Note: Descriptions are shown in the official language in which they were submitted.
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GASOLINE PRODUCTION BY OLEFIN POLYMERIZATION
FIELD OF THE INVENTION
[0001) This invention relates to light olefin polymerization for the
production
of gasoline boiling range motor fuel.
CROSS REFERENCE TO RELATED APPLICATIONS
[0002] This application is related to co-pending applications Serial Nos.
, , and , of even date, claiming priority,
respectively from Applications Serial Nos. 60/656,955, 60/656,945, 60/656,946
and 60/656,947, all filed 28 February 2005 and entitled respectively, "Process
for Making High Octane Gasoline with Reduced Benzene Content, "Vapor
Phase Aromatics Alkylation Process", "Liquid Phase Aromatics Alkylation
Process" and "Olefins Upgrading Process".
BACKGROUND OF THE INVENTION
[0003] Following the introduction of catalytic cracking processes in
petroleum refining in the early 1930s, large amounts of olefins, particularly
light
olefins such as ethylene, propylene, butylene, became available in copious
quantities from catalytic cracking plants in refineries. While these olefins
may
be used as petrochemical feedstock, many conventional petroleum refineries
producing petroleum fuels and lubricants are not capable of diverting these
materials to petrochemical uses. Processes for producing fuels from these
crack-
ing off gases are therefore desirable and from the early days, a number of
different processes evolved. The early thermal polymerization process was
rapidly displaced by the superior catalytic processes of which there was a
number. The first catalytic polymerization process used a sulfuric acid
catalyst
to polymerize isobutene selectively to dimers which could then be hydrogenated
to produce a branched chain octane for blending into aviation fuels. Other
processes polymerized isobutylene with normal butylene to form a co-dimer
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which again results in a high octane, branched chain product. An alternative
process uses phosphoric acid as the catalyst, on a solid support and this
process
can be operated to convert all the C3 and C4 olefins into high octane rating,
branched chain polymers. This process may also operate with a C4 olefin feed
so
as to selectively convert only isobutene or both n-butene and isobutene. This
process has the advantage over the sulfuric acid process in that propylene may
be polymerized as well as the butenes and at the present time, the solid
phosphoric acid [SPA] polymerization process remains the most important
refinery polymerization process for the production of motor gasoline.
[0004] In the SPA polymerization process, feeds are pretreated to remove
hydrogen sulfide and mercaptans which would otherwise enter the product and
be unacceptable, both from the view point of the effect on octane and upon the
ability of the product to conform to environmental regulations. Typically, a
feed
is washed with caustic to remove hydrogen sulfide and mercaptans, after which
it is washed with water to remove organic bases and any caustic carryover.
Because oxygen promotes the deposition of tarry materials on the catalyst,
both
the feed and wash water are maintained at a low oxygen level. Additional pre-
treatments may also be used, depending upon the presence of various
contaminants in the feeds. With the most common solid phosphoric acid
catalyst, namely phosphoric acid on kieselguhr, the water content of the feed
needs to be controlled carefully because if the water content is too high, the
catalyst softens and the reactor may plug. Conversely, if the feed is too dry,
coke tends to deposit on the catalyst, reducing its activity and increasing
the
pressure drop across the reactor. As noted by Henckstebeck, the distribution
of
water between the catalyst and the reactants is a function of temperature and
pressure which vary from unit to unit, and for this reason different water
concentrations are required in the feeds to different units. Petroleum
Processing
Principles And Applications, R. J. Hencksterbeck McCrraw-Hill, 1959.
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[0005] There are two general types of units used for the SPA process, based
on the reactor type, the unit may be classified as having chamber reactors or
tubular reactors. The chamber reactor contains a series of catalyst beds with
bed
volume increasing from the inlet to the outlet of the reactor, with the most
common commercial design having five beds. The catalyst load distribution is
designed to control the heat of conversion.
[0006] Chamber reactors usually operate with high recycle rates. The recycle
stream, depleted in olefin content following polymerization, is used to dilute
the
olefin at the inlet of the reactor and to quench the inlets of the following
beds.
Chamber reactors usually operate at pressure of approximately 3500-5500 kPag
(about 500-800 psig) and temperature betweeri 180 'to 200 C (about 350 -
400 F). The conversion, per pass of the unit, is determined by the olefin
specification in the LPG product stream. Fresh feed LHSV is usually low,
approximately 0.4 to 0.8 hr""1. The cycle length for chamber reactors is
typically
between 2 to 4 months.
[0007] The tubular reactor is basically a shell-and-tube heat exchanger in
which the polymerization reactions take place in a number of parallel tubes
immersed in a cooling medium and filled with the SPA catalyst. Reactor
temperature is controlled with the cooling medium, invariably water in
commercial units, that is fed on the shell side of the reactor. The heat
released
from the 'reactions taking place inside the tubes evaporates the water on the
shell
side. Temperature profile in a tubular reactor is close to isothermal. Reactor
temperature is primarily controlled by means of the shell side water pressure
(controls temperature of evaporation) and secondly by the reactor feed
temperature. Tubular reactors usually operate at pressure between 5500 and
7500 kPag (800-1100 psig) and temperature of around 200 C (about 400 F).
Conversion per pass is usually high, around 90 to 93% and the overall conver-
sion is around 95 to 97%. The space velocity in tubular reactors is typically
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high, e.g., 2 to 3.5 hr"1 LHSV. Cycle length in tubular reactors is normally
between 2 to 8 weeks.
[0008] For the production of motor gasoline only butene and lighter olefins
are employed as feeds to polymerization processes as heavier olefins up to
about
Clo or C11 can be directly incorporated into the gasoline. With the SPA
process,
propylene and butylene are satisfactory feedstocks and ethylene may also be
included, to produce a copolymer product in the gasoline boiling range.
Limited
amounts of butadiene may be permissible although this diolefin is undesirable
because of its tendency to produce higher molecular weight poiymers and to
accelerate deposition of coke on the catalyst. The process generally operates
under relatively mild conditions, typically between 150 and 200 C, usually at
the lower end of this range between 150 and 180 C, when all butenes are
polymerized. Higher temperatures may be used when propylene is included in
the feed. In a well established commercial SPA polymerization process, the
olefin feed together with paraffinic diluent, is fed to the reactor after
being
preheated by exchange with the reaction effluent.
[0009] The solid phosphoric acid catalyst used is non-corrosive, which
permits extensive use of carbon steel throughout the unit. The highest octane
product is obtained by using a butene feed, with a product octane rating of
[R+M]/2 of 89 to 91 being typical. With a mixed propylene/butene feed,
product octane is typically about 91 and with propylene as the primary feed
component, product octane drops to typically 87.
