for the production of olefins has

been in use since the late 1930s. During World War II,

catalytic dehydrogenation of butanes over a chromia/alu-

mina catalyst was practiced for the production of butenes

which were then dimerized to octenes and hydrogenated to

octanes to yield high-octane aviation fuel.

Dehydrogenation of butanes over a chromia/alumina

catalyst was first developed and commercialized at Leuna

in Germany and also independently by Ipatieff and co-

workers at UOP [2, 3]. The first UOP-designed plant came

Prepared for Dr. Jeff Bricker 2011 ACS Award Symposium.

B. V. Vora (&)

UOP LLC, A Honeywell Company, Des Plaines, IL, USA

e-mail: bvvora@uop.com

123

Top Catal (2012) 55:1297–1308

DOI 10.1007/s11244-012-9917-9

on stream at ICI in Billingham in 1940 and was soon fol-

lowed by two other units in Heysham in 1941 [4]. The

production of octenes by butene dimerization made use of

UOP’s catalytic condensation process in which olefins are

dimerized, oligomerized, or alkylated over a solid phos-

phoric acid (SPA) catalyst, discovered by Schaad and

Ipatieff, and first commercialized in 1933 [5].

These pioneering efforts were soon followed by other

companies (e.g., Phillips Petroleum’s multitubular dehy-

drogenation reactor near Borger, Texas, in 1943 [4]).

However, the most significant development was that made

by Houdry using dehydrogenation under vacuum for higher

per-pass conversions and commercialized two stage butane

dehydrogenation chromia/alumina system, known as Cat-

adiene process, for the production of butadiene [4].

The Houdry Catadiene process was used extensively for

the production of butadiene, either by itself (n-butane to

butadiene) or in conjunction with the Oxo-D catalytic

process for the oxydehydrogenation of n-butene to buta-

diene that was commercialized by what was then known as

Petro-Tex Chemical Corp. [6]. A similar oxydehydrogen-

ation approach for the production of butadiene was fol-

lowed by Phillips Petroleum in their O-X-D process [6].

Large quantities of butadiene have become available in

the market over the past 30 years, mostly as a byproduct

from the thermal cracking of naphtha and other heavy

hydrocarbons. The result from this market shift has been

the shutdown of all ad-hoc catalytic dehydrogenation units

for butadiene production in North America, Western Eur-

ope, and the Far East. However, over the past two decades

because of changing to lighter (ethane and LPG) feedstocks

for the ethylene plant crackers, rate of butadiene produc-

tion growth has slowed below the demand growth creating

regional shortage of butadiene. Butadiene price has jumped

from $800/MT in 2005 to over $2000/MT in 2011. As a

result of this there is again increasing interest in on-purpose

butadiene production. In August 2011 Texas Petrochemical

(TPC Group Inc) announced detail engineering study of on-

purpose butadiene production utilizing their idled Oxo-D

plant. In 2011 Mitsubishi announced a development of

oxidative dehydrogenation process for butane dehydroge-

nation for the production of butadiene [7].

In the late 1980s, application of chromia/alumina cata-

lysts was extended by Houdry to the dehydrogenation of

propane to propylene and isobutane to isobutene, and over

the two decades there were several new units were built for

this purpose. These units again operate on the same cyclic

principles as in the former Catadiene process, and the new

process application is named Catofin [8, 9]. The Catofin

process technology is currently owned by Sud-Chemie and

is offered by Lummus, A CBI Company.

In about 1959 an alternative chromia/alumina catalytic

dehydrogenation process was developed in the former

Soviet Union that avoided the use of the cyclic operation

by using a fluidized bed reactor configuration similar to the

fluid catalytic cracking (FCC) process used in refineries

[4]. However, back mixing common to dense fluidized bed

operations results in poor selectivity and increases the

formation of heavies, sometimes called ‘‘green oils’’. Cir-

culating regenerated catalyst is used to provide the heat of

reaction in the riser and spent catalyst is reheated by carbon

burn in a regenerator. A larger scale isobutane dehydro-

genation unit using this principle was commercialized by

Snam Progetti in Saudi Arabia based on technology from

Yarsintez in Russia [10].

Lestor, Carson, and others at UOP during late 1960s and

early 1970s also worked on development of fluid-bed cat-

alytic dehydrogenation employing chromia alumina cata-

lyst but the technology was not commercialized [11–14].

Chromia/alumina catalysts pose a significant health risk in

case of spillage or by exposure to the plant operators during

maintenance or catalyst changeover. Chromia/alumina

catalysts always contain a significant proportion of Cr(VI),

principally in the regenerated catalyst; Cr(VI) is a well-

known carcinogen and its adverse health effects have been

well documented [15, 16]. Chromia/alumina catalysts sin-

ter much more rapidly than alumina when exposed to high

temperatures; the replacement of the mass of spent catalyst

often requires strenuous and lengthy use of manual labor

for such an operation [17, 18].

