U.S. patent application number 10/586024 was filed with the patent office on 2008-08-14 for process for the preparation of hydrogenated hydrocarbon compounds.
Invention is credited to William M. Castor, Susan B. Domke, Simon J. Hamper, Matthew T. Pretz.
Application Number | 20080194891 10/586024 |
Document ID | / |
Family ID | 34860358 |
Filed Date | 2008-08-14 |
United States Patent
Application |
20080194891 |
Kind Code |
A1 |
Pretz; Matthew T. ; et
al. |
August 14, 2008 |
Process for the Preparation of Hydrogenated Hydrocarbon
Compounds
Abstract
A process for the dehydrogenation of a paraffinic hydrocarbon
compound, such as an alkane or alkylaromatic hydrocarbon compound
to produce an unsaturated hydrocarbon compound, such as an olefin
or vinyl aromatic compound or mixture thereof, in which a
dehydrogenation catalyst contacts gaseous reactant hydrocarbons in
a reactor at dehydrogenation conditions.
Inventors: |
Pretz; Matthew T.; (Lake
Jackson, TX) ; Domke; Susan B.; (Rosharon, TX)
; Castor; William M.; (Lake Jackson, TX) ; Hamper;
Simon J.; (Lake Jackson, TX) |
Correspondence
Address: |
The Dow Chemical Company
Intellectual Property Section, P.O. Box 1967
Midland
MI
48641-1967
US
|
Family ID: |
34860358 |
Appl. No.: |
10/586024 |
Filed: |
February 4, 2005 |
PCT Filed: |
February 4, 2005 |
PCT NO: |
PCT/US2005/003772 |
371 Date: |
July 14, 2006 |
Related U.S. Patent Documents
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Application
Number |
Filing Date |
Patent Number |
|
|
60543006 |
Feb 9, 2004 |
|
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Current U.S.
Class: |
585/252 |
Current CPC
Class: |
C07C 5/3332 20130101;
C07C 5/3332 20130101; C07C 5/3332 20130101; C07C 5/3332 20130101;
C07C 5/3332 20130101; Y02P 20/584 20151101; C07C 5/3332 20130101;
C07C 5/3337 20130101; Y02P 20/52 20151101; C07C 11/08 20130101;
C07C 5/3332 20130101; C07C 11/09 20130101; C07C 15/44 20130101;
C07C 11/04 20130101; C07C 15/46 20130101; C07C 11/06 20130101 |
Class at
Publication: |
585/252 |
International
Class: |
C07C 5/02 20060101
C07C005/02 |
Claims
1. A process for dehydrogenating a hydrocarbon selected from at
least one of i) paraffinic hydrocarbons, and ii) alkylaromatic
hydrocarbons; comprising contacting a gaseous stream containing at
least one of the hydrocarbons with a dehydrogenation catalyst at
reaction temperature and in concurrent flow through a
dehydrogenation reactor wherein the average contact time between
the hydrocarbon and catalyst within the dehydrogenation reactor is
from about 0.5 to about 10 seconds.
2. A process of claim 1 wherein the catalyst has an average
residence time within the dehydrogenation reactor from about 0.5 to
about 40 seconds.
3. A process of claim 2 wherein the hydrocarbon and catalyst are
transferred from the dehydrogenation reactor to a separation device
wherein the average contact time between the hydrocarbon and
catalyst while at reaction temperature in the separation device is
less than about 10 seconds.
4. A process of claim 3 wherein the total average contact time
between the hydrocarbon, catalyst and resulting hydrocarbons while
at reaction temperature is less than about 20 seconds.
5. A process of claim 4 wherein the total average contact time is
less than about 10 seconds.
6. A process of claim 1 wherein the paraffinic hydrocarbons are
selected from ethane, propane, isopropane, and butane; and the
alkylaromatic hydrocarbons are selected from ethylbenzene,
propylbenzene and methylethylbenzene.
7. A process of claim 1 wherein the dehydrogenation reactor is a
riser reactor.
8. A process of claim 1 wherein the hydrocarbon is introduced to
the dehydrogenation reactor at multiple points of entry.
9. A process of claim 8 wherein both a paraffinic hydrocarbon and
alkylaromatic hydrocarbon are introduced into the same
dehydrogenation reactor with the paraffinic hydrocarbon being
introduced at a relatively lower point of entry than the
alkylaromatic hydrocarbon.
10. A process of claim 1 wherein the temperature within the
dehydrogenation reactor is from about 500 to about 800.degree. C.,
and the pressure is from about 3.7 to about 64.7 psia.
11. A process of claim 10 wherein the pressure is from about 3.7 to
about 14.7 psia.
12. A process of claim 10 wherein the temperature is from about 570
to about 750.degree. C.
13. A process of claim 3 wherein catalyst from the separation
device is transferred to one of: a catalyst regenerator wherein the
catalyst is regenerated and returned to the dehydrogenation
reactor, and a recycle loop wherein catalyst is recycled from the
separation device back to the dehydrogenation reactor.
14. A process of claim 13 wherein the catalyst from the recycle
loop and regenerator are combined and introduced into the
dehydrogenation reactor.
15. The process of claim 1 wherein the dehydrogenation catalyst is
comprises gallium carried by an alumina or alumina silica
support.
16. The process of claim 15 wherein the catalyst comprises a
comprises an alkali or alkaline earth metal selected from at least
one of: sodium, lithium, potassium, rubidium, magnesium, calcium,
strontium and barium, and further comprises promoter selected from
at least one of: manganese and platinum.
17. A process for dehydrogenating a hydrocarbon selected from at
least one of: i) paraffinic hydrocarbons selected from ethane,
propane, and butane; and ii) alkylaromatic hydrocarbons are
selected from ethylbenzene, propylbenzene and methylethylbenzene;
comprising contacting a gaseous stream containing at least one of
the hydrocarbons with a dehydrogenation catalyst comprising gallium
carried by an alumina or alumina silica support, at reaction
temperature and in concurrent flow through a dehydrogenation
reactor wherein the average contact time between the hydrocarbon
and catalyst within the dehydrogenation reactor is from about 1 to
about 4 seconds; the catalyst has a average residence time within
the dehydrogenation reactor from about 1 to about 10 seconds; and
the temperature and pressure in the dehydrogenation reactor is from
about 570 to about 750.degree. C., and from about 6.0 to about 44.7
psia; and transferring the hydrocarbon and catalyst from the
dehydrogenation reactor to a separation device wherein the average
contact time between the hydrocarbon and catalyst while at reaction
temperature in the separation device is less than about 3 seconds
and the total average contact time between the hydrocarbon,
catalyst and resulting hydrocarbons while at reaction temperature
is less than about 7 seconds.
