Note: Descriptions are shown in the official language in which they were submitted.
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SPLIT-FEED TWO-STAGE PARALLEL AROMATIZATION
FOR MAXIMUM PARA-XYLENE YIELD
The present invention relates to a process for reforming a full-boiling
range hydrocarbon feed to enhance para-xylene and benzene production:-
BACKGROUND OF THE INVENTION
The reforming of petroleum hydrocarbon streams is an important
petroleum refining process that is employed to provide high octane
hydrocarbon blending components for gasoline. The process is usually
practiced on a straight run naphtha fraction that has been hydrodesulfurized.
Straight run naphtha is typically highly paraffinic in nature, but may contain
7 5 significant amounts of naphthenes and minor amounts of aromatics or
olefins.
In a typical reforming process, the reactions include dehydrogenation,
isomerization, and hydrocracking. The dehydrogenation reactions typically
will be the dehydroisomerization of alkylcyclopentanes to aromatics, the
dehydrogenation of paraffins to olefins, the dehydrogenation of cyclohexanes
to aromatics, and the dehydrocyclization of paraffins to aromatics. The
aromatization of the n-paraffins to aromatics is generally considered to be
the
most important because of the high octane of the resulting aromatic product
compared to the low octane ratings for n-paraffins. The isomerization
reactions include isomerization of n-paraffins to isoparaffins, and the
isomerization of substituted aromatics. The hydrocracking reactions include
the hydrocracking of paraffins and hydrodesulfurization of any sulfur that is
remaining in the feedstock.
It is well known in the art that several catalysts are capable of
reforming petroleum naphthas and hydrocarbons that boil in the gasoline
boiling range. Examples of known catalysts useful for reforming include
platinum and optionally rhenium or iridium on an alumina support, platinum on
zeolite X and zeolite Y, platinum on intermediate pore size zeoiites as
described in U.S. Patent No. 4,347,394, and platinum on cation exchanged
zeolite L. U.S. Patent No. 4,104,320 discloses the dehydrocyclization of
aliphatic hydrocarbon to aromatics by contact with a catalyst comprising a
zeolite L containing alkali metal ions and a Group VIII metal such as
platinum.
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The conventional reforming catalyst is a bifunctional catalyst that
contains a metal hydrogenation-dehydrogenation component, which is usually
dispersed on the surface of a porous inorganic oxide support, usually
alumina. Platinum has been widely used commercially in the production of
reforming catalysts, and platinum on atumina catalysts have been
commercially employed in refineries for the past few decades. More recently,
additional metallic components have been added to the platinum to further
promote the activity or selectivity, or both. Examples of such metallic
components are iridium, rhenium, tin and the like. Some catalysts possess
superior activity, or selectivity, or both as contrasted with other catalysts.
Platinum-rhenium catalysts, for example, possess high selectivity in
comparison to platinum catalysts. Selectivity is generally defined as the
ability of the catalyst to produce high yields of desirable products with
concurrent low production of undesirable products, such as gaseous
hydrocarbons.
It is desirable to maximize xytene and benzene production and
ultimately para-xylene and benzene production. The problem of how to do
this has not been previously solved. The prior art has dealt with the problem
of maximizing only benzene production when processing a wide boiling CS-C"
naphtha but has not addressed how to maximize first para-xylene production
and secondly benzene production. Note that maximizing benzene production
should not occur by downgrading Ce and C9 aromatics to benzene. This is
especially important as para-xylene has historically commanded a premium
above benzene.
There exist several processes for dividing naphtha feedstreams into a
higher boiling cut and a lower boiling cut and reforming these cuts
separately.
U.S. Patent No. 2,867,576 discloses separating straight run naphtha into
tower and higher boiling cuts, in which the higher boiling cuts are reformed
with a hydrogenation-dehydrogenation catalyst with the liquid reformate
produced being routed to an aromatics separation process. The paraffinic
fraction ~btaine from the separation process is blevi:. r ~ with the lower
boiling
naphth ~ractieand the resulting blend is reformed mth a reforming catalyst,
which may or may not be the same type employed in reforming the high
boiling cut.