[0010] In spite of the advantages of the SPA polymerization process, which
have resulted in over 200 units being built since 1935 for the production of
gasoline fuel, a number of disadvantages are encountered, mainly from the
nature of the catalyst. Although the catalyst is non-corrosive, so that much
of
the equipment may be made of carbon steel, it does lead it to a number of draw-
backs in operation. First, the catalyst life is relatively short as a result
of pellet
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disintegration which causes an increase in the reactor pressure drop. Second,
the
spent catalyst encounters difficulties in handling from the environmental
point of
view, being acidic in nature. Third, operational and quality constraints limit
flexible feedstock utilization. Obviously, a catalyst which did not have these
disadvantages would offer considerable operating and economic advantages.
[0011] The Mobil Olefins-to-Gasoline [MOG] process employs a proprietary
shape selective zeolite catalyst in a fluidized bed reactor to produce high
octane
motor gasoline by the conversion of reactive olefins such as ethylene and
propylene in FCC off-gas; butenes as well as higher olefins may also be
included
and converted to form a high octane, branched chain gasoline product. The feed
is converted over the catalyst into C5+ components by mechanisms including
oligomerization, carbon number redistribution hydrogen transfer,
aromatization,
alkylation and isomerization. Based on olefins converted, MOG yields 60 to 75
weight percent of high-octane gasoline blend stock with specific qualities of
the
product depending of the processing severity selected and the character of the
feed olefins. Typically, the octane rating for the product is in the range of
88 to
91 [R+M]/2. The zeolite catalyst used in the process is environmentally safe
and
its attrition rate is low, and as an alternative to disposal, the spent
catalyst can be
reused in the FCC unit to increase octane quality.
[0012] The MOG process has, however, the economic disadvantage relative
to the SPA process in that new capital investment may be required for the
fluidized bed reactor and regenerator used to operate the process. If an
existing
SPA unit is available in the refinery, it may be difficult to justify
replacement of
the equipment in spite of the drawbacks of the SPA process, especially in view
of current margins on fuel products. Thus, although the MOG process is
technically superior, with the fluidized bed operation resolving heat problems
and the catalyst presenting no environmental problems, displacement of
existing
SPA polymerization units has frequently been economically unattractive. What
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. is required, therefore, is an economically attractive alternative to the SPA
process for the condensation of light olefins to form motor fuels. Desirably,
the
process should be capable of operation in existing refinery equipment,
especially
as a "drop in" type replacement for the solid phosphoric acid catalyst used in
the
SPA process so that existing SPA polymerization units can be directly used
with
the new catalyst. This implies that the process should use a non-corrosive,
solid
catalyst in fixed bed catalyst operation. Furthermore, the catalyst should
present
fewer handling, operational and disposal problems than solid phosphoric acid
and, for integration into existing refineries, the product volumes and
distribu-
tions should be comparable to those of'the SPA process.
SUMMARY OF THE INVENTION
[0013] We have now devised a process for the conversion of light olefins
such as ethylene, propylene, and butylene to gasoline boiling range motor
fuels
which is capable of being used as a replacement for solid phosphoric acid
catalyst in process units which have previously been used for the SPA process.
The catalyst used in the present process is a solid, particulate catalyst
which is
non-corrosive, which is stable in fixed bed operation, which exhibits the
capability of cycle durations before regeneration is necessary and which can
be
readily handled and which can be finally disposed of simply and economically
without encountering significant environmental problems. Accordingly, the
catalyst used in the present process commends itself as a "drop in"
replacement
for the solid phosphoric acid catalyst used in the SPA catalytic condensation
process for the production of motor fuels.
[0014] According to the present invention, a light olefin stream such as
ethylene, propylene, optionally with butylene and possibly other light
olefins, is
polymerized to form a gasoline boiling range [C5 + - 200 C] [C5+ - 400 F]
product in the presence of a catalyst which comprises a member of the MWW
family of zeolites, a family which includes zeolites PSH 3, MCM-22, MCM-49,
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MCM-56, SSZ 25, ERB-1 and ITQ-1. The term "polymerized" is used here
consistent with the petroleum refinery usage although, in fact, the process is
one
of oligomerization (which term will be used in this specification
interchangeably
with the conventional term) in which a low molecular weight polymer is the
desired product. The process is carried out in a fixed bed of the catalyst
with
feed dilution, normally a hydrocarbon diluent, or added quench to control the
heat release which takes place. In additional to their easy handling and
amenability to regeneration, the solid catalysts used in the present process
exhibit better activity and selectivity than solid phosphoric acid; compared
to
SPA, MCM-22 itself is three to seven times more active and significantly more
stable for the production of motor gasoline by the polymerization of light
olefin
feeds. The catalytic performance of regenerated MCM-22 catalyst is compar-
able to that of the fresh MCM-22 catalyst, demonstrating that the catalyst is
amenable to conventional oxidative regeneration techniques.
[0015] The conversion of an SPA process unit to operation with the present
molecular sieve based catalysts therefore comprises, in principle, withdrawing
the solid phosphoric acid [SPA] catalyst from the unit and loading an olefin
condensation catalyst comprising an MWW zeolite material into the reactor of
the process unit. Following the conversion to operation with the MWW zeolite
catalyst, the unit may be used for production of the gasoline and, if desired,
other
liquid hydrocarbon fuels by polymerization of the refinery olefins using the
appropriate conditions as described below.
DRAWING
[0016] Figure 1 shows a process schematic for the olefin polymerization unit
for converting light refinery olefins to motor gasoline by the present
process.
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DETAILED DESCRIPTION OF THE INVENTION
SPA Unit Conversion
[0017] The present process is for the condensation of light cracking olefins
to
produce motor gasoline and other motor fuels, for example, road diesel blend
stock and is intended to provide a replacement for the SPA polymerization
process. It provides a catalyst which can be used as a direct replacement for
SPA and so enables existing SPA units to be used directly with the new
catalyst,
so allowing the advantages of the new catalyst and process to be utilized
while
retai_n.ing the economic benefit of existing refinery equipment.
[0018] The process units used for the operation of the SPA process for the
catalytic condensation of light olefins to produce motor fuels are well known:
These units typically comprise a feed surge drum to which the olefins and any
diluent are supplied, followed by a heat exchanger in which the feed is
preheated
by exchange with the reactor effluent, after which it is charged to the
reactor
where the polymerization (condensation) takes place. Control of the heat
release
in the reactor is accomplished both by feed dilution and by the injection of
recycled quench between catalyst beds in the reactor. The reactor effluent,
cooled in exchange with the feed, is directed to a flash drum where the flash
vapor is condensed and the condensate cooled. Some of the condensate is
recycled for use as feed diluent and quench. Flash drum liquid flows to a
stabilizer where the polymer gasoline product at the desired Reid Vapor
Pressure
[RVP] and light ends are separated. The light ends may be sent to a C3-C4
splitter depending on the refinery needs. SPA units of this kind can be
directly
converted to use the catalysts of the present process without significant
changes
since the present catalysts are a straight forward "drop in" replacement for
the
solid phosphoric acid [SPA] catalyst used in the conventional process
technology.