18. An integrated process for making a vinyl aromatic compound
comprising: dehydrogenating a paraffinic hydrocarbon and an
alkylaromatic hydrocarbon with a dehydrogenation catalyst at
reaction temperature and in concurrent flow through a
dehydrogenation reactor wherein the average contact time between
the hydrocarbon and catalyst within the dehydrogenation reactor is
from about 0.5 to about 10 seconds; transferring the hydrocarbon
and catalyst from the dehydrogenation reactor to a separation
device and separating the catalyst from the resulting hydrocarbon
wherein the average contact time between the hydrocarbon and
catalyst while at reaction temperature in the separation device is
less than about 10 seconds; recovering vinyl aromatic hydrocarbons
resulting from the dehydrogenation; combining olefins resulting
from dehydrogenation of paraffinic hydrocarbons with aromatic
hydrocarbons in an alkylation reactor under conditions to produce
alkylaromatic hydrocarbons; transferring the alkylaromatic
hydrocarbons from the alkylation reactor to the dehydrogenation
reactor to produce vinyl aromatic compounds.
19. A process according to claim 18 wherein the dehydrogenation of
the paraffinic hydrocarbon and alkylaromatic hydrocarbon is
conducted concurrently in the same dehydrogenation reactor.
20. A process according to claim 19 wherein the dehydrogenation of
the paraffinic hydrocarbon and alkylaromatic hydrocarbon are
conducted in separate dehydrogenation reactors.
Description
BACKGROUND OF THE INVENTION
[0001] 1) Field of the Invention
[0002] This invention generally relates to the field of hydrocarbon
conversion and particularly to the dehydrogenation of paraffinic
hydrocarbons to olefinic hydrocarbons, and/or lower alkylaromatic
hydrocarbons to vinyl aromatic hydrocarbons. In several preferred
embodiments, the invention relates to the dehydrogenation of lower
alkanes, for example ethane, isopropane, propane and butanes to
their corresponding olefins, for example ethylene, propylene and
butylenes; and/or to the dehydrogenation of lower alkylaromatic
hydrocarbon compounds, for example ethylbenzene, propylbenzene and
methylethylbenzene to their corresponding vinyl aromatic (that is
"alkenylaromatic") hydrocarbon compounds, for example styrene,
cumene and alpha-methyl styrene, respectively. The invention
further includes an integrated process for making olefinic and
vinyl aromatic hydrocarbons including alkylation and
dehydrogenation steps.
[0003] 2) Description of Related Art
[0004] U.S. Pat. No. 6,031,143 and its corresponding EP 0 905 112
describe an integrated process for producing styrene by feeding
benzene and recycled ethylene to an alkylation reactor to produce
ethylbenzene, mixing the alkylation effluent with ethane and
feeding the mixture to a dehydrogenation reactor containing a
catalyst capable of contemporaneously dehydrogenating ethane and
ethylbenzene. The resulting product is separated to produce a
stream of styrene and ethylene, with ethylene being recycled to the
alkylation reactor. The dehydrogenation reactor is preferably a
fluidized bed reactor connected to a fluidized bed regenerator from
which the catalyst is circulated between the regenerator and the
dehydrogenation reactor in countercurrent flow. That is, catalyst
is introduced to the dehydrogenation reactor from the top and
slowly descends to the bottom in countercurrent with the gas phase
reactants which are rising through the reactor. During this
descent, the catalyst is deactivated. The deactivated catalyst is
removed from the bottom of the dehydrogenation reactor and
transported to the top of the regenerator where it descends to the
bottom in countercurrent flow with hot air which is rising. During
this descent, the carbonaceous residue present on the catalyst is
burnt and the regenerated catalyst is collected at the bottom of
the regenerator where it is subsequently circulated back to the top
of the dehydrogenation reactor.
[0005] WO 02/096844 describes an improvement to this process where
the dehydrogenation catalyst is transported from the regenerator to
the dehydrogenation reactor by way of a lower alkyl hydrocarbon
carrier, for example ethane. During transport, a portion of the
carrier is dehydrogenated, (for example ethane converted to
ethylene), and the catalyst is cooled.
[0006] EP 1 255 719 (and corresponding co-pending US patent
publication no. US 2003/0028059, both filed by the assignee of the
present application) describes a similar integrated process of
preparing styrene using benzene and ethane as raw materials.
However, the process includes additional separation and recycling
steps that are designed to improve efficiency. For example, the
dehydrogenated effluent exiting the dehydrogenation reactor is
separated into its aromatic and non-aromatic constituents. The
non-aromatic constituents, namely ethane, ethylene and hydrogen are
recycled to an alkylation reactor were they are combined with
benzene. The aromatic constituents are further separated, for
example styrene is recovered and ethylbenzene is recycled to the
dehydrogenation reactor. The alkylation effluent is separated into
its constituents with hydrogen being removed, and ethane and
ethylbenzene being directed to the dehydrogenation reactor. The
dehydrogenation reactor may have a variety of conventional designs
including fixed, fluidized, and transport bed.
[0007] The described dehydrogenation processes are effective at
integrating the production of styrene and ethylene using ethane and
benzene as the starting materials. Thus, these processes
effectively de-coupled the production of styrene from the presence
or proximity of a light hydrocarbon steam cracker as a source for
ethylene. However, the dehydrogenation processes described employ
relatively long contact times between the hydrocarbons and catalyst
while at reaction temperature, resulting in thermal cracking,
undesired side reactions and the formation of tars and other heavy
products.
