BACKGROUND
[0001] When oligomerizing light olefins within a refinery, there is frequently a desire
to have the flexibility to make high octane gasoline, high cetane diesel, or combination
of both. However, catalysts that make high octane gasoline typically make product
that is highly branched and within the gasoline boiling point range. This product
is very undesirable for diesel. In addition, catalysts that make high cetane diesel
typically make product that is more linear and in the distillate boiling point range.
This results in less and poorer quality gasoline due to the more linear nature of
the product which has a lower octane value.
[0002] The oligomerization of butenes is often associated with a desire to make a high yield
of high quality gasoline product. There is typically a limit as to what can be achieved
when oligomerizing butenes. When oligomerizing butenes, dimerization is desired to
obtain gasoline range material. However, trimerization and higher oligomerization
can occur which can produce material heavier than gasoline such as diesel. Efforts
to produce diesel by oligomerization have failed to provide high yields except through
multiple passes.
[0003] When oligomerizing olefins from a fluid catalytic cracking (FCC) unit, there is often
the desire to maintain a liquid phase within the oligomerization reactors. A liquid
phase helps with catalyst stability by acting as a solvent to wash the catalyst of
heavier species produced. In addition, the liquid phase provides a higher concentration
of olefins to the catalyst surface to achieve a higher catalyst activity. Typically,
this liquid phase in the reactor is maintained by hydrogenating some of the heavy
olefinic product and recycling this paraffinic product to the reactor inlet.
[0004] To maximize propylene produced by the FCC unit, refiners may contemplate oligomerizing
FCC olefins to make heavier oligomers and recycling heavier oligomers to the FCC unit.
However, some heavy oligomers may be resistant to cracking down to propylene.
[0005] The products of olefin oligomerization are usually mixtures of, for example, olefin
dimers, trimers, and higher oligomers. Further, each olefin oligomer is itself usually
a mixture of isomers, both skeletal and in double bond location. Highly branched isomers
are less reactive than linear or lightly branched materials in many of the downstream
reactions for which oligomers are used as feedstocks. This is also true of isomers
in which access to the double bond is sterically hindered. Olefin types of the oligomers
can be denominated according to the degree of substitution of the double bond, as
follows:
Table 1
Olefin Type |
Structure |
Description |
I |
R-HC=CH2 |
Monosubstituted |
II |
R-HC=CH-R |
Disubstituted |
III |
RRC=CH2 |
Disubstituted |
IV |
RRC=CHR |
Trisubstituted |
V |
RRC=CRR |
Tetrasubstituted |
wherein R represents an alkyl group, each R being the same or different. Type I compounds
are sometimes described as α- or vinyl olefins and Type III as vinylidene olefins.
Type IV is sometimes subdivided to provide a Type IVA, in which access to the double
bond is less hindered, and Type IVB where it is more hindered.
[0006] US2005/0121361 relates to a process for conversion of a hydrocarbon feedstock comprising for the
most part a heavy fraction, in particular a vacuum distillate fraction, for the combined
production of gasoline and propylene.
US 2012/0029255 relates to a process for conversion of a heavy hydrocarbon feed, having high flexibility
regarding the production of gasoline and propylene.
SUMMARY OF THE INVENTION
[0007] We have discovered that distillate cracks to propylene more readily than VGO. Additionally,
less branched hydrocarbons crack to propylene more readily than more branched hydrocarbons.
Consequently, oligomerate product can be fed or recycled to an FCC unit to increase
production of propylene.
[0008] An object of the invention, as disclosed in claim 1, is to produce oligomers that
can be readily cracked down to propylene molecules.
BRIEF DESCRIPTION OF THE DRAWINGS
[0009]
FIG. 1 is a schematic drawing of the present invention.
FIG. 2 is an alternative schematic drawing of the present invention.
FIG. 3 is a plot of C8-C11 olefin selectivity versus normal butene conversion.
FIG. 4 is a plot of C12+ olefin selectivity versus normal butene conversion.
FIG. 5 is a plot of reactant conversion versus total butene conversion.
FIG. 6 is a plot of normal butene conversion versus reactor temperature.
FIGS. 7 and 8 are plots of butene conversion versus total butene conversion.
FIG. 9 is a plot of selectivity versus maximum reactor bed temperature.
FIGS. 10-12 are bar graphs of conversion and yield for three different catalysts.
FIG. 13 is a plot of C3 olefin yield versus VGO conversion.
DEFINITIONS
[0010] As used herein, the term "stream" can include various hydrocarbon molecules and other
substances. Moreover, the term "stream comprising Cx hydrocarbons" or "stream comprising
Cx olefins" can include a stream comprising hydrocarbon or olefin molecules, respectively,
with "x" number of carbon atoms, suitably a stream with a majority of hydrocarbons
or olefins, respectively, with "x" number of carbon atoms and preferably a stream
with at least 75 wt% hydrocarbons or olefin molecules, respectively, with "x" number
of carbon atoms. Moreover, the term "stream comprising Cx+ hydrocarbons" or "stream
comprising Cx+ olefins" can include a stream comprising a majority of hydrocarbon
or olefin molecules, respectively, with more than or equal to "x" carbon atoms and
suitably less than 10 wt% and preferably less than 1 wt% hydrocarbon or olefin molecules,
respectively, with x-1 carbon atoms. Lastly, the term "Cx- stream" can include a stream
comprising a majority of hydrocarbon or olefin molecules, respectively, with less
than or equal to "x" carbon atoms and suitably less than 10 wt% and preferably less
than 1 wt% hydrocarbon or olefin molecules, respectively, with x+1 carbon atoms.
[0011] As used herein, the term "zone" can refer to an area including one or more equipment
items and/or one or more sub-zones. Equipment items can include one or more reactors
or reactor vessels, heaters, exchangers, pipes, pumps, compressors, controllers and
columns. Additionally, an equipment item, such as a reactor, dryer, or vessel, can
further include one or more zones or sub-zones.
[0012] As used herein, the term "substantially" can mean an amount of at least generally
70%, preferably 80%, and optimally 90%, by weight, of a compound or class of compounds
in a stream.
[0013] As used herein, the term "gasoline" can include hydrocarbons having a boiling point
temperature in the range of 25° to 200°C at atmospheric pressure.
[0014] As used herein, the term "diesel" or "distillate" includes hydrocarbons having a
boiling point temperature in the range of 200° to 400°C.
[0015] As used herein, the term "vacuum gas oil" (VGO) can include hydrocarbons having a
boiling temperature in the range of from 343° to 552°C.
[0016] As used herein, the term "vapor" can mean a gas or a dispersion that may include
or consist of one or more hydrocarbons.
[0017] As used herein, the term "overhead stream" can mean a stream withdrawn at or near
a top of a vessel, such as a column.
[0018] As used herein, the term "bottom stream" can mean a stream withdrawn at or near a
bottom of a vessel, such as a column.
[0019] As depicted, process flow lines in the figures can be referred to interchangeably
as, e.g., lines, pipes, feeds, gases, products, discharges, parts, portions, or streams.
[0020] As used herein, "bypassing" with respect to a vessel or zone means that a stream
does not pass through the zone or vessel bypassed although it may pass through a vessel
or zone that is not designated as bypassed.
[0021] The term "communication" means that material flow is operatively permitted between
enumerated components.
[0022] The term "downstream communication" means that at least a portion of material flowing
to the subject in downstream communication may operatively flow from the object with
which it communicates.
[0023] The term "upstream communication" means that at least a portion of the material flowing
from the subject in upstream communication may operatively flow to the object with
which it communicates.
[0024] The term "direct communication" means that flow from the upstream component enters
the downstream component without undergoing a compositional change due to physical
fractionation or chemical conversion.
[0025] The term "column" means a distillation column or columns for separating one or more
components of different volatilities. Unless otherwise indicated, each column includes
a condenser on an overhead of the column to condense and reflux a portion of an overhead
stream back to the top of the column and a reboiler at a bottom of the column to vaporize
and send a portion of a bottom stream back to the bottom of the column. Feeds to the
columns may be preheated. The top pressure is the pressure of the overhead vapor at
the outlet of the column. The bottom temperature is the liquid bottom outlet temperature.
Overhead lines and bottom lines refer to the net lines from the column downstream
of the reflux or reboil to the column.
[0026] As used herein, the term "boiling point temperature" means atmospheric equivalent
boiling point (AEBP) as calculated from the observed boiling temperature and the distillation
pressure, as calculated using the equations furnished in ASTM D1160 appendix A7 entitled
"Practice for Converting Observed Vapor Temperatures to Atmospheric Equivalent Temperatures".
[0027] As used herein, "taking a stream from" means that some or all of the original stream
is taken.
DETAILED DESCRIPTION
[0028] The present invention is a process that can be used in a first mode to primarily
make gasoline, in a second mode to primarily make diesel and in a third mode to make
primarily propylene. Gasoline, diesel and propylene are produced in all three modes,
but each mode maximizes the primary product intended. The process may be described
with reference to four components shown in FIG. 1: a fluid catalytic cracking (FCC)
zone 20, an FCC recovery zone 100, a purification zone 110, an oligomerization zone
130, and an oligomerization recovery zone 200. Many configurations of the present
invention are possible, but specific embodiments are presented herein by way of example.
All other possible embodiments for carrying out the present invention are considered
within the scope of the present invention.
[0029] The fluid catalytic cracking zone 20 may comprise a first FCC reactor 22, a regenerator
vessel 30, and an optional second FCC reactor 70.
[0030] A conventional FCC feedstock and higher boiling hydrocarbon feedstock are a suitable
FCC hydrocarbon feed 24 to the first FCC reactor. The most common of such conventional
feedstocks is a VGO. Higher boiling hydrocarbon feedstocks to which this invention
may be applied include heavy bottom from crude oil, heavy bitumen crude oil, shale
oil, tar sand extract, deasphalted residue, products from coal liquefaction, atmospheric
and vacuum reduced crudes and mixtures thereof. The FCC feed 24 may include a recycle
stream 280 to be described later.
[0031] The first FCC reactor 22 may include a first reactor riser 26 and a first reactor
vessel 28. A regenerator catalyst pipe 32 delivers regenerated catalyst from the regenerator
vessel 30 to the reactor riser 26. A fluidization medium such as steam from a distributor
34 urges a stream of regenerated catalyst upwardly through the first reactor riser
26. At least one feed distributor injects the first hydrocarbon feed in a first hydrocarbon
feed line 24, preferably with an inert atomizing gas such as steam, across the flowing
stream of catalyst particles to distribute hydrocarbon feed to the first reactor riser
26. Upon contacting the hydrocarbon feed with catalyst in the first reactor riser
26 the heavier hydrocarbon feed cracks to produce lighter gaseous cracked products
while coke is deposited on the catalyst particles to produce spent catalyst.
[0032] The resulting mixture of gaseous product hydrocarbons and spent catalyst continues
upwardly through the first reactor riser 26 and are received in the first reactor
vessel 28 in which the spent catalyst and gaseous product are separated. Disengaging
arms discharge the mixture of gas and catalyst from a top of the first reactor riser
26 through outlet ports 36 into a disengaging vessel 38 that effects partial separation
of gases from the catalyst. A transport conduit carries the hydrocarbon vapors, stripping
media and entrained catalyst to one or more cyclones 42 in the first reactor vessel
28 which separates spent catalyst from the hydrocarbon gaseous product stream. Gas
conduits deliver separated hydrocarbon cracked gaseous streams from the cyclones 42
to a collection plenum 44 for passage of a cracked product stream to a first cracked
product line 46 via an outlet nozzle and eventually into the FCC recovery zone 100
for product recovery.
[0033] Diplegs discharge catalyst from the cyclones 42 into a lower bed in the first reactor
vessel 28. The catalyst with adsorbed or entrained hydrocarbons may eventually pass
from the lower bed into a stripping section 48 across ports defined in a wall of the
disengaging vessel 38. Catalyst separated in the disengaging vessel 38 may pass directly
into the stripping section 48 via a bed. A fluidizing distributor delivers inert fluidizing
gas, typically steam, to the stripping section 48. The stripping section 48 contains
baffles or other equipment to promote contacting between a stripping gas and the catalyst.
The stripped spent catalyst leaves the stripping section 48 of the disengaging vessel
38 of the first reactor vessel 28 stripped of hydrocarbons. A first portion of the
spent catalyst, preferably stripped, leaves the disengaging vessel 38 of the first
reactor vessel 28 through a spent catalyst conduit 50 and passes into the regenerator
vessel 30. A second portion of the spent catalyst may be recirculated in recycle conduit
52 from the disengaging vessel 38 back to a base of the first riser 26 at a rate regulated
by a slide valve to recontact the feed without undergoing regeneration.