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U.S. Patent No. 2,944,959 discloses fractionating a full straight run
gasoline into a light paraffinic fraction, CS and Cs, that is hydroisomerized
with
hydrogen and a platinum-alumina catalyst, a middle fraction that is
catalytically reformed with hydrogen and a platinum-alumina catalyst, and a
heavy fraction that is catalytically reformed with a molybdenum oxide catalyst
and recovering the liquid products. U.S. Patent Nos. 3,003,949, 3,018,244
and 3,776,949 also disclose fractionating a feed into a C5 and Cs fraction,
that
is isomerized, and a heavier fraction that is reformed.
Other processes for dividing feedstocks and separately treating them
include: U.S. Patent Nos. 3,172,841 and 3,409,540 disclose separating
fraction of a hydrocarbon feedstock and catalytically reforming various
fractions of the feed; U.S. Patent No. 4,167,472 discloses separating straight
chain from non-straight chain C6-C,a hydrocarbons and separately converting
to aromatics; and U.S. Patent No. 4,358,364 discloses catalyticaily reforming
a C6 fraction and producing additional benzene by hydrogasifying a
C5_ fraction, a fraction with a boiling point above 300°F and the gas
stream
produced from catalytic reforming.
U.S. Patent No. 3,753,891 discloses fractionating a straight run
naphtha into a tight naphtha fraction containing the Cs and a substantial
portion of the C, hydrocarbons and a heavy naphtha fraction boiling from
about 200° to 400°F; then reforming the light fraction to
convert naphthenes to
aromatics over a platinum-aiumina catalyst or a bimetallic reforming catalyst;
separately reforming the heavy faction, then upgrading the reformer effluent
of the low boiling fraction over a ZSM-5 type zeolite catalyst to crack the
paraffins and recovering an effluent with improved octane rating.
U.S. Patent No. 4,645,586 discloses parallel reforming of a
hydrocarbons feed. In one stream, the hydrocarbons are reformed with an
acidic catalyst. In the second stream, the hydrocarbons are reformed with a
non-acidic catalyst. That patent is silent as to the composition of each
fraction. Preferably, the acidic bi-functional reforming catalyst is not
presulfided.
U.S. Patent No. 4,897,177 discloses using a monofunctional catalyst to
reform a hydrocarbon fraction having less than 10% by volume of
C9+ hydrocarbons. This fraction is either a C6, C,, C8, C6-C,, C,-C8, or Cs-Ca
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fraction, with the most preferred being a C6-C8 fraction. That fraction can
contain up to 15 vol.% hydrocarbons outside the named range (col. 3,
line 44-49). A heavier fraction can be reformed using a bifunctional catalyst
on an acidic metal oxide. That bifunctional catalyst can be a PdSn/alumina
catalyst.
U.S. Reissue Patent No. 33,323 discloses solvent extraction of a light
fraction of a reformats. The goal of that patent is to maximize benzene
production only. A hydrocarbon feed is separated into a lighter fraction (a C6
cut that contains 15-35 Iv% C,+) and a heavier fraction (all remaining C, and
heavier components). The lighter fraction is reformed in the presence of a
non-acidic catalyst to maximize benzene yield. The heavier fraction is
reformed in the presence of an acidic catalyst. The reformats from the non-
acidic catalyst is introduced into an extraction where an aromatic extract
stream and a non-aromatic raffinate stream are recovered. The raffinate
stream can be recycled to the feed.
The paper entitled "New Options For Aromatics Production" presented
to the 20th Annual 1995 Dewitt Petrochemical Review (Houston, Texas,
March 21-23, 1995) by J. D. Swift et al. related recent improvements in UOP's
process for the production of benzene and para-xylene. Case studies were
presented to demonstrate the benefits of using that process to increase total
aromatics production from a fixed quantity of naphtha. One configuration of
that process involved a split-feed process, but it is unclear what the
composition of each feed was.