The dehydrogenation of ethylbenzene to styrene reac-

tion proceeds over an iron or an iron–chromium catalyst

that usually also contains potassium in the form of potas-

sium carbonate, so that at elevated temperatures various

complex mixed carbonates and oxides are formed; e.g.,

KFeO2. Temperatures are elevated, in the order of 630 �C,

and steam dilution is practiced to lower the partial pressure

of the reactants. Because the reaction is strongly endo-

thermic various reaction stages are normally employed

with interheat and interstage addition of superheated steam.

120

0

Dem

and,

MM

MT

A

1995 2000 2005

100

80

60

40

20

EthylenePropylene

2010

Fig. 1 Ethylene and propylene demand

1298 Top Catal (2012) 55:1297–1308

123

2 Noble Metal Dehydrogenation Catalysts

A different approach to catalytic dehydrogenation was first

introduced in the mid-1960s, for the production of linear

olefins of C10–C13 carbon range for the production of

biodegradable detergents.

Synthetic detergents based on the use of alkylbenzene

sulfonates had been introduced in the 1940s. The manu-

facture of these early detergents made use of UOP’s cata-

lytic condensation process to oligomerize propylene to a

mixture of branched dodecene isomers. The branched do-

decenes were then alkylated with benzene in the presence

of a strong acid, usually HF, followed by sulfonation and

neutralization to yield dodecylbenzenesulfonate, the active

ingredient.

By the early 1960s it became apparent that branched

dodecylbenzene-based detergents, though very active and

offering excellent detergency, did not biodegrade readily

and were accumulating in our lake and river waters. The

need for biodegradable detergents prompted the develop-

ment of catalytic dehydrogenation of linear paraffins to

linear olefins.

The work on catalytic reforming of naphtha with noble

metal (Pt) catalysts done in the 1940s by Haensel [19] had

clearly demonstrated that Pt-based catalysts had high

activity for the dehydrogenation of paraffins to the corre-

sponding olefins. In the 1960s Dr. Bloch [20] further

extended this thinking by developing Pt-based catalysts

that could selectively dehydrogenate heavy linear paraffins

to the corresponding internal mono-olefins with high

activity and stability and with minimum cracking. This was

the basis for the UOP PacolTM process for the production

of linear olefins for the manufacture of biodegradable

detergents [21]. Independently Roth [22] at Monsanto

developed a paraffin dehydrogenation catalyst and com-

mercialized dehydrogenation alkylation combination for

production of linear alkylbenzenes. There was only one

plant built at Louisiana and now owned by Huntsmann

Chemicals

3 Light Olefins

In view of the successful application of Pt-based noble

metal catalysts to the dehydrogenation of heavy paraffins it

would seem that their extension to the production of light

olefins would be a trivial undertaking. In reality, this is not

the case.

Heavy paraffins are both valuable and highly prone to

cracking. Therefore, in order to maintain a high selectivity

and yield it is necessary to operate at relatively mild con-

ditions, typically below 500 �C and at relatively low per-

pass conversions and higher H2 partial pressure. While this

is economical for the production of heavy linear olefins, it

is not for the production of light olefins.

Paraffin dehydrogenation is an endothermic reaction that

is limited by chemical equilibrium and, according to Le

Chatelier’s principle, higher conversion will require either

higher temperatures or lower pressures. In a somewhat

abbreviated form for the production of mono-olefins, this

can be expressed as follows:

xe ¼Kp

Kp þ P

where, xe is the equilibrium conversion, P is the total

pressure in atmospheres absolute and Kp is the equilibrium

constant for the dehydrogenation reaction. The equilibrium

constant can be easily calculated from Gibbs free energies

as tabulated in the API 44 report or in similar sources of

thermodynamic data. Figures 2 and 3 illustrate the equi-

librium conversion levels that can be obtained for propane

at 1 and at 0.23 atm. abs. (175 torr), respectively.

The equilibrium constant for paraffin dehydrogenation

increases significantly as the carbon number increases. The

temperature required for the dehydrogenation of light

paraffins is much higher than for heavy paraffins. For

example, at 1 atm absolute pressure for 40 % conversion,

the dehydrogenation of propane requires a temperature of

at least about 580 �C, while dodecane can be theoretically

dehydrogenated to the same extent at only 450 �C. The

equilibrium conversion increases at higher temperatures,

but side reactions, coke formation, and catalyst deactiva-

tion also are accelerated. Thus, we cannot extrapolate

directly from heavy olefins to light olefins without taking

other factors into consideration.

Light olefins, especially ethylene and propylene, are made

in very large steam cracking units using a variety of

100

90

80

70

60

50

40

30

20

10

0627 727 827 927 1027 1127

Temperature, °C

C-C=CPropylene

C-C CMethyl

Acetylene

C=C=CAllene

C-C-CPropane

Mol

e %

Fig. 2 Propane dehydrogenation equilibrium at 1.00 atm abs pressure

Top Catal (2012) 55:1297–1308 1299

123

feedstocks that may range from ethane to vacuum gas oil.

Modern facilities have capacities for up to 1.5 million MTA

ethylene using a plurality of furnaces in parallel, each for a

capacity of about 200,000 MTA ethylene, at very high

severities in excess of 800 �C and at high per-pass conver-

sions and at rather low selectivities. For example, while the

yield of ethylene from ethane cracking is close to 80 wt%,

the yield drops to about 30 wt% for a naphtha feedstock.