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[0019] Like SPA, the molecular sieve catalysts used in the present process are
non-corrosive but possess significant advantages with respect to SPA, in that
they are more stable, less subject to break down and are largely unaffected by
the amount of water in the feed. The present catalysts are readily regenerable
using conventional hydrogen stripping or oxidative regeneration, after which
complete catalytic activity is substantially restored. Cycle times before
regeneration or reactivation is required may be six months, one year, or even
longer, representing a significant improvement over SPA. Since conventional
SPA condensation units necessarily include facilities for the discharge and
reloading of catalysts as a result of the short life of a catalyst, these
units may
readily accommodate the present molecular sieve catalysts. The SPA units do
not, however, include facilities for in-situ regeneration since the SPA
catalyst is
used on a once-through basis before it requires disposal. The molecular sieve
catalysts used in the present process, however, are fully regenerable and for
this
purpose, will need to be withdrawn from the reactors for ex-situ regeneration.
This will typically be a simple matter to arrange using the spent catalyst
discharge equipment of the SPA unit. Similarly, the SPA charging equipment
lends itself directly to the charging of the zeolite catalysts into the
reactors.
Unit Conversion
[0020] A schematic for a converted olefin condensation unit made by the
conversion of an existing SPA unit is shown in simplified from in Figure 1. A
light olefin feed, typically C2, C3 or C4 olefins or mixtures of these olefins
from
an FCC gas plant, is led into the unit through line 10 and combined with
recycled hydrocarbon as diluent before passing through heat exchanger 12 in
which it picks up heat from the reactor effluent before being brought to
reaction
temperature in heater 13. The olefin charge plus diluent passes through a
guard
bed reactor 14a to remove contaminants such as organic nitrogen and sulfur-
containing impurities. The guard bed may be operated on the swing cycle with
two beds, 14a, 14b, one bed being used on stream for contaminant removal and
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the other on regeneration in the conventional manner. If desired, a three-bed
guard bed system may be used with the two beds used in series for contaminant
removal and the third bed on regeneration. With a three guard system used to
achieve low contaminant levels by the two-stage series sorption, the beds will
pass sequentially through a three-step cycle of: regeneration, second bed
sorption, first bed sorption.
[0021] The olefins in the charge stream are polymerized or condensed in
reactor 15 to form the desired olefin polymer product during its passage over
a
sequence of catalyst beds in the reactor. Additional diluent is injected as
quench
from line 16 between the beds in order to control the exotherm. Effluent
passes
out of the reactor through heat exchanger 12 and then to flash drum 20 in
which
the diluent is separated from the olefin polymer product. The diluent which is
suitably a light paraffin such as propane, butane and a portion of the
polymeriza-
tion product, is passed to recycle drum 21 and from there by way of recycle
pump 23 and line 24 to feed line 10 for feed dilution and to recycle line 16
for
injection as interbed quench in reactor 15. The olefin polymer product passes
out of flash drum 20 through line 22 to the fractionator 25 to provide the
final
stabilized gasoline blend component in line 26 with reboil loop 28 providing
column heat; light ends including unreacted olefins pass out through line 27
from reflux loop 29. As noted below, there is the potential for iso-paraffinic
components to undergo reaction with the olefins in the feed to produce highly
desirable branched chain reaction products of high octane value in the
gasoline
boiling range. The use of the recycle as feed diluent is therefore desirable
not
only for controlling reaction temperatures but also since it may also result
in an
increase in product octane.
Catalyst
[0022] The catalysts used in the present process contain, as their essential
catalytic component, a molecular sieve of the MWW type. The MWW family of
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zeolite materials has achieved recognition as having a characteristic
framework
structure which presents unique and interesting catalytic properties. The MWW
topology consists of two independent pore systems: a sinusoidal ten-member
ring [10 MR] two dimensional channel separated from each other by a second,
two dimensional pore system comprised of 12 MR super cages connected to
each other through 10 MR windows. The crystal system of the MWW frame-
work is hexagonal and the molecules diffuse along the [100] directions in the
zeolite, i.e., there is no communication along the c direction between the
pores.
In the hexagonal plate-like crystals of the MWW type zeolites, the crystals
are
formed of relatively small number of units along the c direction as a result
of
which, much of the catalytic activity is due to active sites located on the
external
surface of the crystals in the form of the cup-shaped cavities. In the
interior
structure of certain members of the family such as MCM-22, the cup-shaped
cavities combine together to form a supercage. The MCM-22 family of zeolites
has attracted significant scientific attention since its initial announcement
by
Leonovicz et al. in Science 264, 1910-1913 [1994] and the later recognition
that
the family is currently known to include a number of zeolitic materials such
as
PSH 3, MCM-22, MCM-49, MCM-56, SSZ-25, ERB-1, ITQ-1, and others.
Lobo et al. AIChE Annual Meeting 1999, Paper 292J.
[0023] The relationship between the various members of the MCM-22 family
have been described in a number of publications. Four significant members of
the family are MCM-22, MCM-36, MCM-49, and MCM-56. When'initially
synthesized from a mixture including sources of silica, alumina, sodium, and
hexamethylene imine as an organic template, the initial product will be MCM-22
precursor or MCM-56, depending upon the silica: alumina ratio of the initial
synthesis mixture. At silica:alumina ratios greater than 20, MCM-22. precursor
comprising H-bonded vertically aligned layers is produced whereas randomly
oriented, non-bonded layers of MC-56 are produced at lower silica:alumina
ratios. Both these materials may be converted to a swollen material by the use
of
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a pillaring agent and on calcination, this leads to the laminar, pillared
structure
of MCM-36. The as-synthesized MCM-22 precursor can be converted directly
by calcination to MCM-22 which is identical to calcined MCM-49, an
intermediate product obtained by the crystallization of the randomly oriented,
as-synthesized MCM-56. In MCM-49, the layers are covalently bonded with an
interlaminar spacing slightly greater than that found in the calcined MCM-
22/1VICM 49 materials. The unsynthesized MCM-56 may be calcined itself to
form calcined MCM 56 which is distinct from calcined MCM-22/MCM-49 in
having a randomly oriented rather than a laminar structure. In the patent
litera-
ture MCM-22 is described in U.S. Patent No. 4,954,325 as well as in U.S.
5,250,777; 5,284,643 and 5,382,742. MCM-49 is described in U.S. 5,236,575;
MCM-36 in U.S. 5,229,341 and MCM-56 in U.S. 5,362,697.
[0024] The preferred zeolitic material for use in the catalyst of the present
process is MCM-22 although zeolite MCM-49 may be found to have certain
advantages relative to MCM-22. It has been found that the MCM-22 may be
either used fresh, that is, not having been previously used as a catalyst or
alternatively, regenerated MCM-22 may be used. Regenerated MCM-22 may be
used after it has been used in any of the catalytic processes for which it is
suitable, including the present process in which the catalyst has shown itself
remain active after even multiple regenerations. It may also be possible to
use
MWW catalysts which have previously been used in other commercial processes
and for which they are no longer acceptable, for example, MCM-22 catalyst
which has previously been used for the production of aromatics such as ethyl-
benzene or cumene, normally using reactions such as alkyaltion and transalkyla-
tion. The cumene production (alkylation) process is described in U.S. Patent
No.
