NZ751807B2 - Fluid catalytic cracking process and apparatus for maximizing light olefin yield and other applications - Google Patents
Fluid catalytic cracking process and apparatus for maximizing light olefin yield and other applications Download PDFInfo
- Publication number
- NZ751807B2 NZ751807B2 NZ751807A NZ75180717A NZ751807B2 NZ 751807 B2 NZ751807 B2 NZ 751807B2 NZ 751807 A NZ751807 A NZ 751807A NZ 75180717 A NZ75180717 A NZ 75180717A NZ 751807 B2 NZ751807 B2 NZ 751807B2
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- catalyst
- reactor
- particles
- separator
- hydrocarbon
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- 150000001336 alkenes Chemical class 0.000 title abstract description 62
- 238000004231 fluid catalytic cracking Methods 0.000 title description 69
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- ASCUXPQGEXGEMJ-GPLGTHOPSA-N [(2R,3S,4S,5R,6S)-3,4,5-triacetyloxy-6-[[(2R,3R,4S,5R,6R)-3,4,5-triacetyloxy-6-(4-methylanilino)oxan-2-yl]methoxy]oxan-2-yl]methyl acetate Chemical compound CC(=O)O[C@@H]1[C@@H](OC(C)=O)[C@@H](OC(C)=O)[C@@H](COC(=O)C)O[C@@H]1OC[C@@H]1[C@@H](OC(C)=O)[C@H](OC(C)=O)[C@@H](OC(C)=O)[C@H](NC=2C=CC(C)=CC=2)O1 ASCUXPQGEXGEMJ-GPLGTHOPSA-N 0.000 description 1
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- RWSOTUBLDIXVET-UHFFFAOYSA-N dihydrogen sulfide Chemical compound S RWSOTUBLDIXVET-UHFFFAOYSA-N 0.000 description 1
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Classifications
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-
- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
- C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
- C10G2300/10—Feedstock materials
- C10G2300/1037—Hydrocarbon fractions
-
- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
- C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
- C10G2300/10—Feedstock materials
- C10G2300/1077—Vacuum residues
-
- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
- C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
- C10G2300/20—Characteristics of the feedstock or the products
- C10G2300/201—Impurities
- C10G2300/202—Heteroatoms content, i.e. S, N, O, P
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- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
- C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
- C10G2400/02—Gasoline
-
- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
- C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
- C10G2400/20—C2-C4 olefins
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- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
- C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
- C10G2400/22—Higher olefins
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- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
- C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
- C10G2400/30—Aromatics
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- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
- C10G51/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only
- C10G51/02—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only plural serial stages only
-
- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
- C10G51/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only
- C10G51/02—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only plural serial stages only
- C10G51/026—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only plural serial stages only only catalytic cracking steps
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- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
- C10G51/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only
- C10G51/06—Treatment of hydrocarbon oils, in the absence of hydrogen, by two or more cracking processes only plural parallel stages only
-
- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
- C10G55/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one refining process and at least one cracking process
- C10G55/02—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one refining process and at least one cracking process plural serial stages only
- C10G55/06—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one refining process and at least one cracking process plural serial stages only including at least one catalytic cracking step
Abstract
Apparatus and processes herein provide for converting hydrocarbon feeds to light olefins and other hydrocarbons. The problem being targeted is to provide enhanced productivity and/or flexibility of mixed catalyst systems. The processes and apparatus include, in some embodiments, feeding a hydrocarbon, a first catalyst and a second catalyst to a reactor, wherein the first catalyst has a smaller average particle size and is less dense than the second catalyst. A first portion of the second catalyst may be recovered as a bottoms product from the reactor, and a cracked hydrocarbon effluent, a second portion of the second catalyst, and the first catalyst may be recovered as an overhead product from the reactor. The second portion of the second catalyst may be separated from the overhead product, providing a first stream comprising the first catalyst and the hydrocarbon effluent and a second stream comprising the separated second catalyst, allowing return of the separated second catalyst in the second stream to the reactor. The contribution to the art is the separation of the particles for the goal of catalyst concentration in the second reactor which provides, in particular, enhanced operability, improved flexibility and ability to enhance second catalyst concentrations within the reactor. n, a first catalyst and a second catalyst to a reactor, wherein the first catalyst has a smaller average particle size and is less dense than the second catalyst. A first portion of the second catalyst may be recovered as a bottoms product from the reactor, and a cracked hydrocarbon effluent, a second portion of the second catalyst, and the first catalyst may be recovered as an overhead product from the reactor. The second portion of the second catalyst may be separated from the overhead product, providing a first stream comprising the first catalyst and the hydrocarbon effluent and a second stream comprising the separated second catalyst, allowing return of the separated second catalyst in the second stream to the reactor. The contribution to the art is the separation of the particles for the goal of catalyst concentration in the second reactor which provides, in particular, enhanced operability, improved flexibility and ability to enhance second catalyst concentrations within the reactor.
Description
/350WO1
FLUID CATALYTIC CRACKING PROCESS AND
APPARATUS FOR MAXIMIZING LIGHT OLEFIN YIELD
AND OTHER APPLICATIONS
FIELD OF THE DISCLOSURE
Embodiments herein lly relate to systems and processes for enhancing
the productivity and/or flexibility of mixed catalyst systems. Some ments
disclosed herein relate to a fluid catalytic cracking apparatus and process for
maximizing the conversion of a heavy hydrocarbon feed, such as vacuum gas oil
and/or heavy oil residues into very high yield of light olefins, such as propylene and
ethylene, aromatics and gasoline with high octane number.
OUND
In recent times, production of light s via fluid catalytic cracking (FCC)
processes has been considered one of the most attractive itions. Additionally,
there is an ever increasing demand for petrochemical building blocks such as
propylene, ethylene, and aromatics (benzene, toluene, xylenes, etc.). Further,
integration of petroleum refineries with a hemicals complex has become a
preferred option for both economic and nmental reasons.
Global trends also show that there is sed demand for middle distillates
(diesel) than that of gasoline product. In order to maximize middle lates from
FCC process, it is required to operate FCC at lower reactor temperature and a
different catalyst formulation. The downside of such change is decreased light
olefins yield because of FCC unit operating at much lower reactor temperature. This
will also reduce feedstock for Alkylation units.
Several fluidized bed catalytic ses have been developed over the last
two decades, adapting to the changing market demands. For example, US7479218
discloses a fluidized catalytic reactor system in which a riser-reactor is divided into
two sections of different radii in order to improve the selectivity for light olefins
production. The first part of' the riser reactor with lesser radii is employed for
cracking heavy feed molecules to naphtha range. The enlarged radii portion, the
second part of the riser r is used for further cracking of naphtha range products
into light s such as propylene, ethylene, etc. Though the reactor system concept
17344/350WO1
is fairly simple, the degree of selectivity to light olefins is limited for the following
reasons: (1) the naphtha range feed streams contact partially coked or deactivated
catalyst; (2) the temperature in the second part of the reaction section is much lower
than the first zone because of the endothermic nature of the on in both sections;
and (3) lack of the high activation energy required for light feed cracking as
compared to that of heavy hydrocarbons.
US6106697, US7128827, and 099 employ two stage fluid catalytic
cracking (FCC) units to allow a high degree of control for selective cracking of heavy
hydrocarbons and naphtha range feed streams. In the 1st stage FCC unit, consisting
of a riser reactor, stripper and regenerator for converting gas oil / heavy hydrocarbon
feeds into naphtha boiling range products, in the presence of Y-type large pore
zeolite st. A 2nd stage FCC unit with a similar set of vessels / configuration is
used for catalytic cracking of recycled naphtha streams from the 1st stage. Of course,
the 2nd stage FCC unit employs a ZSM-5 type (small pore zeolite) st to improve
the selectivity to light olefins. Though this scheme provides a high degree of control
over the feed, catalyst and operating window ion and optimization in a broad
sense, the 2nd stage processing of naphtha feed es very little coke that is
insufficient to maintain the heat balance. This demands heat from external sources
to have adequate temperature in the regenerator for achieving good combustion and
to supply heat for feed vaporization and endothermic reaction. Usually, torch oil is
burned in the 2nd stage FCC rator, which leads to excessive catalyst
deactivation due to higher catalyst le atures and hot spots.
US7658837 discloses a process and device to optimize the yields of FCC
products by utilizing a part of a conventional stripper bed as a reactive stripper. Such
reactive ing concept of second reactor compromises the stripping efficiency to
some extent and hence may lead to increased coke load to regenerator. The product
yield and selectivity is also likely to be ed due to contact of the feed with coked
or deactivated catalyst. Further, reactive stripper temperatures cannot be changed
independently because the riser top temperature is ly controlled to maintain a
desired set of conditions in the riser.
US2007/0205139 discloses a process to inject hydrocarbon feed through a
first distributor located at the bottom section of the riser for maximizing gasoline
17344/350WO1
yield. When the objective is to maximize light olefins, the feed is injected at the
upper section of the riser through a similar feed distribution system with an intention
to decrease the residence time of hydrocarbon vapors in the riser.
WO2010/067379 aims at increasing propylene and ethylene yields by
injecting C4 and olefinic naphtha streams in the lift zone of the riser below the heavy
hydrocarbon feed injection zone. These streams not only e the light olefins
yield but also act as media for catalyst transport in place of steam. This concept helps
in reducing the degree of l deactivation of the st. However, this lacks in
flexibility of varying operating conditions such as ature and WHSV in the lift
zone, which are critical for cracking of such light feed steams. This is likely to result
in inferior selectivity to the desired light s.
US6869521 ses that contacting a feed derived from FCC product
(particularly naphtha) with a catalyst in a second reactor operating in fast fluidization
regime is useful for promoting hydrogen transfer reactions and also for controlling
catalytic cracking reactions.
622 discloses an FCC process employing dual risers for converting
a C3/C4 containing feedstock to ics. The first and second hydrocarbon feeds
are supplied to the respective 1st and 2nd risers in the presence of gallium enriched
catalyst and the 2nd riser operates at higher reaction temperature than the first.
US5944982 discloses a catalytic process with dual risers for producing low
sulfur and high octane ne. The second riser is used to process recycle the heavy
naphtha and light cycle oils after treatment to maximize the gasoline yield and
octane number.
0231461 discloses a process that maximizes production of light cycle
oil (LCO) or middle distillate product and light olefins. This process employs a two
reactor system where the first reactor (riser) is used for cracking gas oil feed into
predominantly LCO and a second concurrent dense bed reactor is used for cracking
of naphtha recycled from the first reactor. This process is limited by catalyst
selectivity and lacks in the desired level of olefins in naphtha due to operation of the
first reactor at substantially lower on temperatures.
17344/350WO1
US6149875 deals with removal of feed contaminants such as concarbon and
metals with ent. The FCC catalyst is separated from adsorbent using the
differences between transport/ terminal velocity of the FCC catalyst and adsorbent.
US7381322 disclosed an apparatus and process to separate catalyst from a
metal adsorbent in stripper cum separator, before a regeneration step for eliminating
the adverse effects of contaminant metals deposited on the adsorbent. This patent
employs the difference in m / bubbling velocity differences and the
application is mainly to segregate FCC catalyst from adsorbent.
SUMMARY
It has been found that it is possible to use a two-reactor scheme to crack
hydrocarbons, including cracking of a C4, lighter C5 on, naphtha on,
methanol, etc. for the production of light olefins, where the two-reactor scheme does
not have limitations on selectivity and operability, meets heat balance requirements,
and also maintains a low piece count. Select embodiments sed herein use a
conventional riser reactor in combination with a mixed flow (e.g., including both
counter-current and co-current catalyst flows) zed bed reactor designed for
maximizing light olefins tion. The effluents from the riser reactor and mixed
flow reactor are processed in a common catalyst disengagement vessel, and the
catalysts used in each of the riser reactor and the mixed flow reactor may be
regenerated in a common catalyst ration vessel. This flow scheme is effective
for maintaining a high cracking activity, overcomes the heat balance problems, and
also improves yield and selectivity of light olefins from various hydrocarbon
streams, yet fies the product quenching and unit hardware, as will be bed
in more detail below.
