U.S. patent application number 13/712540 was filed with the patent office on 2013-06-27 for process for increased production of fcc gasoline.
This patent application is currently assigned to EXXONMOBIL RESEARCH AND ENGINEERING COMPANY. The applicant listed for this patent is ExxonMobil Research and Engineering Company. Invention is credited to Michel Daage, Ajit Bhaskar Dandekar, David Lee Johnson, Jiunn-Shyan Liou, Steven S. Lowenthal, Stephen John McCarthy, Rohit Vijay.
Application Number | 20130165717 13/712540 |
Document ID | / |
Family ID | 48655229 |
Filed Date | 2013-06-27 |
United States Patent
Application |
20130165717 |
Kind Code |
A1 |
McCarthy; Stephen John ; et
al. |
June 27, 2013 |
PROCESS FOR INCREASED PRODUCTION OF FCC GASOLINE
Abstract
This invention relates to methods and processes for increasing
the production of FCC (Fluid Catalytic Cracking) gasoline products,
and optionally distillate products, from refinery feedstocks. In
particular, the processes include hydrotreating and further
hydroisomerizing a typical FCC range feedstream prior to
catalytically cracking the feedstream in the FCC unit. The methods
herein result in higher FCC naphtha yields and lower FCC cat
bottoms yields thereby significantly increasing the overall FCC
gasoline production for a given operating unit and increasing the
profit margin of such FCC unit operations.
Inventors: |
McCarthy; Stephen John;
(Center Valley, PA) ; Lowenthal; Steven S.;
(Flanders, NJ) ; Johnson; David Lee; (Doylestown,
PA) ; Daage; Michel; (Hellertown, PA) ; Liou;
Jiunn-Shyan; (Bridgewater, NJ) ; Vijay; Rohit;
(Annandale, NJ) ; Dandekar; Ajit Bhaskar; (Vienna,
VA) |
|
Applicant: |
Name |
City |
State |
Country |
Type |
ExxonMobil Research and Engineering Company; |
Annandale |
NJ |
US |
|
|
Assignee: |
EXXONMOBIL RESEARCH AND ENGINEERING
COMPANY
Annandale
NJ
|
Family ID: |
48655229 |
Appl. No.: |
13/712540 |
Filed: |
December 12, 2012 |
Related U.S. Patent Documents
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Application
Number |
Filing Date |
Patent Number |
|
|
61579812 |
Dec 23, 2011 |
|
|
|
Current U.S.
Class: |
585/310 |
Current CPC
Class: |
C10G 45/58 20130101;
C10G 69/04 20130101; C10G 11/18 20130101; C10L 1/06 20130101 |
Class at
Publication: |
585/310 |
International
Class: |
C10G 69/04 20060101
C10G069/04 |
Claims
1. A process for increasing Fluid Catalytic Cracking ("FCC")
gasoline production comprising: a) contacting a
hydrocarbon-containing hydroisomerization feedstream with a
hydroisomerization catalyst under hydroisomerization conditions to
produce at least one hydroisomerized liquid product stream that has
a higher iso-paraffin content than the hydroisomerization
feedstream; b) contacting in the reaction zone of an FCC reactor
riser an FCC feedstream comprising at least a portion of the
hydroisomerized liquid product stream of step a) with a fluid
catalytic cracking catalyst thereby catalytically cracking the FCC
feedstream into an FCC product that has an average lower boiling
point than the FCC feedstream, and producing a spent catalyst; c)
the FCC product from the spent catalyst; d) cooling the FCC
product; and e) fractionating the FCC product into multiple FCC
product streams, wherein at least one of the FCC product streams is
a naphtha boiling-range product stream; and f) utilizing at least a
portion of the naphtha boiling-range product stream for gasoline
production.
2. The process of claim 1, wherein at least 50 wt % of the normal
paraffins in the hydroisomerization feedstream are converted to
iso-paraffins in the hydroisomerized liquid product stream in step
a).
3. The process of claim 2, wherein the hydroisomerization catalyst
comprises at least one Group VIIIA metal, and further comprises a
zeolite selected from EU-1, ZSM-35, ZSM-11, ZSM-57, NU-87, SAPO-11,
ZSM-22, and ZSM-48.
4. The process of claim 3, wherein the Group VIIIA metal of the
hydroisomerization catalyst is selected from Pt and Pd.
5. The process of claim 3, wherein the hydroisomerization catalyst
further comprises at least one Group VIA metal, wherein the Group
VIA of the hydroisomerization catalyst is selected from Mo and W,
and the Group VIIIA metal of the hydroisomerization catalyst is
selected from Ni and Co.
6. The process of claim 5, wherein the Group VIA of the
hydroisomerization catalyst is W, the Group VIIIA metal of the
hydroisomerization catalyst is Ni, and the zeolite in the
hydroisomerization catalyst is ZSM-48.
7. The process of claim 5, wherein tire hydroisomerization
feedstream contains over 300 ppmw of sulfur.
8. The process of claim 3, wherein the hydroisomerization
conditions include a temperature of from 400 to 850.degree. F. (204
to 454.degree. C.), a hydrogen partial pressure of from 1.8 to 34.6
mPa (250 to 5000 psi), a liquid hourly space velocity of from 0.2
to 10 v/v/hr, and a hydrogen circulation rate of from 35.6 to 1781
m.sup.3/m.sup.3 (200 to 10,000 scf/B).
9. The process of claim 8, wherein the conditions in the reaction
zone of the FCC reactor include a temperature from about 900 to
about 1060.degree. F. (482 to 57.degree. C.), a hydrocarbon partial
pressure from about 10 to 50 psia (70-345 kPa), and a catalyst to
feed (wt/wt) ratio from about 3 to 8, where the catalyst weight is
total weight of the fluid catalytic cracking catalyst.
10. The process of claim 9, wherein the fluid catalytic cracking
catalyst comprises at least one large-pore size faujasite zeolite
and at least one medium-pore size zeolite selected from ZSM-5,
ZSM-12, ZSM-22, ZSM-23, ZSM-34, ZSM-35, ZSM-38, ZSM-48, ZSM-50,
silicalite, and silicalite 2.
11. The process of claim 10, wherein feed residence time in the
reaction zone of the FCC reactor riser is less than about 5
seconds.
12. The process of claim 1, wherein the at least one
hydroisomerized liquid product stream of step a) is sent to a
distillation column of to produce the at least a portion of the
hydroisomerized liquid product stream of step b), as well as
producing a distillation column overhead vapor stream and at least
a first distillate product stream from the distillation column,
wherein the distillation column overhead vapor stream and the first
distillate product stream are not sent to the reaction zone of the
FCC reactor riser.
