12 PDH_Process Vora UOP.pdf
Transcript of 12 PDH_Process Vora UOP.pdf
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: [email protected]
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
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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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