Note: Descriptions are shown in the official language in which they were submitted.
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ISOMERIZATION AND CATALYTIC ACTIVATION OF PENTANE-
ENRICHED HYDROCARBON MIXTURES
CROSS-REFERENCE TO RELATED APPLICATIONS
[0001] This application is a PCT International application which claims the
benefit of and
priority to U.S. Provisional Application Ser. No. 62/742,749 titled
"Isomerization and Catalytic
Activation of Pentane-Enriched Hydrocarbon Mixtures" filed October 8, 2018,
U.S. Provisional
Application Ser. No. 62/742,765 titled "Systems for the Catalytic Activation
of Pentane-
Enriched Hydrocarbon Mixtures" filed October 8, 2018, U.S. Application Serial
No. 16/593,238
filed October 4, 2019, titled "Isomerization and Catalytic Activation of
Pentane-Enriched
Hydrocarbon Mixtures", and U.S. Application Serial No. 16/593,476 filed
October 4, 2019,
titled "Systems for the Catalytic Activation of Pentane-Enriched Hydrocarbon
Mixtures" all of
which are hereby incorporated by reference in its entirety.
FIELD OF THE INVENTION
[0002] The present disclosure generally relates to processes and systems
that converts at
least a portion of the n-pentane in a light hydrocarbon feed stream to
isopentane, followed by an
activation step and subsequent upgrading to larger hydrocarbons in either an
alkylation reactor
or oligomerization reactor. The processes and systems produce hydrocarbons
suitable for use as
a blend component of a liquid transportation fuel.
BACKGROUND
[0003] A large surplus of pentanes are available in the petroleum refining
industry, arising
predominantly from the increased production of light hydrocarbons from U.S.
shale formations,
and also from limits on the quantity of volatile components that can be
blended into finished
transportation fuels, which must adhere to regulations on minimum vapor
pressure.
Unfortunately, conventional processes for upgrading light alkanes to value-
added products are
not well-suited for hydrocarbon feed streams that primarily comprise pentanes
(i.e., isopentane
and n-pentane). Therefore, it would be beneficial to find improved processes
and systems for
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efficiently converting pentanes to more valuable products, including
transportation fuels and
chemicals, while minimizing the production of C1-C4 light paraffins.
[0004] The inventive processes disclosed herein provide an improved
upgrading route for
pentane-rich fuel blend-stocks and other pentane-rich streams that do not meet
government
specifications for a transportation fuel. The inventive processes and systems
provide enhanced
yields of upgraded products that may be suitable for use as transportation
fuels or other value-
added chemical products.
BRIEF SUMMARY OF THE DISCLOSURE
[0005] Certain embodiments comprise an method for converting a feed stream
comprising
pentanes to produce a liquid transportation fuel, the method comprising: a.)
providing a
hydrocarbon feed stream comprising at least 50 wt.% pentanes, including both n-
pentane and
isopentane; b.) contacting the hydrocarbon feed stream with one or more
isomerization catalysts
in a first reaction zone that is maintained at a temperature and a pressure
that facilitates the
isomerization of at least a portion of the n-pentane in the hydrocarbon feed
stream to isopentane,
thereby producing an isomerization effluent characterized by an increased
ratio of isopentane to
n-pentane relative to the hydrocarbon feed stream; c.) contacting the
isomerization effluent with
an activation catalyst in a second reaction zone that is maintained at a
temperature and pressure
that facilitates at least one reaction selected from dehydrogenation, cracking
and aromatization,
thereby converting at least a portion of hydrocarbons present in the
isomerization effluent to
produce an activation effluent comprising olefins containing from two to five
carbon atoms,
monocyclic aromatics and unconverted alkanes containing from two to five
carbon atoms; d.) at
least partially condensing the activation effluent to produce a liquid
hydrocarbons fraction and a
gaseous light hydrocarbons fraction, where the liquid hydrocarbons fraction
comprises
monocyclic aromatics and unreacted alkanes containing at least five carbon
atoms, wherein the
gaseous light hydrocarbons fraction comprises at least 80 wt.% hydrocarbons
containing four or
fewer carbon atoms and hydrogen.
[0006] Some embodiments further comprise separating the mixed liquid
hydrocarbons into
an aromatics fraction and an unreacted C5/C6 hydrocarbons fraction, where the
aromatics
fraction comprises monocyclic aromatics suitable for use as a blend component
of gasoline and
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the unreacted C5/C6 hydrocarbons fraction comprises alkanes and olefins
containing from five
to six carbons that may be mixed with the hydrocarbon feed stream of part a.).
[0007] In some embodiments, the hydrocarbon feed stream comprises at least
5 wt.% of
hydrocarbons containing four or fewer carbon atoms. In some embodiments, the
hydrocarbon
feed stream comprises at least 60 wt.% pentanes. In some embodiments, the
hydrocarbon feed
stream is contacted with one or more isomerization catalysts contained within
in multiple
reaction zones that are arranged in a series configuration.
[0008] In some embodiments, the activation catalyst comprises one or more
zeolites
characterized by Si/A1 ratio ranging from 12 to 80. In some embodiments, the
activation catalyst
comprises ZSM-5 zeolite. In some embodiments, the activation catalyst
facilitates at least one
reaction selected from the group consisting of oligomerization,
dehydrogenation, and
aromatization.
[0009] In some embodiments, the temperature in the first reaction zone is
maintained at a
temperature in the range from 500 C to 625 C and a pressure in the range from
15 psig to 100
psig. In some embodiments, the temperature in the first reaction zone is
maintained at a
temperature in the range from 525 C to 600 C and a pressure in the range from
15 psig to 75
psig. In some embodiments, the temperature in the second reaction zone is
maintained at a
temperature in the range from 550 C to 600 C and a pressure in the range from
20 psig to 60
psig. In some embodiments, the temperature in the second reaction zone is
maintained at a
temperature in the range from 575 C to 600 C and a pressure in the range from
20 psig to 50
psig.
[0010] Some embodiments further comprise adding a diluent to at least one
of the
hydrocarbon feed stream and the isomerization effluent prior to the contacting
with the
activation catalyst, wherein the diluent is characterized as less likely to
react with the activation
catalyst than the hydrocarbon feed stream at the conditions of temperature and
pressure that are
maintained in the first reaction zone, and wherein the diluent is
characterized as less likely to
react with the activation catalyst than molecules present in the isomerization
effluent at the
conditions of temperature and pressure that are maintained in the second
reaction zone.
[0011] Some embodiments further comprise adding a diluent to at least one
of the
hydrocarbon feed stream and the isomerization effluent prior to the contacting
with the
activation catalyst, wherein the diluent does not react with the isomerization
catalyst at the
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conditions of temperature and pressure that are maintained in the first
reaction zone, and wherein
the diluent does not react with the activation catalyst at the conditions of
temperature and
pressure that are maintained in the second reaction zone. In some embodiments,
the diluent is
added in an amount that alters the specificity of the activation catalyst to
increase the production
of olefins, decrease the production of aromatics, or combinations thereof,
thereby increasing the
ratio of olefins to aromatics in the activation effluent. In some embodiments,
the diluent is added
in an amount that is effective to produce an activation effluent characterized
by an olefins to
aromatics ratio in the range of 0.5 to 2Ø In some embodiments, the diluent
is added in an
amount that is effective to produce an activation effluent characterized by an
olefins to aromatics
ratio in the range of 0.5 to 1Ø In some embodiments, the diluent is selected
from methane,
ethane, propane, butanes, and combinations thereof
BRIEF DESCRIPTION OF THE DRAWINGS
[0012] A more complete understanding of the present invention and benefits
thereof may be
acquired by referring to the follow description taken in conjunction with the
accompanying
drawings in which:
[0013] Figure 1 is a diagram depicting a first embodiment of the inventive
processes and
systems.