13. The process of claim 12, wherein at least one of the FCC
product streams is an FCC distillate boiling-range product stream
and at least a portion of the first distillate product stream is
combined with at least a portion of the FCC distillate
boiling-range product stream to form a combined distillate product
stream.
14. The process of claim 13, wherein at least a portion of the
combined distillate product stream is utilized for diesel product
blending.
15. The process of claim 1, further comprising: contacting a
hydrocarbon-containing hydrotreater feedstream containing at least
250 ppmw of sulfur with a hydrotreating catalyst under
hydrotreating conditions to produce at least one hydrotreated
liquid product stream and one hydrotreated vapor stream, wherein
the hydrotreated liquid product stream has a lower sulfur content
than the sulfur-containing hydrocarbon feedstream; separating the
hydrotreated liquid product stream from the hydrotreated vapor
stream; and utilizing at least a portion of the hydrotreated liquid
product stream as the hydroisomerization feedstream in step a).
16. The process of claim 15, wherein the hydrotreated liquid
product stream contains less than 30 ppmw of sulfur.
17. The process of claim 15, wherein the hydrotreating catalyst
comprises at least one Group VIA metal and at least one Group VIIIA
metal on a refractory oxide support, wherein the refractory oxide
support comprises silica, alumina, or silica-alumina; and the
hydroisomerization catalyst is comprised of at least one Group
VIIIA metal, and a zeolite selected from EU-1, ZSM-35, ZSM-11,
ZSM-57, NU-87, SAPO-11, ZSM-22, and ZSM-48.
18. The process of claim 17, wherein the hydrotreating catalyst has
a has an average pore size of from about 100 .ANG. to about 1000
.ANG., and a surface area of from about 100 to 350 m.sup.2/g.
19. The process of claim 15, wherein the hydrotreating conditions
include a temperature in the range 450.degree. F. to 750.degree. F.
(232.degree. C. to 399.degree. C.), pressure in the range of 1480
to 20786 kPa (200 to 3000 psig), a space velocity of from 0.1 to 10
LHSV, and a hydrogen treat gas rate of from 18 to 890
m.sup.3/m.sup.3 (100 to 5000 scf/B).
20. The process of claim 17, wherein the Group VIIIA metal of the
hydroisomerization catalyst is selected from Pt and Pd.
21. The process of claim 17, wherein the hydroisomerization
catalyst further comprises at least one Group VIA metal, wherein
the Group VIA of the hydroisomerization catalyst is selected from
Mo and W, and the Group VIIIA metal of the hydroisomerization
catalyst is selected from Ni and Co.
22. The process of claim 17, wherein the at least one
hydroisomerized liquid product stream of step a) is sent to a
distillation column of to produce the at least a portion of the
hydroisomerized liquid product stream of step b), as well as
producing an distillation column overhead vapor stream and at least
a first distillate product stream from the distillation column,
wherein the distillation column overhead vapor stream and the first
distillate product stream are not sent to the reaction zone of the
FCC reactor riser.
23. The process of claim 22, wherein at least one of tire FCC
product streams is an FCC distillate boiling-range product stream
and at least a portion of the first distillate product stream is
combined with at least a portion of the FCC distillate
boiling-range product stream to form a combined distillate product
stream.
24. The process of claim 17, wherein the hydrotreater feedstream
has a T5 boiling point of at least 400.degree. F. and a T95 boiling
point of less than about 1150.degree. F.
25. The process of claim 24, wherein the hydrotreater feedstream is
comprised of at least 75 wt % of a hydrocarbon feedstream derived
from a fossil-based oil material, and is further comprised of from
5 to 25 wt % of oil derived from renewable biofuel sources.
26. The process of claim 1, further comprising: contacting a
hydrocarbon-containing hydrotreater feedstream containing at least
250 ppmw of sulfur with a hydrotreating catalyst under
hydrotreating conditions to produce the hydroisomerization
feedstream.
27. The process of claim 26, wherein the hydrotreating catalyst and
the hydroisomerization catalyst are in a single reactor.
28. The process of claim 26, wherein the hydrotreating catalyst
comprises at least one Group VIA metal and at least one Group VIIIA
metal on a refractory oxide support, wherein the refractory oxide
support comprises silica, alumina, or silica-alumina; and the
hydroisomerization catalyst is comprised of at least one Group
VIIIA metal, and a zeolite selected from EU-1, ZSM-35, ZSM-11,
ZSM-57, NU-87, SAPO-11, ZSM-22, and ZSM-48.
29. The process of claim 28, wherein the hydroisomerization
catalyst further comprises at least one Group VIA metal, wherein
the Group VIA of the hydroisomerization catalyst is selected from
Mo and W, and the Group VIIIA metal of the hydroisomerization
catalyst is selected from Ni and Co.
30. The process of claim 29, wherein the zeolite in the
hydroisomerization catalyst is ZSM-48.
31. The process of claim 29, wherein the at least one
hydroisomerized liquid product stream of step a) is sent to a
distillation column of to produce the at least a portion of the
hydroisomerized liquid product stream of step b), as well as
producing an distillation column overhead vapor stream and at least
a first distillate product stream from the distillation column,
wherein the distillation column overhead vapor stream and the first
distillate product stream are not sent to the reaction zone of the
FCC reactor riser.
32. The process of claim 31, wherein at least one of the FCC
product streams is an FCC distillate boiling-range product stream
and at least a portion of the first distillate product stream is
combined with at least a portion of the FCC distillate
boiling-range product stream to form a combined distillate product
stream.
33. The process of claim 28, wherein the hydrotreater feedstream
has a T5 boiling point of at least 400.degree. F. and a T95 boiling
point of less than about 1150.degree. F.
34. The process of claim 33, wherein the hydrotreater feedstream is
comprised of at least 75 wt % of a hydrocarbon feedstream derived
from a fossil-based oil material, and is further comprised of from
5 to 25 wt % of oil derived from renewable biofuel sources.