[0034] The first riser 26 can operate at any suitable temperature, and typically operates
at a temperature of 150° to 580°C at the riser outlet 36. The pressure of the first
riser is from 69 to 517 kPa (gauge) (10 to 75 psig) but typically less than 275 kPa
(gauge) (40 psig). The catalyst-to-oil ratio, based on the weight of catalyst and
feed hydrocarbons entering the riser, may range up to 30:1 but is typically between
4:1 and 10:1. Steam may be passed into the first reactor riser 26 and first reactor
vessel 28 at a rate between 2 and 7 wt% for maximum gasoline production and 10 to
15 wt% for maximum light olefin production. The average residence time of catalyst
in the riser may be less than 5 seconds.
[0035] The catalyst in the first reactor 22 can be a single catalyst or a mixture of different
catalysts. Usually, the catalyst includes two catalysts, namely a first FCC catalyst,
and a second FCC catalyst. Such a catalyst mixture is disclosed in, e.g.,
US 7,312,370 B2. Generally, the first FCC catalyst may include any of the well-known catalysts that
are used in the art of FCC. Preferably, the first FCC catalyst includes a large pore
zeolite, such as a Y-type zeolite, an active alumina material, a binder material,
including either silica or alumina, and an inert filler such as kaolin.
[0036] Typically, the zeolites appropriate for the first FCC catalyst have a large average
pore size, usually with openings of greater than 0.7 nm in effective diameter defined
by greater than 10, and typically 12, member rings. Suitable large pore zeolite components
may include synthetic zeolites such as X and Y zeolites, mordenite and faujasite.
A portion of the first FCC catalyst, such as the zeolite portion, can have any suitable
amount of a rare earth metal or rare earth metal oxide.
[0037] The second FCC catalyst may include a medium or smaller pore zeolite catalyst, such
as exemplified by at least one of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48,
and other similar materials. Other suitable medium or smaller pore zeolites include
ferrierite, and erionite. Preferably, the second component has the medium or smaller
pore zeolite dispersed on a matrix including a binder material such as silica or alumina
and an inert filler material such as kaolin. These catalysts may have a crystalline
zeolite content of 10 to 50 wt% or more, and a matrix material content of 50 to 90
wt%. Catalysts containing at least 40 wt% crystalline zeolite material are typical,
and those with greater crystalline zeolite content may be used. Generally, medium
and smaller pore zeolites are characterized by having an effective pore opening diameter
of less than or equal to 0.7 nm and rings of 10 or fewer members. Preferably, the
second FCC catalyst component is an MFI zeolite having a silicon-to-aluminum ratio
greater than 15. In one exemplary embodiment, the silicon-to-aluminum ratio can be
15 to 35.
[0038] The total catalyst mixture in the first reactor 22 may contain 1 to 25 wt% of the
second FCC catalyst, including a medium to small pore crystalline zeolite, with greater
than or equal to 7 wt% of the second FCC catalyst being preferred. When the second
FCC catalyst contains 40 wt% crystalline zeolite with the balance being a binder material,
an inert filler, such as kaolin, and optionally an active alumina component, the catalyst
mixture may contain 0.4 to 10 wt% of the medium to small pore crystalline zeolite
with a preferred content of at least 2.8 wt%. The first FCC catalyst may comprise
the balance of the catalyst composition. The high concentration of the medium or smaller
pore zeolite as the second FCC catalyst of the catalyst mixture can improve selectivity
to light olefins. In one exemplary embodiment, the second FCC catalyst can be a ZSM-5
zeolite and the catalyst mixture can include 0.4 to 10 wt% ZSM-5 zeolite excluding
any other components, such as binder and/or filler.
[0039] The regenerator vessel 30 is in downstream communication with the first reactor vessel
28. In the regenerator vessel 30, coke is combusted from the portion of spent catalyst
delivered to the regenerator vessel 30 by contact with an oxygen-containing gas such
as air to regenerate the catalyst. The spent catalyst conduit 50 feeds spent catalyst
to the regenerator vessel 30. The spent catalyst from the first reactor vessel 28
usually contains carbon in an amount of from 0.2 to 2 wt%, which is present in the
form of coke. An oxygen-containing combustion gas, typically air, enters the lower
chamber 54 of the regenerator vessel 30 through a conduit and is distributed by a
distributor 56. As the combustion gas enters the lower chamber 54, it contacts spent
catalyst entering from spent catalyst conduit 50 and lifts the catalyst at a superficial
velocity of combustion gas in the lower chamber 54 of perhaps at least 1.1 m/s (3.5
ft/s) under fast fluidized flow conditions. In an embodiment, the lower chamber 54
may have a catalyst density of from 48 to 320 kg/m
3 (3 to 20 lb/ft
3) and a superficial gas velocity of 1.1 to 2.2 m/s (3.5 to 7 ft/s). The oxygen in
the combustion gas contacts the spent catalyst and combusts carbonaceous deposits
from the catalyst to at least partially regenerate the catalyst and generate flue
gas.
[0040] The mixture of catalyst and combustion gas in the lower chamber 54 ascends through
a frustoconical transition section to the transport, riser section of the lower chamber
54. The mixture of catalyst particles and flue gas is discharged from an upper portion
of the riser section into the upper chamber 60. Substantially completely or partially
regenerated catalyst may exit the top of the transport, riser section. Discharge is
effected through a disengaging device 58 that separates a majority of the regenerated
catalyst from the flue gas. The catalyst and gas exit downwardly from the disengaging
device 58. The sudden loss of momentum and downward flow reversal cause a majority
of the heavier catalyst to fall to the dense catalyst bed and the lighter flue gas
and a minor portion of the catalyst still entrained therein to ascend upwardly in
the upper chamber 60. Cyclones 62 further separate catalyst from ascending gas and
deposits catalyst through dip legs into a dense catalyst bed. Flue gas exits the cyclones
62 through a gas conduit and collects in a plenum 64 for passage to an outlet nozzle
of regenerator vessel 30. Catalyst densities in the dense catalyst bed are typically
kept within a range of from 640 to 960 kg/m
3 (40 to 60 lb/ft
3).
[0041] The regenerator vessel 30 typically has a temperature of 594° to 704°C (1100° to
1300°F) in the lower chamber 54 and 649° to 760°C (1200° to 1400°F) in the upper chamber
60. Regenerated catalyst from dense catalyst bed is transported through regenerated
catalyst pipe 32 from the regenerator vessel 30 back to the first reactor riser 26
through the control valve where it again contacts the first feed in line 24 as the
FCC process continues. The first cracked product stream in the first cracked product
line 46 from the first reactor 22, relatively free of catalyst particles and including
the stripping fluid, exit the first reactor vessel 28 through an outlet nozzle. The
first cracked products stream in the line 46 may be subjected to additional treatment
to remove fine catalyst particles or to further prepare the stream prior to fractionation.
The line 46 transfers the first cracked products stream to the FCC recovery zone 100,
which is in downstream communication with the FCC zone 20. The FCC recovery zone 100
typically includes a main fractionation column and a gas recovery section. The FCC
recovery zone can include many fractionation columns and other separation equipment.
The FCC recovery zone 100 can recover a propylene product stream in propylene line
102, a gasoline stream in gasoline line 104, a light olefin stream in light olefin
line 106 and an LCO stream in LCO line 107 among others from the cracked product stream
in first cracked product line 46. The light olefin stream in light olefin line 106
comprises an oligomerization feed stream having C
4 hydrocarbons including C
4 olefins and perhaps having C
5 hydrocarbons including C
5 olefins.
[0042] An FCC recycle stream in recycle line 280 delivers an FCC recycle stream to the FCC
zone 20. The FCC recycle stream is directed into a first FCC recycle line 202 with
the control valve 202' thereon opened. In an aspect, the FCC recycle stream may be
directed into an optional second FCC recycle line 204 with the control valve 204'
thereon opened. The first FCC recycle line 202 delivers the first FCC recycle stream
to the first FCC reactor 22 in an aspect to the riser 26 at an elevation above the
first hydrocarbon feed in line 24. The second FCC recycle line 204 delivers the second
FCC recycle stream to the second FCC reactor 70. Typically, both control valves 202'
and 204' will not be opened at the same time, so the FCC recycle stream goes through
only one of the first FCC recycle line 202 and the second FCC recycle line 204. However,
feed through both is contemplated.
[0043] The second FCC recycle stream may be fed to the second FCC reactor 70 in the second
FCC recycle line 204 via feed distributor 72. The second FCC reactor 70 may include
a second riser 74. The second FCC recycle stream is contacted with catalyst delivered
to the second riser 74 by a catalyst return pipe 76 to produce cracked upgraded products.
The catalyst may be fluidized by inert gas such as steam from distributor 78. Generally,
the second FCC reactor 70 may operate under conditions to convert the second FCC recycle
stream to second cracked products such as ethylene and propylene. A second reactor
vessel 80 is in downstream communication with the second riser 74 for receiving second
cracked products and catalyst from the second riser. The mixture of gaseous, second
cracked product hydrocarbons and catalyst continues upwardly through the second reactor
riser 74 and is received in the second reactor vessel 80 in which the catalyst and
gaseous, second cracked products are separated. A pair of disengaging arms may tangentially
and horizontally discharge the mixture of gas and catalyst from a top of the second
reactor riser 74 through one or more outlet ports 82 (only one is shown) into the
second reactor vessel 80 that effects partial separation of gases from the catalyst.
The catalyst can drop to a dense catalyst bed within the second reactor vessel 80.
Cyclones 84 in the second reactor vessel 80 may further separate catalyst from second
cracked products. Afterwards, a second cracked product stream can be removed from
the second FCC reactor 70 through an outlet in a second cracked product line 86 in
downstream communication with the second reactor riser 74. The second cracked product
stream in line 86 is fed to the FCC recovery zone 100, preferably separately from
the first cracked products to undergo separation and recovery of ethylene and propylene.
Separated catalyst may be recycled via a recycle catalyst pipe 76 from the second
reactor vessel 80 regulated by a control valve back to the second reactor riser 74
to be contacted with the second FCC recycle stream.
[0044] In some embodiments, the second FCC reactor 70 can contain a mixture of the first
and second FCC catalysts as described above for the first FCC reactor 22. In one preferred
embodiment, the second FCC reactor 70 can contain less than 20 wt%, preferably less
than 5 wt% of the first FCC catalyst and at least 20 wt% of the second FCC catalyst.
In another preferred embodiment, the second FCC reactor 70 can contain only the second
FCC catalyst, preferably a ZSM-5 zeolite.
[0045] The second FCC reactor 70 is in downstream communication with the regenerator vessel
30 and receives regenerated catalyst therefrom in line 88. In an embodiment, the first
FCC reactor 22 and the second FCC reactor 70 both share the same regenerator vessel
30. Line 90 carries spent catalyst from the second reactor vessel 80 to the lower
chamber 54 of the regenerator vessel 30. The catalyst regenerator is in downstream
communication with the second FCC reactor 70 via line 90.
[0046] The same catalyst composition may be used in both reactors 22, 70. However, if a
higher proportion of the second FCC catalyst of small to medium pore zeolite is desired
in the second FCC reactor 70 than the first FCC catalyst of large pore zeolite, replacement
catalyst added to the second FCC reactor 70 may comprise a higher proportion of the
second FCC catalyst. Because the second FCC catalyst does not lose activity as quickly
as the first FCC catalyst, less of the second catalyst inventory must be forwarded
to the catalyst regenerator 30 in line 90 from the second reactor vessel 80, but more
catalyst inventory may be recycled to the riser 74 in return conduit 76 without regeneration
to maintain a high level of the second FCC catalyst in the second reactor 70.
[0047] The second reactor riser 74 can operate in any suitable condition, such as a temperature
of 425° to 705°C, preferably a temperature of 550° to 600°C, and a pressure of 140
to 400 kPa, preferably a pressure of 170 to 250 kPa. Typically, the residence time
of the second reactor riser 74 can be less than 3 seconds and preferably is than 1
second. Exemplary risers and operating conditions are disclosed in, e.g.,
US 2008/0035527 A1 and
US 7,261,807 B2.
[0048] Before cracked products can be fed to the oligomerization zone 130, the light olefin
stream in light olefin line 106 may require purification. Many impurities in the light
olefin stream in light olefin line 106 can poison an oligomerization catalyst. Carbon
dioxide and ammonia can attack acid sites on the catalyst. Sulfur containing compounds,
oxygenates, and nitriles can harm oligomerization catalyst. Acetylenes and diolefins
can polymerize and produce gums on the catalyst or equipment. Consequently, the light
olefin stream which comprises the oligomerization feed stream in light olefin line
106 may be purified in an optional purification zone 110.