[0014] Figure 2 is a diagram depicting a second embodiment of the inventive
processes and
systems.
[0015] Figure 3 is a diagram depicting a third embodiment of the inventive
processes and
systems.
[0016] Figure 4 is a bar graph depicting product selectivity resulting from
catalytic
activation of either n-pentane or iso-pentane at two different temperatures.
[0017] Figure 5 is a bar graph showing the effect of isomerization of the
feed stream on the
total conversion and product yield for a first feed stream comprising a 1:1
ratio of n-05 to i-05,
and a second feed stream comprising a 7:3 ratio of n-05 to i-05.
[0018] Figure 6 is a bar graph showing the effect of isomerization of the
feed stream on the
total conversion and product selectivity for a first feed stream comprising a
1:1 ratio of n-05 to
i-05, and a second feed stream comprising a 7:3 ratio of n-05 to i-05.
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[0019] The invention is susceptible to various modifications and
alternative forms, specific
embodiments thereof are shown by way of example in the drawings. The drawings
may not be to
scale. It should be understood that the drawings are not intended to limit the
scope of the
invention to the particular embodiment illustrated.
DETAILED DESCRIPTION
[0020] The present disclosure provides processes and systems for converting
a mixture of
light hydrocarbons to liquid transportation fuels. More specifically, it
pertains to the conversion
of any hydrocarbon mixture that predominantly comprises pentanes to generate
upgraded
products that may be sold as a value-added chemical or utilized as a blend
component of a liquid
transportation fuel.
[0021] Generally speaking, the inventive processes and systems described
herein utilize a
hydrocarbon feed stream comprising both isopentane and n-pentane and performs
an initial
isomerization of the hydrocarbon feed stream to convert at least a portion of
the n-pentane (n-
05) in the hydrocarbon feed stream to isopentane (i-05). The resulting
isomerization effluent is
then catalytically activated under conditions of temperature and pressure
(typically measured at
the inlet of the activation reactor) that maximize the catalytic conversion of
the isomerization
effluent to olefins and aromatics, while minimizing the undesirable production
of C1-C4 light
hydrocarbons, often referred to as fuel gas.
[0022] The resulting activation effluent is optionally further upgraded in
a third reactor by
contact with an oligomerization and/or alkylation catalyst at a temperature
and pressure that
facilitates conversion of the activation effluent to value-added chemicals
and/or products
suitable for use as a liquid transportation fuel blend component.
[0023] The present inventive processes and systems take advantage of the
differing
reactivity of pentane isomers to catalytic activation. Isopentane (i-05)
exhibits catalyst-
dependent reactivity that is typically different from n-pentane (n-05), and
the optimal reactor
conditions for the two isomers are therefore distinct. Experimentally,
isopentane (i-05) is more
reactive than n-pentane (n-05), and thus, can be activated at lower
temperatures while
maintaining high yields of desired products (such as olefins and aromatics)
and decreasing the
yield of C1-C4 paraffins. The inventive system takes advantage of this
difference by isomerizing
a portion of the n-05 to i-05 and n-C6 to iC6 in a first isomerization step in
order to maximize
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both the conversion yield and selectivity of the activation step to form
useful products, including
(but not limited to) olefins and aromatics. Additional advantages will become
evident from the
detailed disclosure provided below.
[0024] As mentioned, the hydrocarbon feed stream generally comprises a
stream of light
hydrocarbons that comprises a mixture of pentane isomers (C5), although
certain embodiments
may additionally comprise C1-C4 hydrocarbons, C6-C7 hydrocarbons, or both. The
hydrocarbon
feed stream comprises at least 10 wt.% of a mixture of pentane isomers;
optionally, at least 20
wt. %, optionally, at least 30 wt.%, optionally, at least 40 wt.%, optionally
at least 50 wt.%,
optionally, at least 60 wt.%, or optionally, at least 70 wt.%. of a mixture of
pentane isomers. In
certain embodiments, the hydrocarbon feed stream may be obtained by processing
a stream of
natural gas liquids to remove lighter components (i.e., C1-C4) by way of
conventional natural
gas processing technologies that are well-characterized, such as de-
methanizer, de-ethanizer, de-
propanizer and de-butanizer fractionation columns. A typical result of such
processes is
commonly characterized as natural gasoline, comprising about 72 wt.% pentanes,
with the
remainder mostly comprising C6.
[0025] A first embodiment of the inventive processes and systems is
illustrated by the
process flow-diagram of Figure 1. A hydrocarbon feed stream 101 that comprises
both n-pentane
and isopentane is converted in a system 10. Typically, the hydrocarbon feed
stream 101
comprises at least 50 wt.% of pentane isomers, although in certain
embodiments, the pentane
isomers may comprise at least 60 wt.%, or at least 70 wt.% of the hydrocarbon
feed stream 101.
Further, the hydrocarbon feed stream 101 typically comprises less than 30
wt.%, optionally, less
than 20 wt.%, optionally, less than 10 wt.% of hydrocarbons containing four or
fewer carbon
atoms.
[0026] The hydrocarbon feed stream 101 is received by an isomerization
reactor 110 (that
may optionally contain more than one isomerization catalyst, or may optionally
comprise more
than one isomerization reactor arranged in series configuration) that contains
an isomerization
catalyst 115 and comprises a first reaction zone (not depicted) that is
maintained at a temperature
and pressure that facilitates the isomerization of at least a portion of the n-
pentane in the
hydrocarbon feed stream to isopentane. The isomerization reaction occurring in
the first reaction
zone produces an isomerization effluent 113 that is characterized by an
increased ratio of
isopentane to n-pentane (relative to the corresponding ratio of the
hydrocarbon feed stream 101).
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[0027] Speaking generally, the isomerization process is designed for
continuous catalytic
isomerization of the n-pentane present in the mixture. The process is
conducted in a first reaction
zone that is contained within an isomerization reactor in the presence of an
isomerization
catalyst. The reactor maintains a partial pressure of hydrogen and operating
conditions of
temperature and pressure in the first reaction zone that promote isomerization
while minimizing
hydrocracking.
[0028] Ideally, the isomerization catalyst (or catalysts) facilitates the
conversion of n-
pentane to the higher octane-number isopentane, while any C6 hydrocarbons
present may be
converted to higher octane 2-3 dimethyl butane (and similar molecules). The
isomerization
reaction is equilibrium-limited. For this reason, any n-pentane that is not
converted on its first
pass through the isomerization reactor may optionally be recycled to the
isomerization reactor,
or converted in multiple isomerization reactors, arranged in series
configuration, thereby further
increasing the ratio of i-05 to n-05 in the product. The relative efficiency
of separation of
pentane isomers by distillation is poor. Thus, recycling may be more
effectively accomplished
by a molecular sieve, which selectively adsorbs n-pentane due to its smaller
pore diameter
relative to isopentane.
[0029] In certain embodiments, the activity of the isomerization catalyst
may be decreased in
the presence of sulfur, thereby decreasing the isomerization rate and,
consequently, the octane
number of the final product. In such embodiments, the hydrocarbon feed stream
is hydrotreated
to remove sulfur prior to being conveyed to the isomerization reactor.