4,992,606 (Kushnerick et al). Ethylbenzene production processes are described
in U.S. Pat. Nos. 3,751,504 (Keown); 4,547,605 (Kresge); and 4,016,218
(Haag); U.S. Pat. Nos. 4,962,256; 4,992,606; 4,954,663; 5,001,295; and
5,043,501 describe alkylation of aromatic compounds with various alkylating
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agents over catalysts comprising MWW zeolites such as PSH-3 or MCM-22.. US
Patent No. 5,334,795 describes the liquid phase synthesis of ethylbenzene with
MCM-22.
[0025] As noted above, MCM-22 catalysts may be regenerated after catalytic
use in these processes and other aromatics production processes by
conventional
air oxidation techniques similar to those used with other zeolite catalysts.
Conventional air oxidation techniques are also suitable when regenerating the
catalysts after use in the present process.
[0026] In addition to the MWW active component, the catalysts for use in the
present process will often contain a matrix material or binder in order to
give
adequate strength to the catalyst as well as to provide the desired porosity
characteristics in the catalyst. High activity catalysts may, however, be
formulated in the binder-free form by the use of suitable extrusion
techniques,
for example, as described in U.S. 4,908,120. When used, matrix materials
suitably include alumina, silica, silica alumina, titania, zirconia, and other
inorganic oxide materials commonly used in the formulation of molecular sieve
catalysts. For use in the present process, the level of MCM-22 in a finished
matrixed catalyst will be typically from 20 to 70% by weight, and in most
cases
from 25 to 65% by weight. In manufacture of a matrixed catalyst, the active
ingredient will typically be mulled with the matrix material using an aqueous
suspension of the catalyst and matrix, after which the active component and
the
matrix are extruded into the desired shape, for example, cylinders, hollow
cylinders, trilobe, quadlobe, etc. A binder material such as clay may be added
during the mulling in order to facilitate extrusion, increase the strength of
the
final catalytic material and to confer other desirable solid state properties.
The
amount of clay will not normally exceed 10% by weight of the total finished
catalyst. Self-bound catalysts (alternatively referred to as unbound or binder-
free catalysts), that is, catalysts which do not contain a separately added
matrix
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or binder material, are useful and may be produced by the extrusion method
described in U.S. Pat. No. 4,582,815, to which reference is made for a
description of the method and of the extruded products obtained by its use.
The
method described there enables extrudates having high constraining strength to
be produced on conventional extrusion equipment and accordingly, the method
is eminently suitable for producing the high activity self-bound catalysts.
The
catalysts are produced by mulling the zeolite, as described in U.S. Pat. No.
4,582,815, with water to a solids level of 25 to 75 wt% in the presence of
0.25 to
wt% of basic material such as sodium hydroxide. Further details are to be
found in U.S. Pat. No. 4,582,815. Generally, the self-bound catalysts can be
characterized as particulate catalysts in the form, for instance, of
extrudates or
pellets, containing at least 90 wt. pct., usually at least 95 wt. pct., of the
active
zeolite component with no separately added b,inder material e.g. alumina,
silica-
alumina, silica, titania, zirconia etc. Extrudates may be in the conventional
shapes such as cylinders, hollow cylinders, trilobe, quadlobe, flat platelets
etc.
[0027] The catalyst used in the guard bed will normally be the same catalyst
used in the alkylation reactor as a matter of operating convenience but this
is not
required: if desired another catalyst or sorbent to remove contaminants from
the
feed may used, typically a cheaper guard bed sorbent, e.g a used catalyst from
another process or alumina. The objective of the guard bed is to remove the
contaminants from the feed before the feed comes to the reaction catalyst and
provided that this is achieved, there is wide variety of choice as to guard
bed
catalysts and conditions useful to this end. The volume of the guard bed will
normally not exceed about 20% of the total catalyst bed volume of the unit.
Olefin Feed
[0028] The light olefins used as the feed for the present process are normally
obtained by the catalytic. cracking of petroleum feedstocks to produce
gasoline
as the major product. The catalytic cracking process, usually in the form of
fluid
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catalytic cracking (FCC) is well established and, as is well known, produces
large quantities of light olefins as well as olefinic gasolines and by-
products such
as cycle oil which are themselves subject to further refining operations. The
olefins which are primarily useful in the present process are the lighter
olefins
from ethylene up to butene; although the heavier olefins may also be included
in
the processing, they can generally be incorporated directly into the gasoline
product where they provide a valuable contribution to octane. The present
process is highly advantageous in that it will operate readily not only with
butene and propylene but also with ethylene and thus provides a valuable route
for the conversion of this cracking by-product to the desired gasoline
product.
For this reason as well as their ready availability in large quantities in a
refinery,
mixed olefin streams such a FCC Off-Gas streams (typically containing
ethylene, propylene and butenes) may be used. Conversion of the C3 and C4
olefin fractions from the cracking process provides a direct route to the
branch
chain C6, C7 and C8 products which are so highly desirable in gasoline from
the
view point of boiling point and octane. Besides the FCC unit, the mixed olefin
streams may be obtained from other process units including cokers, visbreakers
and thermal crackers. The presence of diolefins which may be found in some of
these streams is not disadvantageous since catalysis on the 1VIWW family of
zeolites takes place on surface sites rather than in the interior pore
structure as
with more conventional zeolites so that plugging of the pores is less
problematic
catalytically. Appropriate adjustment of the process conditions will enable co-
condensation products to be produced when ethylene, normally less reactive
than
its immediate homologs, is included in the feed. The compositions of two
typical FCC gas streams is given below in Tables 1 and 2, Table 1 showing a
light FCC gas stream and Table 2 a stream from which the ethylene has been
removed in the gas plant for use in the refinery fuel system.
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Table 1
FCC Light Gas Stream
Component Wt. Pct. Mol. Pct.
Ethane 3.3 5.1
Ethylene 0.7 1.2
Propane 14.5 15.3
Propylene 42.5 46.8
Iso-butane 12.9 10.3
n-Butane 3.3 2.6
Butenes 22.1 18.32
Pentanes 0.7 0.4
Table 2
C3-C4 FCC Gas Stream
Component Wt. Pct.