35. A process for increasing Fluid Catalytic Cracking ("FCC")
gasoline production comprising: a) contacting a
hydrocarbon-containing hydroisomerization feedstream with a
hydroisomerization catalyst under hydroisomerization conditions to
produce at least one hydroisomerized product stream that has a
higher iso-paraffin content than the hydroisomerization feedstream;
b) contacting at least a portion of the hydroisomerized product
stream with a hydrotreating catalyst under hydrotreating conditions
to produce at least one hydrotreated liquid product stream and one
hydrotreated vapor stream, wherein the hydrotreated liquid product
stream has a lower sulfur content than the sulfur-containing
hydrocarbon feedstream; c) separating the hydrotreated liquid
product stream from the hydrotreated vapor stream; d) contacting in
the reaction zone of an FCC reactor riser an FCC feedstream
comprising at least a portion of the hydrotreated liquid product
stream of step c) with a fluid catalytic cracking catalyst thereby
catalytically cracking the FCC feedstream into an FCC product that
has an average lower boiling point than the FCC feedstream, and
producing a spent catalyst; e) separating the FCC product from the
spent catalyst; f) cooling the FCC product; and e) fractionating
the FCC product into multiple FCC product streams, wherein at least
one of the FCC product streams is a naphtha boiling-range product
stream; and g) utilizing at least a portion of the naphtha
boiling-range product stream for gasoline production.
Description
CROSS-REFERENCE TO RELATED APPLICATION
[0001] This application claims priority to U.S. Provisional
Application Ser. No. 61/579,812 filed Dec. 23, 2011, herein
incorporated by reference in its entirety.
FIELD OF THE INVENTION
[0002] This invention relates to methods and processes for
increasing the production of FCC (Fluid Catalytic Cracking)
gasoline products, and optionally distillate products, from
refinery feedstocks.
BACKGROUND OF THE INVENTION
[0003] An important process to the overall gasoline production in
the world is the refining Fluid Catalytic Cracking ("FCC") related
processes. FCCs utilize very small particulate catalysts which are
raised to very high temperatures and subsequently fluidized. These
fluidized particles contact high molecular weight petroleum feeds
and catalytically "crack" these larger hydrocarbon molecules to
lower boiling products which are more valuable products. Most FCC
processes contact heavy feed oils (such as vacuum gas oils,
atmospheric gas oils, and often petroleum resids) with the
fluidized catalysts typically with the goal to maximize naphtha
production volumes.
[0004] In the FCC process these low-value, high boiling point
hydrocarbon feedstocks are catalytically converted into more
valuable products by contacting the feeds with fluidized catalyst
particles in the process. In modern "short contact time" fluidized
catalytic cracking (FCC) units, the hydrocarbon feedstocks are
typically contacted with the fluidized catalyst particles in the
riser section of the FCC reactor. The contacting between feed and
catalyst is controlled according to the type of product desired. In
catalytic cracking of the feed, reactor conditions such as
temperature and contact time are controlled to maximize the
products desired, such as naphthas, and minimize the formation of
less desirable products such as light gases and coke.
[0005] The FCC naphthas derived from such processes are very
valuable products as they are used as a component in final gasoline
production. FCC naphthas can often account for about 50% or more of
the overall "gasoline blending feedstock" in a refinery.
Additionally FCC naphthas typically have a relatively high octane
value as compared to "straight run" naphthas that are typically
produced by a refinery's crude unit. This high octane value of the
FCC naphthas is in large part due to the high olefin content of the
FCC naphthas. As is such, maximizing the total of production of FCC
naphthas suitable for gasoline blending is of significant
importance to the commercial operations and economics of any
petroleum refinery.
[0006] While FCC units may target the maximization of other
hydrocarbon products, such as distillates used in diesel
production, or much smaller quantities of petrochemical production,
such as propylene, most FCC units, particularly in the United
States, target to maximize the overall naphtha production. The
overall naphtha production (i.e., light cat naphtha "LCN" and heavy
cat naphtha "HCN") can typically exceed over 50 vol % of the
overall products obtained from an FCC unit. As such, even small
percentages of increases in the naphtha (or gasoline) production
from an FCC unit result in significant cost improvements for the
average refinery. For example, with a 1% increase in naphtha
production, an FCC unit operating at 100,000 barrels per day
(bbl/day) would see an approximate increase in gasoline production
of about 500 bbl/day or about 20,000 gallons per day of increase in
gasoline production. As can be seen, even small improvements in
naphtha/gasoline yields from the FCC unit result in significant
financial benefit.
[0007] As such, what is needed in the industry are new methods and
processes for improving the naphtha yields from a refinery FCC unit
that do not require significantly changing the overall FCC
process.
SUMMARY OF THE INVENTION
[0008] The processes of present invention are designed to increase
the overall naphtha production from a Fluid Catalytic Cracking
("FCC") unit for maximizing gasoline production. The processes
herein may also be utilized to optionally increase distillate
production, preferably for increased diesel and/or jet fuel
production. The processes herein are aimed at modifying the
properties is of a typical FCC feedstream in order to increase the
naphtha, or optionally distillate, production obtained from the FCC
process.
[0009] In the processes herein, the iso-paraffin content of the
feed to the FCC unit is increased resulting in higher naphtha
(i.e., gasoline) production from the FCC unit. It has also been
found that distillate (i.e., diesel) production can also be
increased with a corresponding decrease in heavier cat bottoms
products. Preferably, at least 50 wt %, more preferably at least 75
wt %, even more preferably at least 90 wt %, and most preferably
substantially all of the normal paraffins in the feed to the
hydroisomerization unit are converted to isoparaffins within the
process. At least a portion of the hydroisomerized product is sent
to an FCC unit for further processing into naphtha and distillate
products.
[0010] A first non-limiting embodiment of the invention relates to
a process for increasing Fluid Catalytic Cracking ("FCC") gasoline
production comprising:
[0011] a) contacting a hydrocarbon-containing hydroisomerization
feedstream with a hydroisomerization catalyst under
hydroisomerization conditions to produce at least one
hydroisomerized liquid product stream that has a higher
iso-paraffin content than the hydroisomerization feedstream;
[0012] b) contacting in the reaction zone of an FCC reactor riser
an FCC feedstream comprising at least a portion of the
hydroisomerized liquid product stream of step a) with a fluid
catalytic cracking catalyst thereby catalytically cracking the FCC
feedstream into an FCC product that has an average lower boiling
point than the FCC feedstream, and producing a spent catalyst;
[0013] c) separating the FCC product from the spent catalyst;
[0014] d) cooling the FCC product; and
[0015] e) fractionating the FCC product into multiple FCC product
streams, wherein at least one of the FCC product streams is a
naphtha boiling-range product stream; and
[0016] f) utilizing at least a portion of the naphtha boiling-range
product stream for gasoline production.