In one aspect, ments sed herein relate to a process for the
conversion or catalytic cracking of hydrocarbons. The process may include feeding
a hydrocarbon, a first particle and a second particle to a reactor, where the first
particle has a r average le size and/or is less dense than the second
particle, and where the first and second particles may be catalytic or non-catalytic.
A first portion of the second particle may be recovered as a bottoms product from
the reactor; and a cracked hydrocarbon effluent, a second portion of the second
particle, and the first particle may be recovered as an overhead product from the
17344/350WO1
reactor. The second portion of the second particle may be separated from the
overhead product to e a first stream comprising the first le and the
arbon effluent and a second stream comprising the separated second particle,
ng return of the separated second particle in the second stream to the reactor.
In another aspect, embodiments disclosed herein relate to a system for the
catalytic cracking of hydrocarbons. The system may include a first reactor for
contacting a first and a second cracking st with a hydrocarbon feedstock to
convert at least a portion of the hydrocarbon feedstock to lighter hydrocarbons. An
overhead product line provides for recovering from the first reactor a first stream
comprising first ng catalyst, a first portion of the second cracking catalyst, and
hydrocarbons. A s t line provides for recovering from the first reactor
a second stream comprising a second portion of the second cracking catalyst. A
separator may be used for ting second cracking catalyst from the first stream,
producing a hydrocarbon effluent comprising hydrocarbons and the first cracking
catalyst. A feed line is provided for ing separated second cracking catalyst
from the separator to the first reactor.
The system for catalytic cracking of hydrocarbons may also include a riser
reactor for contacting a mixture of the first cracking catalyst and the second cracking
catalyst with a second hydrocarbon feedstock to convert at least a portion of the
second hydrocarbon feedstock to lighter hydrocarbons and recover a riser reactor
effluent comprising the lighter hydrocarbons and the mixture of the first cracking
catalyst and the second ng catalyst. A second separator may be provided for
separating the second cracking catalyst from the hydrocarbon effluent and for
separating the mixture of first and second cracking catalysts from the riser reactor
effluent. A catalyst regenerator for regenerating first and second ng catalyst
recovered in the second separator and the second portion of the first cracking catalyst
recovered in the bottoms product line may also be used.
In another aspect, embodiments disclosed herein relate to a process for the
conversion of hydrocarbons. The process may include: feeding a first catalyst to a
reactor; feeding a second st to the reactor, wherein the first st has a
smaller e particle size and/or is less dense than the first catalyst, and feeding
a arbon feedstock to the reactor. An overhead effluent may be recovered from
17344/350WO1
the reactor, the nt including cracked hydrocarbon, the first catalyst, and the
second catalyst. The second catalyst may be separated from the ad product to
provide a first stream comprising the first catalyst and the hydrocarbon effluent and
a second stream comprising the separated second catalyst, allowing return of the
ted second st in the second stream to the r.
In another aspect, embodiments herein are directed toward a separator for
separating catalysts or other particles based on size and/or density difference. The
tor may have a minimum of one inlet and may also have a minimum of two
outlets for separating particles from carrier gases. The carrier gas enters the separator
with the particles whereupon inertial, centrifugal and/or gravitational forces may be
exerted on the particles such that a portion of the particles and carrier gas are
collected in the first outlet and a portion of the particles along with the carrier gas
are collected in the second outlet. The combination of forces in the separator may
have the effect of enriching an outlet stream in particle size and/or density versus the
inlet concentration. The separator may have additional carrier gas distribution or
fluidization inside of the vessel/chamber to exert additional forces on the particles
which may tate enhanced classification.
In r , embodiments herein are directed toward an inertial
separator for separating catalysts or other particles based on size and/or density. The
inertial separator may include an inlet for receiving a mixture comprising a carrier
gas, a first particle type, and a second particle type. Each particle type may have an
average particle size and a particle size distribution, which may be different or
pping, and an average density. The second particle type may have an average
particle size and/or average density greater than the first particle type. The inertial
separator may include a U-shaped conduit including a first vertical leg, a base of the
U-shape, and a second al leg. The U-shaped conduit may fluidly connect the
inlet via the first vertical leg to a first outlet and a second outlet, the first outlet being
connected ate the base of the ed conduit and the second outlet being
connected to the second vertical leg. The U-shaped inertial separator may be
configured to: separate at least a portion of the second particle type from the carrier
gas and the first particle type, recover the second particle type via the first , and
recover the carrier gas and the first particle type via the second outlet. The separator
17344/350WO1
may also include a distributor disposed within or proximate the second outlet for
introducing a fluidizing gas, facilitating additional separation of the first le type
from the second le type. The separator, in some ments, may be
configured such that a cross-sectional area of the U-shaped conduit or a portion
thereof is adjustable. For example, in some embodiments the separator may include
a movable baffle disposed within one or more sections of the U-shaped conduit.
In r aspect, embodiments herein are directed toward an inertial
separator for separating catalysts or other particles based on size and/or density as
above. The al separator may include an inlet horizontal t which traverses
a chamber before being deflected by a baffle. The chamber is connected to a first
vertical outlet and a first horizontal . The baffle may be located in the middle,
proximate the inlet, or proximate the outlet of the chamber. The baffle may be at an
angle or moveable such that to t more or less catalyst particles. The baffle
chamber separator may be configured to: separate at least a portion of the second
particle type from the carrier gas and the first particle type, recover the second
particle type via the first vertical outlet and recover the carrier gas and the first
particle type via the first horizontal outlet. The separator may also include a
distributor disposed within or proximate the first vertical outlet for introducing a
fluidizing gas, facilitating additional separation of the first particle type from the
second particle type.
In another aspect, embodiments herein are directed toward an al
separator for separating catalysts or other particles based on size and/or density as
above. The inertial separator may include a vertical inlet connected to a chamber
where one or more vertical sides of the chamber are equipped with narrow slot
outlets, which may be described as louvers. The number of louvers may vary
depending on the application and the angle of the louver may be adjustable in order
to control the amount of vapor leaving the louver s. The chamber is also
ted to a first al outlet at the bottom of the chamber. The louver separator
may be configured to: separate at least a portion of the second particle type from the
carrier gas and the first particle type, recover the second particle type via the first
vertical outlet and recover the carrier gas and the first particle type via the louver
s. The separator may also include a distributor disposed within or proximate
17344/350WO1
the first vertical outlet for introducing for ucing a fluidizing gas, tating
additional separation of the first particle type from the second particle type.
The above described separators may also be used in association with reactors,
regenerators, and catalyst feed systems to enhance system performance and
flexibility.
In one aspect, embodiments disclosed herein relate to a process for the
conversion of hydrocarbons. The s may include regenerating a cat alyst
e comprising a first catalyst and a second particle in a regenerator, wherein
the first catalyst has a smaller average particle size and/or is less dense than the
second particle, and wherein the second particle may be catalytic or talytic.
The catalyst mixture and hydrocarbons may be fed to a riser reactor to convert at
least a portion of the hydrocarbons and recover a first effluent comprising the catalyst
mixture and converted hydrocarbons. The catalyst mixture may also be fed to a
second reactor. Feeding a hydrocarbon feedstock to the second reactor and fluidizing
the catalyst e may contact the hydrocarbon feedstock with the catalyst mixture
to convert the hydrocarbons and provide for recovering an overhead product from
the second r comprising the second particle, the first catalyst, and a reacted
hydrocarbon t. The second particle may then be separated from the overhead
product to provide a first stream comprising the first catalyst and the reacted
hydrocarbon t and a second stream comprising the separated second particle,
returning the separated second particle in the second stream to the reactor.
In another aspect, ments disclosed herein relate to a process for the
conversion of hydrocarbons. The pr ocess may include withdrawing a mixture
comprising a first catalyst and a second catalyst from a catalyst regenerator and
feeding the mixture and hydrocarbons to a riser reactor to convert at least a portion
of the hydrocarbons and recover a first effluent comprising the catalyst mixture and
converted hydrocarbons, wherein the first catalyst has a smaller average particle size
and/or is less dense than the second catalyst. The process may also include
awing the mixture comprising a first catalyst and a second catalyst from the
catalyst regenerator and feeding the mixture to a catalyst tion system,
fluidizing the mixture comprising the first catalyst and the second catalyst with a
zation medium, and separating the first catalyst from the second catalyst in the
17344/350WO1
catalyst separation system to recover a first stream comprising the first catalyst and
the fluidization medium and a second stream comprising the second catalyst. A
hydrocarbon feedstock and either the first stream or the second stream may then be
fed to a reactor to react at least a portion of the hydrocarbon to produce a converted
hydrocarbon.
In another aspect, embodiments disclosed herein relate to a process for the
sion of hydrocarbons. The s may include feeding a hyd rocarbon
feedstock and a st mixture comprising a first catalyst and a second catalyst to
a riser reactor, wherein the first st has a smaller average particle size and/or is
less dense than the second catalyst. An effluent from the riser reactor may then be
separated to recover a first stream comprising the first st and converted
hydrocarbon feedstock and a second stream comprising the second catalyst, and the
second stream may be fed to the riser reactor.
In another aspect, embodiments disclosed herein relate to a process for the
conversion of hydrocarbons. The process may include withdrawing a mixture
comprising a first catalyst and a second catalyst from a catalyst rator and
feeding the mixture to a catalyst feed / separation system, wherein the first catalyst
has a smaller e particle size and/or is less dense than the second catalyst. The
first catalyst may be separated from the second catalyst in the st feed /
separation system to e a first stream sing the first catalyst and a second
stream comprising the second catalyst. A hydrocarbon feedstock and either the first
stream or the second stream may then be fed to a riser reactor to react at least a
portion of the hydrocarbon to produce a converted hydrocarbon.
In another aspect, embodiments disclosed herein relate to a system for the
conversion of arbons. The system may include a st regenerator, and a
first catalyst feed line for withdrawing a e comprising a first catalyst and a
second catalyst from the catalyst regenerator and feeding the mixture to a riser
reactor, wherein the first catalyst has a smaller average particle size and/or is less
dense than the second catalyst. The system may also include a second catalyst feed
line for withdrawing the e comprising a first catalyst and a second catalyst
from the catalyst regenerator and feeding the mixture to a catalyst separation system,
and a fluidization medium feed line for fluidizing the mixture withdrawn via the
17344/350WO1
second catalyst feed line with a fluidization medium and separating the first catalyst
from the second catalyst in the catalyst separation system to recover a first stream
comprising the first catalyst and the fluidization medium and a second stream
comprising the second catalyst. A reactor may be provided for contacting a
hydrocarbon feedstock and either the first stream or the second stream to react at
least a portion of the arbon to produce a converted hydrocarbon.
In another aspect, embodiments disclosed herein relate to a system for the
conversion of hydrocarbons. The system may include a riser reactor for contacting
a hydrocarbon feedstock with a catalyst mixture comprising a first catalyst and a
second catalyst, wherein the first catalyst has a smaller average particle size and/or
is less dense than the second st. A catalyst separation system is provided for
separating a riser reactor effluent to recover a first stream comprising the first
catalyst and converted hydrocarbon feedstock and a second stream comprising the
second catalyst. A flow line feeds the second stream to the riser reactor.
In another aspect, ments disclosed herein relate to a system for the
sion of hydrocarbons. The system may include a catalyst awal line for
withdrawing a mixture comprising a first catalyst and a second catalyst from a
catalyst rator and feeding the mixture to a catalyst feed / separation ,
wherein the first catalyst has a r average particle size and/or is less dense than
the second catalyst. The catalyst feed / separation system separates the first catalyst
from the second catalyst in the catalyst feed / separation system to produce a first
stream comprising the first catalyst and a second stream comprising the second
catalyst. A riser reactor contacts a hydrocarbon feedstock and either the first stream
or the second stream to react at least a n of the hydrocarbon to produce a
converted hydrocarbon.
The apparatus and processes disclosed herein use significantly different
technique than disclosed in the above s (such as US6149875 and US7381322)
to separate particulate mixtures. The purpose of the present disclosure is also
different; the prior art disclosures focus on removing the contaminants from the
catalyst by introducing an adsorbent. However, the present invention aims at
improving the sion, selectivity and heat balance by concentrating a selected
catalyst in a reactor, such as concentrating the 11 in the second reactor.