1- Propene 18.7
Propane 18.1
Isobutane 19.7
2-Me-l-propene 2.1
1-Butene 8.1
n-Butane 15.1
Trans-2-Butene 8.7
Cis-2-butene 6.5
Isopentane 1.5
C3 Olefins 18.7
C4 Olefins 25.6
Total Olefins 44.3
[0027] While the catalysts used in the present process are robust they do have
sensitivity to certain contaminants (the conventional zeolite deactivators),
especially organic compounds with basic nitrogen as well as sulfur-containing
organics. It is therefore preferred to remove these materials prior to
entering the
unit if extended catalyst life is to be expected. Scrubbing with contaminant
removal washes such as caustic, MEA or other amines or aqueous wash liquids
will normally reduce the sulfur level to an acceptable level of about 10-20
ppmw
and the nitrogen to trace levels at which it can be readily tolerated. One
attractive feature about the present process is that although activity
benefits are
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achieved by the use of low or very low water levels in.the feed, it is not
other-
wise unduly sensitive to water, making it less necessary to control water
entering
the reactor than it is in SPA units. Unlike SPA, the zeolite catalyst does not
require the presence of water in order to maintain activity and therefore the
feed
may be dried before entering the unit, for example, to below 200 ppmw water or
lower, e.g. below 50 or even 20 ppmw. In conventional SPA units, the water
content typically needs to be held between 300 to 500 ppmw at conventional
operating temperatures for adequate activity while, at the same time,
retaining
catalyst integrity. The present zeolite catalysts, however, may readily
tolerate
higher levels of water up to about 1,000 ppmw water although levels above
about 800 ppmw may reduce activity, depending on temperature. Thus, with
converted units, the olefin feed may contain from 300 or 500 to 1,000 ppmw
water, although 300-800 ppmw should be regarded as a workable range for
activity with existing feed treatment equipment. As noted, however, activity
benefits are secured with markedly lower feed water levels and these benefits
may justify feed pre-treatment to operate at lower water contents.
[0028] The use of the guard bed is particularly desirable in the operation of
the
present process since the refinery feeds customarily routed to polymerization
units (as distinct from petrochemical unit feeds which are invariably high
purity
feeds for which no guard bed is required) may have a contaminant content,
especially of polar catalyst poisons, such as the polar organic nitrogen and
organic sulfur compounds, which is too high for extended catalyst life. The
use
of a cheaper catalyst in the guard bed reactors which can be readily
regenerated
in swing cycle operation or, alternatively disposed of on a once-through
basis, is
therefore desirable in ensuring extended cycle duration for the active
polymerization catalyst.
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Process Parameters
[0029] The present process is notable for its capability of being operated at
low
temperatures and under moderate pressures. In general, the temperature will be
from about 120 to 250 C (about 250 to 480 F) and in most cases between 150
and 250 C (about 300 -480 C). Temperatures of 170 to 205 C (about 340 to
400 F) will normally be found optimum for feeds comprising butene while
higher temperatures will normally be appropriate for feeds with significant
amounts of propene. Pressures may be those appropriate to the type of unit
from
which the conversion was made, so that pressures up to about 7500 kPag (about
1100 psig) will be typical but normally lower pressures will be sufficient,
for
example, below about 7,000 Kpag (about 1,000 psig) and lower pressure opera-
tion may be readily utilized, e.g. up to 3500 kPag (about 500 psig). Ethylene,
again, will require higher temperature operation to ensure that the products
remain in the gasoline boiling range. Space velocity may be quite high, for
example, up to 50 WHSV (hr-1) but more usually in the range of 5 to 30 WHSV.
Gasoline Product Formation
[0030] With gasoline as the desired product, a high quality product is
obtained
from the polymerization step, suitable for direct blending into the refinery
gasoline pool after fractionation as described above. With clean feeds, the
product is correspondingly low in contaminants. The product is high in octane
rating with RON values of 95 being regularly obtained and values of over 97
being typical; MON is normally over 80 and typically over 82 so that
(RON+MON)/2 values of at least 89 or 90 are achievable with mixed
propylene/butene feeds. Of particular note is the composition of the octenes
in
the product with a favorable content of the higher-octane branched chain
components. The linear octenes are routinely lower than with the SPA product,
typically being below 0.06 wt. pct. of all octenes except at the highest
conver-
sions and even then, the linears are no higher than those resulting from SPA
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catalyst. The higher octane di-branched octenes are noteworthy in consistently
.
being above 90 wt. pct. of all octenes, again except at the highest
conversions
but in all cases, higher than those from SPA; usually, the di-branched octenes
will be at least 92 wt. pct of all octenes and in favorable cases at least 93
wt. pct.
The levels of tri-branched octenes are typically lower than those resulting
from
the SPA process especially at high conversions, with less than 4 wt. pct being
typically except at the highest conversions when 5 or 6 wt. pct. of all
octenes
may be achieved, approximately half that resulting from SPA processing. In the
C5-200 C product fraction, high levels of di-branched C8 hydrocarbons may be
found, with at least 85 weight percent of the octene components being di-
branched C8 hydrocarbons, e.g. 88 to 96 weight percent di-branched C8
hydrocarbons.
[0031 ] Depending on feed composition, reactions other than direct olefin poly-
merization may take place. If branch chain paraffins are present, for example,
from the paraffinic diluent stream, olefin-isoparaffin alkylation reactions
may
take place, especially at the higher conversion levels to which the process is
well
suited, leading to the production of branched-chain, gasoline boiling range
products of high octane rating. The reaction between butene and iso-butane and
between propylene and iso-butane is of particular value since it will result
in the
product containing very desirable, high octane gasoline components. At low to
moderate olefin conversion levels, the isoparaffin-olefin alkylation reaction
is
not significant but at higher conversions above about 75% (olefin conversion),
e.g. at conversion levels of 80% or more (olefin conversion), particularly at
90%
or higher, this reaction will increase markedly with the production of high
octane gasoline components.
[0032] The following examples are given by way of illustration.
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Example 1 2-Butene Oligomerization with Solid Phosphoric Acid
[0033] A commercial solid phosphoric acid (SPA) catalyst was sized to 14-24
mesh in a glove box. In a glove bag, one gram of this sized catalyst was
diluted
with sand to 3 cc and charged to an isothermal, down-flow, 9 mm (outside
diameter) fixed-bed reactor. The catalyst was dried at 150 C and atmospheric
pressure with 100 cc/min flowing N2 for 2 hours. The N2 was turned off and
reactor pressure was set to 5270 kPaa (750 psig) by a Grove loader. A 2-butene
feed (containing 57.1% cis-butene, 37.8% trans-butene, 2.5% n-butane, 0.8%
isobutene and 1-butene, 1.8% others) was introduced into the reactor at 60
cc/hr
until desire reactor pressure of 5270 kPaa (750 psig) was reached. The reactor
temperature was then ramped at 2 C per minute to 170 C. During the tempera-
ture ramp, the feed flow was reduced to a desired level and kept at this level
for
12 hours before data collection. Liquid products were collected at 50%, 70%,
then 90% conversion (not necessary in this order) in a cold-trap and analyzed
off-line.
[0034] Product carbon number distribution was determined with a HP-5890 GC
equipped with a 60 meter DB-1 column (0.25 mm id and 1000 nm film thick-
ness). Product branching was determined with an H2-GC using a HP-5890 GC
equipped with (a) a 100 meter DB-1 column (0.25 mm id and 500 nm film thick-
ness); (b) hydrogen as the carrier gas; and (c) 0.1 g of 0.5% Pt/alumina
catalyst
in the GC insert for in-situ hydrogenation. Both GC's use the same temperature
program: 2 min at -20 C, 8 C/min to 275 C, hold at 275 C for 35 min.