[0049] The light olefin stream in light olefin line 106 may be introduced into an optional
mercaptan extraction unit 112 to remove mercaptans to lower concentrations. In the
mercaptan extraction unit 112, the light olefin feed may be prewashed in an optional
prewash vessel containing aqueous alkali to convert any hydrogen sulfide to sulfide
salt which is soluble in the aqueous alkaline stream. The light olefin stream, now
depleted of any hydrogen sulfide, is contacted with a more concentrated aqueous alkali
stream in an extractor vessel. Mercaptans in the light olefin stream react with the
alkali to yield mercaptides. An extracted light olefin stream lean in mercaptans passes
overhead from the extraction column and may be mixed with a solvent that removes COS
in route to an optional COS solvent settler. COS is removed with the solvent from
the bottom of the settler, while the overhead light olefin stream may be fed to an
optional water wash vessel to remove remaining alkali and produce a sulfur depleted
light olefin stream in line 114. The mercaptide rich alkali from the extractor vessel
receives an injection of air and a catalyst such as phthalocyanine as it passes from
the extractor vessel to an oxidation vessel for regeneration. Oxidizing the mercaptides
to disulfides using a catalyst regenerates the alkaline solution. A disulfide separator
receives the disulfide rich alkaline from the oxidation vessel. The disulfide separator
vents excess air and decants disulfides from the alkaline solution before the regenerated
alkaline is drained, washed with oil to remove remaining disulfides and returned to
the extractor vessel. Further removal of disulfides from the regenerated alkaline
stream is also contemplated. The disulfides are run through a sand filter and removed
from the process. For more information on mercaptan extraction, reference may be made
to
US 7,326,333 B2.
[0050] In order to prevent polymerization and gumming in the oligomerization reactor that
can inhibit equipment and catalyst performance, it is desired to minimize diolefins
and acetylenes in the light olefin feed in line 114. Diolefin conversion to monoolefin
hydrocarbons may be accomplished by selectively hydrogenating the sulfur depleted
stream with a conventional selective hydrogenation reactor 116. Hydrogen may be added
to the purified light olefin stream in line 118.
[0051] The selective hydrogenation catalyst can comprise an alumina support material preferably
having a total surface area greater than 150 m
2/g, with most of the total pore volume of the catalyst provided by pores with average
diameters of greater than 600 angstroms, and containing surface deposits of 1.0 to
25.0 wt% nickel and 0.1 to 1.0 wt% sulfur such as disclosed in
US 4,695,560. Spheres having a diameter between 0.4 and 6.4 mm (1/64 and 1/4 inch) can be made
by oil dropping a gelled alumina sol. The alumina sol may be formed by digesting aluminum
metal with an aqueous solution of 12 wt% hydrogen chloride to produce an aluminum
chloride sol. The nickel component may be added to the catalyst during the sphere
formation or by immersing calcined alumina spheres in an aqueous solution of a nickel
compound followed by drying, calcining, purging and reducing. The nickel containing
alumina spheres may then be sulfided. A palladium catalyst may also be used as the
selective hydrogenation catalyst.
[0052] The selective hydrogenation process is normally performed at relatively mild hydrogenation
conditions. These conditions will normally result in the hydrocarbons being present
as liquid phase materials. The reactants will normally be maintained under the minimum
pressure sufficient to maintain the reactants as liquid phase hydrocarbons which allow
the hydrogen to dissolve into the light olefin feed. A broad range of suitable operating
pressures therefore extends from 276 (40 psig) to 5516 kPa gauge (800 psig). A relatively
moderate temperature between 25°C (77°F) and 350°C (662°F) should be employed. The
liquid hourly space velocity of the reactants through the selective hydrogenation
catalyst should be above 1.0 hr
-1. Preferably, it is between 5.0 and 35.0 hr
-1. The ratio of hydrogen to diolefinic hydrocarbons may be maintained between 0.75:1
and 1.8:1. The hydrogenation reactor is preferably a cylindrical fixed bed of catalyst
through which the reactants move in a vertical direction.
[0053] A purified light olefin stream depleted of sulfur containing compounds, diolefins
and acetylenes exits the selective hydrogenation reactor 116 in line 120. The optionally
sulfur and diolefin depleted light olefin stream in line 120 may be introduced into
an optional nitrile removal unit (NRU) such as a water wash unit 122 to reduce the
concentration of oxygenates and nitriles in the light olefin stream in line 120. Water
is introduced to the water wash unit in line 124. An oxygenate and nitrile-rich aqueous
stream in line 126 leaves the water wash unit 122 and may be further processed. A
drier may follow the water wash unit 122. Other NRU's may be used in place of the
water wash unit. A NRU can consist of a group of regenerable beds that adsorb the
nitriles and other nitrogen components from the diolefin depleted light olefin stream.
Examples of NRU's can be found in
US 4,831,206,
US 5,120,881 and
US 5,271,835.
[0054] A purified light olefin oligomerization feed stream perhaps depleted of sulfur containing
compounds, diolefins and/or oxygenates and nitriles is provided in oligomerization
feed stream line 128. The light olefin oligomerization feed stream in line 128 may
be obtained from the cracked product stream in lines 46 and/or 86, so it may be in
downstream communication with the FCC zone 20. The oligomerization feed stream need
not be obtained from a cracked FCC product stream but may come from another source.
The selective hydrogenation reactor 116 is in upstream communication with the oligomerization
feed stream line 128. The oligomerization feed stream may comprise C
4 hydrocarbons such as butenes, i.e., C
4 olefins, and butanes. Butenes include normal butenes and isobutene. The oligomerization
feed stream in oligomerization feed stream line 128 may comprise C
5 hydrocarbons such as pentenes, i.e., C
5 olefins, and pentanes. Pentenes include normal pentenes and isopentenes. Typically,
the oligomerization feed stream will comprise 20 to 80 wt% olefins and suitably 40
to 75 wt% olefins. In an aspect, 55 to 75 wt% of the olefins may be butenes and 25
to 45 wt% of the olefins may be pentenes. As much as 10 wt%, suitably 20 wt%, typically
25 wt% and most typically 30 wt% of the oligomerization feed may be C
5 olefins.
[0055] The oligomerization feed line 128 feeds the oligomerization feed stream to an oligomerization
zone 130 which may be in downstream communication with the FCC recovery zone 100.
The oligomerization feed stream in oligomerization feed line 128 may be mixed with
recycle streams from line 226 or 246 prior to entering the oligomerization zone 130
to provide an oligomerization feed stream in an oligomerization feed conduit 132.
An oligomerization reactor zone 140 is in downstream communication with the oligomerization
feed conduit 132.
[0056] In an aspect, an oligomerate return stream in oligomerate return line 231 to be described
hereinafter may be mixed with the oligomerization feed stream in oligomerization feed
conduit 132 in a first mixed oligomerization feed line 133. The oligomerization feed
stream in line 133 may comprise 10 to 50 wt% olefins and suitably 25 to 40 wt% olefins
if the oligomerate return stream from oligomerate return line 231 is mixed with the
oligomerization feed stream. Accordingly, the oligomerization feed stream may comprise
no more than 38 wt% butene and in another aspect, the oligomerization feed stream
may comprise no more than 23 wt% pentene. The oligomerization feed stream to the oligomerization
zone 130 in mixed oligomerization feed conduit 133 may comprise at least 10 wt% butene,
at least 5 wt% pentene and preferably no more than 1 wt% hexene. In a further aspect,
the oligomerization feed stream may comprise no more than 0.1 wt% hexene and no more
than 0.1 wt% propylene. At least 40 wt% of the butene in the oligomerization feed
stream may be normal butene. In an aspect, it may be that no more than 70 wt% of the
oligomerization feed stream is normal butene. At least 40 wt% of the pentene in the
oligomerization feed stream may be normal pentene. In an aspect, no more than 70 wt%
of the oligomerization feed stream in the mixed oligomerization feed conduit 133 may
be normal pentene.
[0057] The oligomerization reactor zone 140 comprises a first oligomerization reactor 138.
The first oligomerization reactor may be preceded by an optional guard bed for removing
catalyst poisons that is not shown. The first oligomerization reactor 138 contains
the oligomerization catalyst. An oligomerization feed stream may be preheated before
entering the first oligomerization reactor 138 in an oligomerization reactor zone
140. The first oligomerization reactor 138 may contain a first catalyst bed 142 of
oligomerization catalyst. The first oligomerization reactor 138 may be an upflow reactor
to provide a uniform feed front through the catalyst bed, but other flow arrangements
are contemplated. In an aspect, the first oligomerization reactor 138 may contain
an additional bed or beds 144 of oligomerization catalyst. C
4 olefins in the oligomerization feed stream oligomerize over the oligomerization catalyst
to provide an oligomerate comprising C
4 olefin dimers and trimers. C
5 olefins that are present in the oligomerization feed stream oligomerize over the
oligomerization catalyst to provide an oligomerate comprising C
5 olefin dimers and trimers and co-oligomerize with C
4 olefins to make C
9 olefins. The oligomerization produces other oligomers with additional carbon numbers.
[0058] We have found that adding C
5 olefins to the feed to the oligomerization reactor reduces oligomerization to heavier,
distillate range material. This is counterintuitive since one may expect heavier C
5 olefins to lead to the formation of more distillate range material. However, when
C
5 olefins dimerize with themselves or co-dimerize with C
4 olefins, the C
9 olefins and C
10 olefins produced do not continue to oligomerize as quickly as C
8 olefins produced from C
4 olefin dimerization. Thus, the amount of net gasoline produced can be increased.
In addition, the resulting C
9 olefins and C
10 olefins in the product have a very high octane value.
[0059] Oligomerization effluent from the first bed 142 may optionally be quenched with a
liquid such as recycled oligomerate before entering the additional bed 144, and/or
oligomerization effluent from the additional bed 144 of oligomerization catalyst may
also be quenched with a liquid such as recycled oligomerate to avoid excessive temperature
rise. The liquid oligomerate may also comprise oligomerized olefins that can react
with the C
4 olefins and C
5 olefins in the feed and other oligomerized olefins if present to make diesel range
olefins. Oligomerized product, also known as oligomerate, exits the first oligomerization
reactor 138 in line 146.
[0060] In an aspect, the oligomerization reactor zone may include one or more additional
oligomerization reactors 150. The oligomerization effluent may be heated and fed to
the optional additional oligomerization reactor 150. It is contemplated that the first
oligomerization reactor 138 and the additional oligomerization reactor 150 may be
operated in a swing bed fashion to take one reactor offline for maintenance or catalyst
regeneration or replacement while the other reactor stays online. In an aspect, the
additional oligomerization reactor 150 may contain a first bed 152 of oligomerization
catalyst. The additional oligomerization reactor 150 may also be an upflow reactor
to provide a uniform feed front through the catalyst bed, but other flow arrangements
are contemplated. In an aspect, the additional oligomerization reactor 150 may contain
an additional bed or beds 154 of oligomerization catalyst. Remaining C
4 olefins in the oligomerization feed stream oligomerize over the oligomerization catalyst
to provide an oligomerate comprising C
4 olefin dimers and trimers. Remaining C
5 olefins oligomerize over the oligomerization catalyst to provide an oligomerate comprising
C
5 olefin dimers and trimers and co-oligomerize with C
4 olefins to make C
9 olefins. Over 90 wt% of the C
4 olefins in the oligomerization feed stream can oligomerize in the oligomerization
reactor zone 140. Over 90 wt% of the C
5 olefins in the oligomerization feed stream can oligomerize in the oligomerization
reactor zone 140. If more than one oligomerization reactor is used, conversion is
achieved over all of the oligomerization reactors 138, 150 in the oligomerization
reactor zone 140.
[0061] Oligomerization effluent from the first bed 152 may be quenched with a liquid such
as recycled oligomerate before entering the additional bed 154, and/or oligomerization
effluent from the additional bed 154 of oligomerization catalyst may also be quenched
with a liquid such as recycled oligomerate to avoid excessive temperature rise. The
recycled oligomerate may also comprise oligomerized olefins that can react with the
C
4 olefins and C
5 olefins in the feed and other oligomerized olefins to increase production of diesel
range olefins.
[0062] An oligomerate conduit 156, in communication with the oligomerization reactor zone
140, withdraws an oligomerate stream from the oligomerization reactor zone 140. The
oligomerate conduit 156 may be in downstream communication with the first oligomerization
reactor 138 and the additional oligomerization reactor 150.
[0063] The oligomerization reactor zone 140 contains an oligomerization catalyst. The oligomerization
catalyst comprises a zeolitic catalyst. The zeolite may comprise between 5 and 95
wt% of the catalyst. The oligomerization catalyst comprises a zeolite with a framework
having a ten-ring pore structure. Examples of zeolites having a ten-ring pore structure
include those comprising TON, MTT, MFI, MEL, AFO, AEL, EUO and FER. The oligomerization
catalyst comprises a zeolite having a ten-ring pore structure may comprise a uni-dimensional
pore structure. A uni-dimensional pore structure indicates zeolites containing non-intersecting
pores that are substantially parallel to one of the axes of the crystal. The pores
preferably extend through the zeolite crystal. The oligomerization catalyst used in
the present invention comprises an MTT zeolite.
[0064] The oligomerization catalyst may be formed by combining the zeolite with a binder,
and then forming the catalyst into pellets. The pellets may optionally be treated
with a phosphoric reagent to create a zeolite having a phosphorous component between
0.5 and 15 wt% of the treated catalyst. The binder is used to confer hardness and
strength on the catalyst. Binders include alumina, aluminum phosphate, silica, silica-alumina,
zirconia, titania and combinations of these metal oxides, and other refractory oxides,
and clays such as montmorillonite, kaolin, palygorskite, smectite and attapulgite.