17344/350WO1
In summary, most of the state of the art included dual reactor
urations or two stage fluid catalytic cracking process schemes/ apparatus. The
second / parallel reactor used for processing light feed (naphtha or/and C4 streams)
are either concurrent tic flow riser type or dense bed reactors. It is well
known in the art that ZSM-5 is preferable catalyst/ additive to convert naphtha / C4
streams into propylene and ethylene. However, in processes employing two reactors,
the second reactor also receives Y-zeolite catalyst with small fractions of ZSM-5
additive. In other process schemes, FCC type r-regenerator concepts are
employed for maximizing light s from naphtha/ C4 streams. Such s
pose heat balance problems due to insufficient coke production. The ses and
systems disclosed herein considers separating sts, such as ZSM-5 or ZSM-11
additive from Y-zeolite & ZSM-5/ ZSM-11, in a mixture, so as to have optimal
concentration of ZSM-5 or 11 in the second reactor processing light feed. In addition,
integration of said additional/ second reactor with a conventional FCC unit
essentially helps overcoming these cks (product selectivity and heat balance
in ular) of the prior part and substantially increases the overall conversion and
light olefins yield and increases the capability to process heavier feedstocks.
Other aspects and advantages will be apparent from the following description
and the appended claims.
BRIEF DESCRIPTION OF DRAWINGS
Figure 1 is a simplified process flow diagram of a system for cracking
hydrocarbons and producing light olefins according to one or more embodiments
disclosed herein.
Figures 2-5 are simplified process flow diagrams of separators useful in
systems according to one or more embodiments disclosed herein.
Figure 6 is a simplified process flow diagram of a system for cracking
hydrocarbons and producing light olefins according to one or more embodiments
disclosed herein.
Figure 7 is a simplified process flow diagram of a system for cracking
arbons and producing light olefins according to one or more embodiments
disclosed herein.
17344/350WO1
Figure 8A is a simplified process flow diagram of a system for cracking
hydrocarbons and producing light olefins according to one or more embodiments
disclosed herein.
Figure 8B is a simplified process flow diagram of a system for ng
hydrocarbons and producing light olefins ing to one or more embodiments
disclosed herein.
Figure 9A is a simplified process flow diagram of a system for cracking
hydrocarbons and producing light olefins ing to one or more embodiments
disclosed herein.
Figure 9B is a simplified process flow diagram of a system for cracking
hydrocarbons and producing light olefins ing to one or more embodiments
disclosed herein.
Figure 10 is a simplified process flow diagram of a system for cracking
hydrocarbons and producing light olefins according to one or more embodiments
disclosed herein.
Figure 11 is a simplified process flow m of a system for cracking
hydrocarbons and producing light olefins according to one or more embodiments
disclosed herein.
DETAILED DESCRIPTION
As used herein, the terms “catalyst” and cle” and like terms may be
used interchangeably. Summarized above, and as r described below,
embodiments herein separate mixed particulate als based on size and/or
density to achieve an advantageous effect in a reactor system. The particles or
particulate materials used to facilitate catalytic or thermal reaction may include
catalysts, ents, and/or heat transfer materials having no catalytic activity, for
In one aspect, ments herein relate to a fluid catalytic cracking
apparatus and process for maximizing the conversion of a heavy hydrocarbon feed,
such as vacuum gas oil and/or heavy oil residues into very high yield of light olefins,
such as propylene and ethylene, aromatics and gasoline with high octane number or
middle lates, while concurrently minimizing the yield of heavier bottom
product. To accomplish this goal, a secondary/second reactor, which may be a mixed
17344/350WO1
flow reactor (including both co-current and counter-current flow of particles with
respect to vapor flow) or a st-concentrating reactor, can be integrated with a
conventional fluid catalytic cracking reactor, such as a riser reactor. A heavy
hydrocarbon feed is catalytically d to naphtha, middle distillates and light
olefins in the riser reactor, which is a pneumatic flow co-current type reactor. To
e the yields and selectivity to light olefins (ethylene and propylene), cracked
hydrocarbon products from the riser reactor, such as C4 and a range
arbons (olefins and ins), may be recycled and processed in the
secondary reactor (the mixed flow reactor or the catalyst-concentrating reactor).
Alternatively, or additionally, external feed s, such as C4, naphtha, or other
hydrocarbon ons from other processes such as a steam cracker, metathesis
reactor, or delayed coking unit, and naphtha range streams, such as straight run
naphtha or from delayed coking, visbreaking or natural gas condensates, among
other hydrocarbon feedstocks, may be sed in the secondary r to produce
light olefins, such as ne and propylene. The integration of the secondary
r with a conventional FCC riser reactor according to embodiments disclosed
herein may overcome the drawbacks of prior processes, may substantially increase
the overall conversion and light olefins yield, and/or may increases the capability to
process heavier feedstocks.
Integration of the secondary reactor with a conventional FCC riser reactor
according to ments disclosed herein may be facilitated by (a) using a common
catalyst regeneration vessel, (b) using two types of catalyst, one being selective for
cracking heavier hydrocarbons and the other being selective for the cracking of C4
and naphtha range hydrocarbons for the production of light olefins, and (c) using a
mixed flow reactor or a catalyst-concentrating reactor in a flow regime that will
partially separate the two types of catalysts, ng the t of the C4s or
naphtha feed with the catalyst selective for cracking the same and producing light
olefins.
To enhance the operation window of the secondary reactor, and to provide
greater process flexibility, the secondary reactor may be operated in a flow regime
to entrain the catalyst selective for cracking heavier hydrocarbons, and to entrain a
portion of the catalyst selective for the cracking of C4 and naphtha range
17344/350WO1
hydrocarbons. The cracked hydrocarbon products and the entrained catalysts are
then fed to a separator to separate the catalyst selective for the cracking of C4 and
naphtha range hydrocarbons from the cracked hydrocarbon ts and the catalyst
selective for cracking heavier hydrocarbons. This solids separation vessel is an
external vessel to the reactor and is operated at hydrodynamic ties that enhance
the separation of the two types of catalyst based on their physical properties, such as
particle size and/or density. The separated catalyst, selective for the cracking of C4
and naphtha range arbons, may then be returned to the reactor for continued
reaction and providing an enhanced concentration of the catalyst selective for the
cracking of C4 and naphtha range hydrocarbons within the reactor, improving
ivity of the overall process while also improving the overall process flexibility
due to the enhanced operating window.
As noted above, the cracking system may utilize two types of catalysts, each
favoring a different type of hydrocarbon feed. The first cracking st may be a
Y-type zeolite catalyst, an FCC catalyst, or other similar catalysts useful for cracking
heavier hydrocarbon feedstocks. The second cracking catalyst may be a ZSM-5 or
ZSM-11 type catalyst or similar catalyst useful for cracking C4s or naphtha range
hydrocarbons and selective for producing light olefins. To facilitate the two-reactor
scheme disclosed , the first cracking st may have a first average particle
size and density, and may be smaller and/or lighter than those for the second cracking
catalyst, such that the catalysts may be separated based on density and/or size (e.g.,
based on terminal velocity or other teristics of the catalyst particles).
In the st regeneration vessel, spent catalyst recovered from both the
riser reactor and the ary reactor is rated. Following regeneration, a first
portion of the mixed st may be fed from the regeneration vessel to a riser
reactor (co-current flow reactor). A second portion of the mixed catalyst may be fed
from the ration vessel to the secondary reactor.
In the co-current flow reactor, a first hydrocarbon feed is contacted with a
first portion of the regenerated catalyst to crack at least a n of the hydrocarbons
to form lighter hydrocarbons. An effluent may then be recovered from the co-current
flow r, the effluent comprising a first cracked hydrocarbon product and a spent
mixed catalyst fraction.
17344/350WO1
In some embodiments, the secondary reactor is operated in a fluidization
regime sufficient to entrain the first cracking st, and the second ng
st with the hydrocarbon products recovered as an effluent from the secondary
reactor overhead outlet. The nt is then fed to a separator to separate the cracked
hydrocarbon products and the first cracking catalyst from the second cracking
catalyst.
The vapor / first ng st stream recovered from the separator may
then be forwarded for tion. The second cracking catalyst recovered from the
separator may be recycled back to the secondary reactor for continued reaction, as
noted above.
The first effluent (cracked hydrocarbons and spent mixed catalyst from the
riser reactor) and the second nt (cracked hydrocarbons and separated first
cracking catalyst from the secondary reactor) may both be fed to a disengagement
vessel to separate the spent mixed catalyst fraction and the separated first cracking
catalyst from the first and second cracked hydrocarbon products. The cracked
hydrocarbon products, including light olefins, C4 hydrocarbons, naphtha range
hydrocarbons, and heavier hydrocarbons may then be separated to recover the
desired products or product fractions.
Thus, processes disclosed herein integrate a secondary flow or
catalyst-concentrating reactor, al solids tor, and a riser reactor, with
common product separations and catalyst regeneration, where the catalysts used in
the secondary reactor is highly selective for cracking C4 and a range
hydrocarbons to produce light olefins. The common catalyst regeneration provides
for heat balance, and the common product separation (disengagement vessel, etc.)
provides for simplicity of operations and reduced piece count, among other
advantages.
Referring now to Figure 1, a fied process flow diagram of systems for
cracking hydrocarbons and producing light olefins according to embodiments
disclosed herein is illustrated. The system es a two-reactor uration for
maximizing yield of propylene and ethylene from petroleum residue feedstocks or
other hydrocarbon streams. The first reactor 3 may be a riser reactor for cracking
heavier hydrocarbon feeds, for example. The second reactor 32 is a fluidized bed
17344/350WO1
reactor, which may be equipped with baffles or internals. The C4 olefins and/or light
naphtha ts from the first reactor 3 or similar feed streams from external
sources may be processed in the second reactor 32 to enhance the yield of light
olefins, including propylene and ethylene, and aromatics / high octane gasoline.
A heavy petroleum residue feed is injected through one or more feed injectors
2 located near the bottom of first r 3. The heavy petroleum feed contacts hot
regenerated catalyst introduced through a J-bend 1. The catalyst fed to the first
reactor 3 is a catalyst mixture, ing a first st selective for cracking heavier
hydrocarbons, such as a Y-type zeolite based catalyst, and a second catalyst selective
for the cracking of C4 and naphtha range arbons for the production of light
olefins, such as a ZSM-5 or ZSM-11, which may also be used in combination with
other catalysts. The first and second catalysts may be different in one or both particle
size and density. A first catalyst, such as the Y-type based e, may have a
particle size in the range of 20 – 200 microns and an apparent bulk density in the
range of 0.60 – 1.0 g/ml. A second catalyst, such as ZSM-5 or ZSM-11, may have
a particle size in the range of 20 – 350 microns and an apparent bulk density in the
range of 0.7 – 1.2 g/ml.
The heat required for vaporization of the feed and/or raising the temperature
of the feed to the desired reactor temperature, such as in the range from 500°C to
about 700°C, and for the endothermic heat (heat of reaction) may be provided by the
hot regenerated catalyst coming from the regenerator 17. The pressure in first reactor
3 is typically in the range from about 1 barg to about 5 barg.
After the major part of the cracking reaction is completed, the e of
products, unconverted feed vapors, and spent catalyst flow into a two stage cyclone
system housed in cyclone containment vessel 8. The two -stage cyclone system
includes a primary cyclone 4, for separating spent st from vapors. The spent
catalyst is discharged into stripper 9 through primary cyclone dip leg 5. Fine catalyst
particles entrained with the separated vapors from y e 4 and t
vapors from second reactor 32, uced via flow line 36a and a single stage
e 36c, are separated in second stage cyclone 6. The catalyst mixture collected
is discharged into stripper 9 via dip leg 7. The vapors from second stage cyclone 6
are vented through a secondary cyclone outlet 12b, which may be connected to
17344/350WO1
plenum 11, and are then routed to a main fractionator / gas plant (not shown) for
recovery of products, including the desired olefins. If necessary, the product vapors
are further cooled by introducing light cycle oil (LCO) or steam via distributor line
12a as a quench media.
The spent catalyst recovered via dip legs 5, 7 undergoes ing in stripper
bed 9 to remove interstitial vapors (the hydrocarbon vapors trapped between catalyst
particles) by countercurrent contacting of steam, introduced to the bottom of stripper
9 through a steam distributor 10. The spent catalyst is then transferred to regenerator
17 via the spent catalyst standpipe 13a and lift line 15. Spent catalyst slide valve
l3b, located on spent catalyst standpipe 13a is used for controlling catalyst flow from
stripper 9 to regenerator 17. A small portion of combustion air or en may be
introduced h a distributor 14 to help smooth transfer of spent catalyst.