[0008] WO 02/096844 introduces the concept of a split "riser-type"
dehydrogenation reactor operating in concurrent or "equicurrent"
mode wherein catalyst is carried upwards pneumatically through the
dehydrogenation reactor by the gas phase reactants. The space
velocity (GHSV) for such a reactor is greater than 500.sup.h-1. The
catalyst is introduced into the reactor with an alkyl hydrocarbon
such as ethane whereas the alkylaromatic compound, for example
ethylbenzene, is introduced at a suitable height along the riser
after much of the alkyl hydrocarbon has be dehydrogenated and the
temperature of the catalyst has been reduced. While no specific
examples or operating conditions are provided, the use of such a
riser reactor presumably leads to reduced contact times between
reactants and catalyst while in the reactor.
[0009] Dehydrogenation temperatures and residence times are
typically optimized to balance the reaction kinetics of both
catalytic and gas-phase (thermal) reactions. The catalytic reaction
produces a high selectivity to the desired products while the gas
phase reaction produces many undesired products and impurities.
That is, while the catalytic reaction kinetics to the desired
products increases exponentially with temperature so does the gas
phase reaction kinetics; therefore, the proper residence time and
reaction temperature profile must be selected to drive both the
catalytic reaction to the desired conversion while not allowing the
non-selective gas phase reactions to overwhelm the total product
selectivity. It would be useful to provide an apparatus and process
which minimizes the time period in which reactants and catalyst are
in contact with one another while at reaction temperature. This is
particularly the case when utilizing highly reactive catalyst which
can quickly deactivate.
[0010] While not directed toward a "dehydrogenation process" as
described in the aforementioned references, WO 03/050065 describes
an integrated process for making styrene where benzene and
"recycled" ethylene are combined in an alkylation unit with the
resulting product stream of ethylbenzene being combined with
ethane. Unlike the previously described references, this process
utilizes an oxidative dehydrogenation (oxodehydrogenation)
reaction. That is, the product stream from the alkylation unit is
combined with ethane and oxygen and then contemporaneously
oxidatively dehydrogenated to provided ethylene and styrene. The
resulting ethylene is recycled to the alkylation unit. The
oxodehydrogenation reactor is described as a fluid-bed reactor
operating at a temperature range of from 300 to 550.degree. C., a
pressure range from 1 to 30 bar, a gas hourly space velocity of
2000 to 6000.sup.h-1, with a residence time of the catalyst in the
fluid-bed zone of from 1 to 60 seconds.
BRIEF SUMMARY OF THE INVENTION
[0011] The above described deficiencies of prior art can be
overcome by the subject invention which comprises contacting a
gaseous stream of hydrocarbon with a dehydrogenation catalyst at
reaction temperature for relatively short "contact times." In a
preferred embodiment, lower alkanes, for example ethane, propane
and butanes are dehydrogenated to their corresponding olefins, for
example ethylene, propylene and butylenes; and/or lower
alkylaromatic hydrocarbon compounds, for example ethylbenzene,
propylbenzene and methylethylbenzene are dehydrogenated to their
corresponding vinyl aromatic hydrocarbon compounds, for example
styrene, cumene and alpha-methyl styrene, respectively.
[0012] In another embodiment, the aforementioned dehydrogenation
process is combined with an alkylation step, as part of an
integrated process. Many additional embodiments are also
described.
BRIEF DESCRIPTION OF THE DRAWINGS
[0013] FIG. 1 shows a schematic block flow diagram of an embodiment
of the present invention in which a riser reactor is employed in a
single hydrocarbon feed point which may be used for: 1) paraffinic
hydrocarbon (for example ethane) feed only, 2) alkylaromatic
hydrocarbon (for example ethylbenzene) feed only, or 3) mixed feed
(for example ethane and ethylbenzene), including catalyst
regeneration.
[0014] FIG. 2 shows a schematic block flow diagram of another
embodiment of the present invention in which a riser reactor is
employed with a multiple feed point configuration, that is a split
ethylbenzene and ethane feed configuration, including catalyst
regeneration.
[0015] FIG. 3 shows a schematic block flow diagram of another
embodiment of the present invention including multiple riser
reactors with a catalyst regeneration in a series
configuration.
[0016] FIG. 4 shows a schematic block flow diagram of another
embodiment of the present invention including multiple riser
reactors with catalyst regeneration in a parallel
configuration.
[0017] FIG. 5 shows a schematic block follow diagram of another
embodiment of the present invention similar to FIG. 4, but further
including a catalyst recycle configuration.
DETAILED DESCRIPTION OF THE INVENTION
[0018] The present invention is directed toward the dehydrogenation
of at least one and preferably both of: 1) a paraffinic hydrocarbon
compounds, preferably a lower alkane having from 2 to 6 carbon
atoms but more preferably less than 5 carbon atoms, for example
ethane, propane, isopropane and butanes, to the corresponding
olefin, namely, ethylene, propylene, and butylenes, respectively,
and 2) an alkylaromatic hydrocarbon compound, preferably a lower
alkylaromatic hydrocarbon compound, such as for example,
ethylbenzene, propylbenzene, isopropyl benzene, and methyl
ethylbenzene, to the corresponding vinyl aromatic hydrocarbon
compound, (that is "alkenylaromatic"), namely, styrene, cumene or
alpha-methyl styrene. Several embodiments of the present invention
are described including both the simultaneous and separate
dehydrogenation of lower alkanes and alkylaromatics. The invention
is useful to prepare styrene and ethylene from ethylbenzene and
ethane, respectively. Likewise, cumene and propylene can be
prepared from propylbenzene and propane, respectively.
[0019] The dehydrogenation reaction in the present invention is
conducted under a relatively short contact times in order to
prevent undesirable side reactions and product degradation. The
term "average contact time" or "contact time" as used herein is
intended to refer to the time in which the molar average of gaseous
hydrocarbon molecules are in contact with catalyst while at
reaction temperature, regardless of whether the reactants are
converted to desired products. The term "reaction temperature" is
intended to mean a temperature at which a significant amount of
chemical reaction occurs, regardless of whether such reactions are
the desired dehydrogenation of reactants to their corresponding
olefin and vinyl aromatic products. Said another way, the reaction
temperature is the temperature at which the hydrocarbons are no
longer stable. The term "significant amount" in intended to mean a
detectable amount having in an economic impact on the process. In
most embodiments of the invention, the reaction temperature is
greater than about 500 and preferably 550.degree. C. The average
contact time needs to be sufficiently long to dehydrogenate
acceptable amounts of hydrocarbon reactants but not so long as to
result in unacceptable amounts of by-products. While the required
contact time is related to the specific reactants, catalysts and
reaction temperatures, in preferred embodiments of the invention
the contact time within the dehydrogenation reactor is less than 60
seconds, preferably from about 0.5 to about 10 seconds, more
preferably from about 1 to about 8 seconds, and still more
preferably from about 1 to about 4 seconds.