SUMMARY OF THE INVENTION
The present invention provides a process for reforming a full boiling
hydrocarbon feed to enhance para-xylene and benzene yields.
This invention is based upon the realization that a non-acidic catalyst
has an adverse effec~ c>n production of para-xylenes. It is th a;sght that the
catalyst actually dealkyiates those xylenes. Thus the C8+ fraction should not
be subjected to a non-acidic catalyst if one is trying to recover xylenes.
In that process, the hydrocarbon feed is separated into a C5. cut, a
Cs-C~ cut, and a C8+ cut, wherein the C6-C~ cut has less than 5 Iv. % of
U.S. Patent No. 4,897,177
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C8+ hydrocarbon, and wherein the Cue. cut has less than 10 Iv. % of
C~_ hydrocarbon. The C6-C~ cut is subjected to catalytic aromatization at
elevated temperatures in a first reformer in the presence of hydrogen and
using a non-acidic catalyst comprising at least one Group VIII metal and a
non-acidic zeolite support, preferably platinum on a non-acidic zeolite L
support, to produce a first reformats stream. The C8+ cut is subjected to J
catalytic aromatization at elevated temperatures in a second reformer in the
presence of hydrogen and using an acidic catalyst comprising at least one
Group VIII metal and a metallic oxide support, preferably a non-presulfided
acidic catalyst comprising platinum and tin on an alumina support, to produce
a second reformats stream. Less than 20 wt. % of the total amount of
C8 aromatics produced in the first and second reformer is ethylbenzene, and
more than 20 wt. % of the total amount of xylenes produced in the first and
second reformer are para-xylenes.
Preferably, the first reformats stream and the second reformats stream
are combined to form a combined reformats stream, the combined reformats
stream is separated into a light fraction and a heavy fraction, and at least
part
of the light fraction is recycled either to the hydrocarbon feed or to at
least one
of the reformers.
From our experimental studies where we have investigated the
aromatization of a wide-boiling range naphtha over a nonacidic zeolite such
as Pt/K-Ba L zeolite or Pt/K L zeolite with F and CI, we have found that these
non-acidic catalysts are more efficient than the standard bi-functional
catalysts at aromatizing Cs s and C,'s to the corresponding aromatic.
However, we have also found that the standard reforming bi-functional
catalysts such as PdSn/CI on alumina are more efficient than the non-acidic
zeolites at aromatizing CB s and C9's to the corresponding aromatic.
For example, at Ce paraffin and napththene (P+N) conversions of
92.9%, the selectivity to CB aromatics is about 50% with the non acidic
zeolite
when processing a C6-C,o paraffinic naphtha. When the same naphtha is
processed over a bi-functional aromatization catalyst such as Pt/Sn/CI on
alumina the selectivity to Ce aromatics is about 80% at CB (P+N) conversions
of 90+%. The lower C8 aromatics yield with the non-acidic zeolite is due to
hydro-dealkylation of the Cg aromatics to benzene and toluene.
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Furthermore, when the Cs-C,o naphtha is processed over a non-acidic
zeolite, not only is the yield of Ce aromatics lower, 19 wt% versus 24 wt%
with
a bi-functional catalyst, but also the Ce aromatics stream is of a poorer
quality.
The Ce aromatics stream made with the non-acidic zeolite contains 30%
ethylbenzene compared to about 16% produced with the bi-functional
catalyst. Thus the xylene yield is lower, 13 wt% versus 20 wt% with the
bi-functional catalyst. In other words, the bi-functional catalyst makes 50%
more xylenes.
fn addition, with the non-acidic zeolite, the para-xylene concentration
on a xylene basis is low, 12% compared to 20% with the bi-functional catalyst.
This latter value is very close to the equilibrium value of 23% at the
operating
temperature of the aromatization stage.
Thus from a Ce aromatization standpoint, the bi-functional catalyst, has
a higher CB aromatics yield, a higher xylene yield, and a lower yield of
ethylbenzene than the non-acidic zeolite. Also, the bifunctional catalyst
makes a xylene stream with a higher concentration of para-xylene than the
non-acidic zeolite. All of these are advantages to the para-xyiene producer as
they minimize capital and operating cost.