A preferred binder is an aluminum-based binder, such as alumina, aluminum phosphate,
silica-alumina and clays.
[0065] One of the components of the catalyst binder utilized in the present invention is
alumina. The alumina source may be any of the various hydrous aluminum oxides or alumina
gels such as alpha-alumina monohydrate of the boehmite or pseudo-boehmite structure,
alpha-alumina trihydrate of the gibbsite structure, beta-alumina trihydrate of the
bayerite structure, and the like. A suitable alumina is available from UOP LLC under
the trademark Versal. A preferred alumina is available from Sasol North America Alumina
Product Group under the trademark Catapal. This material is an extremely high purity
alpha-alumina monohydrate (pseudo-boehmite) which after calcination at a high temperature
has been shown to yield a high purity gamma-alumina.
[0066] A suitable oligomerization catalyst is prepared by mixing proportionate volumes of
zeolite and alumina to achieve the desired zeolite-to-alumina ratio. In an embodiment,
5 to 80, typically 10 to 60, suitably 15 to 40 and preferably 20 to 30 wt% MTT zeolite
and the balance alumina powder will provide a suitably supported catalyst. A silica
support is also contemplated.
[0067] Monoprotic acid such as nitric acid or formic acid may be added to the mixture in
aqueous solution to peptize the alumina in the binder. Additional water may be added
to the mixture to provide sufficient wetness to constitute a dough with sufficient
consistency to be extruded or spray dried. Extrusion aids such as cellulose ether
powders can also be added. A preferred extrusion aid is available from The Dow Chemical
Company under the trademark Methocel.
[0068] The paste or dough may be prepared in the form of shaped particulates, with the preferred
method being to extrude the dough through a die having openings therein of desired
size and shape, after which the extruded matter is broken into extrudates of desired
length and dried. A further step of calcination may be employed to give added strength
to the extrudate. Generally, calcination is conducted in a stream of air at a temperature
from 260°C (500°F) to 815°C (1500°F). The MTT catalyst is not selectivated to neutralize
surface acid sites such as with an amine.
[0069] The extruded particles may have any suitable cross-sectional shape, i.e., symmetrical
or asymmetrical, but most often have a symmetrical cross-sectional shape, preferably
a spherical, cylindrical or polylobal shape. The cross-sectional diameter of the particles
may be as small as 40 µm; however, it is usually 0.635 mm (0.25 inch) to 12.7 mm (0.5
inch), preferably 0.79 mm (1/32 inch) to 6.35 mm (0.25 inch), and most preferably
0.06 mm (1/24 inch) to 4.23 mm (1/6 inch).
[0070] In an embodiment, the oligomerization catalyst may be a solid phosphoric acid catalyst
(SPA). The SPA catalyst refers to a solid catalyst that contains as a principal ingredient
an acid of phosphorous such as ortho-, pyro- or tetraphosphoric acid. SPA catalyst
is normally formed by mixing the acid of phosphorous with a siliceous solid carrier
to form a wet paste. This paste may be calcined and then crushed to yield catalyst
particles or the paste may be extruded or pelleted prior to calcining to produce more
uniform catalyst particles. The carrier is preferably a naturally occurring porous
silica-containing material such as kieselguhr, kaolin, infusorial earth and diatomaceous
earth. A minor amount of various additives such as mineral talc, fuller's earth and
iron compounds including iron oxide may be added to the carrier to increase its strength
and hardness. The combination of the carrier and the additives preferably comprises
15-30 wt% of the catalyst, with the remainder being the phosphoric acid. The additive
may comprise 3-20 wt% of the total carrier material. Variations from this composition
such as a lower phosphoric acid content are possible. Further details as to the composition
and production of SPA catalysts may be obtained from
US 3,050,472,
US 3,050,473 and
US 3,132,109. Feed to the oligomerization reactor zone 140 containing SPA catalyst should be kept
dry except in an initial start-up phase.
[0071] The oligomerization reaction conditions in the oligomerization reactors 138, 150
in the oligomerization reactor zone 140 are set to keep the reactant fluids in the
liquid phase. With liquid oligomerate recycle, lower pressures are necessary to maintain
liquid phase. Operating pressures include between 2.1 MPa (300 psia) and 10.5 MPa
(1520 psia), suitably at a pressure between 2.1 MPa (300 psia) and 6.9 MPa (1000 psia)
and preferably at a pressure between 2.8 MPa (400 psia) and 4.1 MPa (600 psia). Lower
pressures may be suitable if the reaction is kept in the liquid phase.
[0072] For the zeolite catalyst, the temperature of the oligomerization reactor zone 140
expressed in terms of a maximum bed temperature is in a range between 150°C and 300°C.
If diesel oligomerate is desired, the maximum bed temperature should between 200°C
and 250°C and preferably between 215° or 225°C and 245°C or between 220° and 240°C.
The space velocity should be between 0.5 and 5 hr
-1.
[0073] For the SPA catalyst, the oligomerization temperature in the oligomerization reactor
zone 140 should be in a range between 100°C and 250°C and suitably between 150°C and
200°C. The liquid hourly space velocity (LHSV) should be between 0.5 and 5 hr
-1.
[0074] Across a single bed of oligomerization catalyst, the exothermic reaction will cause
the temperature to rise. Consequently, the oligomerization reactor should be operated
to allow the temperature at the outlet to be over 25°C greater than the temperature
at the inlet.
[0075] The oligomerization reactor zone 140 with the oligomerization catalyst can be run
in high conversion mode of greater than 95% conversion of feed olefins to produce
a high quality diesel product and gasoline product. Normal butene conversion can exceed
80%. Additionally, normal pentene conversion can exceed 80%.
[0076] We have found that when C
5 olefins are present in the oligomerization feed stream, they dimerize or co-dimerize
with other olefins, but tend to mitigate further oligomerization over the zeolite
with a 10-ring uni-dimensional pore structure. Best mitigation of further oligomerization
occurs when the C
5 olefins comprise between 15 or 30 and 70 wt% and preferably between 20 or 40 and
50 or 60 wt% of the olefins in the oligomerization feed. The oligomerate stream in
oligomerate conduit 156 may comprise less than 80 wt% C
9+ hydrocarbons because C
5 olefins are present in the oligomerization feed at these proportions. Moreover, said
oligomerate may comprise less than 60 wt% C
12+ hydrocarbons when C
5 olefins are present in the oligomerization feed at these proportions. Furthermore,
the net gasoline yield may be at least 40 wt% when C
5 olefins are present in the oligomerization feed.
[0077] If diesel is desired, however, the oligomerization zone with the oligomerization
catalyst can be operated to oligomerize light olefins; i.e., C
4 olefins, to distillate-range material by over 70 wt% yield per pass through the oligomerization
reactor zone 140. In an aspect, at least 70 wt% of the olefins in the oligomerization
feed convert to C
9+ product oligomers boiling above 150°C (302°F) cut point in a single pass through
the oligomerization zone. The C
12+ oligomer from the oligomerization zone boiling above 200°C (392°F) may have a cetane
of at least 30 and preferably at least 40.
[0078] The composition of the oligomerate in oligomerate line 156 may be an olefinic hydrocarbon
composition having C
8 olefins. The olefinic hydrocarbon composition may include gasoline. In an embodiment,
the composition may be moderate in Type 2 disubstituted olefins and high in Type 4
trisubstituted olefins. In an aspect, the oligomerate composition may have a ratio
of Type 2 disubstituted Cg olefins to Type 1 monosubstituted C
8 olefins of greater than 2. In a further aspect, a fraction of Type 2 disubstituted
C
8 olefins in the total Cg olefins in the oligomerate may be no less than 7 and less
than 18 wt%. In an even further aspect, a fraction of Type 4 trisubstituted C
8 olefins in the total C
8 olefins may be no less than 40 wt%. In a still further aspect, an average branch
per C
8 hydrocarbon molecule in the oligomerate may be less than 2. The oligomerate may have
a cetane of greater than 30 and preferably greater than 40. The oligomerate may have
a density of less than 0.83 kg/m
3, preferably less than 0.81 kg/m
3, less than 20 ppmw sulfur or less than 1 vol% aromatics. The oligomerate may have
a ratio of trimethyl pentene to the total C
8 olefins of no more than 50 and preferably no more than 40.
[0079] An oligomerization recovery zone 200 is in downstream communication with the oligomerization
zone 130 and the oligomerate conduit 156. The oligomerate conduit 156 removes the
oligomerate stream from the oligomerization zone 130.
[0080] The oligomerization recovery zone 200 may include a debutanizer column 210 which
separates the oligomerate stream between vapor and liquid into a first vaporous oligomerate
overhead light stream comprising C
4 olefins and hydrocarbons in a first overhead line 212 and a first liquid oligomerate
bottom stream comprising C
5+ olefins and hydrocarbons in a first bottom line 214. When maximum production of
distillate is desired, either to obtain diesel product or to recrack the diesel in
the FCC zone 20 to make more propylene, the overhead pressure in the debutanizer column
210 may be between 300 and 350 kPa (gauge) and the bottom temperature may be between
250° and 300°C. When maximum production of gasoline is desired, the overhead pressure
in the debutanizer column 210 may be between 525 and 575 kPa (gauge) and the bottom
temperature may be between 90° and 140°C. The first vaporous oligomerate overhead
light stream comprising C
4 hydrocarbons may be rejected from the process and subjected to further processing
to recover useful components.
[0081] It is desired to maintain liquid phase in the oligomerization reactors. This is typically
achieved by saturating product olefins and recycling them to the oligomerization reactor
as a liquid. However, if olefinic product is being recycled to either the FCC zone
20 or the oligomerization zone 130, saturating olefins would inactivate the recycle
feed. The oligomerization zone 130 can only further oligomerize olefinic recycle and
the FCC zone 20 prefers olefinic feed to be further cracked to form propylene.
[0082] Liquid phase may be maintained in the oligomerization zone 130 by incorporating into
the feed a C
5 stream from the oligomerization recovery zone 200. The oligomerization recovery zone
200 may include a depentanizer column 220 to which the first liquid oligomerate bottom
stream comprising C
5+ hydrocarbons may be fed in line 214. The depentanizer column 220 may separate the
first liquid oligomerate bottom stream between vapor and liquid into an intermediate
stream comprising C
5 olefins and hydrocarbons in an intermediate line 222 and a liquid oligomerate bottom
product stream comprising C
6+ olefins in a bottom product line 224. When maximum production of distillate is desired,
either to obtain diesel product or to recrack the diesel in the FCC zone 20 to make
more propylene, the overhead pressure in the depentanizer column 220 may be between
10 and 60 kPa (gauge) and the bottom temperature may be between 225° and 275°C. When
maximum production of gasoline is desired, the overhead pressure in the depentanizer
column 220 may be between 250 and 300 kPa (gauge) and the bottom temperature may be
between 150° and 200°C.
[0083] The intermediate stream in intermediate line 222 may comprise at least 30 wt% and
suitably at least 40 wt% C
5 hydrocarbons which can then act as a solvent in the oligomerization reactor zone
140 to maintain liquid phase therein. The overhead intermediate stream comprising
C
5 hydrocarbons should have less than 10 wt% C
4 or C
6 hydrocarbons and preferably less than 1 wt% C
4 or C
6 hydrocarbons.
[0084] The intermediate stream may be condensed and recycled to the oligomerization zone
130 as a first intermediate recycle stream in an intermediate recycle line 226 to
maintain the liquid phase in the oligomerization reactors 138, 150 operating in the
oligomerization zone 130. The C
5 overhead stream may comprise C
5 olefins that can oligomerize in the oligomerization zone. The C
5 hydrocarbon presence in the oligomerization zone maintains the oligomerization reactors
at liquid phase conditions. The pentanes are easily separated from the heavier olefinic
product such as in the depentanizer column 220. The pentane recycled to the oligomerization
zone also dilutes the feed olefins to help limit the temperature rise within the reactor
due to the exothermicity of the reaction.
[0085] We have found that dimethyl sulfide boils with the C
5 hydrocarbons and deactivates the unidimensional, 10-ring pore structured zeolite
which may be the oligomerization catalyst. The mercaptan extraction unit 112 does
not remove sufficient dimethyl sulfide to avoid deactivating the oligomerization catalyst.
Consequently, recycle of C
5 hydrocarbons to the oligomerization reactor zone 140 with oligomerization catalyst
should be avoided by keeping valve 226' shut unless dimethyl sulfide can be successfully
removed from the oligomerate stream or the oligomerization catalyst is not a unidimensional,
10-ring pore structured zeolite. However, the dimethyl sulfide does not substantially
harm the solid phosphoric acid catalyst, so recycle of C
5 hydrocarbons to the oligomerization reactor zone 140 is suitable if SPA is the oligomerization
catalyst.