[0020] Due to the active nature of the preferred catalyst, the
average residence time of the catalyst within the dehydrogenation
reactor is preferably less than about 60 seconds, preferably from
about 0.5 to about 40 seconds, more preferably about 1.0 to about
12.0 seconds, and still more preferably from about 1.0 to about 10
seconds.
[0021] At such short catalyst residence times and average contact
times in the dehydrogenation reactor, the temperature of the
reaction mixture, which may be supplied in major part by the hot
fresh or regenerated catalyst, is preferably from about 500 to
about 800.degree. C. With respect to lower alkanes, the reaction
mixture is preferably from about 600 to about 750.degree. C., and
with respect to alkylaromatics from about 550 to 700.degree. C. but
more preferably from about 570 to about 660.degree. C. In general,
the highest temperature in the reactor will be found at its lower
end and as reaction proceeds and the catalyst and reaction mixture
ascends, the temperature will decrease toward the upper end of the
reactor.
[0022] The applicable operating pressure of the dehydrogenation
reactor is quite broad, that is from about 3.7 to about 64.7 psia.
The pressure at which the reaction proceeds is typically from about
14.7 to about 64.7 psia, and preferably from about 14.7 to about
44.7 psia. However, in several preferred embodiments of the
invention, the operating pressure of the dehydrogenation reactor
may be below atmospheric, that is from about 3.7 to 14.7 psia, more
preferably about 6.0 to about 14.7 psia.
[0023] The gas hourly space velocity (GHSV) for the present process
has been found to range from about 1,000 to about 150,000 normal
cubic meters/hr of hydrocarbon feed per cubic meter of catalyst at
bulk density. The superficial gas velocity of about 5 to about 80
ft/sec, preferably about 15 to about 70 ft/sec. The catalyst flux
is preferably about 10 to about 120 lbs/ft.sup.2-sec with a
catalyst to feed ratio of about 5 to about 100 on a weight to
weight basis. The catalyst is preferably pneumatically moved
through the reaction system by a carrier fluid, which is preferably
either an inert diluent fluid or one of the reactants in gaseous
form. Alternatively, the catalyst may be transported through the
reactor under sub atmospheric pressure without diluent. Examples of
inert diluent carrier gases are nitrogen, volatile hydrocarbons for
example methane, and other carriers which do not interfere with the
reaction, steam, carbon dioxide, argon and the like. The paraffinic
hydrocarbon compounds useful as reactants in the process of the
present invention are also preferred carrier fluids and, most
preferred are ethane, propane, and butane. Steam is preferably not
used in the present invention. The amount of carrier gas required
is only that amount necessary to maintain the catalyst particles in
fluidized state and transport the catalyst from the regenerator to
the reactor. Preferably, the amount of carrier gas employed can
range from about 0 to about 0.2 kg gas/kg catalyst. Injection
points for carrier gas, especially reactant feed material carrier
gas can be made at multiple points along the fresh or regenerated
catalyst transfer line connecting the regenerator with the lower
end of the riser reactor. The carrier gas will exit the reactor
with the product gas or through the vent stream of the regenerator.
In the case where the carrier gas is also a reactant, a
considerable portion of the carrier gas may be reacted and leave
with the product gas stream from the reactor.
[0024] The short contact time required by the present invention can
be accomplished by way of a number of known reactor designs
including fast fluidized, riser and downer reactors. Riser reactors
are well known and commonly employed in conversion of certain
petroleum fractions into gasoline in fluidized bed catalytic
cracking (FCC) processes. See for example U.S. Pat. No. 3,888,762
which describes a short-time dilute-phase riser reactor designed
for contact times of about 10 seconds, and which further includes
catalyst regeneration and recycle configurations--incorporated
herein by reference. See also: US Publication No. 2004/0082824; WO
2001/85872 and WO 2004/029178. In an FCC process, a solid
particulate catalyst, usually an acidic clay, silica-alumina or
synthetic or natural zeolite type of catalyst, is introduced with a
carrier fluid to the lower end of a long, cylindrical or tubular
reaction vessel together with a petroleum fraction at elevated
temperature and moderate pressure. The cracking process occurs in
the petroleum as the liquid petroleum is vaporized by the hot
catalyst and both rise in the reactor cylinder. At the top of the
riser reactor, the catalyst and hydrocarbon product are separated
and the gasoline product stream exits via a vent pipe for
separation and further processing into gasoline and heating oil
fractions. The catalyst settles in an annular space between the
outside wall of the riser tube and an inner wall of the reactor
housing through which a stripper gas contacts the catalyst, at a
rate which does not prevent settling of the catalyst, and strips
off additional petroleum product from the catalyst surface. The
catalyst is then sent to a regenerator/reactivator in which the
catalyst is contacted with a regeneration fluid, usually an
oxygen-containing gas for combustion of any remaining hydrocarbons,
heavy residuals or tars, and the regenerated catalyst is sent back
to the lower end of the riser reactor to contact additional
petroleum for cracking. Spent catalyst may also be directly
recycled to the lower end of the reactor without regeneration.
[0025] In a similar manner, in a preferred embodiment of the
present invention the alkylaromatic hydrocarbon compound and/or the
paraffinic hydrocarbon compound are introduced to the lower end of
a reactor and contacted by the hot fresh or regenerated catalyst
which is pneumatically moved by a carrier gas. As the hydrocarbon
compound(s) rise in the cylindrical reactor with the catalyst, the
dehydrogenation reaction takes place and at the top or upper end of
the riser, the vinyl aromatic hydrocarbon compound and/or lower
olefin is separated from the catalyst. The riser reactor can be
constructed from conventional materials used in FCC or
petrochemical processing and is conveniently a steel vessel using
an alloy sufficient for containing the hydrocarbon materials of the
reaction, considering the temperature, pressure and flow rates
employed and may be refractory lined. The dimensions of the riser
reactor are dependent on the process design of a processing
facility, including the proposed capacity or throughput, gas hourly
space velocity (GHSV), temperature, pressure, catalyst efficiency
and unit ratios of feed converted to products at a desired
selectivity.