[0017] A second non-limiting embodiment of the invention relates to
the first embodiment, further comprising:
[0018] contacting a hydrocarbon-containing hydrotreater feedstream
containing at least 250 ppmw of sulfur with a hydrotreating
catalyst under hydrotreating conditions to produce at least one
hydrotreated liquid product stream and one hydrotreated vapor
stream, wherein the hydrotreated liquid product stream has a lower
sulfur content than the sulfur-containing hydrocarbon
feedstream;
[0019] separating the hydrotreated liquid product stream from the
hydrotreated vapor stream; and
[0020] utilizing at least a portion of the hydrotreated liquid
product stream as the hydroisomerization feedstream in step a),
[0021] A third non-limiting embodiment of the invention relates to
the first embodiment, further comprising:
[0022] contacting a hydrocarbon-containing hydrotreater feedstream
containing at least 250 ppmw of sulfur with a hydrotreating
catalyst under hydrotreating conditions to produce the
hydroisomerization feedstream.
[0023] A fourth non-limiting embodiment of the invention relates to
a process for increasing Fluid Catalytic Cracking ("FCC") gasoline
production comprising:
[0024] a) contacting a hydrocarbon-containing hydroisomerization
feedstream with a hydroisomerization catalyst under
hydroisomerization conditions to produce at least one
hydroisomerized product stream that has a higher iso-paraffin
content than the hydroisomerization feedstream;
[0025] b) contacting at least a portion of the hydroisomerized
product stream is with a hydrotreating catalyst under hydrotreating
conditions to produce at least one hydrotreated liquid product
stream and one hydrotreated vapor stream, wherein the hydrotreated
liquid product stream has a lower sulfur content than the
sulfur-containing hydrocarbon feedstream;
[0026] c) separating the hydrotreated liquid product stream from
the hydrotreated vapor stream;
[0027] d) contacting in the reaction zone of an FCC reactor riser
an FCC feedstream comprising at least a portion of the hydrotreated
liquid product stream of step c) with a fluid catalytic cracking
catalyst thereby catalytically cracking the FCC feedstream into an
FCC product that has an average lower boiling point than the FCC
feedstream, and producing a spent catalyst;
[0028] e) separating the FCC product from the spent catalyst;
[0029] f) cooling the FCC product; and
[0030] e) fractionating the FCC product into multiple FCC product
streams, wherein at least one of the FCC product streams is a
naphtha boiling-range product stream; and
[0031] g) utilizing at least a portion of the naphtha boiling-range
product stream for gasoline production.
[0032] In other more preferred embodiments of the first through
fourth embodiments, at least 50 wt % of the normal paraffins in the
hydroisomerization feedstream are converted to iso-paraffins in the
hydroisomerized liquid product stream in step a). In other more
preferred embodiments, the hydroisomerization catalyst comprises at
least one Group VIIIA metal, and further comprises a zeolite
selected from EU-1, ZSM-35, ZSM-11, ZSM-57, NU-87, SAPO-11, ZSM-22,
and ZSM-48. In other preferred embodiments, the Group VIIIA metal
of the hydroisomerization catalyst is selected from Pt and Pd. In
yet other more preferred embodiments, the hydroisomerization
catalyst further comprises at least one Group VIA metal, wherein
the Group VIA of the hydroisomerization catalyst is selected from
Mo and W, and the Group VIIIA metal of the hydroisomerization
catalyst is selected from Ni and Co; even more preferably, the
Group VIA of the hydroisomerization catalyst is W, the Group VIIIA
metal of the hydroisomerization catalyst is Ni, and the zeolite in
the hydroisomerization catalyst is ZSM-48.
[0033] In other more preferred embodiments of the first through
fourth embodiments, the hydroisomerization feedstream contains over
300 ppmw of sulfur.
[0034] In other more preferred embodiments of the first through
third embodiments, the at least one hydroisomerized liquid product
stream of step a) is sent to a distillation column of to produce
the at least a portion of the hydroisomerized liquid product stream
of step b), as well as producing a distillation column overhead
vapor stream and at least a first distillate product stream from
the distillation column, wherein the distillation column overhead
vapor stream and the first distillate product stream are not sent
to the reaction zone of the FCC reactor riser.
[0035] In another more preferred embodiment of the fourth
embodiment, the separation of step c) is performed in a
distillation column to produce the hydrotreated liquid product
stream and the hydrotreated vapor stream, as well as producing a
distillation column overhead vapor stream and at least a first
distillate product stream from the distillation column, wherein the
distillation column overhead vapor stream and the distillate
product stream are not sent to the reaction zone of the FCC reactor
riser.
DETAILED DESCRIPTION OF PREFERRED EMBODIMENTS
[0036] An object of the present processes of invention is to
increase the is overall naphtha production from a Fluid Catalytic
Cracking ("FCC") unit for maximizing gasoline production. The
processes herein may also be utilized to optionally increase
distillate production, preferably for increased diesel and/or jet
fuel production. The processes herein are aimed at modifying the
properties of a typical FCC feedstream in order to increase the
naphtha, or optionally distillate, production obtained from the FCC
process. The present invention does not significantly alter the FCC
process and as such, it can be used with most existing FCC units
and hardware. Additionally, since the overall FCC process is not
significantly changed, the FCC pretreatment processes herein can be
used to prepare any amount of the FCC feedstock from a small
portion of the overall FCC feedstream (e.g., <25 wt % or <10
wt % of the overall feedstream) to preferably a significant portion
of the overall FCC feedstream >75 wt % or >85 wt % of the
overall feedstream) without significantly altering the overall
effects of the FCC unit processing on the remainder of the overall
feedstream to FCC unit.
[0037] In an embodiment of the present invention, the composition
of a typical FCC feedstream can be utilized as a starting feed to
the present processes. Preferably, the feedstream utilized has a T5
boiling point of at least 400.degree. F., more preferably of at
least 450.degree. F., and a T95 boiling point of less than about
1150.degree. F., more preferably less than about 1100.degree. F.
The "T5 boiling point" is defined as the temperature under
atmospheric pressure at which 5 wt % of the product sample boils.
Similarly, "T95 boiling point" is defined as the temperature under
atmospheric pressure at which 95 wt % of the product sample boils.
Such feeds may be derived from many sources within the refinery but
typically are comprise of an atmospheric gas oil ("AGO"), vacuum
gas oil ("VGO") or both. These feeds may also contain
hydroprocessed feed components such a product stream from a
hydrocracking unit that falls within the boiling points noted
above.
[0038] It should also be noted that the term "naphtha" as used
herein shall mean a hydrocarbon-based stream that has a T5 boiling
point of at least 80.degree. F. (27.degree. C.) and a T95 boiling
point of less than 450.degree. F. (232.degree. C.). The term
"distillate" as used herein shall mean a hydrocarbon-based stream
that has a T5 boiling point of at least 350.degree. F. (177.degree.