[0086] In an aspect, the intermediate stream in the intermediate line 222 comprising C
5 hydrocarbons may be split into a purge stream in purge line 228 and the first intermediate
recycle stream comprising C
5 hydrocarbons in the first intermediate recycle line 226. In an aspect, the first
intermediate recycle stream in first intermediate recycle line 226 taken from the
intermediate stream in intermediate line 222 is recycled to the oligomerization zone
130 downstream of the selective hydrogenation reactor 116. The intermediate stream
in intermediate line 222 and the first intermediate recycle stream in intermediate
recycle line 226 should be understood to be condensed overhead streams. The intermediate
recycle stream comprising C
5 hydrocarbons may be recycled to the oligomerization zone 130 at a mass flow rate
which is at least as great as and suitably no greater than three times the mass flow
rate of the oligomerization feed stream in the oligomerization feed line 128 fed to
said oligomerization zone 130 absent the addition of any recycle streams such as in
line 246 to be explained hereinafter. The recycle rate may be adjusted as necessary
to maintain liquid phase in the oligomerization reactors and to control temperature
rise, and to maximize selectivity to gasoline range oligomer products.
[0087] The purge stream comprising C
5 hydrocarbons taken from the intermediate stream may be purged from the process in
line 228 to avoid C
5 build up in the process. The purge stream comprising C
5 hydrocarbons in line 228 may be subjected to further processing to recover useful
components or be blended in the gasoline pool.
[0088] Three streams may be taken from the liquid oligomerate bottom product stream in bottom
product line 224. A recycle oligomerate product stream comprising C
6+ olefins in recycle oligomerate product line 230 may be taken from the liquid oligomerate
bottom product stream in bottom product line 224. The liquid oligomerate bottom product
stream in the bottom product line 224 may have the same composition as described for
the C
8 olefins of the oligomerate in oligomerate line 156. The liquid oligomerate bottom
product stream in the bottom product line 224 may have greater than 10 wt% C
10 isoolefins. Flow through recycle line 230 can be regulated by control valve 230'.
In another aspect, a distillate separator feed stream in distillate feed line 232
may be taken from the liquid oligomerate bottom product stream in the bottom product
line 224. Flow through distillate feed line 232 can be regulated by control valve
232'. In a further aspect, a gasoline oligomerate product stream in a gasoline oligomerate
product line 250 can be taken from the liquid oligomerate bottom product stream in
bottom product line 224. Flow through gasoline oligomerate product line 250 can be
regulated by control valve 250'. Flow through recycle oligomerate product line 230,
distillate feed line 232 and gasoline oligomerate product line 250 can be regulated
by control valves 230', 232' and 250', respectively, such that flow through each line
can be shut off or allowed irrespective of the other lines.
[0089] In an embodiment designed to bolster production of heavier oligomerate and maintain
liquid phase conditions in the oligomerization reactor zone 140, an oligomerate return
stream in oligomerate return line 231 may be taken from the recycle oligomerate product
stream comprising C
6+ olefins in the recycle oligomerate product line 230 and be recycled to the oligomerization
reactor zone 140 comprising oligomerization catalyst. In this case, a control valve
231' on oligomerate return line 231 is open, so that recycle oligomerate product is
recycled to the oligomerization reactor zone 140 in the oligomerization zone 130.
The oligomerization catalyst is resistant to excess oligomerization of heavier olefins,
so recycling heavier olefins to the oligomerization catalyst will not result in excess
oligomerization to heavier olefins than diesel. The recycle oligomerate product stream
comprising C
6+ olefins serves to maintain liquid phase in the oligomerization reactor zone 140
and provides olefins that can oligomerize to heavier diesel range olefins. In this
embodiment, the oligomerization zone 130 is in downstream communication with the first
bottom line 214 of the debutanizer column 210 and the bottom product line 224 of the
depentanizer column 220. In a further aspect, the recycle oligomerate product line
230 and the oligomerate return line 231 are in downstream communication with the oligomerization
zone 130. Consequently, the oligomerization zone 130 is in upstream and downstream
communication with the first bottom line 214, the bottom product line 224, the recycle
oligomerate product line 230 and the oligomerate return line 231.
[0090] The concentration of dimethyl sulfide in the oligomerate return stream in the oligomerate
return line 231 should be no more than 5 wppm sulfur as dimethylsulfide. Consequently,
if the recycle oligomerate product stream in recycle oligomerate product line 230
is taken from the oligomerate bottom product stream comprising C
6+ olefins in the bottom product line 224 to be recycled to the oligomerization reactor
zone 140, it should comprise no more than 5 wppm sulfur as dimethyl sulfide. Accordingly,
the oligomerization recovery zone 200 should be operated to produce an oligomerate
bottom product stream that has no more than 5 wppm sulfur as dimethyl sulfide and/or
less than 1 wt% C
5 hydrocarbons.
[0091] If a refiner desires to make additional propylene in the FCC unit, an embodiment
may be used in which an FCC recycle oligomerate stream taken from the recycle oligomerate
product stream in the recycle oligomerate product line 230 from the oligomerate bottom
product stream comprising C
6+ olefins in the bottom product line 224 may be recycled to the FCC recycle line 280.
An FCC recycle oligomerate line 233 may take an FCC recycle oligomerate stream from
the recycle oligomerate stream in recycle oligomerate line 230 and forward it to the
FCC zone 20 through FCC recycle line 280. The FCC recycle oligomerate line 233 communicates
the recycle oligomerate line with FCC recycle line 280 and the FCC reaction zone 20.
A control valve 233' on the FCC recycle oligomerate line 233 may be open if recycle
oligomerate product to the FCC zone 20 is desired. The FCC recycle line 280 will carry
the FCC recycle oligomerate stream as feed to the FCC zone 20. In an aspect, the recycle
oligomerate product stream in the recycle oligomerate product line 230 is in downstream
communication with the FCC recovery zone 200. In a further aspect, the FCC recycle
oligomerate line 233 is in downstream communication with the oligomerization zone
130. Hence, in an aspect, the FCC reaction zone 20 is in upstream and downstream communication
with oligomerization zone 130 and/or FCC recovery zone 100. In a still further aspect,
FCC recycle oligomerate line 233 and recycle oligomerate product line 230 are in upstream
communication with the FCC reaction zone 20 to recycle oligomerate for fluid catalytic
cracking down to propylene or other light olefins. One or both of valves 231' and
233' may be opened or closed depending on the refiner's desire for recycle to the
oligomerization zone 130 or the FCC zone 20, respectively.
[0092] In an embodiment in which the oligomerization catalyst is SPA in the oligomerization
reactor zone 140 for oligomerizing C
4 olefins or a mixed C
4 and C
5 olefin stream, we have found that a gasoline product stream can be provided by the
oligomerate bottom product stream in bottom product line 224. The SPA catalyst minimizes
the formation of C
12+ species with either a C
4 olefin or C
4 and C
5 olefin feed. Consequently, even when heavier olefins than C
4 olefins are present in the oligomerization feed stream, the SPA catalyst manages
to keep C
12+ olefins present in the liquid oligomerate bottom product stream in the bottom product
line 224 below less than 20 wt% even when over 85 wt% of feed olefins are converted
and particularly when over 90 wt% of C
4 olefins are converted to oligomerate.
[0093] Accordingly, the liquid oligomerate bottom product stream in bottom product line
224 provides gasoline range material that meets the Engler T90 gasoline specification
of 193°C (380°F) using the ASTM D-86 Test Method without further treatment when SPA
is the second oligomerization catalyst in the oligomerization reactor zone 140. That
is, 90 wt% of the resulting liquid oligomerate bottom product stream, for example,
in bottom product line 224 will boil before its temperature is raised to 193°C (380°F).
Consequently, a gasoline oligomerate product stream can be collected from the liquid
oligomerate bottom product stream in a gasoline oligomerate product line 250 and blended
in the gasoline pool without further treatment such as separation or chemical upgrading.
The gasoline oligomerate product line 250 may be in upstream communication with a
gasoline tank 252 or a gasoline blending line of a gasoline pool. However, further
treatment such as partial or full hydrogenation to reduce olefinicity may be contemplated.
In such a case, control valves 232' and 230' may be all or partially closed and control
valve 250' on oligomerate liquid product line 250 may be opened to allow C
6+ gasoline product to be sent to the gasoline tank 252 or the gasoline blending line.
[0094] The oligomerization recovery zone 200 may also include a distillate separator column
240 to which the distillate separator oligomerate feed stream comprising oligomerate
C
6+ hydrocarbons may be fed in distillate feed line 232 taken from the liquid oligomerate
bottom product stream in line 224 for further separation. The distillate separator
column 240 is in downstream communication with the first bottom line 214 of the debutanizer
column 210 and the bottom product line 224 of the depentanizer column 220.
[0095] The distillate separator column 240 separates the distillate separator oligomerate
feed stream into an gasoline overhead stream in an overhead line 242 comprising C
6, C
7, C
8, C
9, C
10 and/or C
11 olefins and a bottom distillate stream comprising C
8+, C
9+, C
10+, C
11+, or C
12+ olefins in a diesel bottom line 244. When maximum production of distillate is desired,
either to obtain diesel product or to recrack the diesel in the FCC zone 20 to make
more propylene, the overhead pressure in the distillate separator column 240 may be
between 10 and 60 kPa (gauge) and the bottom temperature may be between 225° and 275°C.
When maximum production of gasoline is desired, the overhead pressure in the distillate
separator column 240 may be between 10 and 60 kPa (gauge) and the bottom temperature
may be between 190° and 250°C. The bottom temperature can be adjusted between 175°
and 275°C to adjust the bottom product between a C
9+ olefin cut and a C
12+ olefin cut based on the heaviness of the diesel cut desired by the refiner. The
gasoline overhead stream in gasoline overhead line 242 may have the same composition
as described for the Cg olefins of the oligomerate in oligomerate line 156. The diesel
bottoms stream in diesel bottoms line 244 may have greater than 30 wt% C
9+ isoolefins.
[0096] For refiners who are interested in distillate production at a particular time, the
gasoline overhead stream comprising C
8 olefins in the gasoline overhead line 242 of the distillate separator column can
be recycled to the oligomerization zone 130 to increase the production of distillate.
For example, a gasoline overhead recycle stream in gasoline overhead recycle line
246 may be taken from the gasoline overhead stream in gasoline overhead line 242 and
mixed with the fresh oligomerization feed stream in oligomerization feed line 128.
A control valve 246' may be used to completely shut off flow through gasoline overhead
recycle line 246 or allow partial or full flow therethrough. The gasoline overhead
recycle line 246 may be in downstream communication with the oligomerization recovery
zone 200 to generate diesel range material.
[0097] Preferably, the gasoline recycle gasoline stream in line 246, which may be taken
from the gasoline overhead in line 242, may be recycled to the oligomerization reactors,
138 and 150 of the oligomerization reactor zone 140 with oligomerization catalyst.
The gasoline overhead stream may comprise C
6-C
11 olefins and preferably C
7-C
9 olefins and most preferably C
8 olefins that can oligomerize with C
4-C
5 olefins in the oligomerization feed stream in the oligomerization zone 130 to diesel
range material comprising C
10-C
16 diesel product. C
4 olefins continue to oligomerize with C
4 olefins and C
5 olefins if present in the feed.
[0098] The oligomerization catalyst, and particularly, the uni-dimensional, 10-ring pore
structured zeolite, converts a significant fraction of the gasoline-range olefins,
such as C
8 olefins, to distillate material by oligomerizing them with feed olefins, such as
C
4 and/or C
5 olefins. Additionally, the presence of the gasoline-range olefins also encourages
oligomerization of the feed olefins with each other over the zeolite catalyst. Surprisingly,
the isobutene conversion is lower than normal butene conversion at high overall butene
conversion such as over 90% C
4 olefin conversion. When gasoline is recycled from the gasoline overhead line 242
to the oligomerization reactor zone 140 for oligomerization over uni-dimensional,
10-ring pore structured zeolite, oligomerate from the oligomerization zone in oligomerate
line 156 may comprise greater than 30 wt% C
9+ olefins. Under these circumstances, oligomerate from the oligomerization reactor
zone in oligomerate line 156 may comprise greater than 50 wt% or even greater than
60 wt% C
9+ olefins.
[0099] In an aspect, the gasoline overhead stream in gasoline overhead line 242 may be recovered
as product in product gasoline line 248 in downstream communication with the recovery
zone 200. A control valve 248' may be used to completely shut off flow through gasoline
product line 248 or allow partial or full flow therethrough. The gasoline product
stream may be subjected to further processing to recover useful components or blended
in the gasoline pool. The gasoline product line 248 may be in upstream communication
with a gasoline tank 252 or a gasoline blending line of a gasoline pool. In this aspect,
the overhead line 242 of the distillate separator column may be in upstream communication
with the gasoline tank 252 or the gasoline blending line.