[0026] The separation of gaseous hydrocarbon and catalyst is
conveniently accomplished by means of a centrifugal impingement
separator, such as a cyclone separator, but the separation can by
done by any conventional means for solid-gas separations, including
filtration and liquid suspension. It is important to minimize the
average contact time between the catalyst and hydrocarbon once they
exit the dehydrogenation reactor. This is preferably accomplished
by at least one of two means; physical separation of catalyst from
hydrocarbon, and cooling the catalyst and/or hydrocarbon to a
temperature below the reaction temperature of hydrocarbon present.
The average contact time of the catalyst and hydrocarbon at
reaction temperature in the separation device is typically less
than 60 seconds, preferably less than about 10 seconds, and more
preferably less than about 5 seconds, and still more preferably
less than about 3 seconds. The separation device may be a
conventional solid-gas impingement separator, such as cyclone
separators commonly used in FCC applications. Preferred cyclone
separators include two staged or "coupled" designs including both
positive and negative pressure designs. Further examples are
provided in U.S. Pat. Nos. 4,502,947; 4,985,136 and 5,248,411. Once
separated, the catalyst is either recycled to the dehydrogenation
reactor or transferred to a regenerator.
[0027] In addition to separating the catalyst and hydrocarbon, the
separation device may include a heat exchanger and/or quenching
unit for delivering a fluid to cool the catalyst and/or
hydrocarbons to a temperature below the reaction temperature. Such
fluid may be delivered via a conventional quenching design
including pressurized nozzles for delivering quenching fluid, for
example liquid styrene, water, and the like. Such quenching
technology is available from Stone & Webster and BP Amoco.
[0028] The average contact time between the catalysts and
hydrocarbons while at reaction temperature through the entire
dehydrogenation reactor and separation device is preferably less
than 60 seconds more preferably less than about 20 seconds, and
still more preferably less than about 10 seconds, and event more
preferably less than about 7 seconds.
[0029] Once separated, the gaseous hydrocarbon is further
separated, that is aromatics and non-aromatics, etc., which may be
part of an integrated process as described in U.S. Pat. No.
6,031,143; WO 02/096844; and US 2003/0028059. The spent catalyst
may then optionally be sent to a stripper, and then either to a
regenerator or recycle loop, after which the catalyst is returned
to the dehydrogenation reactor. During regeneration the catalyst is
contacted with a regeneration fluid, usually an oxygen-containing
gas and optionally a fuel source such as methane or natural gas
where remaining hydrocarbons, coke, heavy residues, tar, etc. are
removed from the catalyst, and the resulting regenerated catalyst
is cycled back to the dehydrogenation reactor. A portion of the
spent catalyst may be cycled back to the dehydrogenation reactor
without regeneration via a recycle loop. Recycled spent catalyst
may be combined with regenerated catalyst as a means of controlling
temperature and catalyst activity within the dehydrogenation
reactor. The combination of recycled and regenerated catalyst may
be optimized based upon feedback from the output of the
dehydrogenation reactor. An example of a means for controlling this
combination is described in WO 03/083014, incorporated herein by
reference. Examples of both regeneration and recycle configurations
are provided in U.S. Pat. No. 3,888,762 and US 2003/0196933, which
are also incorporated herein by reference.
[0030] Preferred catalysts for use in the present invention are
very active and are capable of dehydrogenating paraffin and
alkylaromatic hydrocarbons in less than a few seconds at ideal
reaction temperatures. Preferred catalyst include solid particulate
type which are capable of fluidization and, preferably, a those
which exhibit Geldart A properties, as known in the industry.
Gallium-based catalyst described in U.S. Pat. No. 6,031,143 and WO
2002/096844 and are particularly preferred in the present process
and are incorporated herein by reference. One class of preferred
catalyst for the dehydrogenation reaction is based on gallium and
platinum supported on alumina in the delta or theta phase, or in a
mixture of delta plus theta phases, or theta plus alpha phases, or
delta plus theta plus alpha phases, modified with silica, and
having a surface area preferably less than about 100 m.sup.2/g, as
determined by the BET method known to those skilled in the field.
More preferably, the catalyst comprises: [0031] i) from 0.1 to 34
percent by weight, preferably 0.2 to 3.8 percent by weight of
gallium oxide (Ga.sub.2O.sub.3); [0032] ii) from 1 to 200 parts per
million (ppm), preferably 100 to 150 ppm by weight of platinum;
[0033] iii) from 0.05 to 5 percent be weight, preferably 0.1 to 1
percent by weight of an alkaline and/or earth-alkaline such as
potassium; [0034] iv) from 0.08 to 3 percent by weight silica;
[0035] v) the balance to 100 percent being alumina. Similar
gallium-based catalyst are described in WO 2003/053567 which
further includes manganese; and US 2004/02242945 which further
includes zinc, and EP-B1-0,637,578. The description of the catalyst
from these documents is expressly incorporated herein by
reference.
[0036] Another suitable catalyst for the dehydrogenation reaction
is based on chromium and comprises: [0037] i) from 6 to 30 percent,
preferably, from 13 to 25 percent, by weight of chromium oxide
(Cr.sub.2O.sub.3); [0038] ii) from 0.1 to 3.5 percent, most
preferably, from 0.2 to 2.8 percent, by weight stannous oxide
(SnO); [0039] iii) from 0.4 to 3 percent, most preferably, from 0.5
to 2.5 percent, by weight of an alkaline oxide, for example,
potassium oxide; [0040] iv) from 0.08 to 3 percent by weight
silica; [0041] v) the balance to 100 percent being alumina in the
delta or theta phase, or a mixture of delta plus theta phases, or
theta plus alpha phases, or delta plus theta plus alpha phases.