More significantly, the yield of propylene is only 13–17 wt%

of the feed for practically all feedstocks, from propane to gas

oil. Likewise, the yield of the C4 product from a steam

cracking unit with a light naphtha feed is 8–11 wt% and

about one half of that is butadiene [4].

Production of light olefins by the catalytic dehydroge-

nation of light paraffins must be able to maintain reason-

able per-pass conversion levels and high olefin selectivity.

Also, it must be able to produce olefins in high yields over

long periods of times without shutdowns and with high

operating efficiency. From what we saw above, this can

pose a formidable challenge.

In the early 1970s UOP introduced continuous catalyst

regeneration (CCR) technology that enabled noble metal

catalysts to remain at their most desirable stable activity. In

this mode, a small amount of the catalyst from the reactor is

removed and sent to a separate regeneration vessel where

carbon burning and platinum redistribution are completed.

This regenerated catalyst is returned to the top of the

reactor. In this steady state continuous mode of operation

the catalyst maintains constant activity. Use of CCR tech-

nology makes it possible to operate a dehydrogenation

catalyst at high severity without fear of coking and without

the need to frequently shut down the reactor for catalyst

regeneration. While CCR technology was first introduced

for the UOP PlatformingTM catalytic reforming process, it

has proven to be equally useful for the catalytic aromati-

zation of LPG fractions such as in the UOP CyclarTM pro-

cess and for the catalytic dehydrogenation of light paraffins.

The combination of noble metal catalysts operating at

high severity in conjunction with CCR technology has made

it possible to design, build, and operate very large catalytic

dehydrogenation units that can produce light olefins, in

particular propylene and isobutylene, at high selectivities

and economically [23]. Not surprisingly, the world pro-

pylene production capacity based on the use of catalytic

dehydrogenation of propane has increased steadily over the

past 20 years [24] with 2010 production of propylene via

propane dehydrogenation at 3 million metric tons per year.

Use of the Oleflex process for the dehydrogenation of

ethane to ethylene has also been investigated but, to date,

the economics do not appear to be favorable because of the

low equilibrium conversion and the need to operate under

vacuum if a reasonable ethane conversion is to be expected.

The cost of fractionating ethylene in an ethane-ethylene

splitter is otherwise too high. Dow Chemical has recently

been awarded a patent [25] for the dehydrogenation of

ethane over a metal-mordenite catalyst complex at rela-

tively low-conversions in which the product ethylene is

selectively recovered from the dilute ethylene-ethane

stream by alkylating it with benzene.

As mentioned earlier lower partial pressure of reactant

increases conversion achieved. Steam dehydrogenation is

based on the idea that by adding steam the partial pressure

of the paraffin can be lowered such that effectively the

conversion level could be as much as that obtained under

vacuum, but still operating at a superatmospheric level.

Two other potential benefits are that superheated steam can

be used as a heat carrier to supply heat for the dehydroge-

nation reaction and that steam interacts with coke deposits

to maintain the catalyst free of coke and active.

This approach is used by Phillips Petroleum in devel-

oping their STAR technology for isobutene dehydrogena-

tion employing Pt-type catalyst and multi-tubular reactor

design. This reactor design resembles a typical steam

reformer that is operated until the catalyst deactivates as a

result of coke deposition. This reactor is taken out of ser-

vice for catalyst regeneration while a spare muti-tubular

reactor is brought on-stream. The STAR technology is

currently owned and licensed by Krupp-Uhde.

Steam dilution is also employed in dehydrogenation of

ethylbenzene to styrene. Bricker and associates developed

steam stable catalysts for propane and isobutene dehydro-

genation in a high steam environment [26, 27].

3.1 Process Chemistry

The main reaction in catalytic dehydrogenation is the for-

mation of the mono-olefin from the corresponding feed

100

90

80

70

60

50

40

30

20

10

0627 727 827 927 1027 1127

Temperature, °C

C-C=CPropylene

C-C CMethyl

Acetylene

C=C=CAllene

C-C-CPropane

Mol

e %

Fig. 3 Propane dehydrogenation equilibrium at 0.23 atm abs

Pressure

1300 Top Catal (2012) 55:1297–1308

123

paraffin. Other reactions include consecutive and side

reactions. The reaction pathways involved in heavy par-

affin dehydrogenation (e.g., detergent-range C10–C14

n-paraffins) are more complicated than are those in light

paraffin dehydrogenation (e.g., propane and isobutane).

The main difference in reaction pathway is that a signifi-

cant amount of cyclic compounds can form via dehydro-

cyclization from heavy paraffins while this is not the case

for light paraffins. Figures 4 and 5, respectively, illustrate

possible reactions that take place on platinum (Pt) and acid

(A) sites in the dehydrogenation of light and heavy paraf-

fins when the catalyst is not selective, e.g., unmodified

platinum catalysts supported on alumina.