Coked or spent catalyst is discharged through spent catalyst distributor 16 in
the center of the dense regenerator bed 24. tion air is introduced by an air
distributor 18 located at the bottom of regenerator bed 24. Coke ted on the
catalyst is then burned off in regenerator 17 via reaction with the combustion air.
Regenerator 17, for example, may operate at a temperature in the range from about
640°C to about 750°C and a pressure in the range from about 1 barg to about 5 barg.
The catalyst fines entrained along with flue gas are collected in first stage cyclone
19 and second stage cyclone 21 and are discharged into the regenerator catalyst bed
through respective dip legs 20, 22. The flue gas recovered from the outlet of second
stage cyclone 21 is directed to flue gas line 50 via regenerator plenum 23 for
downstream waste heat recovery and/or power recovery.
A first part of the regenerated st mixture is withdrawn via regenerated
catalyst standpipe 27, which is in flow communication with J bend 1. The catalyst
flow from regenerator 17 to reactor 3 may be regulated by a slide valve 28 located
on regenerated st standpipe 27. The opening of slide valve 28 is adjusted to
control the st flow to maintain a desired top temperature in reactor 3.
In addition to lift steam, a provision is also made to inject feed streams such
as C4 olefins and a or similar external streams as a lift media to J bend 1
through a gas butor 1a located at the Y-section for enabling smooth transfer of
regenerated st from J bend 1 to reactor 3. J bend 1 may also act as a dense bed
17344/350WO1
reactor for cracking C4 olefins and naphtha streams into light olefins at conditions
ble for such reactions, such as a WHSV of 0.5 to 50 h-1, a temperature of 640°C
to 750°C, and residence times from 3 to 10 seconds.
A second part of the regenerated catalyst mixture is withdrawn into a second
reactor 32 through a standpipe 30. A slide valve 31 may be used to l the
catalyst flow from regenerator 17 to second reactor 32 based on a vapor outlet
temperature set point. C4 olefins and naphtha streams are injected into the bottom
section of the catalyst bed through one or more feed distributors 34 (34a, 34b), either
in liquid or vapor phase. Second reactor 32 operates in a mixed flow fashion, where
a portion of the regenerated catalyst flows downward (from the top to the bottom of
the reactor bed) and a portion of the regenerated st mixture and the feed
hydrocarbon stream flows upward (from the bottom to the top of the reactor bed).
Second reactor 32 may be equipped with baffles or structured internals (not
shown) that help te contact and mixing of catalyst and feed molecules. These
internals may also help in minimizing channeling, bubble , and/or
coalescence. Second reactor 32 may also be enlarged at ent sections along the
length to maintain a constant or desired superficial gas velocity within the sections.
After the reaction is completed, the catalyst is stripped at the bottommost
portion of second reactor 32 to separate entrained hydrocarbon feed / products using
steam as a stripping media introduced through distributor 35. The spent catalyst
recovered at the bottom of reactor 32 is then transferred to regenerator 17 via
standpipe 37 and lift line 40 through a spent catalyst distributor 41. Combustion air
or nitrogen may be introduced h distributor 39 to enable smooth transfer of
catalyst to rator 17. Slide valve 38 may be used to l the catalyst flow
from second reactor 32 to regenerator 17. Spent catalyst from both reactors 3, 32 is
then regenerated in the common regenerator 17, operating in a complete combustion
mode.
As noted above, second reactor 32 utilizes two different catalysts that may
differ in one or both of particle size and density, such as a r and smaller Y-type
e or FCC catalyst and a larger and/or denser ZSM-5/ ZSM-11 shape-selective
pentacil small pore zeolite. The superficial gas velocity in second r 32 is
maintained such that ially all or a large portion of the lighter, smaller catalyst
17344/350WO1
(e.g., Y-type zeolite / FCC st) and a n of the heavier, larger catalyst (e.g.,
ZSM-5 / ZSM-11) is carried out of the reactor with the cracked hydrocarbons and
steam recovered via flow line 45. A n of the larger and/or denser catalyst may
be retained within the r 32, forming a dense bed toward the lower portion of
the reactor, as noted above.
The effluent from reactor 32 recovered via flow line 45 may thus include
cracked hydrocarbon products, unreacted hydrocarbon ock, steam (stripping
, and a catalyst mixture, including essentially all of the lighter and/or smaller
catalyst and a portion of the larger and/or more dense catalyst introduced to the
reactor. The effluent may then be transported via flow line 45 to a solids separator
47. Separator 47 may be a separator configured to te the two types of catalyst
based on their physical properties, namely particle size and/or density. For example,
separator 47 may use differences in inertial forces or centrifugal forces to separate
FCC catalyst from the ZSM-5. The solids separation vessel 47 is an external vessel
to the second reactor 32 and is operated at hydrodynamic properties that enhance the
separation of the two types of catalyst based on their physical properties.
After separation in separator 47, the smaller and/or lighter catalyst (Y-type
zeolite / FCC catalyst) is then transported from separator 47 to the common
ager or containment vessel 8, housing riser reactor es and/or reaction
termination system, via outlet line 36a. The larger and/or denser catalyst (ZSM-5 /
ZSM-11) may be returned via flow line 49 to the mixed flow reactor 32 for continued
reaction with hydrocarbon feeds introduced through distributors 34.
Entrainment of essentially all of the lighter/smaller catalyst and a n of
the larger and/or more dense catalyst, subsequent separations, and e of the
larger and/or denser catalyst to reactor 32 may allow for a significant accumulation
of the larger and/or denser catalyst in reactor 32. As this catalyst is more selective
for the cracking of C4 and naphtha range hydrocarbons, the accumulation of the
larger and/or denser catalyst may provide a selectivity and yield advantage. r,
operation of the reactor in a fluidization flow regime to entrain both types of catalyst
may provide for improved operability of the reactor or flexibility in operations, as
discussed above.
350WO1
A hydrocarbon feed such as heavy vacuum gas oil or heavy residue feed,
light cycle oil (LCO), or steam may be injected as a quench media in the outlet line
36a through a distributor 36b. The flow rate of such quench media may be controlled
by setting the temperature of the stream entering the containment vessel 8. All the
vapors from second reactor 32, including those fed through distributor 36b, are
discharged into the dilute phase of containment vessel 8 through a single stage
cyclone 36c. Employing a hydrocarbon feed as a quench media is preferred as it
serves dual purpose of cooling the products from second reactor 32 and also
enhances the tion of middle lates.
The first reactor 3, such as a riser reactor, may es in the fast fluidization
regime (e.g., at a gas superficial velocity in the range from about 3 to about 10 m/s
at the bottom section) and tic transport regime (e.g., at a gas superficial
velocity in the range from about 10 to about 20 m/s) in the top section.
WHSV in second reactor 32 is typically in the range from about 0.5 h-1 to
about 50 h-1; vapor and catalyst residence times may vary from about 2 to about 20
seconds. When different feeds are introduced, preferably the C4 feed is injected at an
elevation below naphtha feed injection. However, interchanging of feed injection
locations is possible.
As necessary, make-up catalyst may be introduced via one or more flow lines
42, 43. For example, fresh or p FCC or Y-type zeolite catalyst or a mixture
of these two may be introduced to regenerator 17 via flow line 42 and fresh or makeup
ZSM-5/ ZSM-11 catalyst may be introduced to second r 32 via flow line
43. Overall system catalyst inventory may be maintained by withdrawing mixed
catalyst from regenerator 24, for example. Catalyst inventory and accumulation of
the preferred catalyst within reactor 32 may be controlled, as will be bed
below, via control of the reactor and separator 47 operations.
In some embodiments, a first part of the regenerated catalyst is withdrawn
from regenerator 17 into a Regenerated Catalyst (RCSP) hopper 26 via withdrawal
line 25, which is in flow communication with regenerator 17 and regenerated catalyst
standpipe 27. The catalyst bed in the RCSP hopper 26 floats with regenerator 17 bed
level. The rated st is then transferred from RCSP hopper 26 to reactor 3
via regenerated catalyst standpipe 27, which is in flow communication with J bend
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1. The catalyst flow from regenerator 17 to reactor 3 may be regulated by a RCSP
slide valve 28 d on regenerated catalyst standpipe 27. A pressure equalization
line 29 may also be provided.
A separator bypass line 60 may also be used to facilitate the transfer of
particles from the top of reactor 32 to the vessel 8, such as illustrated in Figure 1. As
described with respect to Figure 1 above, second reactor 32 utilizes two different
catalysts that may differ in one or both of particle size and density, such as a lighter
and/or smaller Y-type e or FCC catalyst and a larger and/or denser ZSM-5/
ZSM-11 shape-selective pentacil small pore zeolite. The superficial gas ty in
second reactor 32 may be maintained such that essentially all of the r, smaller
catalyst (e.g., Y-type zeolite / FCC st) and a portion of larger and /or more
dense catalyst (e.g., ZSM-5 / ZSM-11) is carried out of the reactor with the cracked
arbons and steam recovered via flow line 45.
The effluent from reactor 32 recovered via flow line 45 may thus e
cracked hydrocarbon products, unreacted hydrocarbon feedstock, steam (stripping
media), and a catalyst mixture, including ially all of the lighter, r catalyst
and a portion of the larger and/or more dense catalyst introduced to the reactor. The
effluent may then be transported via flow line 45 to a solids separator 47. Separator
47 may be a separator configured to te the two types of catalyst based on their
physical properties, namely particle size and/or density. The separator 47 is operated
at hydrodynamic properties that e the separation of the two types of catalyst
based on their physical properties.
After separation in separator 47, the smaller/lighter catalyst (Y-type zeolite /
FCC catalyst) is then transported from separator 47 to the common disengager or
containment vessel 8, housing riser r cyclones and/or reaction termination
, via outlet line 36a. The larger and/or denser catalyst (ZSM-5 / )
may be returned to the mixed flow/second reactor 32 for continued reaction with
hydrocarbon feeds introduced through distributors 34.
Continuously or intermittently, a portion of the effluent containing both types
of catalysts being transported via flow line 45 may be diverted to bypass separator
47. The diverted portion of the effluent may flow around separator 47 via flow line
60, which may include a diverter or flow control valve 62. The effluent may then
17344/350WO1
continue via flow line 64 back to ager 8 for separation of the arbon
products from the sts. Flow line 64 may be combined with the effluent and
smaller catalyst recovered from separator 47 via flow line 36a, and may be
introduced either upstream or downstream of quench 36b. Alternatively, the diverted
effluent in line 60 may be fed directly to disengager/ containment vessel 8.
While illustrated in Figure 1 with a diverter valve 62, ments herein
contemplate use of y-shaped flow conduit or similar apparatus to continuously send
a portion of the effluent, containing both catalyst particle types, to disengager 8,
while continuously sending a portion of the effluent to separator 47, thus ng
for the desired accumulation of the larger and/or denser catalyst particles within
reactor 32. As depicted in Figure 1, the catalyst from second reactor can also be
erred via line 37, slide valve 38 and transfer line 40 to the regenerator 17. The
blower air is used as carrier gas 39 to transfer the catalyst to regenerator 17. Such
catalyst er facility will not only help in controlling the catalyst bed level in
reactor 32 but also help in more frequent catalyst regeneration.
The use of increased flow of carrier fluid and/or the use of a flow diverter, as
described above, may beneficially provide for the accumulation of the catalyst
selective for cracking naphtha range arbons in the second reactor 32. In some
embodiments, it has been found that reactor 32 may be operated in a manner to
provide regenerated catalyst and maintain sufficient ty within the catalyst bed
of reactor 32 such that the st transfer line (flow lines 37, 40) and the associated
equipment may be omitted from the flow scheme (as shown in Figure 6) without
detriment to the selectivity and throughput of the reactor and with the added benefits
of reduced mechanical complexity and reduced capital and operating costs.