C.) and a T95 boiling point of less than 650.degree. F.
(343.degree. C.). Both naphthas and distillates typically refer to
intermediate product streams in a petroleum or petrochemical
refinery that may also be utilized for final product blending.
[0039] For the processes herein, the feedstream to the overall
process embodiments as described preferably is comprised
substantially of a hydrocarbon feedstream derived from a
fossil-based oil material such as a crude oil, tar sands, or
bitumens. In preferred embodiments, the feedstream is comprised of
at least 75 wt %, more preferably at least 85 wt %, of a
hydrocarbon feedstream derived from a fossil-based oil material.
The processes herein may also be utilized to process hydrocarbon
streams that are derived from renewable materials (i.e., "biofuel
sources"). However, in preferred embodiments of the processes
herein, from 5 to 25 wt %, more preferably from 10 to 20 wt %, of
the overall feedstream is derived from renewable biofuel sources.
Such biofuel sources include, but are not limited to, vegetables,
animal, fish, and/or algae materials. If such renewable biofuel
sources are utilized as a portion of the feedstream herein, it is
preferred that the materials have been hydroprocessed and
deoxygenated prior to incorporating them with the fossil-based oil
material being supplied to the present processes.
[0040] The process herein has the ability to process feedstreams
with high sulfur contents. In preferred embodiments herein the
feedstream has a sulfur content of at least 250 ppmw sulfur, or at
least 500 ppm sulfur, or at least 1000 ppmw sulfur, or at least
3000 ppmw sulfur. The feedstreams may also contain nitrogen, but it
is preferred that the nitrogen content of the feed be kept to less
is than about, 5000 ppm w, more preferably less than 2000 ppmw,
even more preferably less than 1500 ppmw, and most preferably less
than 100 ppmw of nitrogen.
[0041] In a preferred embodiment of the processes of invention
herein, a hydrocarbon feedstream is sent to a hydroisomerization
unit. In this step of the process, at least a portion of the
hydrocarbon feedstream is hydroisomerized. The increase in the
iso-paraffin content of the feed to the FCC unit is shown to
increase the much desired naphtha (i.e., gasoline) production from
the FCC unit. It has also been found that distillate (i.e., diesel)
production can also be increased with a corresponding decrease in
heavier cat bottoms products. In this step, at least a portion,
preferably most, of the normal paraffins are converted to
isoparaffinic hydrocarbon species. Preferably, at least 50 wt %,
more preferably at least 75 wt %, even more preferably at least 90
wt %, and most preferably substantially all of the normal paraffins
in the feed to the hydroisomerization unit are converted to
isoparaffins within the process. Here, it is preferred that the
hydroisomerization unit is run under conditions to maximize normal
paraffin to isoparaffin conversion, while minimizing the conversion
of naphthalenes and aromatics in the hydroisomerization feed, as
these latter components are valuable to gasoline production in the
FCC unit. With the disclosures herein, one of skill in the art will
be able to make such adjustments to the system so as to achieve
these results. In more preferred embodiments, the liquid product
from the hydroisomerization unit contains at least 10 wt %, more
preferably at least 15 wt %, and even more preferably at least 20
wt % of isoparaffinic species. in all configurations herein, the
liquid product from the hydroisomerization unit will have a higher
isoparaffin content (by wt %) than the hydrocarbon feed.stream to
the hydroisomerization unit. This hydroisomerization will also
result in some additional beneficial isomerization in alky
side-chains of the ringed molecules (such as for example,
aromatics, naphtheno-aromatics, and naphthenes) in the feed which
can be present in significant amounts. Such additional alkyl
side-chain isomerization will also improve the final naphtha and/or
distillate production in the FCC stage of the present processes
which are to be further discussed herein.
[0042] Preferred operating conditions in the hydroisomerization
reaction unit include contacting the hydroisomerization feed
obtained from the first hydrotreating step described above with an
isomerization catalyst at a temperature of from 400 to 850.degree.
F. (204 to 454.degree. C.), preferably 525 to 750.degree. F. (274
to 399.degree. C.), a hydrogen partial pressure of from 1.8 to 34.6
mPa (250 to 5000 psi), preferably 4.8 to 20.8 mPa, a liquid hourly
space velocity of from 0.2 to 10 v/v/hr. preferably 0.5 to 3.0, and
a hydrogen circulation rate of from 35.6 to 1781 m.sup.3/m.sup.3
(200 to 10,000 scf/B), preferably 178 to 890.6 m.sup.3/m.sup.3
(1000 to 5000 scf/B).
[0043] Preferably, the hydroisomerization catalysts utilized in the
processes herein are comprised of at least one zeolite. More
preferably, the zeolites have a unidimensional pore structure.
Preferred catalysts include 10 member ring pore zeolites, such as
EU-1, ZSM-35 (or ferrierite), ZSM-11, ZSM-57, NU-87, SAPO-11,
ZSM-22, and ZSM-48. Other suitable materials are EU-2, EU-11,
ZBM-30, MCM-48, and ZSM-23. Most preferably, the hydroisomerization
catalyst is comprised of ZSM-48. Note that a zeolite having the
ZSM-23 structure with a silica to alumina ratio of from about 20:1
to about 40:11 can sometimes be referred to as SSZ-32. Other
molecular sieves that are isostructural with the above materials
include Theta-1, NU-10, EU-13, KZ-1, and NU-23.
[0044] In various embodiments, the hydroisomerization catalyst
further comprises a metal hydrogenation component. The metal
hydrogenation component is typically a Group VIA and/or a Group
VIIIA metal. Preferably, the hydroisomerization catalyst includes
at least one Group VIIIA metal. More preferably, the
hydroisomerization catalyst includes at least one Group VIIIA metal
and at least one Group VIA metal. In some preferred embodiments,
the metal hydrogenation component of the hydroisomerization
catalyst is a Group VIIIA noble metal. Preferably, the metal
hydrogenation component is Pt, Pd, or a mixture thereof. In an
alternative preferred embodiment, the metal hydrogenation component
can be at least one non-noble Group VIIIA metal optionally coupled
with at least one Group VIA metal. Suitable combinations of this
alternative preferred embodiment can include Ni, Co, or Fe with Mo
or W, preferably Ni with Mo or W. The alternative
hydroisomerization catalysts are particularly preferred in
embodiments wherein the feedstream to the hydroisomerization
catalyst is not first subjected to a desulfurization step with
H.sub.7S removal.