[0030] Generally speaking, the isomerization catalyst may comprise any
known
isomerization catalyst. Currently, three basic families of light naphtha
isomerization catalysts are
known. The first are termed super-acidic catalysts (impregnated acid type),
such as, for example,
chlorinated alumina catalysts with platinum. Super acidic isomerization
catalysts are highly
active and have significant activity at temperatures as low as 265 F (130 C)
using a lower
H2/HC ratio (less than 0.1 at the outlet of the reactor). However, maintaining
the high acidity of
these catalysts requires the addition of a few ppm of chloriding agent to the
feedstock. At the
inlet of the isomerization reactor, this chloriding agent reacts with hydrogen
to form HC1, which
inhibits the loss of chloride from the catalyst. Unlike a zeolitic catalyst,
the acidic sites on a
super-acidic catalyst are irreversibly deactivated by water. These catalysts
are also sensitive to
sulfur and oxygenate contaminants, so the feed stream is generally
hydrotreated and dried to
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remove residual water contamination. Commercially-available examples of
chlorided-alumina
catalysts include, but are not limited to, IS614A, AT-2, AT-2G, AT-10 and AT-
20 ( by Akzo
Nobel) and ATIS-2L (by Axens). Due to their chlorinated nature, these are very
sensitive to feed
impurities, particularly water, elemental oxygen, sulfur, and nitrogen. When
using such super-
acidic catalysts, the reactor operating temperature generally ranges from 14
C to 175 C, while
the operating pressure is generally in the range from 200 psig to 600 psig,
preferably in the range
from 425 psig to 475 psig.
[0031] Zeolitic isomerization catalysts (structural acid type) require a
higher operating
temperature and are effective at isomerization at temperatures ranging from
220 C to about
315 C, preferably at a temperature ranging from 230 C to 275 C. Pressures
utilized for
isomerization with zeolitic isomerization catalysts typically range from 300
psig to 550 psig with
a LHSV from 0.5 to 3.0 hr-1. These catalysts react as bifunctional catalysts
and require hydrogen
at a H2/HC ratio ranging from about 1.5 to about 3. Zeolitic catalysts have
advantages over
chlorided-alumina catalysts due to zeolitic catalyst tolerance for typical
catalyst poisons sulfur,
oxygenates and water. Zeolitic catalysts also do not require the injection of
a chloriding agent in
order to maintain catalyst activity.
[0032] A third type of conventional isomerization catalyst that may be
useful in certain
embodiments comprises sulfated zirconia/metal oxide catalysts. These catalysts
are active at
relatively low temperatures (e.g., 100 C) with the advantage of providing
enhanced isoparaffin
yield. Their biggest drawback is their relative sensitivity to catalyst
poisons, especially water.
Certainly, other examples of isomerization catalysts that are suitable for use
with the present
processes and systems described herein are known by those having experience in
the field, and
thus, require no further disclosure here.
[0033] Again, referring to the embodiment disclosed in Figure 1, the
isomerization effluent
113 is next conveyed to an activation reactor 120 containing a first
activation catalyst 125 and
comprising a second reaction zone (not depicted). The activation reactor 120
is operable to
maintain a temperature and pressure that is suitable to facilitate conversion
of the isomerization
effluent 113 to an activation effluent 128 that comprises olefins containing
from two to five
carbon atoms, monocyclic aromatics and unconverted alkanes containing from two
to five
carbon atoms.
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[0034] Speaking generally, the activation catalyst may comprise a single
catalyst, or a
mixture of different catalysts that contacts the alkanes present in the
isomerization effluent and
facilitates at least one of dehydrogenation, cracking, and aromatization of
the alkanes, thereby
converting at least a portion of hydrocarbons present in the isomerization
effluent to produce the
activation effluent. Moreover, the activation effluent comprises products that
may be utilized as
a commodity chemical, an intermediate amenable to further catalytic upgrading,
or a
transportation fuel (or a component thereof).
[0035] Activation catalysts suitable for use with the processes and systems
described herein
may comprise any catalyst capable of cracking and/or aromatizing hydrocarbons.
Favored
catalysts include supported or unsupported solid acids, metals, metal
chalcogenides, or metal
pnictogenides, including (but not limited to) structured and amorphous silica-
aluminas,
structured and amorphous solid phosphoric acids, clays, other metal oxides,
metal sulfates, or
metal phosphates, and graphite-supported materials. In certain embodiments,
ZSM-5 zeolite
catalysts are utilized that are characterized by Si/A1 ratios ranging from 12-
80, optionally
ranging from 35 to 50. Optionally, one or more elements may be impregnated on
the zeolite
catalyst, including one or more of Ga, Pt, Ni, Mn, Mg, Fe, Cr, P, Cu, La, Sr
and F.
[0036] Generally speaking, dehydrogenation is not a prerequisite for
paraffin activation in
the present inventive process. A sufficient concentration of intermediate
olefins can be
generated through a combination of thermal dehydrogenation and catalytic
cracking such that
typical dehydrogenation catalyst metals (such as platinum, zinc, molybdenum,
or gallium) can
be avoided without significantly decreasing product yield. Conventional
dehydrogenation
catalysts are prone to fouling by sulfur and nitrogen contaminants that are
often present in
hydrocarbon feed streams derived from petroleum, so the ability to operate in
the absence of
these sensitive catalytic materials is highly advantageous to the process.
[0037] The inventive processes generally take advantage of the large
difference in catalytic
reactivity between n-05 and i-05. For example, utilizing a solid acid
activation catalyst at
temperature in excess of 550 C, the measured activation rates differ by up to
4 fold in favor of
i-05, when each isomer is contacted with the same catalyst under identical
conditions (even in
the same reactor simultaneously). Thus, an initial isomerization of the
hydrocarbon feed stream
to increase i-05 content, followed by activating the resulting effluent in
catalytic activation zone,
maximize the yield of value-added, upgraded products (such as olefins and/or
aromatics).
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Increasing conversion of pentane isomers to i-05 also was found to
unexpectedly decrease
selectivity of the activation reaction to Cl-C4 light gases, which typically
have little value other
than as fuel gas. This helps maximize the conversion of the feed to upgraded
products, which is
one of many advantages of the process and systems.
[0038] Table 1 (below) illustrates the difference in the activation
reactivity of i-05 versus n-
05 over a microporous silica-alumina activation catalyst. Feed streams
comprising either 100
wt.% i-05 or 100 wt.% n-05 were each catalytically activated in separate
experiments utilizing
temperatures of either 600 C or 550 C. The conversion and product distribution
for i-05 are
shown in Table 1, columns 2 and 3, while similar results for the activation of
n-05 are shown in
Table 1, columns 4 and 5.
Table 1. Product distributions for i-05 or n-05 isomer feed streams following
conversion by a
1/8" extrudate consisting of 50 wt.% alumina binder and 50 wt.% ZSM-5 zeolite.
Activation was
performed by contacting the ZSM-5 catalyst with a feed stream comprising
either 100 wt.% of i-
05 or 100 wt % of n-05. Results were time-averaged over 16 hours and all
reactions were
performed at 1 atm with a WHSV = 4.0 hr'.