Referring now to Figure 6, a simplified s flow diagram of systems for
cracking hydrocarbons and producing light s according to embodiments
disclosed herein is illustrated, where like numerals represent like parts. Similar to
the process scheme illustrated in Figure 1, described above, the system as illustrated
in Figure 6 will have a two reactor scheme and introduce two kinds of particles (such
as a lighter and/or smaller Y-type or FCC catalyst and a larger and/or denser ZSM-
or ZSM-11 catalyst) in the second reactor 32. The larger and/or denser catalyst
ves (e.g., ZSM-5 or ZSM-11) may be added ly to the second reactor 32
17344/350WO1
via flow line 43. The regenerated st mixture transfers from regenerator 17
through pipe 30 to the second reactor 32.
The catalyst bed in the second reactor vessel 32 is expected to operate in
turbulent bed, bubbling bed or fast fluidization regimes. A light naphtha feed 34a,
such as the light a product from a first reactor or riser reactor 3, as illustrated,
may be fed into the second reactor 32 and converted to light olefins in the presence
of the mixed catalyst. The lifting gas along with product gas in the vessel 32 will lift
the solids, including both catalysts, through the pipe 45 to the solids separation vessel
47, then back to the regenerator 17. Due to the differences in size and/or density of
the two st particles, most of the ZSM-5 or ZSM-11 catalyst particles will be
separated from the Y-type or FCC catalyst in the solids separation vessel 47 and
transferred via return line 49 back to the reactor 32. Most of Y-type or FCC catalyst
particles will be transferred back to the stripper 8 for gas solid separation.
As compared to other embodiments discussed above, a y difference is
the absence of a catalyst return line and related control valves and equipment from
the bottom of the second reactor vessel 32 back to the regenerator vessel 17. As
discussed briefly above, such a process uration may still provide for efficient
catalyst regeneration, as well as lation and concentration of the desired larger
and/or denser ZSM-5 or ZSM-11 st within reactor 32. It is expected that a
higher concentration of the larger and/or denser catalyst may result in a better
performance in the second reactor vessel 32, even when the return line 37 is removed.
This design, with the removal of return line 37, also mitigates the ical
complexity and reduces the capital and operational costs.
The embodiment without a return line 37 (Figure 6) also includes steam as a
lifting gas. As there is no catalyst outlet at the bottom of the reactor 32, the catalyst
will fill up the reactor 32 and in some embodiments no catalyst bed level is observed.
The lifting gas along with product gas in the vessel 32 will lift the , including
both catalysts, through the pipe 45 to the solids separation vessel 47. Due to the
ences in size and/or density of the two catalyst particles, most of the ZSM-5 or
ZSM-11 catalyst particles will be separated from the Y-type or FCC catalyst in the
solids separation vessel 47 and transferred via return line 49 back to the reactor 32.
Most of Y-type or FCC catalyst les will be erred back to the stripper 8 for
17344/350WO1
gas solid separation. As compared to Figure 1, this design without return line 37
may lead to a much higher concentration of the larger and/or denser catalyst, which
will result in a better reaction performance in the reactor 32. Although not illustrated,
vessel 32 may include a bottom flange or outlet ng the vessel to be deinventoried
of catalyst. Such an outlet may also be used to periodically remove
larger and/or heavier catalyst particles that may accumulate within vessel 32, if
ary.
As described above, s ing to embodiments herein may include
a separator 47 configured to separate the two types of catalysts based on their
physical properties, such as particle size and/or density. Separator 47 may be a
cyclone tor, a screen separator, mechanical sifters, a gravity chamber, a
centrifugal separator, a baffle chamber, a louver separator, an in-line or pneumatic
classifier, or other types of separators useful for efficiently separating les based
on size and/or hydrodynamic properties.
Examples of separators or classifiers useful in embodiments herein are
illustrated in Figures 2-5. In some embodiments, separator 47 may be a U-shaped
al separator, as illustrated in Figure 2, to separate two kinds of solid particles
or catalysts with different particle sizes and/or particle density. The separator may
be built in the form of U-shape, having an inlet 70 at the top, a gas outlet 84 at the
other end of the U, and a main solid outlet 80 at the base of ed separator.
A mixture 72 of solid particles or catalysts with different sizes is introduced
along with a carrier gas stream through inlet 70 and inertial separation forces are
applied on the solids by making no more than one turn to separate the different sizes
of solid particles. Larger and/or more dense solid les 78 preferentially go
downward in ns 74/76 to a standpipe or dipleg 80 connected to the base of U-
shape while lighter or smaller solid particles are preferentially carried along with the
gas stream to outlet 82, where the mixture 84 of small particles and gases may be
recovered. The solid outlet 80 at the base of U-shaped separator (the inlet of the
standpipe or dipleg used to flow the larger and/or more dense catalyst particles back
to the second reactor 32) should be large enough to odate the normal
solid/catalyst flow.
17344/350WO1
By controlling the gas flow rates entering the downward standpipe and
exiting the main gas stream , the overall separation efficiency of the U-shape
inertial separator and the selectivity to separate larger and/or more dense particles
from smaller and/or less dense particles can be manipulated. This extends to a fully
sealed dipleg where the only gas stream exiting the dipleg are those entrained by the
exiting solid/catalyst flow. As the U-shaped inertial separator provides the ability to
manipulate the separation efficiency, intermediate sized particles, which have the
potential to late in the system as noted above, may be periodically or
continuously entrained with the hydrocarbon products recovered from separator 47
for separation in vessel 8 and regeneration in regenerator 24.
In some embodiments, a gas sparger 75 or extra steam/inert gas may be
provided proximate a top of outlet section 80, such as near a top of the standpipe
inlet. The additional lift gas provided within the separator may further facilitate the
tion of larger and/or more dense solid particles from less dense and/or smaller
solid particles, as the extra gas may preferentially lift lighter solid particles to gas
outlet 84, resulting in better solid classification.
The cross sectional area of the U-shaped tor at the inlet 70, outlet 82
and throughout the U-shaped separator (including areas 74, 76) may be ed to
manipulate the superficial gas velocity within the apparatus to control the separation
efficiency and the ivity. In some embodiments, a position of one or more of the
separator walls may be adjustable, or a movable baffle may be disposed within one
or more sections of the separator, which may be used to control the separation
efficiency and selectivity. In some ments, the system may include a particle
size analyzer downstream of outlet 82, enabling real-time adjustment of the flow
uration through the U-shaped separator to effect the desired separations.
ation of U-shaped inertial separators connected in series or a
combination of U-shape inertial separators and cyclones may provide flexibility to
allow simultaneously achievement of both target overall separation efficiency and
target selectivity of larger and/or more dense les over smaller and/or less dense
particles.
The second r 32 may also be equipped with s or structured
internals such as modular grids as described in US patent 7,179,427. Other types of
17344/350WO1
internals that enhance contact efficiency and product selectivity / yields may also be
used. The internals may enhance the catalyst distribution across the reactor and
improve the contact of feed vapors with st, leading to an se in the average
reaction rate, enhance the overall activity of the catalyst and optimize the operating
conditions to increase the production of light olefins.
Embodiments disclosed herein use Y-type zeolite or conventional FCC
catalyst, maximizing the conversion of heavy hydrocarbon feeds. The Y-type zeolite
or FCC catalyst is of a smaller and/or lighter particle size than the ZSM-5 or similar
catalysts used to enhance the production of light s in the countercurrent flow
reactor. The ZSM-5 or similar catalysts have a larger particle size and/or are more
dense than the Y-type zeolite or FCC catalysts used to enhance separations of the
catalyst types in each of the mixed flow r and the solids separator. The
superficial gas velocity of vapors in the second reactor is maintained such that it
allows entrainment of the Y-type zeolite or FCC catalyst and a portion of the ZSM-
or ZSM-11 catalyst out of the mixed flow reactor, and the solids separator may
utilize the differences in single particle terminal velocities or differences between
minimum fluidization / minimum ng velocities to separate and return the
ZSM-5 / ZSM-11 to the mixed flow reactor. This t allows the elimination of
two stage FCC systems and hence a simplified and efficient process. The catalysts
ed in the s could be either a combination of Y-type e / FCC
catalyst and ZSM-5 or other r catalysts, such as those mentioned in
US5043522 and US5846402.
The entrainment of both catalysts from the mixed flow reactor, subsequent
separation, and recycle and accumulation of the ZSM-5 / ZSM-11 st in the
mixed flow reactor eliminates any potential restriction on superficial gas velocity in
the second reactor. The use of a solids separation vessel thus provides process
flexibility in the second r, allowing the second r to be operated in
bubbling bed, turbulent bed, or fast fluidization regimes, rather than restricting the
operations to only a bubbling bed regime. The solids separation vessel may be a
cyclone or other vessel where solids and gases are introduced at a common inlet, and
through degassing, inertial and centrifugal forces, the particles are separated based
on size and/or density, with the majority of the smaller FCC type particles entraining
17344/350WO1
with the vapor outlet, and the larger and/or denser ZSM-5 or ZSM-11 type particles
returning via a dense phase standpipe or dipleg back to the second reactor 32.
In addition to the U-type particle separator described in relation to Figure 2,
Figures 3-5 illustrate various additional particle separation devices for use in
embodiments herein. Referring to Figure 3, a baffle chamber separator 900 for
ting catalysts or other les based on size and/or density may e an
inlet 910, such as a horizontal conduit. The vapors and particles contained in the
horizontal conduit then enter a chamber 912, before being deflected by a baffle 914.
The chamber 912 is connected to a first vertical outlet 916 and a first horizontal outlet
918. The baffle 914 may be located in the middle of chamber 912, proximate the
inlet 910, or proximate the horizontal outlet 918 of the chamber. The baffle may be
at an angle or moveable such that the baffle may be used to t more or less
catalyst particles, and may be ured for a particular mixture of les.
Processes herein may utilize the baffle r separator 900 to segregate
larger and/or denser particles from smaller and/or less dense particles contained in a
carrier gas, such as a hydrocarbon reaction effluent. The baffle chamber separator
900 may be configured to: separate at least a portion of a second particle type from
the carrier gas and a first le type, recover the second particle type via the first
vertical outlet 916 and recover a mixture including the carrier gas and the first
particle type via the first horizontal outlet 918. The separator may also include a
distributor (not rated) disposed within or ate the first vertical outlet for
introducing a fluidizing gas, facilitating additional separation of the first particle type
from the second particle type.
Referring now to Figure 4, a louver separator for use in accordance with
embodiments herein is illustrated. Similar to other separators illustrated and
described, the louver separator 1000 may be used for separating catalysts or other
particles based on size and/or density. The louver separator 1000 may include a
vertical inlet 1010 connected to a r 1012 where one or more vertical sides
1014 of the chamber are equipped with narrow slot s 1016, which may be
described as louvers. The number of louvers may vary depending on the application,
such as the desired particle mixture to be separated, and the angle of the louver may
be adjustable in order to control the amount of vapor passing through and leaving
17344/350WO1
the louver outlets. The chamber 1012 is also connected to a first vertical outlet 1014
at the bottom of the chamber.
Processes herein may utilize the louver separator 1000 to segregate larger
and/or denser particles from r and/or less dense les contained in a carrier
gas, such as a hydrocarbon reaction nt. The louver separator 1000 may be
configured to: separate at least a portion of the second particle type from the carrier
gas and the first particle type, recover the second particle type via the first vertical
outlet 1014 and recover the carrier gas and the first particle type via the louver outlets
1016. The tor may also include a butor (not illustrated) ed within
or proximate the first vertical outlet for introducing a fluidizing gas, facilitating
additional separation of the first particle type from the second particle type.
Referring now to Figure 5, an inertial separator 1100 for use in accordance
with embodiments herein is illustrated. Similar to other separators illustrated and
described, the inertial separator 1100 may be used for separating catalysts or other
particles based on size and/or y. The separator may include an inlet 1110 at
the top of and extending into a chamber 1112. In some embodiments, the height or
disposition of inlet 1110 within chamber 1112 may be adjustable. The tor may
also include one or more side outlets 1114, 1116, such as one to eight side outlets,
and a vertical outlet 1118. The separator may also include a distributor (not
illustrated) disposed within or ate the vertical outlet 1118 for introducing a
fluidizing gas.