[0035] Representative data collected at 50%, 70%, and 90% nominal conver-
sion are shown in Tables 1-3.
Average Branching = 0 x % linear + 1 x % mono-branched + 2 x % di-branched
+ 3 x.% tri-branched
where: % linear + % mono-branched + % di-branched +% tri-branched = 100%
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Example 2 2-Butene Oligomerization with A Binder-Free MCM-22
[0036] The catalyst was a binder-free, 1.6 mm quadrulobe extrudate containing
100% MCM-22. A 0.10 gram sample of this catalyst, chopped to 3mm length,
was tested for 2-butene oligomerization using the same procedure described in
Example 1. Representative data are shown in Tables 3-5.
Example 3 2-Butene Oligomerization with Alumina-Bound MCM-22
[0037] The catalyst was a 1.6 mm cylindrical extrudate containing 65% MCM-
22 aiid 35% alumina binder. A 0.20 gram of this catalyst, chopped to 1.6mm
length, was tested for 2-butene oligomerization using the same procedure
described in Example 1. Representative data are shown in Tables 3-5.
Example 4' 2-Butene Oligomerization with a Spent-Regenerated MCM-22
[0038] The same MCM-22 catalyst, as described in Example 3, was used in a
commercial petrochemical catalytic process until it became unsuitable for
further
service in this application; it was then commercially regenerated in full air
at
455 C (850 F) and about 60 torr partial pressure of water. The catalyst was
treated at these conditions for about 1 hour to achieve complete regeneration.
A
0.15 gram of this catalyst, chopped to 1.6 mm length, was evaluated for 2-
butene
oligomerization using the same procedure described in Example 1. Representa-
tive data are shown in Tables 3-5.
Example 5 2-Butene Oligomerization in with a Spent-Regenerated-Steamed
MCM-22
[0039] The same Spent-Regenerated MCM-22 catalyst, as described in
Example 4, was steamed at 510 C and 1 atm for 5 hours with 100% steam. A
0.20 gram of this "spent-regenerated-steamed" catalyst, chopped to 1.6 mm
length, was evaluated for 2-butene oligomerization using the same procedure
described in Example 1. Representative data are shown in Tables 3-5.
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Example 6 2-Butene Oligomerization with a Multiply-Regenerated MCM-22
[0040] The same Spent-Regenerated MCM-22 catalyst, as described in
Example 4, was further treated in the laboratory to simulate multiple
commercial
regeneration (oxidative) conditions. The sample was treated in flowing air at
700 v/v/minute, 455C (850 F) and 60 torr partial pressure of water for 1 hour.
The treatment was repeated 4 additional times on the same sample, with the
intention of simulating 5 additional regenerations of the already commercially
regenerated MCM-22. A 0.20 gram of this multiple regenerated catalyst,
chopped to 1.6 mm length, was evaluated for 2-butene oligomerization using the
same procedure described in Example 1. Representative data are shown in
Tables 3-5.
Comparison of Catalyst Performance in a Fixed-Bed Reactor
[0041] Tables 3-5 compare catalyst performance at 50%, 70%, and 90%
conversion of 2-butene. Table 4 further compares catalysts activity and
deactiva-
tion rate. Relative activity of each catalyst is determined by measuring the
first-
order rate constant for 2-butene oligomerization at 170 C relative to that of
SPA
catalyst. Deactivation rate is given. as conversion drop observed per day per
WHSV.
[0042] The data in Tables 3-5 show that, when compared to SPA at constant
conversion level, MCM-22 provided comparable product selectivity and average
branching. MCM-22 produced slightly more di-branched and less tri-branched
octenes than those of SPA.
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Table 3
Comparison of Catalyst Performance at Nominal 50% Conversion
Binder-free Alumina- Regen Regen. Multiple
Catalyst SPA MCM-22 bound MCM-22 Steamed Regen
MCM-22 MCM-22 MCM-22
Ex. Number 1 2 3 4 5 6
WHSV 5.7 48.7 29.3 33.3 17.8 28.1
TOS, day 13.1 8.1 4.8 5.3 0.8 3.9
Conversion % 53.5 53.4 54.2 48.5 56.26 55.21
Product Selectivity, wt%
C5_7 0.91 1.13 1.44 1.60 1.61 1.61
C8= 80.71 82.94 79.58 78.16 75.27 77.57
C9_11 2.18 2.44 3.65 3.50 3.55 3.44
Cla_ 14.65 11.45 12.94 13.01 14.81 13.11
C16= 1.50 2.05 2.38 3.63 4.49 3.65
C20_ 0.05 0.00 0.02 0.12 0.22 0.62
C24+ 0.00 0.00 0.00 0.00 0.06 0.00
Total 100.00 100.00 100.00 100.00 100.00 100.00
Average Branching
Me/C8 1.98 1.97 1.98 1.99 1.98 1.99
Me/C12 2.47 2.45 2.44 2.44 2.45 2.44
Octene Distribution, %
Linear 0.08 0.04 0.06 0.03 0.02 0.02
Mono-branched 6.28 4.65 4.59 3.09 3.20 2.94
di-branched 89.28 93.63 92.81 94.85 95.09 95.24
tri-branched 4.36 1.68 2.54 2.03 1.68 1.80
Sum 100.00 100.00 100.00 100.00 100.00 100.00
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Table 4
Comparison of Catalyst Performance at Nomina170% Conversion
Binder-free Alumina- Spent-Regen Multiple
Catalyst SPA MCM-22 bound MCM-22 Regen MCM-22
MCM-22
Example No. 1 2 3 4 6
WHSV 4.0 38.1 16.3 6.7 19.0
TOS, day 11.5 5.9 6.3 8.8 0.8
Conversion % 70.7 70.4 68.7 73.3 66.3
Product Selectivity, wt%
C5_7 1.03 1.65 1.91 1.98 2.04
Cg_ 77.98 77.41 71.52 71.14 70.55
C9_11 2.48 3.27 4.52 4.31 4.24
C12_ 16.85 14.25 17.21 16.38 16.34
C16= 1.61 3.42 4.82 5.75 6.21
C20-- 0.02 0.00 0.02 0.40 0.56
C2A+ 0.03 0.00 0.00 0.04 0.06
Total 100.00 100.00 100.00 100.00 100.00
Average Branching
Me/C8 1.99 1.97 1.99 1.99 1.99
Me/C12 2.47 2.44 2.43 2.43 2.44
Octene Distribution, %
Linear 0.08 0.06 0.06 0.03 0.03
mono-branched 6.19 5.38 4.68 3.72 3.50
di-branched 87.92 92.08 91.46 93.51 94.05
tri-branched 5.82 2.48 3.79 2.74 2.42
Sum 100.00 100.00 100.00 100.00 100.00
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Table 5
Comparison of Catalyst Performance at Nominal 90% Conversion
Catalyst SPA Binder-free MCM-22 Alumina-bound MCM-22
Exam le Number 1 2 3
WHSV 2.2 19.2 13.1
TOS, day 20.8 3.9 0.1
Conversion % 92.48 88.1 80.4
Product Selectivity, wt%
C5_7 1.55 3.28 2.54
C8_ 67.62 64.94 61.52
C9_11 3.68 5.11 5.92
C12= 23.44 18.67 21.22
C16= 3.41 7.34 8.21
C20-- 0.30 0.66 0.59
C24+ 0.00 0.00 0.00
Total 100.00 100.00 100.00
Average Branching
Me/C8 2.05 1.98 2.01
Me/C12 2.47 2.41 2.43
Octene Distribution, %
Linear 0.12 0.12 0.09
mono-branched 6.63 6.85 5.08
di-branched 81.67 87.87 88.65
tri-branched 11.58 5.16 6.19
Sum 100.00 100.00 100.00
[0043] Data in Table 6 show that MCM-22 catalysts are more active than SPA:
nine times,more active with binderless MCM-22 and five times with alumina-
bound MCM-22. A comparison of deactivation rate shows that MCM-22
catalysts are significantly more stable than SPA. The data also show that
spent-
regenerated and multiple-regenerated MCM-22 catalysts have similar
performance as fresh MCM-22.