The consecutive reactions, the dehydrogenation of mono-

olefins to diolefins and triolefins, are catalyzed on the same

active site as the dehydrogenation of paraffins to mono-

olefins. Therefore, the consecutive reactions cannot be

eliminated but can be suppressed not only by catalyst mod-

ification but also by controlling the reaction kinetics. The

conversion of triolefins to aromatics is a very fast reaction

and thermodynamically favorable. Thus, the formation of

aromatics from triolefins must also be suppressed

kinetically.

3.2 Dehydrogenation Catalysts and Modifiers

The key role of dehydrogenation catalysts is to accelerate

the main reaction and to control other reactions. Unmodi-

fied alumina-supported platinum catalysts are highly active

but are not selective to dehydrogenation. Various by-

products, as indicated in Figs. 4 and 5, can also form. In

addition, the catalyst rapidly deactivates because of fouling

by heavy carbonaceous materials that forms coke on cat-

alyst and blocks active platinum sites. Therefore, the

properties of platinum and the alumina support need to be

modified to suppress the formation of by-products and to

increase catalytic stability.

The reaction of olefins on platinum is faster than that of

paraffins because olefins interact with platinum more

strongly than do paraffins. The role of platinum modifiers is

to weaken the platinum-olefin interaction selectively

without affecting the platinum–paraffin interaction. The

consecutive dehydrogenation rate of mono-olefins and di-

olefins is decreased by this modification without lowering

the rate of paraffin dehydrogenation significantly. The

modifier also improves the stability against coking by

heavy carbonaceous materials.

Alumina support has acidic sites that accelerate skeletal

isomerization, cracking, oligomerization, and polymeriza-

tion of olefinic materials, and enhances ‘‘coke’’ formation.

Therefore, acidity must be eliminated by a proper modifier

to control these undesirable reactions.

Since first commercial operation of the PacolTMprocess

in 1968, sevral catalytic advances have been made. Imai

[28, 29] developed next generation modified Pt catalyst

showing significantly lower coking and doubling the cat-

alyst stability. For longer carbon chain paraffins diffusion

also becomes a critical parameter and thus large pore low

density support is need to access all Pt sites. Jensen [30, 31]

further advanced the catalysis by introducing finely dis-

persed thin Pt layer over an inert core and thus prepared

next generation dehydrogenation catalyst. No catalyst

development is complete unless a sound process design is

wrapped around it that overcomes the techno-economic

barrier for a commercial success. Vora [32, 33] developed

a new radial flow reactor design.

The modified catalyst described above has high activity

and high selectivity to mono-olefins. The major by-prod-

ucts are diolefins that can be controlled kinetically. The

‘‘coke’’ formation is also suppressed, and therefore, sta-

bility is greatly improved. Over the modified catalyst, the

major reaction pathways for both light and heavy paraffin

dehydrogenation systems are simpler (Fig. 6).

n-Paraffinn-

Olefin n-DienePtPt

APtiso-Paraffin iso-

OlefinPolymers

Cracked ProductsA, Pt A AA

Coke

A, Pt

Pt = Platinum Site A = Acid Site

A, Pt

A, Pt

A

A

Fig. 4 Reactions by platinum and acid sites in light paraffin

dehydrogenation with unmodified catalyst dehydrogenation

n-Paraffin n-Olefin

n-Diene

PtPt

APtiso-

Paraffiniso-Olefin Polymers

Cracked Products

A, Pt A AA

Coke

A, Pt

Pt = Platinum Site A = Acid Site

A, Pt

A, Pt

A

A

n-Triene

Pt

A

Cyclo-Paraffins

A, Pt

Cyclo-OlefinsPt

Aromatics

A, Pt

A, Pt

A, Pt

Fig. 5 Reactions by platinum and acid sites in heavy paraffin

dehydrogenation with unmodified catalyst dehydrogenation

Top Catal (2012) 55:1297–1308 1301

123

3.3 Dehydrogenation Catalyst Supports

Platinum is a highly active catalytic element and is not

required in large quantities to catalyze the reaction when it is

dispersed on a high surface-area support. The high dispersion

is also necessary to achieve high selectivity to dehydrogena-

tion relative to undesirable side reactions, such as cracking.

Among many high surface area materials, alumina is the

support of choice. Alumina is relatively inert when it is

modified properly, as described previously, and when the

undesirable reactions are controlled. Alumina has excellent

thermal stability and mechanical strength under processing,

transport, and catalyst regeneration conditions. However,

the most important reason why alumina is chosen as sup-

port material is its superior capability of maintaining a high

degree of platinum dispersion, which is essential for

maintaining the dehydrogenation activity and selectivity.

The catalytic reaction rate is limited by the intraparticle

mass transfer rate. If the mass transfer rate is relatively

slow, both activity and selectivity are lowered. As a result,

the support must have a low pore diffusional resistance.

The surface area and the strength of the support increase as

the pore diameter decreases, and the pore diffusional

resistance decreases as the pore diameter increases. Surface

area retention with multiple regeneration cycles is also an

important feature in selecting an optimal support material.

Thus, an appropriate pore structure must be determined for

the support to achieve optimal catalytic performance.

The shape, strength, surface area, and surface smooth-

ness are important factors in applying CCR process tech-

nology, which consists of a moving-bed reactor and

catalyst regenerator. A proprietary catalyst manufacturing

technology developed by UOP meets the requirements for

preparing high-strength, low-attrition spherical supports.