A mixture 1172 of solid les or catalysts with different sizes is
introduced along with a carrier gas stream through inlet 1110. The gases in the
mixture 1172 are preferentially directed toward outlets 1114, 1116 based on pressure
differentials, and inertial separation forces are applied on the solids by making the
particles and carrier gas turn from the extended inlet 1110 within chamber 1112 to
flow toward outlets 1114, 1116, the inertial forces separating the different sizes /
densities of particles. Larger and/or heavier solid particles 1174 preferentially go
downward in sections 1118 to a standpipe or dipleg (not shown) connected to the
base of the separator, while lighter or smaller solid particles 1176 are preferentially
carried along with the gas stream to outlets 1114, 1116, where the e of small
particles and gases may be recovered.
17344/350WO1
In each of the separators described herein, by controlling the gas flow rates
entering the rd standpipe / separation chamber and exiting the main gas
stream , the overall separation efficiency of the separator and the selectivity to
separate heavier and/or larger particles from lighter or smaller particles can be
lated. This extends to a fully sealed dipleg where the only gas stream g
the dipleg are those entrained by the exiting solid/catalyst flow.
In some embodiments, a gas sparger or extra steam/inert gas may be ed
proximate a top of the heavy / dense particle outlet section, such as near a top of the
standpipe inlet. The additional lift gas provided wit hin the separator may further
facilitate the separation of heavier and/or larger solid particles from lighter or smaller
solid particles, as the extra gas may preferentially lift lighter solid particles to the gas
outlets, resulting in better solid fication.
The le separators described herein may be disposed external or internal
to a vessel. Further, in some embodiments, the large / dense particle outlets of the
particle separators may be fluidly ted to an external vessel, providing for
selective recycle or feed of the separated particles to the desired r, so as to
maintain a desired catalyst balance, for example.
Embodiments disclosed herein, by the methods described above,
significantly se the tration of desired catalysts in the second reactor
(vessel 32), consequently increasing light olefin yield. In addition, this process also
serves as a method to decouple the withdrawal and addition of the ZSM-5 and ZSM5-
11 with the withdrawal and addition of FCC catalyst. In y, the FCC process
presented in this disclosure creates a desired ZSM-5 or ZSM-11 catalyst additive rich
environment in the second reactor 32, which could preferentially convert light
naphtha products, such as those derived from first reactor, to improve light olefin
yield while simultaneously zing middle distillate yield by applying optimum
operation condition in the first reactor/riser reactor.
Another benefit of embodiments disclosed herein is that the ated tworeactor
scheme overcomes the heat balance limitations in the stand alone C4 / naphtha
catalytic ng processes. The second (mixed flow) reactor acts as a heat sink
due to integration with the catalyst regenerator, minimizing the requirement of
catalyst cooler while processing residue feed stocks.
17344/350WO1
The product vapors from the second reactor are transported into the first
reactor / disengaging vessel or reaction termination device wherein these vapors are
mixed and quenched with the products from the first stage and or external quench
media such as LCO or steam to minimize the unwanted thermal cracking reactions.
atively, the t outlet line of the second reactor / solids tor can also
be used to introduce additional quantity of heavy feed or re-route part of the feed
from the first reactor (the riser reactor). This serves two purposes: (1) the catalyst in
the solids separator vapor outlet line is inantly Y-type zeolite / conventional
FCC catalyst that is preferred to crack these heavy feed molecules into middle
distillates, and (2) such ng reaction is endothermic that helps in ng the
temperature of the outgoing product vapors and also residence time.
In some embodiments disclosed herein, an existing FCC unit may be
retrofitted with a second reactor as described above. For example, a properly sized
reactor may be fluidly connected to an existing catalyst ration vessel to
provide catalyst feed and return from the mixed flow vessel, and y connected
to an existing disengagement vessel to separate the hydrocarbon products and
catalysts. In other embodiments, a mixed flow reactor may be added to a grass-roots
FCC unit that is aimed at operating in gasoline mode, light olefins mode, or diesel
mode.
The reactor system described above with respect to Figures 1 and 6 related
primarily to light olefins production, and advantageous concentration of a catalyst in
a mixed st system to enhance reactivity and selectivity of the system. Such a
r system may also be used for other mixed catalyst systems, where
concentration of one of the catalysts may be advantageous.
For example, in some embodiments, the reaction system may be used for
gasoline desulfurization, where catalyst mixture may include a smaller and/or less
dense FCC catalyst, such as zeolite Y, and a larger and/or denser catalyst, such as a
gasoline desulfurization additive. Such a s is described with respect to Figure
Referring now to Figure 7, a simplified process flow diagram of systems for
cracking and desulfurizing hydrocarbons according to embodiments sed herein
is rated. The system includes a two-reactor configuration for producing olefins,
17344/350WO1
such as propylene and ethylene, from petroleum ocks or other hydrocarbon
streams. The first r 3 may be a riser reactor for cracking heavier hydrocarbon
feeds, for example. The second reactor 32 is a fluidized bed reactor, which may be
equipped with baffles or internals. The cracked hydrocarbon products, including
olefins and/or light naphtha products from the first reactor 3 or similar feed streams
from external sources, may be processed in the second reactor 32 to enhance the
quality of the product, such as decreasing the overall sulfur content of the
hydrocarbons processed in the second reactor.
A heavy eum e feed is injected through one or more feed injectors
2 located near the bottom of first r 3. The heavy petroleum feed contacts hot
regenerated catalyst introduced through a J-bend 1. The catalyst fed to the first
reactor 3 is a catalyst mixture, including a first catalyst selective for cracking heavier
hydrocarbons, such as a Y-type zeolite based catalyst, and a second catalyst selective
for the desulfurization of naphtha range hydrocarbons, which may also be used in
combination with other catalysts. The first and second catalysts may be ent in
one or both particle size and density.
The heat required for vaporization of the feed and/or raising the temperature
of the feed to the desired reactor temperature, such as in the range from 500°C to
about 700°C, and for the endothermic heat (heat of reaction) may be provided by the
hot regenerated catalyst coming from the regenerator 17.
After the major part of the cracking reaction is completed, the mixture of
products, erted feed vapors, and spent catalyst flow into a two stage cyclone
system housed in cyclone nment vessel 8. The two -stage cyclone system
includes a primary cyclone 4, for separating spent catalyst from . The spent
st is discharged into stripper 9 through primary cyclone dip leg 5. Fine catalyst
particles entrained with the separated vapors from primary cyclone 4 and t
vapors from second reactor 32, introduced via flow line 36a and a single stage
cyclone 36c, are separated in second stage cyclone 6. The catalyst mixture collected
is discharged into stripper 9 via dip leg 7. The vapors from second stage cyclone 6
are vented through a secondary cyclone outlet 12b, which may be connected to
plenum 11, and are then routed to a fractionator / gas plant 410 for recovery of
products, including the d olefins. If necessary, the t vapors are further
17344/350WO1
cooled by introducing light cycle oil (LCO) or steam via distributor line 12a as a
quench media.
] The fractionator 410 may be, for example, a main fractionator of an FCC
plant, and may produce various hydrocarbon fractions, including a light olefincontaining
fraction 412, a naphtha fraction 414, and a heavies fraction 416, among
other various arbon cuts. The products routed to fractionator / gas plant 410
may include other light gases, such as hydrogen sulfide that may be produced during
desulfurization; Separators, absorbers, or other unit operations may be included
where such impurities are desired to be separated am of the main fractionator
/ gas plant.
The spent st recovered via dip legs 5, 7 undergoes stripping in stripper
bed 9 to remove interstitial vapors (the hydrocarbon vapors trapped between catalyst
les) by countercurrent ting of steam, introduced to the bottom of er
9 h a steam distributor 10. The spent catalyst is then transferred to regenerator
17 via the spent catalyst standpipe 13a and lift line 15. Spent catalyst slide valve
l3b, located on spent catalyst standpipe 13a, is used for controlling catalyst flow from
er 9 to regenerator 17. A small portion of combustion air or nitrogen may be
uced through a butor 14 to help smooth transfer of spent catalyst.
Coked or spent catalyst is discharged through spent catalyst distributor 16 in
the center of the dense regenerator bed 24. Combustion air is uced by an air
distributor 18 located at the bottom of regenerator bed 24. Coke deposited on the
catalyst is then burned off in regenerator 17 via reaction with the combustion air.
The catalyst fines entrained along with flue gas are collected in first stage cyclone
19 and second stage cyclone 21 and are discharged into the rator catalyst bed
through respective dip legs 20, 22. The flue gas recovered from the outlet of second
stage cyclone 21 is directed to flue gas line 50 via regenerator plenum 23 for
downstream waste heat recovery and/or power recovery.
A first part of the regenerated catalyst mixture is withdrawn via regenerated
catalyst standpipe 27, which is in flow communication with J bend 1. The catalyst
flow from regenerator 17 to reactor 3 may be regulated by a slide valve 28 located
on regenerated catalyst standpipe 27. The opening of slide valve 28 is adjusted to
control the catalyst flow to maintain a desired top temperature in reactor 3.
350WO1
In addition to lift steam, a provision is also made to inject feed streams such
as C4 s and naphtha or similar external streams as a lift media to J bend 1
through a gas distributor 1a located at the ion for enabling smooth transfer of
regenerated catalyst from J bend 1 to reactor 3. J bend 1 may also act as a dense bed
reactor for cracking C4 olefins and naphtha streams into light olefins at ions
favorable for such reactions.
A second part of the regenerated catalyst mixture is withdrawn into a second
reactor 32 through a standpipe 30. A valve 31 may be used to control the catalyst
flow from rator 17 to second reactor 32 based on a vapor outlet temperature
set point. One or more arbon ons, such as naphtha streams, may be
ed into the bottom section of the catalyst bed through one or more feed
distributors 34 (34a, 34b), either in liquid or vapor phase. In some embodiments, the
naphtha feed may include a portion or all of the naphtha 414 from the fractionator
410. Second reactor 32 operates in a mixed flow fashion, where a portion of the
regenerated catalyst flows rd (from the top to the bottom of the reactor bed)
and/or circulates within vessel 32, and a portion of the regenerated catalyst mixture
and the feed hydrocarbon stream flows upward (from the bottom to the top of the
reactor bed, the r / less dense particles carrying out of the top of the reactor
with the effluent hydrocarbons).
Second reactor 32 may be equipped with s or structured internals (not
shown) that help intimate contact and mixing of catalyst and feed molecules. These
internals may also help in minimizing channeling, bubble growth, and/or
coalescence. Second reactor 32 may also be enlarged at different sections along the
length to maintain a constant or desired icial gas velocity within the sections.
After the reaction is completed, the catalyst is stripped at the bottommost
portion of second reactor 32 to separate entrained hydrocarbon feed / products using
steam as a stripping media introduced through distributor 35. The spent catalyst
red at the bottom of reactor 32 may then be withdrawn through catalyst
withdrawal line 418. Alternatively, the spent catalyst red at the bottom of
reactor 32 may be transferred to regenerator 17, as described above with respect to
Figure 1 (via standpipe 37 and lift line 40 through a spent catalyst distributor 41,
where combustion air or nitrogen may be introduced through distributor 39 to enable
17344/350WO1
smooth transfer of catalyst to regenerator 17). A valve (not illustrated) may be used
to control the catalyst flow from second reactor 32.
As noted above, second r 32 utilizes two different catalysts that may
differ in one or both of particle size and/or density, such as a less dense and/or smaller
Y-type zeolite or FCC catalyst and a larger and/or denser desulfurization catalyst.
The icial gas velocity in second r 32 is maintained such that essentially
all or a large portion of the lighter, smaller catalyst and a portion of the larger and/or
denser catalyst is carried out of the r with the hydrocarbon products and steam
recovered via effluent flow line 45. A portion of the larger and/or denser catalyst
may be retained within the reactor 32, forming a dense bed toward the lower n
of the reactor, as noted above.
The effluent from reactor 32 recovered via flow line 45 may thus include
desulfurized hydrocarbon products, unreacted hydrocarbon feedstock, steam
(stripping media), and a catalyst mixture, including essentially all of the lighter
and/or smaller catalyst and a portion of the heavier and/or larger catalyst introduced
to reactor 32. The effluent may then be transported via flow line 45 to a solids
separator 47. Separator 47 may be a separator configured to separate the two types
of catalyst based on their physical properties, namely particle size and/or y.