[0045] Please note that the designation of Group VIA and Group
VIIIA herein corresponds to the older IUPAC designations such as
shown in the Periodic Table of Elements, published by the
Sargent-Welch Scientific Company, 1979, wherein the Group VIA
elements include the column from the periodic table of elements
containing Cr, Mo, and W, and the Group VIIIA elements include the
columns from the periodic table of elements containing Fe, Co, Ni,
Ru, Rh, Pd, Os, Ir, and Pt.
[0046] The amount of metal in the hydroisomerization catalyst can
be at least 0.1 wt % based on catalyst, or at least 0.15 wt %, or
at least 0.2 wt %, or at least 0.25 wt %, or at least 0.3 wt %, or
at least 0.5 wt % based on catalyst. The amount of metal in the
catalyst can be 20 wt % or less based on catalyst, or 10 wt % or
less, or 5 wt % or less, or 2.5 wt % or less, or 1 wt % or less.
For embodiments where the metal is Pt, Pd, another Group VIIIA
noble metal, or a combination thereof, the amount of metal can be
from 0.1 to 5 wt %, preferably from 0.1 to 2 wt %, or 0.25 to 1.8
wt %, or 0.4 to 1.5 wt %, For embodiments where the metal is a
combination of a non-noble Group VIIIA metal with a Group VIA
metal, the combined amount of metal can be from 0.5 wt % to 20 wt
%, or 1 wt % to 15 wt %, or 2.5 wt % to 10 wt %.
[0047] Preferably, the hydroisomerization catalysts used in
processes according to the invention are catalysts with a low ratio
of silica, to alumina. For example, for ZSM-48, the ratio of silica
to alumina in the zeolite can be less than 200:1, or less than
110:1, or less than 100:1, or less than 90:1, or less than 80:1. In
various embodiments, the ratio of silica to alumina can be from
30:1 to 200:1, 60:1 to 110:1, or 70:1 to 100:1.
[0048] The hydroisomerization catalysts useful in processes
according to the invention can also include a hinder. In some
embodiments, the hydroisomerization catalysts used in process
according to the invention are formulated using a low surface area
binder, a low surface area binder represents a binder with a
surface area of 100 m.sup.2/g or less, or 80 m.sup.2/g or less, or
70 m.sup.2/g or less.
[0049] As noted in the alternative preferred embodiment of the
hydroisomerization catalysts used in the processes herein, the
metal hydrogenation component is selected to be a combination of a
non-noble Group VIIIA metal with a Group VIA metal. Such suitable
combinations can include Ni, Co, or Fe with Mo or W, preferably Ni
with Mo or W. In this preferred embodiment, low limits of sulfur
content do not need to be maintained in the feed to the
hydroisomerization unit. In these embodiment, the feedstream to the
hydroisomerization unit can contain over 100 ppmw sulfur or even
over 300 ppmw, or even over 500 ppmw as these catalysts will be
resistant to substantial loss of hydroisomerization activity. These
catalysts should also have some amount of amount of
hydrodesulfurization activity.
[0050] As such, in further preferred configurations of these
embodiments, some sulfur is removed from the hydroisomerization
feedstream in this portion is of the overall process. The sulfur
may be removed from the hydroisomerization liquid product stream by
sending the hydroisomerization liquid product stream to a vapor
separator, which will remove the sulfur primarily in the form of
H.sub.2S. This vapor separation step is preferred to be included
particularly when there is no other additional hydrotreating or
substantial fractionation of the hydroisomerization unit product
prior to the hydroisomerization product being sent to an FCC unit
as described in the following steps. Preferably, this
hydroisomerization stage also removes some portion of the nitrogen
from the hydroisomerization feedstream by additionally removing
NH.sub.3 from the hydroisomerization liquid product stream in the
vapor separator.
[0051] In preferred embodiments of the processes herein wherein
there are no additional substantial intermediate treating steps
between the hydroisotnerization unit and the FCC unit (except for
post-isomerization vapor separation as described), at least a
portion of the hydroisomerized product stream is sent for further
processing in a Fluid Catalytic Cracking ("FCC") unit. Here, the
hydroisomerized product stream is preferably mixed with at least
one other heavy hydrocarbon feedstream (although the addition of
the heavy hydrocarbon feedstream is not required for the invention
embodiments herein) to make an FCC feedstream, which is then
injected through one or more feed nozzles into the feed zone of an
FCC reactor riser. Such heavy hydrocarbon feedstreams can include
heavy hydrocarbon feeds boiling in the range of about 430.degree.
F. to about 1050.degree. F. (221 to 566.degree. C.), such as gas
oils, heavy hydrocarbon oils comprising materials boiling above
1050.degree. F. (566.degree. C.); heavy and reduced petroleum crude
oil; petroleum atmospheric distillation bottoms; petroleum vacuum
distillation bottoms; pitch, asphalt, bitumen, other heavy
hydrocarbon residues; tar sand oils; shale oil; liquid products
derived from coal liquefaction processes; and mixtures thereof. The
FCC feed may also comprise recycled hydrocarbons, such as light or
heavy cycle oils. Preferred heavy hydrocarbon feedstreams for use
in the present process are vacuum gas oils boiling in the range
above about 650.degree. F. (343'C).
[0052] Within this reactor riser, the FCC feedstream, containing at
least a portion of the hydroisomerization products is contacted
with a catalytic cracking catalyst under cracking conditions
thereby resulting in spent catalyst particles containing carbon
deposited thereon and a lower boiling product stream. The cracking
conditions will typically include: temperatures from about 900 to
about 1060.degree. F. (482 to 571.degree. C.), preferably about 950
to about 1040.degree. F. (510 to 560.degree. C.); hydrocarbon
partial pressures from about 10 to 50 psia (70-345 kPa), preferably
from about 20 to 40 psia (140-275 kPa); and a catalyst to feed
(wt/wt) ratio from about 3 to 8, preferably about 5 to 6, where the
catalyst weight is total weight of the catalyst composite. Steam
may be concurrently introduced with the feed into the reaction
zone. The steam may comprise up to about 5 wt % of the feed.
Preferably, the FCC feed residence time in the reaction zone is
less than about 5 seconds, more preferably from about 3 to 5
seconds, and even more preferably from about 2 to 3 seconds.
[0053] Catalysts suitable for use within the FCC reactor herein are
fluid cracking catalysts comprising either a large-pore molecular
sieve or a mixture of at least one large-pore molecular sieve
catalyst and at least one medium-pore molecular sieve catalyst.