A further benefit of the bi-functional catalysts is that the conversion and
selectivity of C9 paraffins and naphthenes to the C9 aromatics is much higher.
Thus the overall C9 aromatics yield is about 10 wt% compared to about 4.0
wt% with the non-acidic zeolite. In addition, the C9 aromatics produced with
the bi-functional catalyst contain about 55% trimethylbenzenes and about 35
methyl-ethylbenzenes. This compares to about 20 % trimethylbenzenes
and about 46% methyl-ethylbenzenes with the non-acidic zeolite. The
C8 aromatics are converted to xylenes and benzene by transalkylation with
toluene. In this process, the trimethylbenzenes are the preferred species, as
they yield two moles of xylenes per mole of trimethylbenzenes and toluene,
whereas methyl-ethylbenzenes can yield one mole of xylenes and
ethylbenzenes, which is undesirable, or alternatively one mole of benzene
and a C,o aromatic. So not only does the bi-functional catalyst make more
C9 aromatics, but they are of a better quality from a xylenes and ultimately
para-xylene production standpoint.
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BRIEF DESCRIPTION OF THE DRAWINGS
In order to assist the understanding of this invention, reference will now
be made to the appended drawings. The drawings are exemplary only, and
should not be construed as limiting the invention.
Figure 1 shows a flow diagram of one embodiment of the present
invention.
pETAILED DESCRIPTION OF THE INVENTIC,~N
In its broadest aspect, the present invention involves a process for
reforming a full boiling hydrocarbon feed to enhance para-xylene and
benzene yields.
In that process the hydrocarbon feed is separated into a C5_ cut, a
Cs-C~ cut, and a C8+ cut. The C6-C~ cut has less than 5 Iv. % of CB.~
hydrocarbon, and the C8+ cut has less than 10 Iv. % of C~_ hydrocarbon.
The C6-C7 cut is subjected to catalytic aromatization at elevated
temperatures in a first reformer in the presence of hydrogen and using a
non-acidic catalyst comprising at least one Group VIII metal and a non-acidic
zeolite support to produce a first reformate stream.
The CB.,. cut is subjected to catalytic aromatization at elevated
temperatures in a second reformer in the presence of hydrogen and using an
acidic catalyst comprising at least one Group VIII metal and a metallic oxide
support to produce a second reformate stream.
Less than 20 wt. % of the total amount of C8 aromatics produced in the
first and second reformer is ethylbenzene, and more than 20 wt. % of the total
amount of xylenes produced in the first and second reformers are
para-xylenes.
To minimize capital investment and maximize aromatics yield, both
reformers operate at a common operating pressure that allows linking of the
two reformers and where possible common equipment can be used such as
recycle gas compressor, net gas booster compressor, separator and
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depentanizer. Thus essentially we have one aromatization plant. This
processing scheme solves the problem of how to maximize benzene and
particularly para-xylene production at low capital cost.
NON-ACIDIC CATALYSTS
One of the catalysts used must be a non-acidic catalyst having a non-
acidic zeolite support charged with one or more dehydrogenating
constituents. Among the zeolites useful in the practice of the present
invention are zeolite L, zeolite X, and zeolite Y. These zeolites have
apparent
pore sizes on the order of 7 to 9 Angstroms.
Zeolite L is a synthetic crystalline zeofitic molecular sieve which may
be written as:
(0.9-1.3)Mv~O:A1203(5.2-6.9)Si02:yti20
wherein M designates a cation, n represents the valence of M, and y may be
any value from 0 to about 9. Zeolite L, its X-ray diffraction pattern, its
properties, and method for its preparation are described in detail in U.S.
Pat.
No. 3,216,789. U.S. Pat. No. 3,216,789 is hereby incorporated by reference
to show the preferred zeolite of the present invention. The real formula may
vary without changing the crystalline structure; for example, the mole ratio
of
silicon to aluminum (Si/A1) may vary from 1.0 to 3.5.