[0042] The catalysts mentioned hereinabove can be used as such or
diluted with an inert material, for example, alpha-alumina,
possibly modified with oxides of alkaline metals and/or silica, at
a concentration of the inert product of between 0 and 50 percent by
weight.
[0043] Details on the preparation of the aforementioned catalysts
and their more preferred species can be found in U.S. Pat. No.
6,031,143 and EP-B1-0,637,578. Typically, the process of preparing
the aforementioned dehydrogenation catalysts comprises dispersing
precursors of the catalytic metals, for example, solutions of
soluble salts of the catalytic metals onto the carrier consisting
of alumina or silica. An example of dispersion can comprise
impregnation of the carrier with one or more solutions containing
the precursors of gallium and platinum, or with one or more
solutions of the precursors of chromium and tin, followed by drying
and calcination. An alternative method comprises ion adsorption
followed by the separation of the liquid portion of the adsorption
solution, drying, and activation of the resultant solid. As another
alternative, the carrier can be treated with volatile species of
the desired metals. In the case of added alkaline or alkaline earth
metals, the addition procedure comprises co-impregnation of the
alkaline or alkaline earth metal with the primary catalytic metals
(that is, Ga and Pt, or Cr and Sn), or alternatively, addition of
the alkali or alkaline earth metal to the carrier prior to
dispersion of the primary catalytic metals, and thereafter,
possible calcination of the solid.
[0044] Other suitable dehydrogenation catalysts, based on iron
oxide, are disclosed in EP 1 216 219. These catalyst comprise:
[0045] (i) from 1 to 60 percent, preferably from 1 to 20 percent,
by weight iron oxide; [0046] (ii) from 0.1 to 20 percent,
preferably from 0.5 to 10 percent, by weight of at least one
alkaline or alkaline earth metal oxide, more preferably, potassium
oxide; [0047] (iii) from 0 to 15 percent, preferably, from 0.1 to 7
percent, by weight of at least one rare earth oxide, preferably,
selected from the group consisting of cerium oxide, lanthanum
oxide, praseodymium oxide, and mixtures thereof; [0048] (iv) the
complement to 100 percent being a carrier consisting of a
microspheroidal alumina with a diameter selected from those in
delta or theta phase, or in a mixture of theta plus alpha phases,
or in a mixture of delta plus theta plus alpha phases, modified
preferably with from 0.08 to 5.0 weight percent of silica. The
carrier in the preferred iron oxide catalyst more preferably has an
average particle diameter and particle density such that the final
product can be classified as Group-A according to Geldart (Gas
Fluidization Technology, D. Geldart, John Wiley & Sons) and a
surface area of less than about 150 m.sup.2/g, as measured by the
BET method known to those skilled in the art. The process of
preparing the iron oxide catalyst is well known and fully described
in EP 1216 219
[0049] Another applicable dehydrogenation catalyst consists of a
mordenite zeolite, optionally, promoted with a metal selected from
gallium, zinc, the platinum group metals, or a combination thereof,
as described in U.S. Pat. No. 5,430,211 and incorporated herein by
reference. The mordenite is preferably acid extracted and
thereafter impregnated or ion-exchanged with one or more metals
selected from gallium, zinc, and the platinum group metals, more
preferably, gallium. In this catalyst, the total metal loading
typically ranges from 0.1 to 20 weight percent, based on the total
weight of the catalyst.
[0050] As mentioned, the preferred catalyst for use with the
present invention are very active and are capable of completing the
dehydrogenation reaction in a relatively short reaction time, for
example in matter of seconds. Consequently, if the catalyst is
allowed to remain in contact with the hydrocarbon mixture at
reaction temperature for a longer period than necessary to complete
the dehydrogenation reaction, undesirable by-products are formed
from unreacted starting materials and/or the desired products are
degraded by a continued exposure to the catalyst at process
conditions. The use of short contact times between the hydrocarbon
and catalyst while at reaction temperature in the dehydrogenation
reactor results in an unexpectedly beneficial conversion,
selectivity and decrease in the amounts of by-products formed. This
unexpected effect is magnified by the use of short contact times
between the hydrocarbon products and catalyst while at reaction
temperature in the separation device. Further, the use of a reactor
with relatively short contact or residence time decreases the
amount of catalyst required for the process. A lower catalyst
inventory provides operating and capital advantages compared with
prior art processes.
BRIEF DESCRIPTION OF THE SEVERAL VIEWS OF THE DRAWINGS
[0051] Several preferred embodiments of the invention are
illustrated in the attached figures. Turning to FIG. 1, a tubular
cylindrical riser reactor 10 having a lower end 12 and an upper end
14 is connected at its lower end 12 to a fresh or regenerated
catalyst transfer line 16 and at its upper end 14 to a product gas
exit line 18. Spent or deactivated catalyst is removed from the
product gas at upper end 14 by a separation device (not shown)
which can be a conventional solid-gas impingement separator, such
as a cyclone separator as previously described, and the catalyst is
sent via spent catalyst transfer line 20 to regenerator 22 which is
a reaction vessel in which combustion air is blown into the
regenerator 22 by means of air line 24. Supplemental fuel may be
added via fuel line 62 to provide the heat of reaction and
necessary sensible heat, including the heat of vaporization in the
case of liquid feed in the riser reactor 10. The combustion
products from the oxidation of hydrocarbon on the catalyst are
removed from the regenerator 22 by means of vent gas line 28. Prior
to being sent for disposal or additional heat recovery, the vent
gas may be filtered for removal of catalyst fines and dust by
conventional equipment which is not shown. As a result of the
combustion and hydrocarbon removal the catalyst is regenerated and
heated to a temperature sufficient to dehydrogenate the hydrocarbon
feed materials and is removed from the regenerator 22 by means of
regenerated catalyst exit line 30. Fluidization is maintained by
injection of a diluent or carrier gas, for example nitrogen, by
means of nitrogen injection lines 26 and 32, and carrier gas
injection lines 34, 36, and 38, so that catalyst is introduced to
the lower end 12 of riser reactor 10 where it contacts ethane which
is introduced via hydrocarbon feed line 40.