Feed Isomer: i-05 i-05 n-05 n-05
Inlet Temperature: 600 C 550 C 600 C 550 C
Conversion (wt.%): 94.5 82.4 78.5 48.3
Product Distribution (wt.%)
Hydrogen 2.4% 1.6% 1.1% 0.4%
Methane 9.8% 7.8% 5.3% 2.3%
Ethane 3.0% 2.6% 11.5% 6.6%
Ethylene 17.4% 15.6% 14.3% 7.7%
Propane 5.3% 4.8% 10.6% 9.9%
Propylene 21.2% 22.1% 16.4% 10.3%
Butane 2.9% 4.2% 0.8% 1.2%
Butene 8.5% 9.9% 5.8% 5.5%
Isopentane 5.5% 17.6% 0.1% 0%
n-Pentane 0% 0% 21.5% 51.7%
Pentene 1.2% 1.8% 0.8% 1.2%
C6+ alkanes 0.0% 0.3% 0.0% 0.0%
Benzene 4.8% 3.3% 4.9% 1.0%
Toluene 11.1% 6.0% 5.3% 1.6%
Xylene 6.4% 2.2% 1.3% 0.6%
Ethylbenzene 0.3% 0.1% 0.1% 0.0%
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Coke 0.2% 0.1% 0.2% 0.1%
[0039] The data indicates that when comparing the activation of pentane
isomers, conversion
of i-05 to olefins and aromatics is possible at a temperature about 50 C less
than is required for
equivalent conversion of n-05. To be clear, we observed that activation of the
i-05 feed stream
at 550 C converted about the same weight percentage of the feed stream as did
activation of n-
05 at 600 C using the same WHSV. Further, utilizing a decreased temperature of
550 C for
activation of the i-05 feed stream advantageously decreased the production of
C1-C4 light
paraffins from 21.0% to 19.4% by increasing the product distribution toward
olefins rather than
aromatic products. Thus, the ability to separate the i-05 isomer from n-05
isomer (and any C6+
hydrocarbons) and activate the i-05 enriched mixture at relatively reduced
temperature, results
in approximately equivalent total conversion of the overall feed stream, while
decreasing the
formation of undesired Cl-C4 light paraffins.
[0040] Speaking generally, the temperature within the activation reactor
(typically measured
at, or proximal to, the inlet of the activation reactor) is maintained in the
range from 500 C to
650 C; optionally, within the range from 525 C to 625 C; optionally, within
the range from
525 C to 600 C; optionally, within the range from 550 C to 600 C; optionally,
within the range
from 550 C to 575 C; optionally, within the range from 575 C to 600 C.
[0041] Referring again to the embodiment depicted in Figure 1, the
activation effluent 128 is
conveyed into a separator 145 that separates hydrogen and light hydrocarbons
148 containing
four or fewer carbons from a mixed liquid hydrocarbons 152 that predominantly
comprises C5
olefins, single-ring aromatics as well as unreacted pentanes and larger C6+
components
originally present in the hydrocarbon feed stock 101. In certain embodiments,
the separator 145
is a two-phase splitter and separation of the activation effluent 128 is
achieved by partial
condensation. The light hydrocarbons 148 can be either combusted for heat
generation, diverted
to other upgrading processes that are outside the scope of this disclosure
(not depicted), or
directed to a separator 150.
[0042] Again, referring to the embodiment depicted in Figure 1, light
hydrocarbons 148
predominantly comprises hydrogen as well as Cl - C4 hydrocarbons that were not
converted in
the activation reactor 125. Light hydrocarbons 148 is conveyed to the
separator 150 that
typically utilizes a conventional separation technology (such as, but not
limited to, pressure
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swing adsorption technology, membrane separation technology, etc.) to separate
hydrogen from
the light hydrocarbons 148 to produce a hydrogen stream 151 and a Cl-C4 light
paraffins stream
155 that may be combusted 157 to provide at least a portion of the heat
required for the process,
or recycled to a point that is upstream from the activation catalyst 125 to
serve as the diluent 115
that is mixed with the isomerization effluent 113.
[0043] Again, referring to the embodiment depicted in Figure 1, the mixed
liquid
hydrocarbons 152 is next conveyed to a conventional naphtha stabilizer 160
that separates the
mixed liquid hydrocarbons 152 into an aromatics fraction 163 (predominantly
comprising
aromatics) and an unreacted C5/C6 components fraction 167 that predominantly
comprises
unreacted pentanes and larger non-aromatic C6+ components. The unreacted C5/C6
components
fraction 167 may be utilized directly as a gasoline blend component 169 or
optionally be
recycled via a C5/C6 components recycle conduit 171 and reintroduced
downstream from the
isomerization reactor 110.
[0044] Certain embodiments of the inventive processes and systems convey an
activation
effluent to an oligomerization reactor containing at least one oligomerization
catalyst. The
activation effluent contacts the oligomerization catalyst and is converted to
larger hydrocarbon
products that can be utilized as a component of a liquid transportation fuel,
such as, but not
limited to: gasoline, diesel and jet fuel.
[0045] A second embodiment of the inventive processes and systems that
includes an
oligomerization reactor and additional inventive features is illustrated by
the process flow-
diagram of Figure 2. A hydrocarbon feed stream 201 that comprises both n-
pentane and
isopentane is converted in a system 20. Typically, the hydrocarbon feed stream
201 comprises at
least 50 wt.% of pentane isomers, although in certain embodiments, the pentane
isomers may
comprise at least 60 wt.%, or at least 70 wt.% of the feed. Further, the
hydrocarbon feed stream
201 typically comprises less than 10 wt.% of hydrocarbons containing four or
fewer carbon
atoms.
[0046] The hydrocarbon feed stream 201 is received by an isomerization
reactor 210 that
contains an isomerization catalyst 215 and comprises a first reaction zone
(not depicted) that is
maintained at a temperature (measured at the isomerization reactor inlet) and
pressure that
facilitates the isomerization of at least a portion of the n-pentane in the
hydrocarbon feed stream
to isopentane. The isomerization reaction occurring in the first reaction zone
produces an
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isomerization effluent 213 that is characterized by an increased ratio of
isopentane to n-pentane
(relative to the corresponding ratio of the hydrocarbon feed stream 201).
Optionally,
isomerization reactor 210 may contain more than one isomerization catalyst or
may optionally
comprise more than one isomerization reactor arranged in series configuration
(not depicted).
[0047] The isomerization effluent 213 is next conveyed to an activation
reactor 220
containing a first activation catalyst 225 and comprising a second reaction
zone (not depicted).
The activation reactor 220 is operable to maintain a temperature and pressure
that is suitable to
facilitate conversion of the isomerization effluent 213 to an activation
effluent 228 that
comprises olefins containing from two to five carbon atoms, monocyclic
aromatics and
unconverted alkanes containing from two to five carbon atoms. In certain
embodiments, a
diluent 215 is added at any point that is upstream from, or optionally within,
the activation
reactor 220, but prior to contacting the activation catalyst 225. The diluent
may comprise any
substance that is less chemically-reactive than the constituents present in
the isomerization
effluent 213 at the conditions of temperature and pressure that are maintained
within the
activation reactor 220.
[0048] The activation effluent 228 leaves the activation reactor 220, and
is conveyed to
condenser 230, which may comprise one or more functions including a condenser,
splitter,
compressor and pump. Condenser 230 is operable to receive and condense at
least a portion of
the activation effluent 228 to produce a liquid hydrocarbons 231 comprising C6
and larger
hydrocarbons including paraffins, olefins and aromatics and a (gas phase)
light activation
effluent 232 comprising Cl-05 alkanes and olefins. The liquid hydrocarbons 231
are removed,
while the light activation effluent 232 is then compressed in compressor 233
located
immediately downstream from the condenser 230. Compressor 233 produces a
compressed
activation effluent 234 that is next conveyed to an oligomerization reactor
235 that contains an
oligomerization catalyst 240 and comprises a second reaction zone (not
depicted).
[0049] Speaking generally, the oligomerization catalyst may comprise any
solid catalyst (or
mixture of catalysts) characterized as possessing either Bronsted or Lewis
acidic properties. In
certain embodiments, the oligomerization catalyst is a zeolite or mixture of
zeolites, or a reactive
transition metal oxide. In certain embodiments, the oligomerization catalyst
is ZSM-5, although
many zeolites are well-characterized as possessing oligomerization properties
and may be
suitable for use (either alone or in combination) with the inventive processes
and systems
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described herein. Other well-characterized oligomerization catalysts include,
but are not limited
to: nickel oxides, aluminum alkyls, aluminum halides, perfluoroaryl boranes,
oligomeric methyl
aluminoxanes (including supported), perfluoroaryl boranes, fluoroarylanes,
trityl borate,
ammonium borate (and aluminate salts thereof), supported PhNMe2H+B(C6F5)4- and
borate
anions and superacidic solid Bronsted acids, among others.