Large-pore molecular sieves suitable for use herein can be any
molecular sieve catalyst having an average pore diameter greater
than 0.7 nm which are typically used to catalytically "crack"
hydrocarbon feeds. It is preferred that both the large-pore
molecular sieves and the medium-pore molecular sieves used herein
be selected from those molecular sieves having a crystalline
tetrahedral framework oxide component. Preferably, the crystalline
tetrahedral framework oxide component is selected from the group
consisting of zeolites, tectosilicates, tetrahedral
aluminophosphates (ALPOs) and tetrahedral silicoaluminophosphates
(SAPOs). More preferably, the crystalline framework oxide component
of both the large-pore and medium-pore catalyst is a zeolite. It
should be noted that when the cracking catalyst comprises a mixture
of at least one large-pore molecular sieve catalyst and at least
one medium-pore molecular sieve, the large-pore component is
typically used to catalyze the breakdown of primary products from
the catalytic cracking reaction into clean products such as naphtha
and distillates for fuels and olefins for chemical feedstocks.
[0054] Large pore molecular sieves that are typically used in
commercial FCC process units are also suitable for use herein. FCC
units used commercially generally employ conventional cracking
catalysts which include large-pore zeolites such as USY or REY.
Additional large pore molecular sieves that can be employed in
accordance with the present invention include both natural and
synthetic large pore zeolites. Non-limiting examples of natural
large-pore zeolites include gmelinite, chabazite, dachiardite,
clinoptilolite, faujasite, heulandite, analcite, levynite,
erionite, sodalite, cancrinite, nepheline, lazurite, scolecite,
natrolite, offretite, mesolite, mordenite, brewsterite, and
ferrierite. Non-limiting examples of synthetic large pore zeolites
are zeolites X, Y, A, L. ZK-4, ZK-5, B, E, F, H, J, M, Q, T, W, Z,
alpha and beta, omega, REY and USY zeolites. It is preferred that
the large pore molecular sieves used herein be selected from large
pore zeolites. The more preferred large-pore zeolites for use
herein are the faujasites, particularly zeolite Y, USY, and
REY.
[0055] Medium-pore size molecular sieves that are suitable for use
herein include both medium pore zeolites and
silicoaluminophosphates (SAPOs). Medium pore zeolites suitable for
use in the practice of the present invention are described in
"Atlas of Zeolite Structure Types", eds. W. H. Meier and D. H.
Olson, Butterworth-Heineman, Third Edition, 1992, which is hereby
incorporated by reference. The medium-pore size zeolites generally
have an average pore diameter less than about 0.7 nm, typically
from about 0.5 to about 0.7 nm and includes for example, MFI, MFS,
MEL, MTW, EUO, MTT, HEU, FER, and TON structure type zeolites
(IUPAC Commission of Zeolite Nomenclature). Non-limiting examples
of such medium-pore size zeolites, include ZSM-5, ZSM-12, ZSM-22,
ZSM-23, ZSM-34, ZSM-35, ZSM-38, ZSM-48, ZSM-50, silicalite, and
silicalite 2. The most preferred medium pore zeolite used in the
present invention is ZSM-5, which is described in U.S. Pat. Nos.
3,702,886 and 3,770,614. ZSM-11 is described in U.S. Pat. No.
3,709,979; ZSM-12 in U.S. Pat. No. 3,832,449; ZSM-21 and ZSM-38 in
U.S. Pat. No. 3,948,758; ZSM-23 in U.S. Pat. No. 4,076,842; and
ZSM-35 in U.S. Pat. No. 4,016,245. As mentioned above SAPOs, such
as SAPO-11, SAPO-34, SAPO-41, and SAPO-42, which are described in
U.S. Pat. No. 4,440,871 can also be used herein. Non-limiting
examples of other medium pore molecular sieves that can be used
herein are chromosilicates; gallium silicates; iron silicates;
aluminum phosphates (ALPO), such as ALPO-11 described in U.S. Pat.
No. 4,310,440; titanium aluminosilicates (TASO), such as TASO-45
described in EP-A No. 229,295; boron silicates, described in U.S.
Pat. No. 4,254,297; titanium aluminophosphates (TAPO), such as
TAPO-11 described in U.S. Pat. No. 4,500,651 and iron
aluminosilicates. All of the above patents are incorporated herein
by reference.
[0056] The medium-pore size zeolites used herein can also include
"crystalline admixtures" which are thought to be the result of
faults occurring within the crystal or crystalline area during the
synthesis of the zeolites. Examples of crystalline admixtures of
ZSM-5 and ZSM-11 are disclosed in U.S. Pat. No. 4,229,424 which is
incorporated herein by reference. The crystalline admixtures are
themselves medium-pore size zeolites and are not to be confused
with physical admixtures of zeolites in which distinct crystals of
crystallites of different zeolites are physically present in the
same catalyst composite or hydrothermal reaction mixtures,
[0057] The large-pore and medium-pore catalysts of the present
invention will typically be present in an inorganic oxide matrix
component that binds the catalyst components together so that the
catalyst product is hard enough to survive inter-particle and
reactor wall collisions. The inorganic oxide matrix can be made
from an inorganic oxide sol or gel which is dried to "glue" the
catalyst components together. Preferably, the inorganic oxide
matrix will be comprised of oxides of silicon and aluminum. It is
also preferred that separate alumina phases be incorporated into
the inorganic oxide matrix. Species of aluminum
oxyhydroxides-.gamma.-alumina, boehmite, diaspore, and transitional
aluminas such as .alpha.-alumina, .beta.-alumina, .gamma.-alumina,
.delta.-alumina, .epsilon.-alumina, .kappa.-alumina, and
.rho.-alumina can be employed. Preferably, the alumina species is
an aluminum trihydroxide such as gibbsite, bayerite, nordstrandite,
or doyelite. The matrix material may also contain phosphorous or
aluminum phosphate. It is within the scope of this invention that
the large-pore catalysts and medium-pore catalysts be present in
the same or different catalyst particles, in the aforesaid
inorganic oxide matrix.
[0058] In the FCC reactor, the cracked FCC product is removed from
the fluidized catalyst particles. Preferably this is done with
mechanical separation devices, such as an FCC cyclone. The FCC
product is removed from the reactor via an overhead line, cooled
and sent to a fractionator tower for separation into various
cracked hydrocarbon product streams. These product streams may
include, but are not limited to, a light gas stream (generally
comprising C.sub.4 and lighter hydrocarbon materials), a naphtha
(gasoline) stream, a distillate (diesel and/or jet fuel) steam, and
other various heavier gas oil product streams, in the present
invention, the gasoline production in increased (maximized) clue to
the series of pretreatment steps described herein to at least one
component stream of the combined FCC feedstream to the FCC
reactor.