[0100] In an embodiment, the diesel bottom stream in a diesel bottom line 244 may be recycled
to the FCC zone 20 in FCC recycle line 280 via a recycle diesel line 260 in downstream
communication with the oligomerization recovery zone 200 to be cracked to propylene
product in the FCC zone. A recycle diesel bottom stream in recycle diesel line 260
taken from the diesel bottom stream in line 244 may be forwarded to the FCC recycle
line 280. The diesel bottom stream may comprise C
9+, C
10+, C
11+ or C
12+ olefins that can crack to propylene. A control valve 260' may be used to completely
shut off flow through recycle diesel line 260 or allow partial or full flow therethrough.
In this embodiment, the FCC zone 20 is in downstream communication with the distillate
separator column 240 and particularly the diesel bottom line 244.
[0101] If the FCC zone 20 comprises a single reactor riser 26, the first reactor riser 26
may be in downstream communication with the hydrocarbon feed line 24 and the diesel
bottom line 244 of the distillate separator column 240. If the FCC zone 20 comprises
the first reactor riser 26 and a second reactor riser 74, the first reactor riser
26 may be in downstream communication with the hydrocarbon feed line 24 and the second
reactor riser 74 may be in downstream communication with the bottom line 244 of the
distillate separator column 240.
[0102] We have found that C
6+ oligomerate and distillate oligomerate subjected to FCC is converted best over a
blend of medium or smaller pore zeolite blended with a large pore zeolite such as
Y zeolite as explained previously with respect to the FCC zone 20. Additionally, oligomerate
produced over the oligomerization catalyst in the oligomerization reactor zone 140
provides an excellent feed to the FCC zone that can be cracked to yield greater quantities
of propylene.
[0103] In an aspect, the diesel bottom stream may be recovered as product in a diesel product
line 262 in downstream communication with the oligomerization recovery zone 200. The
diesel product line in line 262 is taken from the diesel bottom stream in diesel bottom
line 244. A control valve 262' may be used to completely shut off flow through the
diesel product line 262 or allow partial or full flow therethrough. The diesel product
stream may be subjected to further processing to recover useful components or blended
in the diesel pool. The diesel product line 262 may be in upstream communication with
a diesel tank 264 or a diesel blending line of a diesel pool. Additionally, LCO from
LCO line 107 may also be blended with diesel in diesel product line 262.
[0104] FIG. 2 depicts an alternative embodiment of the oligomerization recovery zone 200.
Elements in FIG. 2 with the same configuration as in FIG. 1 will have the same reference
numeral as in FIG. 1. Elements in FIG. 2 which have a different configuration as the
corresponding element in FIG. 1 will have the same reference numeral but designated
with a suffix "a". The configuration and operation of the embodiment of FIG. 2 is
essentially the same as in FIG. 1 with the exceptions noted below.
[0105] In FIG. 2, the oligomerization recovery zone 200a comprises a fractionation debutanizer
column 210a in downstream communication with the oligomerization zone130. The oligomerate
steam in oligomerate line 156 is fed to an inlet 181 to the fractionation debutanizer
column 210a which separates the oligomerate stream between vapor and liquid into a
first vaporous oligomerate overhead light stream in a first overhead line 212 comprising
C
4 hydrocarbons, an intermediate side stream in intermediate line 214a comprising C
5 hydrocarbons and a liquid oligomerate bottom product stream comprising C
6+ olefins in a bottom product line 224a. The intermediate side stream may be taken
from a side outlet 215 of the fractionation debutanizer column 210a. The intermediate
stream may be a liquid collected on a tray in the fractionation debutanizer column
210a.
[0106] The fractionation debutanizer column 210a feeds the intermediate side stream from
the side outlet 215 of the fractionation debutanizer column 210a to a side stripper
column 220a to separate the intermediate side stream into a second overhead stream
in second overhead line 221 comprising C
4- hydrocarbons and a second bottom stream in a second bottom line 228a comprising
C
5 hydrocarbons. The side stripper column 220a may be in downstream communication with
the side outlet 215 of the fractionation debutanizer column 210a. The second overhead
stream 221 is fed to the fractionation debutanizer column 210a at a side inlet 223.
Consequently, the fractionation debutanizer column 210a is in downstream communication
with the overhead line 221 from the side stripper column 220a. Hence, in an aspect,
the fractionation debutanizer column 210a is in upstream and downstream communication
with the side stripper column 220a.
[0107] The feed inlet 181 to the fractionation debutanizer column may be at a lower elevation
than a side inlet 223 from the overhead line 221 from the side stripper 220a. Additionally,
the side inlet 223from the overhead line 221 from the side stripper 220a may be at
a higher elevation than the side outlet 215. Lastly, the side outlet 215 may be at
a higher elevation on the debutanizer column 210a than the feed inlet 181.
[0108] When maximum production of distillate is desired, either to obtain diesel product
or to recrack the diesel in the FCC zone 20 to make more propylene, the overhead pressure
in the debutanizer column 210a may be between 350 and 400 kPa (gauge) and the bottom
temperature may be between 270° and 320°C. When maximum production of gasoline is
desired, the overhead pressure in the debutanizer column 210a may be between 350 and
400 kPa (gauge) and the bottom temperature may be between 170° and 220°C. The side
stripper column 220a may have an overhead pressure of between 400 and 450 kPa and
a bottom temperature of between 60° and 115°C in both modes.
[0109] One or both of the first vaporous oligomerate overhead light stream in first overhead
line 212 comprising C
4 hydrocarbons and the second bottom stream in second bottom line 228a comprising C
5 hydrocarbons may be purged from the process.
[0110] A stream comprising C
5 hydrocarbons may be used to maintain the oligomerization zone 130 in liquid phase
and provide additional C
5 olefins for oligomerization. An intermediate stream comprising C
5 hydrocarbons in intermediate line 222a may be taken from the intermediate side stream
in line 214a before it is further fractionated such as in the side stripper 220a and
recycled to the oligomerization zone 130 through an open control valve 222a' thereon.
Taking a stream of C
5 hydrocarbons from the intermediate side stream removes a large amount of material
from the side stripper column 220a without requiring it to be further reboiled or
condensed thus decreasing its capacity and the expense to operate. Accordingly, the
oligomerization zone 130 is in downstream communication with said side outlet 215.
[0111] A recycle oligomerate stream comprising C
6+ olefins may be used to maintain the oligomerization zone 130 in liquid phase and
provide additional olefins for oligomerization. A recycle oligomerate product stream
may be taken in recycle oligomerate product line 230a through open control valve 230a'
from the liquid oligomerate bottom product stream in the bottom product line 224a
which comprises C
6+ olefins. An oligomerate return stream in oligomerate return line 231 through open
control valve 231' may be taken from the recycle oligomerate product stream in recycle
oligomerate product line 230a and be recycled to the oligomerization zone 130. The
oligomerization zone 130 may be, therefore, in downstream communication with the bottom
product line 224a of the fractionation column. The oligomerate return stream in oligomerate
return line 231 may be recycled to the oligomerization reactor zone 140 having the
oligomerization catalyst.
[0112] Distillate oligomer product may be recycled to the FCC unit to make more propylene.
An FCC recycle oligomerate stream in FCC recycle oligomerate line 233 may be taken
from the recycle oligomerate product stream in recycle oligomerate product line 230a
and be forwarded through open valve 233' to the FCC zone 20 in FCC recycle line 280.
Accordingly, the FCC zone may be in downstream communication with the bottom product
line 224a of the fractionation debutanizer column. Hence, in an aspect, the FCC zone
20 may be in upstream and downstream communication with the oligomerization zone 130
and/or the debutanizer column 210a.
[0113] If the oligomerate bottom product stream has a suitable composition, it may be taken
as gasoline product in line 250a through control valve 250a' to a gasoline pool which
may comprise a gasoline tank 252 or a gasoline blending line. Accordingly, the gasoline
tank 252 or the gasoline blending line may be in downstream communication with an
oligomerate bottom product line 224a of said fractionation debutanizer column 210a.
[0114] If sufficient diesel is provided in bottom product line 224a, the gasoline should
be separated from the diesel. A distillate separator feed stream may be taken from
the oligomerate bottom product stream in bottom product line 224a in line 232a through
open control valve 232a' to a distillate separator column 240. The distillate separator
column 240 can separate the distillate separator feed stream into a gasoline stream
242 and a distillate stream 244 as previously described with respect to FIG. 1. Accordingly,
the distillate separator column 240 is in downstream communication with a bottom product
line of said fractionation debutanizer column 210a.
[0115] In the embodiment of FIG. 2, when maximum production of distillate is desired, either
to obtain diesel product or to recrack the diesel in the FCC zone 20 to make more
propylene, the overhead pressure in the distillate separator column 240 may be between
150 and 200 kPa (gauge) and the bottom temperature may be between 250° and 300°C.
When maximum production of gasoline is desired, the overhead pressure in the debutanizer
column 210 may be between 150 and 200 kPa (gauge) and the bottom temperature may be
between 210° and 260°C.
[0116] The invention will now be further illustrated by the following non-limiting examples.
EXAMPLES
EXAMPLE 1
[0117] Feed 1 in Table 2 was contacted with four catalysts to determine their effectiveness
in oligomerizing butenes.
Table 2
Component |
Fraction, wt% |
propylene |
0.1 |
Iso-C4's |
70.04 |
isobutylene |
7.7 |
1-butene |
5.7 |
2-butene (cis and trans) |
16.28 |
3-methyl-1-butene |
0.16 |
acetone |
0.02 |
Total |
100 |
[0118] Catalyst A is an MTT catalyst purchased from Zeolyst having a product code Z2K019E
and extruded with alumina to be 25 wt% zeolite. Of MTT zeolite powder, 53.7 grams
was combined with 2.0 grams Methocel and 208.3 grams Catapal B boehmite. These powders
were mixed in a muller before a mixture of 18.2g HNO
3 and 133 grams distilled water was added to the powders. The composition was blended
thoroughly in the muller to effect an extrudable dough of 52% LOI. The dough then
was extruded through a die plate to form cylindrical extrudates having a diameter
of 3.18 mm. The extrudates then were air dried, and calcined at a temperature of 550°C.
The MTT catalyst was not selectivated to neutralize surface acid sites such as with
an amine.
[0119] Catalyst B is a SPA catalyst commercially available from UOP LLC.
[0120] Catalyst C is an MTW catalyst with a silica-to-alumina ratio of 36:1. Of MTW zeolite
powder made in accordance with the teaching of
US 7,525,008 B2, 26.4 grams was combined with and 135.1 grams Versal 251 boehmite. These powders
were mixed in a muller before a mixture of 15.2 grams of nitric acid and 65 grams
of distilled water were added to the powders. The composition was blended thoroughly
in the muller to effect an extrudable dough of 48% LOI. The dough then was extruded
through a die plate to form cylindrical extrudates having a diameter of 1/32". The
extrudates then were air dried and calcined at a temperature of 550°C.
[0121] Catalyst D is an MFI catalyst purchased from Zeolyst having a product code of CBV-8014
having a silica-to-alumina ratio of 80:1 and extruded with alumina at 25 wt% zeolite.
Of MFI-80 zeolite powder, 53.8 grams was combined with 205.5 grams Catapal B boehmite
and 2 grams of Methocel. These powders were mixed in a muller before a mixture of
12.1 grams nitric acid and 115.7 grams distilled water were added to the powders.
The composition was blended thoroughly in the muller, then an additional 40 grams
of water was added to effect an extrudable dough of 53% LOI. The dough then was extruded
through a die plate to form cylindrical extrudates having a diameter of 3.18 mm. The
extrudates then were air dried, and calcined at a temperature of 550°C.
[0122] The experiments were operated at 6.2 MPa and inlet temperatures at intervals between
160° and 240°C to obtain different normal butene conversions. Results are shown in
FIGS. 3 and 4. In FIG. 3, C
8 to C
11olefin selectivity is plotted against normal butene conversion to provide profiles
for each catalyst.
[0123] Table 3 compares the RONC ±3 for each product by catalyst and provides a key to FIG.
3. The SPA catalyst B is superior, but the MTT catalyst A is the least effective in
producing gasoline range olefins.
Table 3
Catalyst |
RONC |
A |
MTT |
circles |
92 |
B |
SPA |
diamonds |
96 |
C |
MTW |
triangles |
97 |
D |
MFI-80 |
asterisks |
95 |
[0124] The SPA catalyst was able to achieve over 95 wt% yield of gasoline having a RONC
of >95 and with an Engler T90 value of 185°C for the entire product. The T-90 gasoline
specification is less than 193°C.
[0125] In FIG. 4, C
12+ olefin selectivity is plotted against normal butene conversion to provide profiles
for each catalyst. Table 4 compares the derived cetane number ±2 for each product
by catalyst and provides a key to FIG. 4.