[0050] Speaking generally, the oligomerization reactor is maintained at a
temperature and
pressure suitable to facilitate oligomerization of olefins present in the
gaseous activation
effluent, thereby producing larger hydrocarbons comprising at least six
carbons that are
preferably characterized by a boiling point that is in the boiling point range
of a liquid
transportation fuel (e.g., gasoline or diesel). The oligomerization reactor is
generally maintained
at a total pressure in a range from 14 psia to 800 psia, optionally in the
range from 50 psia to 300
psia. The oligomerization reactor is typically maintained at a temperature
(measured within the
oligomerization reactor inlet) in the range from 200 C to 420 C, optionally in
the range from
200 C to 350 C. Typically, flow thorough the oligomerization reactor is
maintained at a weight
hourly space velocity (WHSV) in the range from 0.5 hr-1 to 10 hr-1.
Optionally, the WHSV is in
the range from 0.5 hr-1- to 2.0 111-1. While higher overall throughput is
desirable, ideally the
chosen WHSV allows for conversion of at least 85% of hydrocarbons present in
the gaseous
activation effluent at the selected operating temperature and pressure.
[0051] The catalytic conversion occurring in the oligomerization reactor
produces an
oligomerization effluent that typically comprises an increased quantity of
hydrocarbon
molecules that are characterized by a boiling-point in the range of a liquid
transportation fuel
(e.g., gasoline and diesel). Preferably, the combination of isomerization,
activation and
oligomerization converts at least 30 wt% of the original feed stream to
hydrocarbon molecules
that are characterized by a boiling point that is in the range of gasoline.
[0052] Referring again to the embodiment depicted in Figure 2, the
oligomerization effluent
242 produced in the second reaction zone (not depicted) that is contained
within the
oligomerization reactor 235 is conveyed to a separator 245 that separates the
oligomerization
effluent 242 into two fractions: a light hydrocarbons fraction 248 comprising
C1-C4
hydrocarbons and hydrogen, and a heavy hydrocarbons fraction 252 comprising
hydrocarbons
containing at least five carbon atoms (C5+) that may be utilized directly as a
blend component of
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a liquid transportation fuel or an intermediate product that may be
additionally processed prior to
blending into a liquid transportation fuel.
[0053] In the embodiment depicted in Figure 2, the heavy hydrocarbons
fraction 252 is
conveyed to a second separator 260 that is optionally a naphtha stabilizer.
The second separator
260 is operable to remove an olefins fraction 267 comprising predominantly
alkanes and olefins
containing five to six carbon atoms from the condensed liquid hydrocarbons 252
in order to
decrease Reid vapor pressure and increase octane rating of the resulting
liquid hydrocarbon
product 263, which predominantly comprises hydrocarbon molecules that are
characterized by a
boiling-point in the range of a liquid transportation fuel, such as, but not
limited to, gasoline,
diesel and jet fuel. The olefins fraction 267 may be used directly as a blend
component 269 of a
liquid transportation fuel or is optionally mixed with hydrocarbon feed stream
201 at a point that
is downstream from the isomerization reactor 210. This recycling not only
increases the overall
yield of fuel-range products, but also serves as a route to indirectly recycle
any benzene present
in the olefins fraction 267 to the alkylation reactor 240, as any such benzene
would be relatively
unreactive in the isomerization reactor 210 and activation reactor 220.
Optionally, a portion of
the liquid hydrocarbons 231 derived from the condenser 230 may be combined
with the liquid
hydrocarbon product 263.
[0054] Speaking generally, in certain embodiments, the liquid hydrocarbon
product of the
process may be hydrotreated in a hydrotreating reactor containing a
hydrotreating catalyst in
order to reduce olefin and aromatic content in the liquid hydrocarbon product,
as well as to
remove nitrogen-containing and sulfur-containing compounds. The hydrotreating
reactor
contains at least one hydrotreating catalyst (such as, for example, NiMo,
CoMo, etc.) or a
precious metal catalyst (such as Pt/A1203, Pd/A1203, or Pd/C, etc) and is
maintained at a pressure
and temperature suitable for facilitating hydrotreating catalytic reactions.
Such processes are
conventional in nature and therefore will not be described in greater detail
here.
[0055] Again, referring to the embodiment depicted in Figure 2, light
hydrocarbons fraction
248 predominantly comprises hydrogen as well as C1-C4 hydrocarbons that
remained
unconverted in the oligomerization reactor 240. Light hydrocarbons fraction
248 leaves the
separator 245 and is optionally conveyed to a third separator 250 that
utilizes a conventional
separation technology (such as, but not limited to, pressure swing adsorption
technology,
membrane separation technology, etc.) to separate hydrogen from the light
hydrocarbons to
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produce a hydrogen stream 251 and a C1-C4 light paraffins stream 255 that may
be combusted
257 to provide at least a portion of the heat required for the process, or
recycled to a point that is
upstream from the activation catalyst 225 to serve as the diluent 215 that is
mixed with the
isomerization effluent 213.
[0056] Certain embodiments of the inventive processes and systems convey
the activation
effluent to an aromatic alkylation reactor containing at least one alkylation
catalyst. This
produces larger hydrocarbon products that can be utilized as either gasoline
or diesel
transportation fuel, or a component thereof
[0057] A third embodiment of the inventive processes and systems that
includes an
alkylation reactor and additional inventive features is illustrated by the
process flow-diagram of
Figure 3. A hydrocarbon feed stream 301 that comprises both n-pentane and
isopentane is
converted in a system 30. Typically, the hydrocarbon feed stream 301 comprises
at least 50 wt.%
of pentane isomers, although in certain embodiments, the pentane isomers may
comprise at least
60 wt.%, or at least 70 wt.% of the feed. Further, the hydrocarbon feed stream
301 typically
comprises less than 10 wt.% of hydrocarbons containing four or fewer carbon
atoms.
[0058] The hydrocarbon feed stream 301 is received by an isomerization
reactor 310 that
contains an isomerization catalyst 315 and comprises a first reaction zone
(not depicted) that is
maintained at a temperature (measured at the isomerization reactor inlet) and
pressure that
facilitates the isomerization of at least a portion of the n-pentane in the
hydrocarbon feed stream
to isopentane. The isomerization reaction occurring in the first reaction zone
produces an
isomerization effluent 313 that is characterized by an increased ratio of
isopentane to n-pentane
(relative to the corresponding ratio of the hydrocarbon feed stream 301).
Optionally,
isomerization reactor 310 may contain more than one isomerization catalyst or
may optionally
comprise more than one isomerization reactor arranged in series configuration
(not depicted).
[0059] The isomerization effluent 313 is next conveyed to an activation
reactor 320
containing a first activation catalyst 325 and comprising a second reaction
zone (not depicted).
The activation reactor 320 is operable to maintain a temperature and pressure
that is suitable to
facilitate conversion of the isomerization effluent 313 to an activation
effluent 328 that
predominantly comprises olefins containing from two to five carbon atoms,
monocyclic
aromatics and unconverted alkanes containing from two to five carbon atoms. In
certain
embodiments, a diluent 315 is added at any point that is upstream from, or
optionally within, the
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activation reactor 320, but prior to contacting the activation catalyst 325.
The diluent may
comprise any substance that is less chemically-reactive than the constituents
present in the
isomerization effluent 313 at the conditions of temperature and pressure that
are maintained
within the activation reactor 320.