Zeoiite X is a synthetic crystalline zeolitic molecular sieve which may
be represented by the formula:
(0.7-1.1 )Mv"O:AI203:(2.0-3.0)Si02:yH20
wherein M represents a metal, particularly alkali and alkaline earth metals, n
is the valence of M, and y may have any value up to about 8 depending on
the ideny~°~~ of M and the degree of hydration of tw a crystalline
zeolit~. Zeolite
X, its X- ! diffraction pattern, its properties, and method for its
preparation
are described in detail in U.S. Pat. No. 2,882,244. U.S. Pat. No. 2,882,244 is
hereby incorporated by reference to show a zeolite useful in the present
invention.
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Zeolite Y is a synthetic crystalline zeolitic molecular sieve which may
be written as:
(0.7-1.1 )Na20:A1203:xSi02:yH20
wherein x is a value greater than 3 up to about 6 and y may be a value up~to
about 9. Zeolite Y has a characteristic X-ray powder diffraction pattern which
may be employed with the above formula for identification. Zeolite Y is
described in more detail in U.S. Pat. No. 3,130,007. U.S. Pat. No. 3,130,007
is hereby incorporated by reference to show a zeolite useful in the present
invention.
The preferred non-acidic catalyst is a type L zeolite charged with one
or more dehydrogenating constituents.
The zeolitic catalysts according to the invention are charged with one
or more Group VI11 metals, e.g., nickel, ruthenium, rhodium, palladium,
iridium
or platinum.
The preferred Group Vlll metals are iridium and particularly platinum,
which are more selective with regard to dehydrocyclization and are also more
stable under the dehydrocyclization reaction conditions than other Group VIII
metals.
The preferred percentage of platinum in the dehydrocyclization catalyst
is between 0.1% and 5%, the lower limit corresponding to minimum catalyst
activity and the upper limit to maximum activity. This allows for the high
price
of platinum, which does not justify using a higher quantity of the metal since
the result is only a slight improvement in catalyst activity.
Group VIII metals are introduced into the large-pore zeolite by
synthesis, impregnation or exchange in an aqueous solution of appropriate
salt. When it is desired to introduce two Group VIII metals into the zeolite,
the
operation may be carried out simultaneously or sequentially.
By way of example, platinum can be introduced by impregnating the
zeolite with an aqueous solution of tetrammineplatinum (II) nitrate,
tetrammineplatinum (II) hydroxide, dinitrodiamino-platinum or
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tetrammineplatinum (Il) chloride. In an ion exchange process, platinum can
be introduced by using cationic platinum complexes such as
tetrammineplatinum (II) nitrate.
A preferred, but not essential, element of the present invention is the
presence of an alkaline earth metal in the dehydrocyclization catalyst. That
alkaline earth metal can be either barium, strontium or calcium. Preferably
the
alkaline earth metal is barium. The alkaline earth metal can be incorporated
into the zeolite by synthesis, impregnation or ion exchange. Barium is
preferred to the other alkaline earths because the resulting catalyst has high
activity, high selectivity and high stability.
An inorganic oxide may be used as a carrier to bind the large-pore
zeolite containing the Group VIII metal. The carrier can be a natural or a
synthetically produced inorganic oxide or combination of inorganic oxides.
Typical inorganic oxide supports which can be used include clays, alumina,
and silica, in which acidic sites are preferably exchanged by cations that do
not impart strong acidity.
The non-acidic catalyst can be employed in any of the conventional
types of equipment known to the art. It may be employed in the form of pills,
pellets, granules, broken fragments, or various special shapes, disposed as a
fixed bed within a reaction zone, and the charging stock may be passed
therethrough in the liquid, vapor, or mixed phase, and in either upward or
downward flow. Alternatively, it may be prepared in a suitable form for use in
moving beds, or in fluidized-solid processes, in which the charging stock is
passed upward through a turbulent bed of finely divided catalyst.