Table 4
Catalyst |
Cetane |
A |
MTT |
circles |
41 |
B |
SPA |
diamonds |
<14 |
C |
MTW |
triangles |
28 |
D |
MFI-80 |
asterisks |
36 |
[0126] FIG. 4 shows that the MTT catalyst provides the highest C
12+ olefin selectivity which reaches over 70 wt%. These selectivities are from a single
pass of the feed stream through the oligomerization reactor. Additionally, the MTT
catalyst provided C
12+ oligomerate with the highest derived cetane. Cetane was derived using ASTM D6890
on the C
12+ fraction at the 204°C (400°F) cut point. Conversely to gasoline selectivity, the
MTT catalyst A is superior, but the SPA catalyst B is the least effective in producing
diesel range olefins.
[0127] The MTT catalyst was able to produce diesel with a cetane rating of greater than
40. The diesel cloud point was determined by ASTM D2500 to be -66°C and the T90 was
319°C using ASTM D86 Method. The T90 specification for diesel in the United States
is between 282 and 338°C, so the diesel product meets the U.S. diesel standard.
EXAMPLE 2
[0128] A comprehensive two-dimensional gas chromatography with flame ionization detection
(GCxGC-FID) method was developed and utilized to analyze the composition of light
olefin oligomerization product streams. To develop the peak identifications, a GCxGC
instrument equipped with a time of flight mass spectrometer (TOFMS) was used. Peak
identifications were checked against a table of C
8 olefin boiling points for consistency and by performing GC-FID of the olefinic sample
with and without hydrogenation catalyst in the GC inlet to ensure that peaks assigned
to a particular C
8 mono-olefin moved to their respective saturated C
8 isoparaffins. The identification of C
8 paraffin isomers can be achieved using the UOP690 method. Careful matching of chromatographic
conditions between GCxGC-FID and GCxGC-TOFMS allows one to translate identifications
made from the TOFMS analysis to the GCxGC-FID for quantitative analysis. The following
48 compounds in the C
8 region were identified and quantified:
[0129] C
8 olefin species identified are listed as follows: 2,3-dimethyl-2-butene, 3,4-dimethyl-2-pentene,
3,4-dimethyl-2-pentene, 2,4,4-trimethyl-1-pentene, 2,2-dimethyl-trans-3-hexene, 2,5-dimethyl-3-hexene,
3,3-dimethyl-1-hexene, 3,4,4-trimethyl-1-pentene, 2,4,4-trimethyl-2-pentene, 4,4-dimethyl-2-hexene,
4,4-dimethyl-1-hexene, 2,3,4-trimethyl-1-pentene, 2,3,3-trimethyl-1-pentene, 2,4-dimethyl-trans-3-hexene,
2,4-dimethyl-cis-3-hexene, 3,3-dimethyl-2-ethyl-1-butene, 2,4-dimethyl-1-hexene, 2,3-dimethyl-1-hexene,
2-methyl-3-heptene, 3,4,4-trimethyl-2-pentene, 2,5-dimethyl-2-hexene, 5-methyl-3-heptene,
3,5-dimethyl-2-hexene, 6-methyl-3-heptene, 4-methyl-1-heptene, 4-methyl-3-ethyl-trans-2-pentene,
2,3-dimethyl-3-hexene, 4-methyl-3-ethyl-cis-2-pentene, 3,4-dimethyl-2-hexene, 3-ethyl-3-hexene,
6-methyl-2-heptene, 2,3,4-trimethyl-2-pentene, 2-methyl-3-ethyl-2-pentene, 5-methyl-2-heptene,
2-n-propyl-1-pentene, 4-methyl-3-heptene, 2-ethyl-1-hexene, 2-methyl-1-heptene, 3-methyl-3-heptene,
trans-3-octene, 2,3-dimethyl-2-hexene, 3-methyl-2-heptene, 3,4-dimethyl-trans-3-hexene,
cis-3-octene, 2-methyl-2-heptene, trans-2-octene, cis-2-octene, and 3,4-dimethyl-cis-3-hexene.
[0130] C
8 olefins in the oligomerate produced by all four catalysts were evaluated by the GCxGC-FID
method to characterize oligomerate product composite by olefin type. GCxGC analysis
on the composite product of the experiment was used to compare the product olefinic
isomers from Catalysts A and B as shown in Table 5.
Table 5
Isomer Species |
Fraction from Catalyst A, wt% |
Fraction from Catalyst B, wt% |
C5 olefins |
0.02 |
3.12 |
C6 olefins |
1.50 |
0.25 |
C7 olefins |
1.13 |
0.92 |
linear C8= |
0.91 |
0.03 |
methyl-heptenes |
10.03 |
1.68 |
dimethyl-hexenes |
13.25 |
18.70 |
trimethyl-pentenes |
7.24 |
52.63 |
C9 olefins |
3.03 |
3.63 |
C10 olefins |
1.92 |
1.40 |
C11 olefins |
5.67 |
7.13 |
C12 olefins |
29.86 |
6.64 |
C13 olefins |
3.13 |
0.56 |
C14 olefins |
1.61 |
0.17 |
C15 olefins |
3.28 |
0.37 |
C16 olefins |
13.37 |
0.25 |
C17 olefins |
1.47 |
0.00 |
C18 olefins |
0.64 |
0.00 |
C19 olefins |
1.12 |
0.00 |
Other Olefins and Polar Unknowns |
0.82 |
2.52 |
Total |
100.00 |
100.00 |
[0131] Catalyst B, SPA, produces over 70 wt% C
8 olefins with over 70 wt% of the C
8 olefins being trimethyl pentenes. However, Catalyst A, MTT, produces only just over
31 wt% C
8 olefins of which only 23 wt% of the C
8 olefins are trimethyl pentenes. Catalyst A produced almost 30 wt% C
12 olefins. It is evident that MTT can produce a more linear and larger product from
light olefins such as butene.
[0132] Comparisons are shown in Table 6. All percentages are in weight percent. "Composition
Fraction" is the fraction of the species in the entire composition. "Olefin Fraction"
is the fraction of the species among the C
8 olefins. "Average Branches" for the C
8 olefins is the average branches or alkyl groups per olefin molecule calculated by
the ratio of the sum of the total weight of each isomer of branched C
8 olefins multiplied by the number of alkyl groups of that isomer in the composition
divided by the total weight of normal octene, methyl heptene, dimethyl hexene and
trimethyl pentene in the composition. "Olefin Isomer Fraction" is fraction of C
8 olefin isomer with the structure among all C
8 olefins. "TMP/C
8 Olefins" is the ratio of trimethyl pentene among linear octene, methyl heptene, dimethyl
hexene and trimethyl pentene.
Table 6
|
Catalysts |
C8 Olefin Species |
A |
B |
C |
D |
Composition Fraction |
|
|
|
|
Type I composition,% |
0.6 |
2.1 |
0.5 |
2.3 |
Type II composition,% |
3.3 |
1.0 |
0.4 |
3.5 |
Type III composition,% |
6.8 |
16.2 |
3.1 |
15.4 |
Type IV composition,% |
14.0 |
14.6 |
3.7 |
21.4 |
Type V composition,% |
4.8 |
19.5 |
3.8 |
13.9 |
Total |
29.5 |
53.3 |
11.5 |
56.3 |
Olefin Fraction |
|
|
|
|
Type I olefin,% |
2.1 |
4.0 |
4.4 |
4.0 |
Type II olefin,% |
11.3 |
1.8 |
3.4 |
6.1 |
Type III olefin,% |
22.8 |
30.3 |
27.1 |
27.3 |
Type IV olefin,% |
47.3 |
27.3 |
31.7 |
38.0 |
Type V olefin,% |
16.3 |
36.6 |
33.4 |
24.7 |
Total |
100 |
100 |
100 |
100 |
|
|
|
|
|
Average Branches |
1.98 |
2.75 |
2.55 |
2.32 |
|
|
|
|
|
C8 Olefin Isomer Fraction |
|
|
|
|
Linear octene,% |
2.8 |
0.0 |
0.1 |
0.7 |
Methyl heptene,% |
27.3 |
2.1 |
5.1 |
11.9 |
Dimethyl hexene,% |
34.9 |
20.0 |
32.9 |
38.7 |
Trimethyl pentene,% |
31.3 |
75.4 |
59.2 |
44.4 |
Other C8 monoolefins,% |
3.8 |
2.5 |
2.8 |
4.4 |
|
|
|
|
|
TMP/C8 Olefins,% |
32.5 |
77.3 |
60.9 |
46.4 |
[0133] The fraction of Type 2 disubstituted C
8 olefins in the total C
8 olefins for Catalyst A was 11.3 wt% which was much higher than all the other catalysts.
The ratio of Type 2 disubstituted C
8 olefins to Type 1 monosubstituted C
8 olefins for catalyst A was 5.3. All the other catalysts had the same ratio of less
than one. The fraction of Type 4 trisubstituted C
8 olefins in the total C
8 olefins in the oligomerate for Catalyst A was 47 wt%. All of the other catalysts
had the same fraction of no more than 38 wt%. The average branch per hydrocarbon molecule
for Catalyst A was 1.98; whereas, the other catalysts were all over 2. The ratio of
trimethyl pentene to the total C
8 olefins in the oligomerate was 32.5; whereas all of the other catalysts had ratios
over 46.
EXAMPLE 3
[0134] Two types of feed were oligomerized over oligomerization catalyst A of Example 1,
MTT zeolite. Feeds 1 and 2 contacted with catalyst A are shown in Table 7. Feed 1
is from Example 1.
Table 7
|
Feed 1 |
Feed 2 |
Component |
Fraction, wt% |
Fraction, wt% |
propylene |
0.1 |
0.1 |
isobutane |
70.04 |
9.73 |
isobutylene |
7.7 |
6.3 |
1-butene |
5.7 |
4.9 |
2-methyl-2-butene |
0 |
9.0 |
2-butene (cis & trans) |
16.28 |
9.8 |
3-met-1-butene |
0.16 |
0.16 |
n-hexane |
0 |
60 |
acetone |
0.02 |
0.01 |
Total |
100 |
100 |
[0135] In Feed 2, C
5 olefin is made up of 2-methyl-2-butene and 3-methyl-1-butene which comprises 9.16
wt% of the reaction mixture representing a third of the olefins in the feed. 3-methyl-1-butene
is present in both feeds in small amounts. Propylene was present at less than 0.1
wt% in both feeds.
[0136] The reaction conditions were 6.2 MPa and a 1.5 WHSV. The maximum catalyst bed temperature
was 220°C. Oligomerization achievements are shown in Table 8.
Table 8
|
Feed 1 |
Feed 2 |
Inlet Temperature, °C |
192 |
198 |
C4 olefin conversion,% |
98 |
99 |
nC4 olefin conversion,% |
97 |
99 |
C5 olefin conversion,% |
n/a |
95 |
C5-C7 selectivity, wt% |
3 |
5 |
C8-C11 selectivity, wt% |
26 |
40 |
C12-C15 selectivity, wt% |
48 |
40 |
C16+ selectivity, wt% |
23 |
16 |
Total C9+ selectivity, wt% |
78 |
79 |
Total C12+ selectivity, wt% |
71 |
56 |
Net gasoline yield, wt% |
35 |
44 |
Net distillate yield, wt% |
76 |
77 |
[0137] Normal C
4 olefin conversion reached 99% with C
5 olefins in Feed 2 and was 97 wt% without C
5 olefins in Feed 1. C
5 olefin conversion reached 95%. With C
5 olefins in Feed 2, it was expected that a greater proportion of heavier, distillate
range olefins would be made. However, the Feed 2 with C
5 olefins oligomerized to a greater selectivity of lighter, gasoline range product
in the C
5-C
7 and C
8-C
11 range and a smaller selectivity to heavier distillate range product in the C
12-C
15 and C
16+ range.
[0138] This surprising result indicates that by adding C
5 olefins to the feed, a greater yield of gasoline can be made over Catalyst A, MTT.
This is confirmed by the greater net yield of gasoline and the lower selectivity to
C
12+ fraction for Feed 2 than for Feed 1. Also, but not to the same degree, by adding
C
5 olefins to the feed, a greater yield of distillate range material can be made. This
is confirmed by the greater net yield of distillate for Feed 2 than for Feed 1 on
a single pass basis. Gasoline yield was classified by product meeting the Engler T90
requirement and distillate yield was classified by product boiling over 150°C (300°F).
EXAMPLE 4
[0139] Three types of feed were oligomerized over oligomerization catalyst B of Example
1, SPA. The feeds contacted with catalyst B are shown in Table 9. Feed 2 is the same
as Feed 2 in Example 3. Isooctane was used as a diluent with Feed 3 because it is
expected to behave inertly just as isobutane. Feed 4 is similar to Feed 2 but has
the pentenes evenly split between iso and normal pentenes, which is roughly expected
to be found in an FCC product, and diluted with isobutane instead of hexane and isobutane.