[0060] The activation effluent 328 leaves the activation reactor 320, and
is conveyed to
condenser 330, which may comprise one or more functions including a condenser,
splitter,
compressor and pump. Condenser 330 is operable to receive and condense at
least a portion of
the activation effluent 328 to produce a liquid hydrocarbons 331 comprising C6
and larger
hydrocarbons including paraffins, olefins and aromatics, and a (gas-phase)
light activation
effluent 332 comprising Cl-05 alkanes and olefins. The liquid hydrocarbons 331
are removed,
while the light activation effluent 332 is then compressed in compressor 333
located
immediately downstream from the condenser 330. Compressor 333 produces a
compressed
activation effluent 334 that is next conveyed to an alkylation reactor 335
that contains an
alkylation catalyst 340 and comprises a second reaction zone (not depicted).
[0061] Speaking generally, the alkylation reactor is maintained at a feed
inlet temperature
and pressure suitable to facilitate the catalytic alkylation of aromatics
present in the mixed
effluent. The aromatics that are alkylated may be produced by aromatization
that takes place in
the activation reactor or may be a constituent of the hydrocarbon feed stream
301. These
aromatics are alkylated by olefins that are largely produced by the activation
of alkanes in the
activation reactor. Alkylation of aromatics in the alkylation reactor produces
an alkylation
effluent comprising larger hydrocarbons comprising at least seven carbons that
are preferably
characterized by a boiling point that is in the boiling point range of a
liquid transportation fuel
(e.g., gasoline or diesel). Typically, the alkylation effluent comprises an
increased percentage of
alkylated aromatic compounds comprising from seven to nine carbon atoms.
Optionally, the
larger hydrocarbons also are characterized by a lower Reid vapor pressure and
an increased
octane rating.
[0062] The alkylation reactor is generally maintained at a pressure in a
range from 14 psia
to 800 psia, optionally in the range from 50 psia to 600 psia. The alkylation
reactor is typically
maintained at a temperature (generally measured within the alkylation reactor
inlet) in a range
from 150 C to 350 C, optionally between 200 C to 350 C. Typically, flow
thorough the
alkylation reactor is maintained at a weighted hourly space velocity (WHSV) in
the range from
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0.5 hr-1 to 10 hr-1 on an olefin basis. Optionally, the WHSV is in the range
from 0.5 hr-1 to 2.0 hr-
1. While higher overall throughput is desirable, ideally the chosen WHSV
allows for conversion
of at least 85% olefinic of hydrocarbons present in the mixed effluent at the
selected operating
temperature and pressure. The catalytic conversion occurring in the alkylation
reactor produces
an aromatic alkylation reactor effluent that typically comprises at least 30
wt.% (preferably, at
least 40 wt%) of hydrocarbon molecules that are characterized by a boiling-
point in the range of
a liquid transportation fuel.
[0063] Speaking generally, the alkylation catalyst may comprise any
catalyst characterized
as either Bronsted or Lewis acidic. A wide variety of catalysts have been
found to promote
aromatic alkylation including, but not limited to, aluminum chloride,
phosphoric acid, sulfuric
acid, hydrofluoric acid, silica, alumina, sulfated zirconia, zeolites
(including, for example, ZSM-
5, ZSM-3, ZSM-4, ZSM-18, ZSM-20, zeolite-beta, H-Y, MCM-22, MCM-36 and MCM-
49). In
certain embodiments, the alkylation catalyst simultaneously promotes
alkylation of aromatics
and oligomerization of olefins present in the mixed effluent.
[0064] Referring again to the embodiment depicted in Figure 3, the
alkylation effluent 342
is conveyed to a separator 345 that separates the alkylation effluent 342 into
two fractions: a
light hydrocarbons fraction 348 comprising C1-C4 hydrocarbons and Hz, and a
condensed liquid
hydrocarbons 352 comprising hydrocarbons containing at least five carbon atoms
(C5+) that
may be utilized directly as a blend component of a liquid transportation fuel
or additionally
processed prior to blending into a liquid transportation fuel. Preferably, the
alkylation effluent
comprises an increased quantity (or increased wt%) of alkylated aromatics
containing from
seven to nine carbon atoms. Preferably, these alkylated aromatics are
monocyclic aromatic
hydrocarbons.
[0065] In the embodiment depicted in Figure 3, the condensed liquid
hydrocarbons 352 is
conveyed to a second separator 360 that is optionally a naphtha stabilizer.
The second separator
360 is operable to remove a olefins fraction 367 (comprising predominantly
alkanes and olefins
containing four to six carbon atoms) from the condensed liquid hydrocarbons
352 in order to
decrease Reid vapor pressure and increase octane rating of the resulting
liquid hydrocarbon
product 363 that predominantly comprises hydrocarbons that are characterized
by a boiling-point
in the range of a liquid transportation fuel, such as, but not limited to,
gasoline, diesel and jet
fuel. The olefins fraction 367 may be used directly as a blend component 369
of a liquid
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transportation fuel or is optionally mixed with hydrocarbon feed stream 302 at
a point that is
downstream from the isomerization reactor 310. This recycling also serves as a
route to
indirectly recycle any benzene present in the olefins fraction to the
alkylation reactor 340, as any
such benzene would be relatively unreactive in the isomerization reactor 310
or activation
reactor 320. Optionally, a portion of the mixed liquid hydrocarbons 331
derived from the
condenser 330 may be combined with the liquid hydrocarbon product 363.
[0066] Speaking more generally, in certain embodiments the liquid
hydrocarbon product
may be hydrotreated in a hydrotreating reactor containing a hydrotreating
catalyst in order to
reduce olefin and aromatic content in the liquid hydrocarbon product, as well
as to remove
nitrogen-containing and sulfur-containing compounds. The hydrotreating reactor
contains at
least one hydrotreating catalyst (such as, for example, NiMo, CoMo, etc.) or a
precious metal
catalyst (such as Pt/A1203, Pd/A1203, or Pd/C, etc) and is maintained at a
pressure and
temperature suitable for facilitating hydrotreating catalytic reactions. Such
processes are
conventional in nature and therefore will not be described in greater detail
here.
[0067] Again, referring to the embodiment depicted in Figure 3, light
hydrocarbons fraction
348 predominantly comprises hydrogen as well as Cl - C4 hydrocarbons that were
not converted
in the alkylation reactor 340. Light hydrocarbons fraction 348 leaves the
separator 345 and is
conveyed to a third separator 350 that utilizes a conventional separation
technology (such as, but
not limited to, pressure swing adsorption technology, membrane separation
technology, etc.) to
separate hydrogen from light hydrocarbons to produce a hydrogen stream and a
C1-C4 light
paraffins stream 355 that may be combusted 357 (not depicted) to provide at
least a portion of
the heat required for the process, or recycled to a point that is upstream
from the activation
catalyst 325 to serve as the diluent 215 that is mixed with the isomerization
effluent 313
[0068] Certain embodiments comprise mixing a diluent with the isomerization
effluent prior
to contacting the resulting mixture with an activation catalyst. The diluent
may be added in a
ratio ranging from 10:1 to 1:10 molar ratio relative to the quantity of
isomerization effluent fed
to the activation reactor. The diluent may be added at any point that is
upstream from, or within,
the activation reactor, but prior to contacting the activation catalyst.
[0069] The diluent may comprise any substance that is less chemically-
reactive than the
constituents present in the isomerization effluent at the conditions of
temperature and pressure
that are maintained within the activation reactor. This is intended to prevent
the diluent from
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reacting with the activation catalyst. Such properties are found in a large
number of substances
that are fully within the grasp of a person who is knowledgeable in the field.
In certain
embodiments, the diluent may comprise a C1-C4 light paraffins, including
recycling C1-C4 light
paraffins produced by the processes and systems described herein. In certain
embodiments, the
diluent may comprise any of methane, ethane, propane, butanes, benzene,
toluene, xylenes,
alkyl- or dialkyl-benzenes, naphthenes, C2-05 olefins, and combinations
thereof.