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Table 6
Comparison of Catalyst Activity and Stability
Binder-free Alumina- Regen Regen Multiple
Catalyst SPA bound Steamed regen.
MCM-22 MCM-22 MCM-22 MCM-22 MCM-22
Example No. 1 2 3 4 5 6
1 st-Order Rate 4.3 40.3 20.9 20.7 14.8 20.6
Constant
Relative Activity 1.0 9.4 4.9 4.8 3.4 4.8
Deact. Rate
(% Conv. 0.90 0.07 0.02 0.03 0.60 0.04
Dro /da /WFiSV)
Example 8 Propylene Oligomerization in with a Spent-Regenerated MCM-22
[0044] A Spent-Regenerated MCM-22 catalyst (SiO2:A1203 25:1) in the form
of a chopped 1.5 mm extrudate 1.5 mm long, alpha value 330, was evaluated for
propylene oligomerization in a 9 mm outside diameter downflow stainless steel
three-zone reactor. Conditions used included a nominal temperature of 170 C,
pressure from a nominal 5270-6550 kPaa (750-950 psig)and a WHSV of 8.0
hr-1. The propylene was first flowed over an alumina guard bed to sorb
impurities.
[0045] Representative data are shown in Table 7 below .
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Table 7
Propylene Oligomerization
Sam le No. 1 2 3 4 5 6
Time on Stream, 31 55 79 103 151 175
hours
Conditions
Tem , C 169 169 169 170 170 170
Press, kPag
Flow Rate, ml.hr 5.43 5.43 5.43 5.43 5.43 5.43
WHSV, hr 8.0 8.0 8.0 8.0 8.0 8.0
Conversion, % 97.83 97.54 96.32 95.85 92.50 90.41
Selectivity,%
C5-C12 79.27 85.48 88.63 89.94 93.01 94.83
C3 0.29 0.48 0.35 0.30 0.27 0.22
C4=, C4== 0.09 0.12 0.11 0.12 0.10 0.10
C4 0.11 0.16 0.13 0.10 0.07 0.06
C5 0.46 0.55 0.50 0.47 0.39 0.34
C6 3.67 3.76 3.98 4.44 5.63 6.59
C7-C8 5.45 6.23 6.06 6.05 5.95 5.80
C9 29.14 34.66 37.80 41.71 45.55 47.77
C10-C11 11.44 11.54 11.70 11.22 11.08 10.08
C12 29.10 28.75 28.60 26.04 24.41 23.44
C15 10.20 8.63 6.71 6.10 4.69 3.66
C16+ 10.04 5.13 4.08 3.45 1.86 1.13
Total 100.00 100.00 100.00 100.00 100.00 100.00
Me/C6 1.01 1.02 1.01 1.00 1.00 0.99
Me/C9 2.47 2.47 2.48 2.48 2.48 2.48
Me/C12 2.55 2.54 2.54 2.61 2.62 2.65
[0046] The results reported in Table 7 show that the catalyst retains its
activity
and selectivity for propylene conversion over extended periods of time with
good selectivity to products below C12 and moderately good selectivity to
gasoline boiling range products no higher than C12.
Exam lp e 9 Butene Oligomerization/Alkylation over MCM-22
[0047] A C4 feed comprising isomeric butenes and isobutane was reacted over
a regenerated MCM-22 catalyst in a laboratory scale reactor unit at 3790 kPag
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(550 psig) at temperatures from 124 to 236 C (256 to 456 F), using one, two
or
three reactors in sequence under isothermal (I) or adiabatic (A) conditions.
The
feed composition was as shown in Table 8 below. The three reactors contained,
respectively, 4.84 g, 17.74 g. and 27.42 g of the catalyst.
Table 8
C4 Feed Composition
1-butene 10%
cis 2-butene 9%
trans-2-butene 9%
isobutene 11%
isobutane balance
bu lmerca tan 15 m
water 250p m
[0048] The results are shown in Table 9 below. Selectivities are reported
consistently on a C8+ basis.
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Table 9
C4 Olefin/Paraffin Conversion
Days on Stream 5.2 7.7 8.9 11.5 15.8 20.0 21.6 25.8
Operation I A A A I A A A
LHSV 15.6 15.6 15.6 15.6 3.74 1.65 1.65 1.65
Reactor 1
Inlet T(C) 124 128 131 139 140 147 160 186
Outlet T(C) 129 136 138 144 146 161 178 209
Reactor 2
Inlet T (C) N/A N/A N/A N/A 144 156 177 207
Outlet T(C) N/A N/A N/A N/A 156 171 189 217
Reactor 3
Inlet T (C) N/A N/A N/A N/A N/A 173 186 213
Outlet T (C) N/A N/A N/A N/A N/A 191 205 236
Conversions
Overall olefin 17% 23% 25% 27% 56% 64% 80% 90%
(C4) convsn
1+iso 31% 40% 43% 45% 87% 93% 95% 97%
conversion
cis 2 conversion 4% 7% 8% 11% 40% 56% 73% 87%
trans 2 0% 1% 1% 1% 2% 8% 50% 78%
conversion
iso-butane 0% 0% 0% 0% 0% 0% 13% 40%
conversion
Product comp
(wt%)
1+ Iso 14.1 12.6 12.0 11.4 2.7 1.5 1.2 0.7
Cis 2 9.0 8.7 8.6 8.3 5.7 4.2 2.7 1.2
Trans 2 9.5 9.3 9.4 9.2 9.3 8.8 5.1 2.1
C8s 7.1 8.4 9.0 9.2 12.6 11.4 12.7 23.5
C12+ 0.6 1.0 1.3 1.8 3.6 4.1 5.8 11.0
C16+ 0.03 0.1 0.1 0.1 1.6 2.2 5.3 4.3
C8 Selectivity 92% 88% 86% 83% 71% 65% 53% 61%
C12 Selectivity 8% 11% 12% 16% 20% 23% 24% 28%
C16 Selectivity 0% 1% 1% 1% 9% 12% 22% 11%
C12-C13 (wt%) 0.60 1.01 1.26 1.76 3.65 4.09 4.67 11.17
C14-C15 (wt%) - - 0.00 0.00 0.12 0.04 0.04 0.04
C 16+ (wt. %) 0.04 0.1 0.1 0.1 1.7 2.1 2.3 4.1
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[0049] The results reported in Table 9 show that C8 selectivity increases at
the
highest olefin conversion levels with isobutane conversion increasing at a
somewhat lower level indicating that alkylation of butene with isobutane is
occurring at such levels.