For example, separator 47 may use differences in inertial forces or centrifugal forces
to separate the smaller and/or lighter catalyst from the larger and/or heavier catalyst.
The solids separation vessel 47 is an external vessel to the second reactor 32 and is
ed at ynamic properties that enhance the separation of the two types of
catalyst based on their physical properties.
After separation in separator 47, the smaller and/or r catalyst (Y-type
zeolite / FCC catalyst) is then transported from separator 47 to the common
disengager or containment vessel 8, housing riser reactor cyclones and/or reaction
termination system, via outlet line 36a. The larger and/or r desulfurization
catalyst may be ed via flow line 49 to the mixed flow reactor 32 for continued
reaction with arbon feeds introduced through distributors 34a/b.
] Entrainment of essentially all of the lighter/smaller catalyst and a portion of
the heavier and/or larger catalyst, subsequent separations, and recycle of the heavier
and/or larger catalyst to reactor 32 may allow for a icant accumulation of the
350WO1
larger and/or heavier desulfurization catalyst in reactor 32. As this catalyst is more
selective for the desulfurization of naphtha range hydrocarbons, the accumulation of
the larger and/or heavier catalyst may provide a ivity and yield advantage.
Further, operation of the reactor in a fluidization flow regime to entrain both types
of catalyst may provide for improved operability of the reactor or flexibility in
operations, as discussed above.
A hydrocarbon feed such as heavy vacuum gas oil or heavy residue feed,
light cycle oil (LCO), or steam may be injected as a quench media in the outlet line
36a through a distributor 36b. The flow rate of such quench media may be controlled
by setting the temperature of the stream entering the containment vessel 8. All the
vapors from second reactor 32, including those fed through distributor 36b, are
discharged into the dilute phase of containment vessel 8 through a single stage
cyclone 36c. ing a hydrocarbon feed as a quench media is preferred as it
serves dual purpose of cooling the products from second reactor 32 and also
enhances the production of middle distillates.
The first reactor 3, such as a riser reactor, may operates in the fast fluidization
regime (e.g., at a gas icial velocity in the range from about 3 to about 10 m/s
at the bottom section) and pneumatic transport regime (e.g., at a gas superficial
velocity in the range from about 10 to about 20 m/s) in the top section.
WHSV in second reactor 32 is typically in the range from about 0.5 h-1 to
about 50 h-1; vapor and st residence times may vary from about 2 to about 20
seconds. As necessary, make-up catalyst may be introduced via one or more flow
lines 42, 43. For example, fresh or make-up FCC or Y-type zeolite catalyst or a
mixture of these two may be uced to regenerator 17 via flow line 42 and fresh
or make-up gasoline desulfurization additive may be introduced to second reactor 32
via flow line 43. Overall system catalyst inventory may be ined by
withdrawing mixed st from regenerator 24, for example, and / or reactor 32.
st inventory and accumulation of the preferred catalyst within reactor 32 may
be lled, such as described above. Additionally, in some embodiments, a
st hopper 26 may be used in conjunction with catalyst withdrawal line 25,
pressure zation line 29, and standpipe 27, as described above.
17344/350WO1
rly, the reactor system of Figure 7 may be used for advantageous
processing of heavy arbon feedstocks, including heavy crudes or virgin
crudes. In such an embodiment, the mixed catalyst system may include, for example,
a r and/or less dense FCC catalyst, such as zeolite-Y, and a larger and/or
denser heavy oil treatment additive. For example, the heavy oil treatment additive
may be one of an active matrix catalyst, a metals trapping additive, a coarse and/or
dense Ecat (equilibrium catalyst), a matrix or binder type catalyst (such as kaolin or
sand) or a high matrix / zeolite ratio FCC catalyst, among . The heavy oil
treatment additive may have minimal catalytic activity towards cracking of heavier
hydrocarbons and may simply supply the surface area necessary for thermal cracking
ons to take place. The heavy hydrocarbon feed may be introduced to reactor
32 via distributors 43 a/b, and the system may be ed as bed above to
enhance the processing of heavy hydrocarbon ocks.
WHSV in the second reactor 32 when operating under heavy hydrocarbon
treatment conditions is typically in the range from 0.1-100 hr-1; vapor and particle
residence times may vary from 1-400 seconds. As ary, makeup particles may
be introduced via one or more lines 42, 43; it may be advantageous to add the FCC
or Y-type catalyst to the regenerator 17 via line 42 and the heavy oil treatment
additive via line 43 to the second reactor 32. Overall system activity is maintained
by withdrawing particles via line 418 from the second reactor 32 and from the
regenerator 24. Solids inventory and the accumulation of the preferred heavy oil
ent additive in second reactor 32 may be controlled by additions through line
43 and withdrawals through line 418. Operating temperature in second reactor 32 is
controlled using catalyst from regenerator 17 line 30 via valve 31 and may range
from 400-700 °C. In some embodiments, the product of second reactor 32 may be
essentially the feed for the first reactor 3. Additionally, in some embodiments, a
catalyst hopper 26 may be used in ction with st withdrawal line 25,
pressure equalization line 29, and standpipe 27, as described above
In general, the process flow ms illustrated in Figures 1, 6, and 7 use the
catalyst / particle separation technology to process additional or recycle hydrocarbon
feedstocks in a secondary vessel. The st mixture circulating through the
system may include catalysts selective to particular reactions, such as cracking,
17344/350WO1
desulfurization, demetalization, denitrogenation, and other, where the catalysts of
the mixture are selected to have ing physical properties, as described above,
such that a desired catalyst may be concentrated in the second reactor. Regenerated
catalyst is fed to the second reactor/vessel which may operate in fast fluidized,
bubbling, or ent bed operation (depending on application). The effluent of the
second reactor/vessel goes to the separator 47, where the primary and secondary
catalysts are separated based on size and/or density and the tor bottoms, which
is enriched in the secondary catalyst, is recycled back to the second reactor/vessel.
The second reactor/vessel has optional catalyst withdrawals which may be
advantageous depending on application as well as different hydrocarbon feeds
depending on application. The concentration of the secondary catalyst may enhance
the operability, ility, and ivity of the overall reaction .
The tor 47 as described above with respect to Figure 2 may be used to
enhance productivity and flexibility of mixed catalyst hydrocarbon processing
systems, where the separator 47 may be located at other advantageous locations
within the system. Such processes and systems are described further below with
t to Figures 8-11, where like numerals ent like parts.
Referring now to Figure 8A, a simplified process flow diagram of s
for converting hydrocarbons and producing olefins according to embodiments
disclosed herein is illustrated, where like numerals represent like parts. The process
scheme of Figure 8A adds a catalyst holding vessel 510 which is fed regenerated
st from the FCC regenerator via catalyst withdrawal line 30 and valve 31. The
holding vessel 510 may be fluidized with a fluidization medium, such as air,
nitrogen, or steam, for example, introduced via flow line 516. The holding vessel
effluent 45 is sent to the separator 47 where the mixture of catalysts is separated.
The separator bottoms 49, which is enriched in the larger and/or heavier st, is
recycled back to catalyst g vessel 510, where the concentration of the larger
and/or denser catalyst will build up. The remaining stream 514 from the tor
510 is returned to the disengagement vessel 8 in this embodiment. The bottoms 512
of the holding vessel may be coupled to a slide valve (not illustrated) which can
control the feed of catalyst to second reactor / vessel 32, which can be operated in a
similar fashion to that described above with respect to Figures 1, 6, and 7.
17344/350WO1
Advantageously, the st concentrated in vessel 510 will not be saturated with
arbon and may allow for lower contact times with catalyst in the second
reactor/vessel 32.
Figure 8B illustrates a system similar to that of Figure 8A, except the st
recovered from separator 47 via flow line 514 is ed to the catalyst regenerator
17 as opposed to being forwarded to the disengagement vessel 8. The vessel to which
the catalyst in flow line 514 is forwarded may depend upon the type of fluidization
gas uced via flow line 516 as well as the capabilities of the systems receiving
flow from either regenerator 17 or vessel 8, via flow lines 50 and 12b, respectively.
Where the fluidization gas is steam, for example, the catalyst in flow line 514 is
preferably forwarded to vessel 8; where the fluidization gas is air or nitrogen, for
example, the catalyst in flow line 514 is preferably forwarded to regenerator 17.
Figures 8A and 8B rate the smaller les recovered via flow line 514
as being forwarded to the regenerator 17 or disengagement vessel 8, and the larger
and/or heavier particles recovered via flow line 512 as being ded to second
reactor 32. Embodiments herein also contemplate forwarding of the smaller and/or
r particles recovered via the separator 47 and flow line 514 to second reactor
32 while recirculating the larger and/or heavier particles to the regenerator 17 or
stripper 9.
Figures 8A and 8B further illustrate a system with a vessel 510 lating
/ concentrating large particles for use in the second reactor. Where a single-pass
separation may suffice, the containment vessel 510 may be excluded from the
system, as illustrated in Figures 9A and 9B, where like numerals represent like parts.
In these embodiments, the catalyst mixture is fed directly from the catalyst
regenerator 17 via dip leg 30 to separator 47. Air or other fluidization gases may be
supplied via flow line 610, provided at a flow rate sufficient for the inertial
separations. The smaller / lighter particles may be recovered via flow line 612 and
the larger and/or heavier particles may be red via flow line 614. Figure 9A
illustrates the larger and/or heavier particles being forwarded to second reactor 32,
whereas Figure 9B illustrates the smaller and/or lighter particles being forwarded to
second reactor 32.
17344/350WO1
Figures 9A and 9B illustrate return of a particle portion to the regenerator 17.
Similar to the above description with respect to Figures 8A and 8B, the particles not
fed to r 32 may be ed to either the regenerator 17 or the agement
vessel 8, and such may depend on the fluidization medium and/or downstream
processing capabilities.
The process schemes illustrated in Figures 9A and 9B use a single pass
version of the separator as opposed to those versions that incorporate recycle to
increase the concentration. In this scheme, the regenerated catalyst is directed to the
tor where either the bottoms or overhead of the separator can be directed to
the second reactor. If the bottoms were to be directed, the catalyst would be enriched
based on the larger and/or denser particles. If the overhead of the separator were to
be directed to the second reactor, the catalyst would be enriched in the smaller and/or
less dense particles. This scheme could also be arranged such that no second reactor
is present, and the separator is between the regenerator and the first reactor,
concentrating a catalyst similar to that described for the process of Figure 11, below.
The embodiments of Figures 8A/B decouple the recycle catalyst from the
second reactor, achieving a higher concentration of the desired catalyst in the second
reactor, however requiring additional capital costs. The embodiments of 6A/B also
decouple the recycle catalyst from the second reactor, achieving a moderate increase
in concentration of the desired catalyst as compared to the flow scheme of Figure 7,
for example, but at a lower capital cost than the embodiment of s 9A/B.
Referring now to Figure 10, a fied process flow diagram of systems for
processing hydrocarbons according to embodiments disclosed herein is rated,
where like ls represent like parts. This process schemes removes the second
reactor and has the separator 47 receiving an effluent from the first reactor 3. The
riser nt, which ns a mixed st, could be directed to the separator 47
where a portion of catalyst is recycled to the riser 3 from the separator bottoms 710,
y ing the concentration of the larger and/or heavier catalyst in the riser
r 3. The overhead 712 of the separator 47 would continue to the er vessel
8, where the hydrocarbon products would be separated from the remaining catalyst.
This configuration could also be used with a catalyst mixture with no degree of
classification as a method of recycling spent catalyst to the riser 3.
350WO1
] The enriched catalyst fraction 710 may be introduced to the riser 3 upstream
or downstream (as illustrated) of the regenerated catalyst feed inlet from standpipe
27, and in some embodiments may be introduced at one or more points along the
length of the riser reactor 3. The inlet point may be based on secondary arbon
feeds, temperature of the recirculating catalyst 710, and other variables that may be
used to ageously process arbons in the riser reactor 3.