ACIDIC CATALYSTS
An acidic catalyst is used in conjunction with the non-acidic catalyst.
The acidic catalyst can comprise a metallic oxide support having disposed
therein a Group VIII metal. Suitable metallic oxide supports include aiu:iina
and silica. Preferably, the acidic catalyst comprises a metallic oxide support
having disposed therein in intimate admixture a Group VIII metal (preferably
platinum) and a Group VIII metal promoter, such as rhenium, tin, germanium,
cobalt, nickel, iridium, rhodium, ruthenium and combinations thereof. More
preferably, the acidic catalyst comprises an afumina support, platinum, and
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rhenium. A preferred acidic catalyst comprises platinum and tin on an
alumina support.
Preferably, the acidic catalyst has not been presulfided before use.
This is important to avoid sulfur contamination of the non-acidic catalyst by
recycle of part of the reformate produced by the acidic catalyst. On the other
hand, if one can insure no sulfur contamination of the non-acidic catalyst
from
the reformate produced by the acidic catalyst, then one might be able to use a
presulfided catalyst, such as Pt/Re on alumina.
REFORMING CONDITIONS
The reforming in both reformers is carried out in the presence of
hydrogen at a pressure adjusted to favor the dehydrocyclization reaction
thermodynamically and to limit undesirable hydrocracking reactions. The
pressures used preferably vary from 1 atmosphere to 500 psig, more
preferably from 50 to 300 psig, the molar ratio of hydrogen to hydrocarbons
preferably being from 1:1 to 10:1, more preferably from 2:1 to 6:1.
In the temperature range of from 400° C. to 600° C., the
dehydrocyclization reaction occurs with acceptable speed and selectivity. If
the operating temperature is below 400° C, the reaction speed is
insufficient
and consequently the yield is too low for industrial purposes. When the
operating temperature of dehydrocyclization is above 600° C.,
interfering
secondary reactions such as hydrocracking and coking occur, and
substantially reduce the yield. It is not advisable, therefore, to exceed the
temperature of 600° C. The preferred temperature range (430° C.
to 550° C.)
of dehydrocyclization is that in which the process is optimum with regard to
activity, selectivity and the stability of the catalyst.
The liquid hourly space velocity of the hydrocarbons in the
dehydrocyclization reaction is preferably between 0.3 and 5.
EXAMPLES
The invention will be further illustrated by following examples, which
set forth particularly advantageous method embodiments. While the
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Examples are provided to illustrate the present invention, they are not
intended to limit it.
EXAMPLE 1
Referring to Figure 1, in one embodiment, a full boiling hydrocarbon
feed 1 is fed to a depentanizer 10 to produce a C~ fraction stream 2 and a C6+
stream 3. The C6+ stream 3 is fed to splitter 15 to produce an overhead Cs-C,
cut 4 with nil Ca+, and a bottoms C~ cut 5 with all the C8+ material. Note
that
no Cs.,. material is in the overhead C6-C, cut 4. The bottoms C8+ cut 5
contains less than 10 Iv. % of C~_ hydrocarbon. The quantity of feed to the
overhead and bottoms cut, as well as the composition of each cut, is shown in
Table 1.