[0052] While FIG. 1 has been described with reference to the
dehydrogenation of ethane, it will be appreciated that the present
invention, along with the embodiment of FIG. 1 is also applicable
for the dehydrogenation of other hydrocarbons, including lower
alkanes such as propane and butane, and lower alkylaromatics, such
as ethylbenzene, propylbenzene and methylethylbenzene.
[0053] In operation, the embodiment shown in FIG. 1 proceeds by
feeding regenerated catalyst at a temperature of from about 600 to
about 800.degree. C. from the regenerator 22 by means of
regenerated catalyst exit line 30 into fresh or regenerated
catalyst transfer line 16 with the catalyst being maintained in a
fluid state of a Geldart A solid particulate material by means of
fluidizing inert gas, such as nitrogen, fed through nitrogen
injection lines 26 and 32, and carrier gas, which may be inert
(again, such as nitrogen) or a reactant gas, such as a paraffinic
hydrocarbon, such as for example, a lower alkane, preferably
ethane, propane, or a butane, via carrier gas injection lines 34,
36, and 38. This catalyst and carrier gas mixture is introduced to
the lower end 12 of riser reactor 10 and contacts a hydrocarbon
feed in liquid or gaseous form, preferably the latter, introduced
by means of hydrocarbon feed line 40. The catalyst and hydrocarbon
feed, for example, a lower alkane, such as ethane, propane or a
butane, or an alkylaromatic hydrocarbon compound, or a mixture of
both lower alkane and an alkylaromatic hydrocarbon compound,
contacts the catalyst and rises in the riser reactor 10 with the
catalyst, feed (which by this time has been transformed into a gas)
and the carrier gas. As the catalyst-feed-carrier gas mixture rises
in the reactor, the dehydrogenation reaction occurs and the feed is
converted into a lower olefin and/or a vinyl aromatic compound,
depending on the feed material. As the reaction mixture containing
gas and catalyst arrives at the upper end 14 of riser reactor 10,
the catalyst and gaseous reaction mixture are separated by a
solid-gas separation device, such as an impingement separation
device which may preferably be a cyclone gas-solid separator, which
is conventional and not shown, but which is well known to those of
skill in the art of the FCC industry. The separated product gas is
sent to recovery and purification and the catalyst is sent for
regeneration and re-heating by means of spent or deactivated
catalyst transfer line 20. As the spent or deactivated catalyst is
introduced into the regenerator 22, it contacts heated combustion
air which is introduced by air line 24 and supplemental fuel
introduced by fuel line 62, such that the hydrocarbon materials
remaining on the surface of the catalyst are burned off and exit
the regenerator via vent gas line 28. The combustion process also
serves a second purpose and that is to heat the catalyst so that
the catalyst can function as a heat transfer agent or medium in the
riser reactor 10. As used in this embodiment, the hydrocarbon feed
40 can be a paraffinic hydrocarbon such as a lower alkane, an
alkylaromatic hydrocarbon compound, or a mixture of the two.
[0054] FIG. 2 illustrates another preferred, non-limiting
embodiment which is a variant on the process of the present
invention using a similar riser reactor 10 configuration as
described with respect to FIG. 1. In this embodiment the paraffinic
hydrocarbon (for example ethane) is fed to the riser reactor 10 at
or adjacent the lower end 12 by means of ethane feed line 44 and
the lower alkylaromatic hydrocarbon compound (for example
ethylbenzene), is fed at a higher point in the riser reactor 10,
for example at ethylbenzene feed line 42. Thus, the type of
reaction illustrated by the process of FIG. 2 is a "split feed"
riser reactor process which produces styrene and by-products, such
as ethylene which can be returned to an alkylation step to react
with additional benzene to produce more ethylbenzene as part of an
integrated process.
[0055] FIG. 3 illustrates yet another preferred, non-limiting
embodiment of the invention. In this embodiment, a "dual riser"
reactor configuration is illustrated in which the riser reactors 10
and 48 are connected in series. As shown in FIG. 3, riser reactor
10 has lower end 12 and upper end 14. Connected to lower end 12 is
fresh or regenerated catalyst line 16 and the catalyst is
maintained in fluidized state by injection of carrier gas via lines
34 and 36. Hydrocarbon feed material, such as ethane, is introduced
to the lower end 12 of riser reactor 10 by means of hydrocarbon
feed line 40. At this stage of the process, the configuration is
much like that of FIG. 1; however, the product gas from riser
reactor 10 in FIG. 3 is fed to a separation and recovery section
(not shown) by means of product gas exit line 18 from which a side
product gas line 46 leads to an alkylaromatic hydrocarbon compound
feed line, such as ethylbenzene feed line 42. Alternatively, both
side product gas line 46, which carries primarily the lower olefin
produced in riser reactor 10 in addition to by-products and carrier
gas, can be fed separately into a second riser reactor, such as at
48, having a lower end 50 and an upper end 52. Also entering the
lower end 50 of second riser reactor 48 is a partially deactivated
catalyst line 54 which leads from the upper end 14 of riser reactor
10 to the lower end 50 of second riser reactor 48. Carrier gas line
38 can be used to introduce fluidizing carrier gas into partially
deactivated catalyst line 54 at one or multiple points along
partially deactivated catalyst line 54. As the ethylene and
ethylbenzene rise in second riser reactor 48 with the catalyst and
carrier gas, the catalyst is at a lower temperature than when
initially introduced to the lower end 12 of riser reactor 10. The
relatively lower temperature than in riser reactor 10 permits
satisfactory reaction rates for the alkylaromatic hydrocarbon
compound and prevents over reaction to undesired by-products, thus
decreasing the yield, conversion and selectivity of the
dehydrogenation reaction. The upper end 52 of second riser reactor
48 is connected to second product gas exit line 56 and can lead the
vinyl aromatic hydrocarbon compound, such as crude styrene monomer
contained in the product gases, into the product gas separation and
recovery section, which is conventional and not further described
or identified herein. Prior to exit from second riser reactor 48,
the reaction mixture must be separated from the deactivated
catalyst and this is done in a solid-gas separation device, such as
a cyclone separator, not shown. The separated and deactivated
catalyst is fed back to the regenerator 22 by means of spent or
deactivated catalyst transfer line 20 which in this embodiment
leads from the upper end 52 of second riser reactor 48 to the
regenerator 22 where the catalyst is regenerated, as previously
described. In operation, the process is much like that described in
relation to the process illustrated in FIGS. 1 and 2, except that
the product gas from the upper end 14 of riser reactor 10 is split
and a portion is introduced into the lower end 50 of second riser
reactor 48. Ethylbenzene is also introduced into the lower end 50
of second riser reactor 48 along with the partially deactivated
catalyst via partially deactivated catalyst line 54 and the
dehydrogenation of the ethylbenzene proceeds at somewhat milder
conditions in second riser reactor 48 than in riser reactor 10. At
the upper end 52 of second riser reactor 48, the product gases are
separated from the catalyst in a solid gas separator device, such
as a cyclone separator (which is conventional and not shown) and
the product gases exit via second product gas exit line 56 and the
catalyst is sent back to regenerator 22 for regeneration and
reheating via spent or deactivated catalyst transfer line 20.