[0070] The presence of diluent during catalytic activation (i.e.,
activation) provides
numerous advantages. First, it effectively decreases the concentration of the
isomerization
effluent within the activation reactor. This results in a small increase in
the total conversion of
alkanes to olefins or aromatics within the activation reactor. However, it
also increases the
selectivity toward the production of olefins, while slightly decreasing the
selectivity toward the
production of aromatics. Adjusting the ratio of diluent to isomerization
effluent changes the ratio
of olefins to aromatics in the resulting activation effluent, thereby
providing a valuable point of
operational control for downstream processes. Typically, the optimal molar
production ratio of
olefins to aromatics ranges from about 0.5:1 to about 1.5:1, in order to
maximize the value
captured in the olefin intermediates during the alkylation in the alkylation
reactor. Mono-
alkylated aromatics exhibit beneficial (increased) octane rating and vapor
pressure for
application as blending components in certain transportation fuels such as
gasoline. In contrast,
di-alkyl and tri-alkyl aromatics comprising more than nine carbon atoms are
not well-suited for
blending into gasoline, and exhibit nonoptimal cetane number for blending into
diesel.
[0071] Addition of a diluent also advantageously favors the production of
value-added
olefins relative to C1-C4 light paraffins, and also mitigates dimerization of
C5 hydrocarbons to
form durene (1,2,4,5-tetramethylbenzene), a byproduct notorious for
precipitating as a solid
from of gasoline blends.
EXAMPLES
[0072] The following examples are representative of certain embodiments of
the inventive
processes and systems disclosed herein. However, the scope of the invention is
not intended to
be limited to the embodiment specifically disclosed. Rather, the scope is
intended to be as broad
as is supported by the complete disclosure and the appending claims.
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EXAMPLE 1:
[0073] This example demonstrates the preliminary rationale for isomerizing
n-05 to i-05 in
a hydrocarbon feed stream prior to contacting an activation catalyst. The
graphs below illustrate
differences in activation reactivity for n-05 and i-05. Feed streams were
utilized that comprised
either 100 wt.% i-05 (i-05) or 100 wt.% of n-05 (n-05). The catalyst was 1/8"
extrudate
consisting of 50 wt.% alumina binder and 50 wt.% ZSM-5 zeolite, and
experiments were
conducted at a WHSV of 1.3 hr' at 1 atm. Results were averaged over the total
time on stream of
16 hr.
[0074] Figure 4 is a bar graph depicting the results of catalytically
activating each fraction at
either 550 C or 600 C. The graph depicts, as percentages, the total catalytic
conversion of each
feed stream (first column), the selectivity to light olefins as product
(second column), the
selectivity to aromatics as product (third column) and the selectivity to C1-
C4 light paraffins
(defined as non-olefin hydrocarbons containing from one to four carbon atoms),
fourth column.
Selectivity was calculated on a % carbon basis, relative to the portion of the
feed stream fraction
that was converted.
[0075] The results demonstrate that total conversion of a 100% n-05
fraction at 600 C was
79%, and a similar 82% conversion was observed when activating a 100% i-05
feed stream at a
50 C cooler temperature (i.e., 550 C). Activating the i-05 fraction at 550 C
(instead of 600 C)
also increased the selectivity towards the production of olefins while
decreasing the selectivity
of conversion toward aromatics. Lastly, these changes in selectivity caused no
significant
increase in the production of byproduct C1-C4 light paraffins. However,
activation of n-05 at
550 C was generally unsuitable, and resulted in a 31% decrease in total
conversion, and a
noticeable increase in the production of C1-C4 light paraffins.
EXAMPLE 2:
[0076] This experiment demonstrates that isopentane (i-05) is
advantageously converted by
both an activation catalyst and a subsequent oligomerization catalyst to
produce a high
percentage of product that is suitable for use as a liquid transportation
fuel. A 100 wt.% i-05
feed stream was upgraded by first contacting it with a zeolite activation
catalyst, followed by
contacting a zeolite oligomerization catalyst. Activation was conducted by
contacting the feed
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stream with 1/8 in. diameter catalyst extrudate consisting of 50 wt.% alumina
binder and 50
wt.% ZSM-5 zeolite catalyst at a temperature of 579 C, and a WHSV of 2.6 hr-1
at 1 atm.
[0077] Oligomerization was conducted by contacting the activation effluent
with a ZSM-5
zeolite catalyst in a reactor where the inlet temperature for the activation
effluent was maintained
at 250 C, the pressure was 1 atm, and the WHSV for the feed stream was 1.3 hr-
1. Results were
time-averaged over 16 hours. The table shows the product distribution
following conversion
along with the selectivity to olefins and liquid product. The term
"selectivity" indicates the
percentage of the catalytically converted feed stream that was converted to a
particular product.
Table 2. Upgrading pentanes by activation alone or activation plus
oligomerization.
Activation Activation +
Oligomerization
Total Conversion (wt.%) 88 87
C1-C4 Light paraffins Yield 32 32
Upgraded Product Yield (wt.%) 55 54
Total Coke Yield (wt.%) 0.1 0.1
Light Olefin Yield (wt.%) 42 16
Light Olefin Selectivity (wt.%) 48 19
Liq. Yield (wt.%) 13 38
Liq. Product Selectivity (wt.%) 15 44
[0078] The data in Table 2 show that the subjecting the effluent from the
first activation
reactor to a subsequent oligomerization step in a second reactor increased the
liquid product
yield from 13 wt.% to 38 wt.%. This liquid product yield represents a liquid
product suitable for
blending into a liquid transportation fuel such as gasoline (up from 13 wt.%
prior to
oligomerization), and that selectivity to liquid product for the portion of
the feed stream that was
converted was 44 wt.%. Undesirable C1-C4 light paraffins production was
limited to 32 wt.% of
the original feed stream, which optionally may be recycled to be either
activated or to serve as a
diluent in at least one of the activation reactors. Further, the final product
only comprised 16
wt.% of light olefins, (primarily ethylene), which may be recycled to the
process, or diverted to
be utilized in any of a variety of conventional processes.
EXAMPLE 3:
[0079] This experiment demonstrates that isopentane (i-05) is
advantageously converted by
both an activation catalyst and a subsequent alkylation catalyst to produce a
high percentage of
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product that is suitable for use as a liquid transportation fuel. A 100 wt.% i-
05 feed stream was
upgraded by first contacting it with a zeolite activation catalyst, followed
by contacting the
activation effluent with a zeolite alkylation catalyst. Activation was
conducted by contacting the
feed stream with a 1/8 in. diameter catalyst extrudate consisting of 50 wt.%
alumina binder and
50 wt.% ZSM-5 zeolite catalyst in an activation reactor. The temperature of
the activation
reactor at the inlet for the feed stream was 579 C, the pressure was 1 atm,
and the WHSV for the
feed stream was 2.6 hr'. Alkylation was then conducted by contacting the
effluent with a ZSM-5
catalyst in a reactor where the temperature at the inlet for the feed stream
was 230 C and the
WHSV of the feed stream was 1.3 hr' at 1 atm. Results were time-averaged over
16 hours. The
table shows the product distribution following conversion along with the
selectivity to olefins
and liquid product. The term "selectivity" indicates the percentage of the
catalytically converted
feed stream that was converted to a particular product.
Table 3. Upgrading isopentane by activation only or activation followed by
alkylation.
Activation Activation +
Alkylation
Total Conversion (wt.%) 87 87
Light paraffins Yield (wt.%) 32 32
Upgraded Product Yield (wt.%) 55 55
Total Coke Yield (wt.%) 0.1 0.2
Light Olefin Yield (wt.%) 42 12
Light Olefin Selectivity (wt.%) 48 14
Liquid Yield (wt.%) 13 42
Liquid Product Selectivity (wt.%) 15 48
[0080] The data in Table 3 show that subjecting the activation effluent to
a subsequent
alkylation step increased the liquid product yield from 13 wt.% to 42 wt.%.