3.4 Preliminary Catalyst Evaluation

In the catalyst screening stage, a short-time test is con-

ducted to determine activity and selectivity in a small

fixed-bed reactor containing 5–100 cc of catalyst.

In paraffin dehydrogenation employing a platinum cat-

alyst, hydrogen is used as a co-feed to increase catalytic

stability by suppressing undesirable consecutive reactions.

Thus, hydrogen also improves selectivity to the desired

mono-olefins.

Reaction conditions, such as temperature, pressure, and

hydrogen-to-paraffin feed ratio, are determined on the basis

of chemical equilibrium analysis and testing of a reference

catalyst. The most important data to be analyzed in the

preliminary evaluation stage are the change of conversion

as a function of time and the change of selectivity as a

function of conversion. For instance, Fig. 7 shows the

conversions of C10–C14 n-paraffins obtained with DeH-5TM

and DeH-7TM PacolTM dehydrogenation catalysts. DeH-7

catalyst is more stable than DeH-5.

The selectivity to the corresponding n-mono-olefins for

DeH-7 catalyst is equivalent to DeH-5 catalyst, as shown in

Fig. 8. The results indicate that DeH-7 has superior sta-

bility and activity to that of DeH-5 and maintains equiva-

lent selectivity to DeH-5 catalyst. Therefore, using DeH-7

n-Paraffin n-Olefin n-Diene

PtPt

A. Light Paraffin Dehydrogenation

n-Paraffin n-Olefin n-Diene

PtPt

A. Heavy Paraffin Dehydrogenation

n-TrienePtPt

Aromatics

Pt

Fig. 6 Paraffin dehydrogenation on modified Pt catalyst

5

-3

0

Hours On-Stream

Dif

fere

nce

of n

-Par

affi

n C

onve

rsio

nan

d T

arge

t C

onve

rsio

n, W

t-%

4

3

2

1

0

-1

-2

20 40 60 80 100 120

DEH-5DEH-7

140

-4

-5

Fig. 7 n-Paraffin conversion versus hours on-stream

5

-3

-5

Difference of Conversion and Target Conversion, %

Dif

fere

nce

of L

inea

r O

lefi

n Se

lect

ivit

yan

d T

arge

t Se

lect

ivit

y, W

t-%

4

3

2

1

0

-1

-2

-4 -3 -2 -1 0 1 2

-4

-53 4 5 6 7

DEH-5DEH-7

Fig. 8 Reactions by platinum and acid sites linear olefin selectivity

as a function of n-paraffin conversion

1302 Top Catal (2012) 55:1297–1308

123

achieves a longer life under the same processing conditions

or higher productivity with the same catalyst life under

more severe conditions.

Selectivity varies as a function of conversion for a given set

of operational variables. Therefore, understanding the rela-

tionship between selectivity and these variables is important.

Selectivity decreases as the conversion increases because n-

mono-olefins are consecutively converted into by-products.

The selectivity decreases sharply as conversion approaches

equilibrium because the main dehydrogenation process is

limited by equilibrium, but other reactions continue to occur.

Therefore, if side reactions are controlled, the selectivity is

improved as the equilibrium conversion becomes higher by

increasing the temperature and by decreasing the pressure and

the feed ratio of hydrogen to paraffin.

Figure 9 shows simulated selectivities to n-heptene and

n-heptadiene for the dehydrogenation of n-heptane. In this

simulation, the relative rate constants used are unity, which

represents that the catalyst possesses perfect selectivity

regarding consecutive dehydrogenation; the dehydrogena-

tion rate of paraffin is equal to that of mono-olefin and

diolefin. Experimental selectivities obtained over a UOP

dehydrogenation catalyst show that the catalyst has virtu-

ally perfect selectivity for consecutive dehydrogenation

steps, as seen in Fig. 9.

3.5 Catalyst Stability and Regeneration

Once an active and selective catalyst is developed, stability

tests are conducted. The dehydrogenation of detergent-

range paraffins is performed under relatively mild

temperature conditions. Thus, the catalyst can maintain a

long life even at high space velocity, and therefore, it is not

economical to regenerate the catalyst. Because of the

equilibrium limitation, the dehydrogenation of light par-

affins requires significantly higher temperatures to achieve

economically attractive conversion. The catalyst deactiva-

tion is accelerated under high-temperature conditions, and

therefore, frequent catalyst regeneration is necessary for

light paraffin dehydrogenation. For the dehydrogenation of

light paraffins, the UOP CCR technology is applied. In this

mode, ultimate catalyst life of several years is achieved.

3.6 Heat of Reaction

The dehydrogenation of n-paraffins is an endothermic reac-

tion with a heat of reaction of about 30 kcal/mol. Therefore, a

significant temperature drop occurs in a commercial adiabatic

reactor, which lowers the equilibrium conversion level. A

multistage reactor system with interstage reheating is needed

for light paraffin dehydrogenation. Figure 10 illustrates con-

version, equilibrium conversion, and temperature along the

catalyst bed in a three-stage reactor system for the dehydro-

genation of isobutane. For propane dehydrogenation, a four-

stage reactor system becomes more economical because

higher average temperatures are needed.