20
30
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Table 1
Overhead Bottoms
Feed wt% wt% feed comp wt% wt% feed comp wt%
n-paraffin
1.21 1.21 2.43 - -
13.49 13.49 27.06 - -
8.99 8.99 18.03 0.47 0.93
C8 10.60 - - 10.60 21.13
C9 3.69 - - 3.69 7.36
i-paraffin
C5 0.21 0.21 0.42 - -
10.06 10.06 20.17 - -
5.76 5.76 11.55 - -
CB 11.28 - - 11.28 22.50
C9 6.12 - - 6.12 12.21
C,o 0.42 - - 0.42 0.84
Olefins 0.64 0.64 1.28 - -
Naphthene
C5 0.40 0.40 0.80 - -
Cs 3.28 3.28 6.58 - -
C7 5.19 4.93 9.89 0.26 0.52
Ce 6.01 - - 6.01 11.99
C9 2.80 - - 2.80 5.58
Aromatics
C6 0.89 0.89 1.79 - -
C7 2.28 - - 2.28 4.35
Ce 5.88 - - 5.88 11.73
C9+ 0.33 - - 0.33 0.66
The overhead Cs-C, cut 4 is passed through a sulfur sorber 20 to
protect against sulfur/H2S contamination, and is processed over a first
reformer 22 which contains a non-acidic zeolite, such as Pt/K-Ba zeolite L, or
Pt/K zeolite L with and without fluorine and/or chlorine. Operating conditions
of the first reformer are 75 psig, 1.0 LHSV-"~-~, a hydrogen/hydrocarbon
(H2/HC) ratio of 5/1 mole/mole and a target Cs+C, normal and iso-paraffin
(n+i) paraffin conversion of 90-93%. The Cs and C, naphthenes as
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cyclohexanes are fully converted while the cyclopentanes are not fully
converted. The individual paraffin, iso-paraffin and naphthene conversion by
carbon number in the first reformer is shown in Table II with the associated
selectivity to the corresponding aromatic. The first reformate stream 24, from
the first reformer 22, has a benzene yield of 21.0 wt.% of splitter feed and a
toluene yield of 14.8 wt.% of splitter feed.
The bottoms C8+ cut 5 is passed through a sulfur sorber 30 to protect
against sulfur/H2S contamination, and is processed over a second
reformer 32 which contains an acidic bi-functional aromatization catalyst
which does not need to be sulfided, such as Pt/Sn/CI on alumina. Operating
conditions of the second reformer are 75 psig, 1.0 LHSV-~-~, H~/HC mole ratio
of 5/1 and a Ce+C9 (n+i)paraffin conversion of 95-100%. The CB and C9
naphthenes are also fully converted. The paraffin and naphthene conversion
7 5 and selectivity used are shown in Table II.
Table II
Conversion% Selectivity%
1 st Reformer
Cs n paraffins 91.0 92.9
C, n paraffins 98.0 84.0
Cs demethylbutane 40.0 -
Cs methylpentane 91.0 92.9
C, iso-parffins 98.0 84.0
C6 napththenes 89.1 92.9
C, napththenes 100.0 84.0
2nd Reformer
C, (n+i) paraffins 88.0 74.0
C8 (n+i) paraffins 100.0 81.0
C9 (n+i) paraffins 100.0 92.0
C~ napththenes 100.0 74.0
Ce napththenes 100.0 81.0
C9 napththenes 100.0 92.0
The first reformate stream 24 from the first reformer 22 is combined
with the second reformate stream 34 from the second reformer 32 and sent to
a common liquid-gas separator 40 where the H2 produced is recovered along
with C,-C3 gas and recycled to each reformer via a common recycle
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compressor 42. Excess H2 and C,-C3 exits the system via line 44 for
subsequent recovery of pure H2, and C,-C3 as fuel gas.
One of the benefits of having a common separator is that it then allows
for a common recycle compressor that operates on the off gas from the
separator. Alternatively we could also have two separate recycle
compressors (one for each reformer) to maintain operating flexibility. A
benefit of a common separator is that it reduces capital cost, which is
further
reduced if a common recycle compressor is used. A further benefit is that the
gas produced in the non-acidic reformer will have a higher hydrogen purity
than the gas produced in the acidic reformer. By combining these off-gases
the acidic reformer will be provided with a gas that has a higher hydrogen
purity. This can be taken advantage of by reducing fouling rate or lowering
recycle compressor capital and operating cost.
The liquid 46 from the separator 40 can be sent to a depentanizer to
recover a C4-CS overhead cut and a C6+ bottoms cut, and the components of
the C6+ stream can be processed to separate the stream into component
streams.
While the present invention has been described with reference to
specific embodiments, this application is intended to cover those various
changes and substitutions that may be made by those skilled in the art
without departing from the spirit and scope of the appended claims.
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