Example 10 Comparison with SPA Polymerization Product
[0050] A refinery FCC gas stream containing olefins up to C4 was polymerized
over an SPA catalyst to produce a gasoline boiling range product which was
then
hydrotreated at 65 C (150 F) and 93 C (200 F) over a Pt/Pd hydrotreating
catalyst at 2410 kPag (3560 psig), 2 hr-1 WHSV, 5:1 hydrogen/liquid ratio, to
saturate olefins. The hydrotreated product was then analyzed for composition
and the octane ratings (RON, MON) were determined by engine test. The same
olefinic stream was also polymerized over an MCM-22 catalyst in a unit of the
same configuration and the same determinations made. The simulated distilla-
tion data (ASTM D 2887) are given in Table 10 below for the polymerized
products and the hydrotreated (93 C) products. The results of the octane
testing
are shown in Table 11 below. The hydrotreating technique was used to
demonstrate more clearly the extent of branching in the two polymerization
products: hydrotreatment saturates the olefins so that their contribution to
product octane is eliminated with the differences remaining then being
attributable to the branching of the paraffin chains, the product being
essentially
free of aromatics.
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Table 10
Simulated Distillation, C
SimDis, sPA MCM-22 Saturation Saturation Product
Pct. Off Product Product Product of sPA of MCM-22 Feed
Feed (93 C) (93 C)
0.5 -0.1 -2.4 16.4 -4.2
55.7 70.8 79.0 85.2
88.3 95.8 92.9 99.2
101.2 111.7 104.8 109.7
113.9 122.7 112.6 114.9
125.4 125.4 130.6 124.4
143.9 139.8 148.0 139.1
161.8 158.6 159.4 159.2
176.4 182.4 174.0 184.0
199.6 225.7 200.1 226.8
99.5 274.8 348.6 329.7 343.0
Table 11
SPA Product Octane Comparison
Sp. API Br2 No RON MON (R+M)/2
Grav.
MCM-22 0.7373 60.4 106.3 96.8 82.7 89.75
Product
Hydrotreated 0.7446 58.5 63.5 92.9 83.2 88.05
65 C
Hydrotreated 0.7426 59 49.3 90 82.8 86.4
93 C
SPA Product 0.7296 62.4 113.7 96.2 82.2 89.2
Hydrotreated 0.7315 61.9 58.9 89.1 82.3 85.7
65 C
Hydrotreated 0.7322 61.8 51.5 86.2 81.5 83.85
93 C
[0051 ] The results in Table 11 show that the product from the MCM-22
polymerization initially has octane numbers which are slightly higher than
those
of the SPA product but upon hydrogenation to remove olefins, a moderate
octane credit relative to the SPA product is consistently maintained,
indicating
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that the MCM-22 product is more highly branched than the product from the
SPA polymerization. In addition, the bromine number is less sensitive.
Example 11 Reaction Product
[0052] A feed containing iso-butane and mixed butene isomers was passed into
a three-reactor unit which could be operated under isothermal or adiabatic
conditions containing a supported MCM-22 zeolite catalyst. The feed
composition is given in Table 12 below. In successive runs, one, two or three
of
the reactors were used to simulate the operation of a chamber type unit.
Reaction temperatures (first reactor inlet) varied from 124 C (256 F) to 186 C
(367 F). The product compositions are given in Table 13 below together with
the conditions for present in the unit.
[0053] The results indicate that at more severe conditions not only is
polymerization taking place but also a degree of isobutene alkylation, as
evidenced by the conversion of the iso-butane towards the end of the run
period
with three series reactors.
Table 12
Mixed C4 Feed
Isobutane wt. pct. 60.2
1-C4= plus iso C4=, wt. pct. 21Ø
Cis 2-C4=, wt. pct. 9.4
Trans 2-C4=, wt. pct. 9.4
Water, ppmw 250.0
Butyl mercaptans, ppmw 15.0
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Table 13
Butane/Butene Conversion
Days on Stream 5.2 7.7 8.9 15.8 20.0 21.6 25.8
LHSV (hr-1) 15.6 15.6 15.6 3.74 1.65 1.65 1.65
Pressure (kPag) 3792 3792 3792 3792 3792 3792 3792
Product Compositions
1C4= + isoC4= 14.1 12.6 12.0 2.7 1.5 1.2 0.7
Cis 2-C4= 9.0 8.7 8.6 5.7 4.2 2.7 1.2
Trans C4= 9.5 9.3 9.4 9.3 8.8 5.1 2.1
C8 7.1 8.4 9.0 12.6 11.4 12.7 23.5
C12+ 0.6 1.0 1.3 3.6 4.1 5.8 11.0
C16+ 0.03 0.1 0.1 1.6 2.2 5.3 4.3
Conditions
Reactor 1
Inlet Tem , C 124 128 131 140 147 160 186
Outlet Tem , C 129 136 138 146 161 178 209
R10T, C 4 7 7 6 14 18 23
Reactor 2
Inlet Tem , C 144 156 177 207
Outlet Tem , C 156 171 189 217
R2 AT, C 12 15 12 10
Reactor 3
Inlet Tem , C 173 186 213
Outlet Temp, C 191 205 236
R3 AT, C 18 19 23
Av. Bed Tem , C 127 133 136 150 174 188 216
Overall OT, C 4 7 7 18 47 51 56
Isothermal/Adiabatic I A A I A A A
Steady State N Y Y N Y Y Y
Conversions
Overall C4 olefin 17 23 25 56 64 80 90
1C4= + isoC4= 31 40 43 87 93 95 97
Cis 2C4= 4 7 8 40 56 73 87
Trans 2C4= 0 1 1 2 8 50 78
Isobutane 0 0 0 0 0 13 40
C8 Selectivity (C8+ basis) 92 88 86 71 65 53 61
C12 Selectivity (C8+ basis) 8 11 12 20 23 24 28
C16 Selectivity (C8+ basis) 0 1 1 9 12 22 11
C12-C13, wt. ct. 0.60 1.01 1.26 3.65 4.09 4.67 11.17
C14-C15, wt. pct. - - 0.00 0.12 0.04 0.04 0.04
C16+, wt. pct. 0.04 0.1 0.1 1.7 2.1 2.3 4.1