Excessively high inlet temperatures can be avoided by

employing the multistage reactor system as opposed to a

single stage reactor system (Fig. 11). Thus, thermal

cracking and catalyst deactivation, which are accelerated at

higher temperatures, can be controlled to low levels.

3.7 Commercial Processes

The UOP Pacol TM process for selective dehydrogenation

of C10–C13 range linear paraffins to produce the corre-

sponding linear mono-olefins is shown in Fig. 12 in com-

bination with the UOP Detergent Alkylation process. The

Pacol process consists of a radial-flow reactor and a

product recovery section. Worldwide more than 3 million

MTA LAB is produced employing this process [34].

As described earlier, the Houdry Catadiene process and

the Catofin process make use of parallel reactor and operate

in a cyclic mode. In this type of dehydrogenation process

using chromia/alumina catalysts, the catalyst is in a fixed

shallow bed located inside a reactor that may be spherical

or often a horizontal cylinder. Significant quantity of inert

materials, such as, various size of alumina or silica balls are

used below and above the catalyst bed to support the cat-

alyst and also to fill up the reactor. A significant amount of

coke is deposited on the catalyst during the dehydrogena-

tion step, such that a number of reactors are used in parallel

with some being used for dehydrogenation while the others

100

50 2

n-C7 Conversion, %

n-C

7=Se

lect

ivit

y, m

ol-% 90

80

70

60

5 10 15

ExperimentalData

20 25 30

100

50

n-C

7==Se

lect

ivit

y, m

ol-%90

80

70

60

Reaction: n-C7 n-C7= n-C7

== TolueneK1 K2 K3

Kinetic Parameters: K2/K1 = K3/K1 = 1, K1 = 0.51, K2 = 0.27

n-Heptene

n-HeptadieneExperimental

Data

Fig. 9 Simulation of selectivity for dehydrogenation of n-heptane

Top Catal (2012) 55:1297–1308 1303

123

are in various stages of purges or catalyst regeneration

(coke burning) steps. The dehydrogenation reactions are

strongly endothermic, and in this process the heat is pro-

vided by the sensible heat stored in the catalyst and the

inert materials during regeneration. The length of the total

reactor cycle is limited by the amount of heat available, and

can be as short as 10–30 min. Figure 13 illustrates a

schematic of such a process.

Onstream timeþ purge timeþ regeneration time

þ purge time ¼ total cycle time

Another cyclic process is Phillips STAR technology

[35]. It uses a fixed-bed, fired-tube reactor operating at

superatmospheric pressure with steam as a diluent to lower

the partial pressure of the reactants in order to achieve

reasonable conversion level. In many respects it is similar

in design to a steam reforming furnace with the heat of

reaction provided by firing outside tubes and thus operating

at near isothermal condition. As of 2010 there are two

small operating plants, 118,000 and 13,000 MTA, for the

production of isobutylene from isobutene. As mentioned

earlier, this technology is currently owned and licensed by

Krupp-Uhde.

The UOP OleflexTM is more widely used process for

propane and isobutene dehydrogenation. Figure 14 illus-

trates the flow diagram of the UOP OleflexTM process. The

process consists of a reactor section and a product recovery

section. The reactor section consists of three or four stages

of radial-flow reactors, charge and interstage heaters,

reactor feed-effluent exchangers, and the CCR continuous

catalyst regeneration unit (Fig. 15). Today more than 3

million MTA of propylene and 2 million MTA isobutene

are produced via this route. In 2011 there are 11 operating

UOP OleflexTM units for propane to propylene and six

more for isobutane to isobutene, with five more in con-

struction and design [36].

4 Ethylbenzene Dehydrogenation

As mentioned earlier, the ethylbenzene dehydrogenation

reaction proceeds over an iron or an iron-chromium catalyst

that usually also contains potassium in the form of potassium

carbonate. The reaction takes place at 630 �C temperature

and pressure is usually subatmospheric for improved per-

pass conversion. Steam dilution is practiced to further lower

the partial pressure of the reactants. Because the reaction is

strongly endothermic, various reaction stages with interheat

and interstage addition of superheated steam are normally

employed. Figure 16 illustrates a typical process scheme for

the dehydrogenation of ethylbenzene to styrene.