This liquid product
is suitable for blending into a liquid transportation fuel such as gasoline,
and possesses an
increased research octane number, a suitable distillation T50 and endpoint,
and low vapor
pressure. Selectivity to liquid product for the portion of the feed stream
that was converted
increased from 15 wt.% to 48 wt.%. Undesirable C1-C4 light paraffins
production was limited to
32 wt.% of the original feed stream. Further, the final product only comprised
14 wt.% of light
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olefins. These olefins may be recycled to the activation reactor, used as a
diluent in the
alkylation reactor, or diverted to be utilized in any of a variety of
conventional processes.
[0081]
Note that the results shown in the above table may underestimate the total
percentage
of a mixed pentanes feed stream that would be available for blending into a
liquid transportation
fuel, as a typical hydrocarbon feed stream (such as, but not limited to,
natural gasoline) may also
include an excess quantity of C5/C6+ that would not be either catalytically
cracked or
introduced into the alkylation reactor. This excess quantity of C5/C6+ is
suitable for direct
blending into the liquid hydrocarbon product. In certain embodiments, a
portion of the nC5/C6+
fraction is diverted when necessary to achieve the desired 0.5:1 to 1.5:1
olefin to aromatic ratio
that maximizes production of mono-alkylated aromatics in the alkylation
reactor.
EXAMPLE 4:
[0082]
Preliminary experimentation revealed that contacting a simulated natural
gasoline
feed stream (containing approximately 1:1 ratio of n-05:i-05) with an
isomerization catalyst in a
single pass increases the ratio of i-05 to n-05 to approximately 7:3. An
experiment was next
performed to assess potential differences in conversion yield and selectivity
to various products
when a feed stream comprising a simulated isomerization effluent (7:3 ratio of
i-05 to n-05) was
contacted with an activation catalyst comprising an 1/8" extrudate consisting
of 50 wt.% alumina
binder and 50 wt.% ZSM-5 zeolite. The reaction was conducted at 600 C, with a
flow rate of
5.0 hr', for a total of 16.5 hr. on stream, and produced an effluent
comprising light olefins,
aromatics and light paraffins. The averaged results are shown in Table 4,
below:
Table 4: ZSM-5 activation and conversion of several feed streams comprising
different amounts
of n-pentane (n-05) and isopentane (i-05) isomers.
100 wt.% 100 wt.% n- 70 wt.% i-05 50 wt.% i-05
iC5 C5 30 wt.% n- 50 wt% n-05
C5
Material Balance 93% 104% 84%
101%
Conversion 95% 79% 88%
84%
Fuel gas yield 20% 27% 21%
24%
Product Yield 74% 51% 67%
59%
Coke Yield 0% 0% 0%
0%
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Lt Olefin Yield 50% 38% 45%
40%
Lt Olefin 53% 49% 51%
47%
Selectivity
Aromatic Yield 24% 13% 22%
20%
Aromatic 26% 16% 25%
23%
Selectivity
Fuel Gas Yield 20% 27% 21%
24%
Fuel Gas Selectivity 21% 35% 24%
29%
[0083]
The results clearly indicate that increasing the percentage of i-05 in the
feed resulted
in a significant increase in both olefin yield (+5%) and selectivity (+4%),
and a lesser increase in
both aromatic yield (+2%) and selectivity (+2%). Simultaneously, selectivity
to fuel gas was
advantageously decreased by 5%.
EXAMPLE 5:
[0084]
This experiment demonstrates the effect that a methane diluent has on
catalytic
activation and conversion of two different hydrocarbon feed streams: 1) a
simulated "natural
gasoline" comprising 50 wt.% i-05 and 50 wt.% n-05 isomers, and 2) a simulated
"isomerization effluent" comprising 70 wt.% i-05 and 30 wt.% n-05. Each of the
two feed
streams were fed at a WHSV of 1.3 hr' to a reactor containing an activation
catalyst comprising
a 1/8" extrudate consisting of 50 wt.% alumina binder and 50 wt.% ZSM-5
zeolite. The
temperature of the reactor (at the inlet for the feed stream) was maintained
at 600 C and 20 psig
(2.4 Bar) and results were time-averaged for 16.5 hr. For certain reactions,
methane diluent was
co-fed along with each feed stream at a methane: feed stream molar ratio of
2:1.
[0085]
The reaction produced an effluent comprising light olefins, aromatics and
light
paraffins. Table 5 (below) shows the effect of the methane diluent on the
total conversion of the
1:1 and 7:3 feed streams, respectively, as well as the selectivity of each
conversion toward light
olefins, aromatics, and byproduct C1-C4 fuel gas.
Table 5: Catalytic activation of a 1:1 1-05:n-05 feed stream and a 7:3 1 1-
05:n-05 feed stream
in both the absence and presence of methane diluent.
Feed Stream 1:1 i-05:n-05 1:1 i-05:n-05 7:3 1 i-05:n-05 7:3 1 i-05:n-
05
+/- Diluent No Diluent CH4 Diluent No Diluent CH4 Diluent
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PCT/US2019/054832
Material Balance 101% 103% 102% 101%
Conversion 92% 80% 93% 81%
Fuel gas yield 37% 22% 34% 21%
Product Yield 54% 58% 58% 60%
Coke Yield 0% 0% 0% 0%
Lt. Olefin Yield 34% 44% 35% 43%
Lt. Olefin 37% 55% 38% 53%
Selectivity
Aromatic Yield 20% 14% 22% 17%
Aromatic 21% 17% 24% 21%
Selectivity
Fuel Gas Yield 37% 22% 34% 21%
Fuel Gas 41% 27% 37% 26%
Selectivity
[0086] The data in Table 5 indicate that adding inert diluent caused slight
loss of overall
conversion, but significantly increased the yield and selectivity to light
olefin production for
both the 1:1 and 7:3 feed streams. Adding inert diluent also greatly
diminished selectivity to
production of C1-C4 fuel gas. Meanwhile, only a small drop in selectivity to
aromatics
production was observed for the 1:1 ratio feed stream in the presence of
diluent, which was
offset by an equivalent increase in aromatics production in the 7:3 ratio feed
stream (in the
presence of diluent). All of these results are advantageous to the process,
particularly in certain
embodiments where the mixed effluent is immediately utilized as feed stream
for either an
oligomerization or alkylation process. In certain embodiments that comprise an
oligomerization
process, diluent is added to the activation feed stream at a ratio that
maximizes light olefin
production, providing an advantageous feed stream for the oligomerization
catalyst. In certain
embodiments that comprise an aromatic alkylation process, diluent can be added
to the
activation feed stream at a ratio that produces a first effluent comprising
olefins and aromatics at
a ratio (typically between 0.5:1 and 1.5:1 by mole) that provides an
advantageous feed stream
for an aromatic alkylation process.
[0087] In closing, it should be noted that the discussion of any reference
is not an admission
that it is prior art to the present disclosure, in particular, any reference
that may have a
publication date after the priority date of this application. Although the
systems and processes
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described herein have been described in detail, it is understood that various
changes, substitutions,
and alterations can be made without departing from the spirit and scope of the
invention as defined
by the following claims.
Definitions:
[0088] In the present disclosure, the term "conversion" is defined as any
of the chemical
reactions that occur during upgrading of hydrocarbons to liquid transportation
fuels. Examples
of such reactions include, but are not limited to: oligomerization,
aromatization,
dehydrogenation, alkylation, hydrogenation and cracking.
27