Table 9
|
Feed 2 |
Feed 3 |
Feed 4 |
Component |
Fraction, wt% |
Fraction, wt% |
Fraction, wt% |
propylene |
0.1 |
0.08 |
0.1 |
1,3-butadiene |
0 |
0.28 |
0 |
isobutane |
9.73 |
6.45 |
69.72 |
isobutylene |
6.3 |
7.30 |
6.3 |
1-butene |
4.9 |
5.07 |
4.9 |
2-methyl-2-butene |
9.0 |
0 |
4.5 |
2-butene (cis & trans) |
9.8 |
11.33 |
9.8 |
3-met-1-butene |
0.16 |
0.16 |
0.16 |
2-pentene |
0 |
0 |
4.5 |
cyclopentane |
0 |
0.28 |
0 |
n-hexane |
60 |
0 |
0 |
isooctane |
0 |
60.01 |
0 |
acetone |
0.01 |
0.01 |
0.02 |
Total |
100 |
100 |
100 |
[0140] The reaction pressure was 3.5 MPa. Oligomerization achievements are shown in Table
10.
Table 10
|
Feed 2 |
Feed 3 |
Feed 4 |
WHSV, hr-1 |
.75 |
1.5 |
.75 |
Pressure, MPa |
3.5 |
3.5 |
6.2 |
Inlet Temperature, °C |
190 |
170 |
178 |
Maximum Temperature, °C |
198 |
192 |
198 |
C4 olefin conversion,% |
95 |
92 |
93 |
nC4 olefin conversion,% |
95 |
90 |
93 |
C5 olefin conversion,% |
90 |
n/a |
86 |
C5-C7 selectivity, wt% |
8 |
5 |
8 |
C8-C11 selectivity, wt% |
77 |
79 |
77 |
C12-C15 selectivity, wt% |
15 |
16 |
15 |
C16+ selectivity, wt% |
0.3 |
0.1 |
.01 |
Total C9+ selectivity, wt% |
35 |
20 |
25 |
Total C12+ selectivity, wt% |
17 |
16 |
15 |
Net gasoline yield, wt% |
94 |
92 |
91 |
Net distillate yield, wt% |
32 |
18 |
23 |
RONC (±3) |
97 |
96 |
96 |
Engler T-90, °C |
182 |
164 |
182 |
[0141] Olefin conversion was at least 90% and normal butene conversion was over 90%. Normal
C
4 olefin conversion reached 90% with C
5 olefins in Feed 2 and was 97% without C
5 olefins in Feed 1. C
5 olefin conversion reached 90% but was less when both iso and normal C
5 olefins were in the feed.
[0142] It can be seen that the SPA catalyst minimized the formation of C
12+ species to below 20 wt% at 16 and 17 wt%, respectively, without and with C
5 olefins in the oligomerization feed stream. When normal C
5 olefins were added, C
12+ formation reduced to 15 wt%. The C
6+ oligomerate produced by all three feeds met the gasoline T-90 spec indicating that
90 wt% boiled at temperatures under 193°C (380°F). The Research Octane Number for
all three products was high, over 95, with and without substantial C
5 olefins present.
EXAMPLE 5
[0143] Feed 2 with C
5 olefins present was subjected to oligomerization with Catalyst B, SPA, at different
conditions to obtain different butene conversions. C
5 olefin is made up of 2-methyl-2-butene and 3-methyl-1-buene which comprises 9.16
wt% of the reaction mixture representing a third of the olefins in the feed. Propylene
was present at less than 0.1 wt%. Table 11 shows the legend of component olefins illustrated
in FIG. 5.
Table 11
Component |
Symbols in FIG. 5 |
isobutylene |
Circle |
1-butene |
Triangle |
2-methyl-2-butene and 3-met-1-butene |
Diamond |
2-butene (cis & trans) |
Asterisk |
[0144] FIG. 5 shows conversions for each of the olefins in Feed 2 over Catalyst B, SPA.
Over 95% conversion of normal C
4 olefins was achieved at over 90% butene conversion. Pentene conversion reached 90%
at over 90% butene conversion. Normal butene conversion actually exceeded isobutene
conversion at high butene conversion over 95%.
EXAMPLE 6
[0145] Three feeds were oligomerized to demonstrate the ability of Catalyst A, MTT, to produce
diesel range oligomerate by recycling gasoline range oligomerate to the oligomerization
zone. Feed 1 from Example 1 with an isobutane diluent was tested along with Feed 5
which had a normal hexane diluent and Feed 6 which had an isobutane diluent but spiked
with diisobutene to represent recycled gasoline range oligomers. The feeds are shown
in Table 12. The symbols in FIG. 6 correspond to those indicated in the last row of
Table 12.
Table 12
|
Feed 1 |
Feed 5 |
Feed 6 |
Component |
Fraction, wt% |
Fraction, wt% |
Fraction, wt% |
propylene |
0.1 |
0.08 |
0.08 |
isobutane |
70.04 |
15.75 |
15.75 |
isobutylene |
7.7 |
7.3 |
7.3 |
1-butene |
5.7 |
5.1 |
5.1 |
2-butene (cis & trans) |
16.28 |
11.6 |
11.6 |
3-met-1-butene |
0.16 |
0.16 |
0.16 |
n-hexane |
0 |
60 |
0 |
acetone |
0.02 |
0.01 |
0.01 |
tert-butyl alcohol |
0 |
0.0008 |
0.0008 |
diisobutene |
0 |
0 |
60 |
Total |
100 |
100 |
100 |
FIG. 6 symbol |
square |
diamond |
asterisk |
[0146] The oligomerization conditions included 6.2 MPa pressure, 0.75 WHSV over Catalyst
A, MTT. Normal butene conversion as a function of temperature is graphed in FIG. 6
for the three feeds.
[0147] FIG. 6 demonstrates that Feed 6 with the diisobutene oligomer has greater normal
butene conversion at equivalent temperatures between 180° and 240°C. Consequently,
gasoline oligomerate recycle to the oligomerization zone will improve normal C
4 conversion. Butene conversion for Feed 5 is shown in FIG. 7 and for Feed 6 is shown
in FIG. 8. The key for FIGS. 7 and 8 is shown in Table 13.
Table 13
Component |
Symbols in FIGS. 7 & 8 |
isobutylene |
Circle |
1-butene |
Triangle |
2-butene (cis & trans) |
Asterisk |
[0148] At higher butene conversions and with diisobutene recycle, isobutene has the lowest
conversion with both 1-butene and 2-butene having greater oligomerization to oligomers.
However, without diisobutene recycle, isobutene undergoes the greatest conversion,
but with 1-butene conversion apparently surpassing isobutene conversion at over 94%
total butene conversion. We expect the same performance for Feed 1 with isobutane
diluent.
[0149] Table 14 gives feed performance for the three feeds at conditions selected to achieve
high butene conversion and high C
12+ yield including 6.2 MPa of pressure.
Table 14
Run |
Feed 1 |
Feed 5 |
Feed 6 |
WHSV, hr-1 |
0.9 |
0.6 |
0.7 |
Maximum Bed Temperature, °C |
240 |
236 |
239 |
C4 olefin conversion,% |
95 |
96 |
95 |
n-C4 olefin conversion,% |
95 |
95 |
97 |
i-C4 olefin conversion,% |
96 |
97 |
91 |
1-C4 olefin conversion,% |
97 |
98 |
97 |
2-C4 olefin conversion,% |
94 |
94 |
97 |
C5-C7 selectivity, wt% |
3 |
3 |
0.8 |
C8-C11 selectivity, wt% |
27 |
27 |
26 |
C12-C15 selectivity, wt% |
49 |
52 |
39 |
C16+ selectivity, wt% |
20 |
19 |
34 |
Total C9+ selectivity, wt% |
76 |
77 |
77 |
Total C12+ selectivity, wt% |
70 |
71 |
73 |
Diesel Yield, wt% |
72 |
74 |
73 |
[0150] C
12+ selectivity increased and C
16+ increased substantially with diisobutene presence over the feeds without diisobutene
presence. Yield calculated by multiplying C
4 olefin conversion by total C
9+ selectivity taken at the 150°C (300°F) cut point was over 70% for all feeds based
on a single pass through the oligomerization reactor.
EXAMPLE 7
[0151] Feed 1 and Feed 5 were reacted over Catalyst A, MTT, at 6.2 MPa and 0.75 WHSV. A
graph of selectivity as a function of maximum catalyst bed temperature in FIG. 9 shows
optimal maximum bed temperature between 220° and 240°C has an apex that corresponds
with maximal C
12+ olefin selectivity and to a minimum C
8-C
11 olefin selectivity and a C
5-C
7 olefin selectivity. Table 15 provides a key for FIG. 9. In FIG. 9, solid points and
lines represent Feed 1; whereas; hollow points and dashed lines represent Feed 5.
Table 15
Symbol |
Solid - Feed 1 |
Hollow - Feed 5 |
C12+ olefin selectivity |
Triangles |
C8-C11 olefin selectivity |
Circles |
C5-C7 olefin selectivity |
Greek Crosses |
Asterisks |
EXAMPLE 8
[0152] Three different feeds representing product oligomerate were subjected to micro reactor
cracking testing over three different catalysts. The three feeds were 2,4,4-trimethyl-1-pentene,
1-octene and mixed C
12 and larger olefins which contained linear molecules. The three catalysts included
a ZSM-5 additive with 40 wt% ZSM-5 crystals , Zeolite Y and a blend of 25 wt% of the
ZSM-5 additive and 75 wt% Zeolite Y such that 10 wt% of the blend was ZSM-5 crystals.
The test conditions included 565°C, 10.3 kPa (gauge) and a residence time of 0.05
seconds at standard feed conditions of 25°C and atmospheric pressure. The feeds were
a mixture of 10 mol-% hydrocarbon, 5 mol-% steam, and the balance nitrogen. Table
16 provides the key for FIGS. 10-12.
Table 16
Component |
Key |
Conversion,% |
Diagonal lines |
C3 olefin yield, wt% |
Dotted fill |
C4 olefin yield, wt% |
Cross Hatch |
C5 olefin yield, wt% |
Diagonal Cross Hatch |
ZSM-5 |
Left |
Zeolite Y |
Middle |
Blend of ZSM-5 and Zeolite Y |
Right |
Trimethyl pentene feed |
FIG. 10 |
1-Octene feed |
FIG. 11 |
Mixed C12 olefins |
FIG. 12 |
[0153] FIG. 10 reveals that achieving high conversion of 2,4,4,-trimethyl-1-pentene over
ZSM-5 alone was very difficult. The same feed over Zeolite Y or the blend of ZSM-5
and Zeolite Y reached high conversion easily. The blend of ZSM-5 and Y zeolite had
the highest propylene yield. FIG. 11 shows that the conversion of 1-octene was very
high over all three catalysts. We saw a similar pattern for methyl heptene in a separate
test. Again, the blend of ZSM-5 and Y zeolite had the highest propylene yield. FIG.
12 shows that conversion of C
12 and larger olefins, propylene tetramer, over the blend of ZSM-5 and Y zeolite had
the highest propylene yield of all the feeds tested. ZSM-5 alone was not able to achieve
much conversion of the C
12 and larger olefin feed.
[0154] This example establishes that feeding oligomerate produced over Catalyst A of Example
1, MTT, which produces less of the trimethyl pentene but more of the linear and less-branched
C
8 olefins and C
12 olefins to an FCC unit will provide the best FCC feed to crack into the most propylene.
EXAMPLE 9
[0155] Three feeds were reacted over FCC equilibrium catalyst comprising 8 wt% ZSM-5. Feed
7 comprised hydrotreated VGO with a hydrogen content of 13.0 wt%. Feed 8 comprised
the same VGO mixed with 25 wt% oligomerate product catalyzed by Catalyst A of Example
1. Feed 9 comprised the same VGO mixed with 25 wt% oligomerate product catalyzed by
Catalyst B of Example 1. The feeds were heated to 260-287°C and contacted with the
FCC catalyst in a riser apparatus to achieve 2.5-3.0 seconds of residence time. FIG.
13 plots C
3 olefin yield versus VGO conversion. The key for FIG. 13 is in Table 17.
Table 17
Feed |
Composition |
Key |
Feed 7 |
VGO |
Solid diamond |
Feed 8 |
VGO/MTT oligomerate |
Square |
Feed 9 |
VGO/SPA oligomerate |
Triangle |
[0156] FIG. 13 shows that recycle of oligomerate product to the FCC zone can boost propylene
production. At the apex of the propylene yield curve of VGO alone, the feed comprising
VGO and oligomerate provided 3.2 wt% more propylene yield from the FCC zone.
[0157] Without further elaboration, it is believed that one skilled in the art can, using
the preceding description, utilize the present invention to its fullest extent. Preferred
embodiments of this invention are described herein, including the best mode known
to the inventors for carrying out the invention. The preceding preferred specific
embodiments are, therefore, to be construed as merely illustrative, and not limitative
of the remainder of the disclosure in any way whatsoever.
[0158] In the foregoing, all temperatures are set forth in degrees Celsius and, all parts
and percentages are by weight, unless otherwise indicated. Pressures are given at
the vessel outlet and particularly at the vapor outlet in vessels with multiple outlets.
Control valves should be opened or closed as consistent with the intent of the disclosure.