In an interesting variant of the conventional process,

part of the hydrogen that is produced in the first stage of

conversion is selectively reacted with oxygen over a sep-

arate bed of oxidation catalyst such that significant

amounts of heat are released internally within the reactor

system. The hydrogen oxidation catalyst is selected such

that there is practically no conversion or degradation of

either ethylbenzene or styrene to CO2 [37]. While this

process could be thought of as an oxydehydrogenation

process, in reality it is just a conventional dehydrogenation

coupled with an oxidative reheat step. The alternative

ethylbenzene dehydrogenation process, known commer-

cially as SMARTTM (Styrene Monomer by Advanced

Reheat Technology), was originally called the Styro-Plus

process and initially demonstrated at Mitsubishi Chemi-

cals, Kashima, Japan The SMART process is now licensed

jointly by UOP LLC and ABB Lummus Global Inc. In

Temperature

Conversion

EquilibriumConversion

Isobutane/H2

Reheat Reheat

Rx1 Rx2 Rx3

Fig. 10 Temperature profile and conversions of three-stage isobutane

dehydrogenation process

50

10

-150

Difference of Temperatureand Target Inlet Temperature, °C

Con

vers

ions

, %

45

40

35

30

25

20

15

-125 -100 -75 -50 25 0 25

5

050 75 100 150 175125

Single-Stage ProcessThree-Stage ProcessInterstage Reheating

Fig. 11 Isobutane dehydrogenation

1304 Top Catal (2012) 55:1297–1308

123

addition to supplying the heat of reaction internally, the

SMART process benefits from the equilibrium displace-

ment that results from the selective removal of one of the

reaction products, hydrogen [38].

5 Other Dehydrogenation Technologies

The processes discussed above are all for the direct cata-

lytic dehydrogenation of paraffins to the corresponding

olefins. Other approaches have also been considered in the

past although none has reached to the level of commer-

cialization. Some of the most notable are:

• Halogenated dehydrogenation

• Oxydehydrogenation

Use of halogens for the dehydrogenation of paraffins has

been proposed in different ways. For example, as noted

earlier, heavy paraffins were first chlorinated and then

dehydrochlorinated to heavy olefins commercially in the

past both by Shell (CDC process) and by Huls, among

others. Pyrolysis of methane in the presence of chlorine has

been proposed by Prof. Benson [39] for the production of

acetylene and ethylene. Other chlorination/dehydro-

chlorination cycles have been proposed for the production

of ethylene from ethane. Propane dehydrogenation in the

presence of iodine via a propyl iodide intermediate has also

Reactor on Purge

Reactor on Stream

Reactor on Reheat

ChargeHeater Air

Heater

ProductCompressor

FlashDrum

Cooler

Cooler

Steam

Air

Exhaust Air

H2 (Optional)

Fuel Gas

C3-C5 Cut

PSA

Cold Box

Drier

C3-C5

Paraffin

Steam

GasifierFuel

Fig. 13 Catofin process flow

diagram

HydrogenRecycle Gas

Hydrogen-RichOff Gas

LinearParaffinCharge

AluminaTreater

Paraffin Recycle

Separator

Stripper

Light EndGas

Light EndLiquid

Benzene

LinearDetergentAlkylene

UOPDetergent Alkylate

Process

ChargeHeater

Reactor

Fig. 12 UOP pacol

dehydrogenation process

Top Catal (2012) 55:1297–1308 1305

123

been proposed. Apart from the apparent corrosion prob-

lems associated with the use of chlorine, other difficulties

readily come to mind owing to the relatively high cost of

chlorine, and even more so of iodine, and the need to either

dispose of or recycle vast quantities of hydrogen chloride

generated as a byproduct.

Oxydehydrogenation or oxidative dehydrogenation can

be considered in at least two different ways:

– Use of oxygen as a way to selectively oxidize the

hydrogen coproduct from dehydrogenation, and thus to

displace the dehydrogenation equilibrium to higher

conversions. As mentioned earlier, this approach has

also been used commercially in the catalytic dehydro-

genation of ethylbenzene to styrene as in the UOP

Styro-PlusTM process or in the ABB Lummus/UOP

SMARTTM process. Vora [40] has shown a similar flow

scheme where oxygen is used in between the

multi-stage reactors to combust hydrogen and thus

improve equilibrium for the propane or butane

dehydrogenation.

– Direct use of oxygen as a means of dehydrogenating, say,

ethane to ethylene. Some interesting work in this area is

currently being done by Prof. Schmidt [41] and his

colleagues at the University of Minnesota in the United

States, and also by Prof. Eliseo Ranzi and coworkers at

the Polytechnic University of Milan, Italy. Ethylene

production at greater than 85 % selectivity and 70 %

ethane conversion has been claimed at a small experi-

mental scale over a platinum–tin catalyst. If these results

can be extrapolated to a large scale commercial operation

it could become an attractive alternative to conventional

steam cracking of ethane since oxydehydrogenation

units in principle could be much smaller than the very

large capacities required in steam cracking furnaces.

Significant concerns exist however over the operating

conditions reported by these research teams: (1) The

relatively low ratio molar ratio of ethane:oxygen:hydro-

gen s of 2:1:2 over a catalyst at about 950 �C is

potentially explosive, (2) the hydrogen feed require-

ments is surprisingly high (although the research team

claims that they can be balanced with the net production

of hydrogen in the reaction), and (3) the operating

conditions are exceedingly severe for most catalyst

compositions all of which contribute to the unlikely

commercial feasibility of such a process.

As indicated earlier, oxydehydrogenation found suc-

cessful commercial application in the conversion of n-

butenes to butadiene, but not yet for the production of

ethylene or propylene.