The hydrocarbon products recovered from disengagement vessel 8 / stripper
9 may be forwarded, as described above, to a onator / gas plant 720, for
separation and ry of one or more hydrocarbon fractions 722, 724, 726, 728,
730. One or more of the recovered hydrocarbon fractions from the fractionator / gas
plant in embodiments herein may be recirculated to the riser reactor 3 or second
reactor 32 for further processing.
a simplified process flow diagram of systems for processing hydrocarbons
according to embodiments disclosed herein is illustrated, where like numerals
represent like parts. In this process scheme, a regenerator catalyst hopper 26 is
fluidly connected to riser reactor 3. Regenerated mixed catalyst, which contains a
smaller and/or less dense catalyst and a larger and/or denser catalyst, flows from the
regenerator 17 to the regen catalyst hopper 26. The hopper 26 is fluidized with steam
and/or air, provided by distributor 810. The overhead effluent 816 of the hopper
flows to the separator 47. In the separator 47, which is a separation device as
described previously, the catalysts are ted, and the bottoms 814, which is
enriched in the larger and/or denser st, may be fed back to the regen catalyst
hopper 26, such as when zed with air, or to disengagement vessel 8, such as
when fluidized with steam. This will increase the concentration of the larger and/or
denser catalyst in the regen catalyst hopper 26. The overhead 812 of the tor
47 may be directed to either the regenerator or the stripper vessel. The bottom 27 of
the regenerator catalyst hopper has a withdrawal with slide valve 28 which controls
the flow of catalyst which is enriched in the larger and/or denser catalyst to the riser
3. In this manner, the riser 3 operates with an effective higher concentration of
catalyst than the inventory in the system, ng preferential products based on the
properties of the catalyst.
17344/350WO1
Concentration of a catalyst in the regen catalyst hopper as described above
with respect to Figure 11 may be performed intermittently. The system may circulate
the catalyst mixture through the riser, stripper, and regenerator, without sufficient
fluidization in the hopper 26 to n catalysts to the separator 47. When there is
a change in the desired product mixture, the hydrocarbon feeds, or other factors,
where it may be advantageous to operate with a higher tration of a particular
catalyst in the catalyst mixture, the catalyst in the regen hopper 26 may be fluidized
and separated using tor 47. When factors again , fluidization of the
catalyst hopper may be discontinued. In this manner, the flexibility of the system
with regard to products and feed may be enhanced.
While Figures 10 and 11 are illustrated with a single riser, the solids
separation device may be used to enhance the performance of a le riser system.
For example, a two-riser system may benefit from the concentration of one catalyst
in a riser, which may be sing different feeds than a second riser.
Embodiments herein may utilize various types of catalysts or particles to
m desired reactions, where a common regenerator may be used to regenerate
the mixture of catalysts, and a separator is advantageously located to enrich one or
more rs with a particular catalyst contained in the mixture of catalysts.
Embodiments herein may be used to improve unit operations, and enhance the
selectivity and flexibility of the reaction systems, such as for applications including
light olefins production, gasoline desulfurization, and heavy oil processing.
Light olefins production may include s light, medium, and heavy
arbon feeds to the riser, as described above. Feeds to the second reactor 32
may include naphtha, such as straight run naphtha or recycle cat naphtha, among
other feeds. The catalyst mixture for light s production may include a smaller
and/or less dense catalyst, such as an FCC catalyst (zeolite Y, for e), and a
heavier / denser catalyst, such as ZSM-5 or ZSM-11, among other combinations.
Other cracking catalysts may also be used Various sts for the cracking of
hydrocarbons are disclosed in U.S. Patent Nos. 257, 7,314,963, 7,268,265,
7,087,155, 6,358,486, 6,930,219, 6,809,055, 5,972,205, 5,702,589, 5,637,207,
,534,135, and 5,314,610, among others.
17344/350WO1
] Embodiments directed toward gasoline desulfurization may include various
light, medium, and heavy hydrocarbon feeds to the riser, as described above. Feeds
to the second reactor 32 may also include naphtha, such as straight run naphtha or
recycle cat naphtha, among other feeds. The catalyst mixture for light olefins
production may include a smaller and/or less dense catalyst, such as an FCC catalyst
(zeolite Y, for example), and a larger and/or denser catalyst, with desulfurization
functionality such as a MgO / Al 2O3 with various metals promotion. Other
desulfurization catalysts may also be used as sed in US Patent Nos. 5,482,617,
6,482,315, 6,852,214, 7,347,929 among . In some embodiments, the catalyst
mixture may include a ng catalyst composition having desulfurization ty,
such as those disclosed in US5376608, among others.
Embodiments directed toward heavy oil processing may include various
light, , and heavy hydrocarbon feeds to the riser, as described above. Feeds
to the second reactor 32 may include hydrocarbons or hydrocarbon mixtures having
g points or a g range above about 340°C. H ydrocarbon feedstocks that
may be used with processes disclosed herein may include various refinery and other
hydrocarbon streams such as petroleum atmospheric or vacuum residua, deasphalted
oils, deasphalter pitch, hydrocracked atmospheric tower or vacuum tower bottoms,
straight run vacuum gas oils, hydrocracked vacuum gas oils, fluid catalytically
cracked (FCC) slurry oils, vacuum gas oils from an ebullated bed racking
process, shale-derived oils, coal-derived oils, tar sands bitumen, tall oils, bio-derived
crude oils, black oils, as well as other similar hydrocarbon streams, or a combination
of these, each of which may be straight run, process derived, racked, partially
urized, and/or partially demetallized streams. In some embodiments, residuum
hydrocarbon fractions may include hydrocarbons having a normal boiling point of at
least 480°C, at least 524°C, or at least 565°C. The catalyst mixture for heavy
hydrocarbon processing may include a smaller and/or less dense catalyst, such as an
FCC catalyst (zeolite Y, for e), and a larger and/or denser catalyst, such as an
active matrix catalyst, a metals trapping catalyst, a coarse / dense Ecat (equilibrium
catalyst), a matrix or binder type catalyst (such as kaolin or sand) or a high matrix /
zeolite FCC catalyst. Other cracking catalysts may also be used, such as, for
example, one or more of those disclosed in US5160601, US5071806, US5001097,
17344/350WO1
US4624773, US4536281, US4431749, US6656347, 757, 132, and
US7591939, among others.
Systems herein may also be utilized for pre-treatment of a heavy crude or
virgin crude, such as a crude oil or bitumen recovered from tar sands. For example,
reactor 32, such as that in Figures 1 or 9, among others, may be used to pre-treat the
bitumen, prior to further processing of the treated heavy crude in downstream
operations, which may include separation in a downstream separation system and
recycle of one or more fractions for further conversion in reactor 3. The ability to
eat the heavy crude with a preferred particle within a particle or catalyst
e may advantageously allow integration of heavy crude processing where it
otherwise would be detrimental to catalyst and overall system performance.
Embodiments herein describe the catalyst mixture being separated by the
separator and the effective preferential concentration of a catalyst within the mixture
in a reactor. As illustrated in the Figures, the catalyst being concentrated in the
reactor is illustrated as being returned from the separator proximate the top of the
reactor or vessel. Embodiments herein also contemplate return of the catalyst from
the separator to a middle or lower portion of the reactor, and where the catalyst is
returned may depend on the hydrocarbon feeds being processed, the st types
in the mixture, and the desired st gradient within the reactor vessel.
Embodiments herein also plate return of the catalyst to multiple ons
within the reactor. While providing the ability to enhance the concentration of a particular catalyst or particle within
a mixture in a given reactor, ments herein may also be used for a one st system; the particle separators and systems
described herein may increase the catalyst/oil ratio, which enhances catalytic t time
As described for embodiments above, a second reactor is integrated with a
FCC riser reactor and separation . This reactor is in flow communication with
other vessels, allowing selective catalytic processing and ated arbon
product quenching, separation and catalyst regeneration. Such an integrated reactor
system offers one or more of the above advantages and features of embodiments of
the processes disclosed herein may provide for an improved or optimal process for
the catalytic cracking of hydrocarbons for light olefin production.
Embodiments herein may employ two types of catalyst particles, such as Y-
zeolite/ FCC st of smaller particle size and/or less density and ZSM-5 particles
larger in size and/or denser than the former. A separator with selective recycle may
17344/350WO1
be utilized to preferentially segregate the ite from ZSM-5 catalyst. Use of such
catalyst system allows entrainment of lighter and r particles, thereby retaining
ZSM-5 type les within the additional new reactor bed. The reactants undergo
selective tic cracking in presence of ZSM-5 type catalyst that is preferred to
maximize the yield of light olefins from C4 and naphtha feed streams. The separator
is a device which can facilitate the separation of two types of catalysts due to the
difference in their particle size and/or density. Examples of separators with selective
recycle may be a cyclone separator, a screen tor, mechanical sifters, a gravity
chamber, a centrifugal separator, an in-line or pneumatic classifier, or other types of
separators useful for efficiently separating particles based on size and/or
hydrodynamic properties. The separator is connected to the top of the second reactor
which is in flow communication with second reactor as well as regenerator and first
reactor/ stripper.
The reactor may be provided with baffles or modular grid internals. This
provides te contact of catalyst with hydrocarbon feed les, helps in
bubble breakage and avoiding bubble growth due to coalescence, channeling or
bypassing of either st or feed.
Conventionally, fresh catalyst make-up for maintaining the catalyst activity
is introduced to the rator bed using plant air. In st, it is proposed to inject
the desired high concentration catalyst/additive directly into the second reactor bed
using steam or en as conveying media. This helps to produce incremental
ses in concentration and favorable selectivity.
The reactor configurations described herein provide enough flexibility and
operating window to adjust operating conditions such as weight hourly space
velocity (WHSV), catalyst and hydrocarbon vapor residence time, reaction
temperature, catalyst/oil ratio, etc. As for example, in some embodiments, the second
reactor top/ bed temperature is controlled by adjusting catalyst flow from rator
which indirectly controls the catalyst/oil ratio. Whereas reactor bed level may be
controlled by manipulating the spent catalyst flow from reactor to regenerator, which
controls the WHSV and catalyst residence time.
] While the disclosure includes a limited number of embodiments, those skilled
in the art, having benefit of this disclosure, will appreciate that other embodiments
17344/350WO1
may be devised which do not depart from the scope of the present disclosure.
Accordingly, the scope should be limited only by the attached .
The term ‘comprise’ and variants of the term such as ‘comprises’ or
‘comprising’ are used herein to denote the ion of a stated integer or stated
integers but not to exclude any other integer or any other integers, unless in the
context or usage an exclusive interpretation of the term is required.
Any reference to publications cited in this ication is not an admission
that the disclosures constitute common general knowledge in New Zealand.
17344/350WO1
Claims (1)
1. A system for ng hydrocarbons, sing: a mixed flow reactor for contacting a mixture comprising first particles and second particles with a hydrocarbon feedstock to convert at least a portion of the hydrocarbon feedstock to lighter hydrocarbons, wherein the first particles have a smaller average particle size and/or are less dense than the second particles; an overhead product line for orting from the mixed flow reactor a first stream comprising first particles, a first portion of the second particles, and arbons; a bottoms product line for transporting from the mixed flow reactor a second stream comprising a second portion of the second particles; a particle separator, external to the mixed flow reactor, configured to separate the first particles and the second particles based on particle size and/or density, for separating second particles from the first stream, producing a hydrocarbon effluent comprising hydrocarbons and the first particles; a feed line for returning separated second particles from the particle separator to the mixed flow reactor; a riser reactor for contacting a e of the first les and the second particles with a second hydrocarbon ock to convert at least a portion of the second hydrocarbon feedstock to r hydrocarbons and recover a riser reactor effluent comprising the lighter hydrocarbons and the mixture of the first and second particles; a separation system for separating the second particles from the hydrocarbon effluent and for separating the mixture of first and second particles from the riser reactor nt; and a regenerator for regenerating first and second particles recovered in the separation system and the second portion of the second les recovered in the bottoms product line.
Applications Claiming Priority (3)
Application Number | Priority Date | Filing Date | Title |
---|---|---|---|
US201662395707P | 2016-09-16 | 2016-09-16 | |
US62/395,707 | 2016-09-16 | ||
PCT/US2017/051537 WO2018053110A1 (en) | 2016-09-16 | 2017-09-14 | Fluid catalytic cracking process and apparatus for maximizing light olefin yield and other applications |
Publications (2)
Publication Number | Publication Date |
---|---|
NZ751807A NZ751807A (en) | 2020-09-25 |
NZ751807B2 true NZ751807B2 (en) | 2021-01-06 |
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