[0056] In a still further preferred embodiment of this invention
shown in FIG. 4, the reactor/regenerator configuration is similar
to that of FIG. 3, except that the second riser reactor 48 has its
own catalyst feed and removal transfer lines, namely second fresh
or regenerated catalyst transfer line 58 and second spent or
deactivated catalyst transfer line 60 which feed active catalyst to
second riser reactor 48 and remove catalyst from it and send the
deactivated or spent catalyst back to regenerator 22. While shown
as utilizing a common regenerator 22, it will be appreciated that
each reactor may include a separate regenerator.
[0057] In operation and as shown in FIG. 4, the catalyst from
regenerator 22 is led by regenerated catalyst exit line 30 to
either riser reactor 10 or second riser reactor 48 via fresh or
regenerated catalyst transfer line 16 or second fresh or
regenerated catalyst transfer line 58, respectively. The feed to
riser reactor 10 is ethane via hydrocarbon feed line 40 and to
second riser reactor 48 is ethylbenzene via ethylbenzene feed line
42. On contact with the catalyst in the riser reactors, the ethane
and ethylbenzene are converted into ethylene and styrene monomer,
respectively, and the crude gaseous products are separated from the
catalyst in gas-solid separators, such as cyclone separators (not
shown) and sent to product gas separation and recovery operations
(not shown) to produce ethylene for recycle to make additional
ethylbenzene and styrene monomer, respectively. In a similar manner
and using propane or butane instead of ethane feed, the process of
this invention would dehydrogenate the feed to propylene or
butylenes, respectively; or using isopropyl benzene or methyl ethyl
benzene as feed material, the process of this invention would
dehydrogenate the feed to cumene or alpha-methyl styrene,
respectively.
[0058] FIG. 5 illustrates yet another embodiment of the invention
similar to that shown in FIG. 4 but with the addition of a catalyst
recycle loop comprising a catalyst transfer line 64, carrier gas
injector line 66 and flow valve 68. Spent catalyst is removed from
the product gas at the upper end 52 of the dehydrogenation reactor
48 via a separation device (not shown) and is recycled back to the
bottom end 50 of the reactor 48 via catalyst transfer line 64.
Fluidization of the spent catalyst is maintained by the injection
of a carrier gas, for example nitrogen by means of injection line
in module 66. In addition to providing a carrier gas, an
oxygen-containing gas may be introduced in order to partially
reactivate the catalyst, in which case module 66 would include a
chamber for reaction and removal of hydrocarbon residue. Flow of
catalyst through the recycle loop is controlled by one or more
valves, for example 68 which may be controlled remotely according
to predetermined performance criteria including reactor 48
temperature, catalyst activity, etc. Recycled catalyst may be
combined with regenerated catalyst prior to introduction in the
bottom of reactor 48, or may be introduced via separate entry
points (not shown).
[0059] Additional configurations of the reactor(s), regenerator and
recycle loop can be envisioned by one skilled in the art. For
example, one skilled in the art will appreciate that multiple
reactors could be arranged to feed into a common separation device
with shared or separate catalyst regenerators and various recycle
loops. The present invention is desired to be limited only by the
lawful scope of the appended claims The present invention does not
preferably include oxidative dehydrogenation, that is
oxodehydrogenation. In fact, oxygenates can poison some types of
catalyst; however, oxygen may used to regenerate or reactive
catalyst during the regeneration process. Moreover, the present
invention preferably does not utilize steam as is typically used in
convention styrene production process.
[0060] Another preferred embodiment of the invention utilizes the
previously described dehydrogenation process as part of an
integrated process for making olefins and vinyl aromatics. More
specifically, the previously described dehydrogenation, (along with
regeneration and/or recycle processes) can be used to replace the
dehydrogenation schemes described in U.S. Pat. No. 6,031,143; WO
02/096844; and co-pending US 2003/0028059. In such an integrated
process, a paraffinic hydrocarbon such as a lower alkane, for
example ethane, and benzene are the primary raw materials.
Ethylene, preferably "recycled" and benzene are feed to a
conventional alkylation reactor as is well known in the art and as
described in the references mentioned above, hereby incorporated by
reference. Alkylation of benzene with ethylene in typically
conducted in the presence of aluminum chloride or zeolites
catalyst. Variations include the use of dilute ethylene and a
catalytic distillation approach where liquid phase alkylation and
product separation take place simultaneously. Specific examples
include the "EBOne Process" available from ABB Lummus/UOP, "EB Max
Process" available from ExxonMobil/Badger and similar alkylation
technology available from CDTECH, a partnership between ABB Lummus
Global Inc. and Chemical Research and Licensing.
[0061] The alkylation affluent is recovered and optionally subject
to separation, that is separation of aromatics from non-aromatics,
removal of hydrogen, etc. Alkylaromatic, for example ethylbenzene,
and paraffinic hydrocarbon, for example ethane, are then
dehydrogenation as previously described. The gaseous products of
dehydrogenation are recovered and separated, for example aromatics
from non-aromatics, with vinyl aromatics, for example styrene being
recovered, olefins, for example ethylene (and possibly paraffinic
hydrocarbons, for example ethane) being recycled to the alkylation
reactor, and alkylaromatics being recycled to the dehydrogenation
reactor.
* * * * *