DustCollector

Lift GasBlower

DisengagingHopper

RegenerationTower

CatalystLift Lines

LockHopper

FlowControlHopper

SurgeHopper

LiftEngager

NitrogenLift Gas

H2Lift Gas

LiftEngager

LockHopper

R

R

Fig. 15 Oleflex regeneration section

Heater Cells

NetSeparatorOff Gas

To SHP

Dryer

TurboExpander

Fresh& Recycle

Feed

H2 Recycle

CCR

Rx EffluentCompressor

Regeneration SectionReactor Section

Product Recovery Section

Fig. 14 UOP oleflex process

1306 Top Catal (2012) 55:1297–1308

123

6 Reactor Design Options

The choice of reactor design plays a very important role in

the success of catalytic processes. The following types of

reactor design are commercially proven for endothermic

catalytic dehydrogenation processes:

• Downflow adiabatic fixed bed

• Radial flow fixed bed or moving bed adiabatic

• Tubular isothermal

• Fluidized bed

The conventional choice in reactor design is the down-

flow packed bed. Reactants, either vapor or vapor plus

liquid, enter at the top of the catalyst bed and flow down

through the catalyst. The main characteristics of the

downflow reactor are:

• Plug flow

• Tendency for imperfect distribution of process flow

• Relatively high pressure drop

• Adiabatic: no temperature control

• Inefficient use of the catalyst

Pressure drop is the main disadvantage of the downflow

reactor in many process applications, particularly in processes

that require a low operating pressure and a large catalyst

inventory. This pressure-drop concern was addressed by the

introduction of the radial-flow reactor. Reactants in a radial-

flow reactor normally enter the vessel from the top and then

flow radially either inward or outward through an annular

catalyst bed. The advantage of this design is that the flow path

is short and the cross-sectional area is large, allowing a rea-

sonable ratio of vessel length to diameter. Low pressure drops

can be achieved because of the flow path through the packed

bed can be kept short. Also, catalyst can easily be added and

removed while the reactor is in operation. Characteristics of

the radial flow reactor are:

• Plug flow

• Low pressure drop

• Easy catalyst replacement

• Adiabatic: no temperature control

• Limited to single-phase operation

The tubular reactor was developed to allow for heat

transfer in the reactor bed. This reactor is essentially a shell

and tube heat exchanger that has catalyst in the tubes and a

heat transfer medium in the shell. Processes with high heats

of reaction require long tube lengths to increase the nec-

essary surface area. Tubular reactors loaded with conven-

tional-shaped catalysts, spherical or extruded, have high

pressure drops. Consequently, specialized catalyst shapes

were developed to increase voidage and reduce pressure

drop. However, reduced voidage increased the catalyst

volume requirements. Large catalyst volumes can require

many tubes (on the order of tens of thousands). Loading

tubular reactors is a time-consuming and labor-intensive

operation. Tubular reactors have these attributes:

• Plug flow

• Intrinsically high pressure drop that requires special

catalyst shapes to minimize pressure drop.

• Fabrication constraints of large tube sheets

• Difficult to load and unload

• Limited control of the reactor temperature profile

The fluidized-bed reactor can mitigate the problem of

pressure drop to an extent, but the catalyst recovery

equipment can cause higher pressure drop. This reactor is

approximately isothermal as a result of the high degree of

mixing. Heat can be added or removed by heat exchange

coils in the fluidized bed. Its main disadvantage is that the

reaction is not plug flow. In addition, catalyst losses can be

prohibitive if the catalyst is costly or is environmentally

unsafe. A fluidized-bed reactor has these characteristics:

SMColumn

EB/SMSplitter

Separator

Benzene/TolueneSplitter

EB RecoveryColumn

Condenser

WasteHeat

Exchanger

SteamSuperheater

Steam

Steam

Fresh EB Recycle EB

Fuel Gas

Water

Tar

SM Product

TolueneOff Gas

Recovery

SM ReactorSection

BenzeneFig. 16 Typical ethylbenzene

dehydrogenation unit for the

production of styrene monomer

(SM)

Top Catal (2012) 55:1297–1308 1307

123

• Not plug flow and highly backmixed

• Pseudoisothermal

• Possibly low pressure drop

• Care required to retain catalyst in the reactor

• Suitable for large capacities and high heats of reaction.

The following table summarizes the main characteristics

of the five reactor systems discussed.

Downflow Radial

flow

Tubular Fluidized

bed

Low pressure drop • •Plug flow • • •Catalyst addition or

removal

• •

High heat transfer, near

isothermal

• •

7 Conclusions

Catalytic dehydrogenation of paraffins and of ethylbenzene

is a commercial reality in widespread use for numerous

applications. These include the production of light olefins,

heavy olefins, and alkenylaromatics. Oxydehydrogenation,

on the other hand, is still in the developmental stage but if

successful holds great promise due to the potential energy

savings of this process. For the production of heavy olefins,

selective paraffin dehydrogenation over noble metal cata-

lysts has proven to be the preferred and dominant route.

When only one or two light olefins are desired, in par-

ticular propylene or isobutylene or perhaps a mixture of

propylene and isobutylene, catalytic dehydrogenation over

noble metal catalysts has acquired a significant and grow-

ing market share.

Finally, the choice of reactor plays an important role in

securing the success of a catalytic process. Pressure drop,

heat transfer, and the ability to move or to regenerate the

catalyst all must be taken into account in the process

development and design stages.

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