[0059] In the FCC reactor, after removing most of the cracked FCC
product through mechanical means, the majority of, and preferably
substantially all of, is the spent catalyst particles are conducted
to a stripping zone within the FCC reactor. The stripping zone will
typically contain a dense bed (or "dense phase") of catalyst
particles where stripping of volatiles takes place by use of a
stripping agent such as steam. There will also be space above the
stripping zone wherein the catalyst density is substantially lower
and which space can be referred to as a "dilute phase". This dilute
phase can be thought of as either a dilute phase of the reactor or
stripper in that it will typically be at the bottom of the reactor
leading to the stripper.
[0060] The majority of, and preferably substantially all of, the
stripped catalyst particles are subsequently conducted to a
regeneration zone wherein the spent catalyst particles are
regenerated by burning coke from the spent catalyst particles in
the presence of an oxygen containing gas, preferably air thus
producing regenerated catalyst particles. This regeneration step
restores catalyst activity and simultaneously heats the catalyst to
a temperature from about 1200.degree. F. to about 1400.degree. F.
(649 to 760.degree. C.). The majority of, and preferably
substantially all of the hot regenerated catalyst particles are
then recycled to the FCC reaction zone where they contact injected
FCC feed.
[0061] In additional preferred embodiments of the processes of
invention, a hydrotreating step is included in the overall process
either prior to or following the hydroisomerization step. In the
hydrotreating step, a feedstream is contacted with a hydrotreating
catalyst under hydrotreating conditions which include temperatures
in the range 450.degree. F. to 750.degree. F. (232.degree. C. to
399.degree. C.), preferably 550.degree. F. to 700.degree. F.
(288.degree. C. to 371.degree. C.) at pressures in the range of
1480 to 20786 kPa (200 to 3000 psig), preferably 2859 to 13891 kPa
(400 to 2000 psig), a space velocity of from 0.1 to 10 LHSV,
preferably 0.1 to 5 LHSV, and a hydrogen treat gas rate of from 18
to 890 m.sup.3/m.sup.3 (100 to 5000 scf/B), preferably 44 to 178
m.sup.3/m.sup.3 (250 to 1000 scf/B).
[0062] Hydrotreating catalysts suitable for use herein are those
containing at least one Group VIA metal and at least one of a Group
VIIIA metal, including mixtures thereof. Preferred metals include
Ni, W, Mo, Co and mixtures thereof, with CoMo, NiMoW, or NiW being
preferred. These metals or mixtures of metals are typically present
as oxides or sulfides on refractory metal oxide supports. The
mixture of metals may also be present as bulk metal catalysts
wherein the amount of metal is 30 wt % or greater, based on
catalyst,
[0063] Suitable metal oxide supports for the hydrotreating
catalysts include oxides such as silica, alumina, silica-alumina,
titania, or zirconia; preferably alumina. Preferred aluminas are
porous aluminas such as gamma or eta. When a porous metal oxide
support is utilized, the catalyst has an average pore size (as
measured by nitrogen adsorption) of preferably in the range of
about 100 .ANG. to about 1000 .ANG., more preferably from about 200
.ANG. to about 500 .ANG.; and the catalyst has a surface area (as
measured by the BET method) of about 100 to 350 m.sup.2/g, more
preferably about 150 to 250 m.sup.2/g. The amount of metals for
supported hydrotreating catalysts, either individually or in
mixtures, ranges from 0.5 to 35 wt %, based on catalyst. In the
case of preferred mixtures of Group VIA and Group VIIIA metals, the
Group VIIIA metals are present in amounts of from 0.5 to 5 wt %
based on catalyst, and the Group VIA metals are present in amounts
of from 5 to 30 wt % based on the catalyst.
[0064] In an embodiment, the hydrotreating step may comprise a unit
separate from the hydroisomerization step, such unit comprising at
least one hydrotreating reactor, and in an alternate embodiment
comprising two hydrotreating reactors arranged in series flow.
Here, a vapor separation drum is preferably oriented after each
hydrotreating reactor and removes the vapor phase reaction products
from the reactor effluent(s). This vapor phase is primarily
comprised of hydrogen, H.sub.2S, NH.sub.3, and hydrocarbons
containing four (4) or less carbon atoms (i.e.,
"C.sub.4-hydrocarbons"). In the hydrotreating process, is
preferably at least 70 wt %, more preferably at least 80 wt %, and
even more preferably at least 90 wt % of the sulfur content in the
feedstream is removed from the resulting liquid products.
Additionally, preferably at least 50 wt %, more preferably at least
75 wt %, of the nitrogen content in the feedstream is removed from
the resulting liquid products. Preferably, the final liquid product
from the hydrotreating unit has less than about 100 ppmw sulfur,
more preferably less than about 50 ppmw sulfur, and most
preferably, less than about 30 ppmw sulfur. However, as will be
described more fully in select embodiments below, the liquid
product from the hydrotreating unit may contain over 100 ppmw
sulfur or even over 300 ppmw depending on the catalyst and
conditions in the isomerization stage of the process as will be
described next.
[0065] In embodiments or the processes herein wherein the
hydrocarbon feedstream is first hydrotreated, then hydroisomerized,
or wherein the hydrocarbon feedstream is first hydroisomerized and
then hydrotreated prior to processing in the FCC unit, an
intermediate vapor separator drum, such as described above, may
optionally be employed. If employed, the vapor separation step
would be utilized to remove at least a portion of the sulfur
species and nitrogen species in the feed as gases (e.g., H.sub.2S
and NH.sub.3) prior to the treated product from the former stage
being processed in the latter stage.
[0066] In an alternate and improved embodiment, at least a portion,
preferably all, of the hydrotreating catalyst and at least a
portion of the, preferably all, of the hydroisomerization catalyst
are located in the same reactor. Here, it is preferred that the
hydrotreating catalyst is located in at least one separate catalyst
bed within the reactor and that the hydroisomerization catalyst is
located in at least one separate catalyst bed within the same
reactor. In this embodiment, the hydrocarbon feedstream can first
contact, or flow through, the hydrotreating catalyst bed and then
contact, or flow through, the hydroisomerization catalyst bed, or
visa versa.
[0067] In this single reactor embodiment of the present processes
however, it is preferred that the hydrocarbon feedstream first
contacts the hydrotreating catalyst bed prior to the
hydroisomerization catalyst bed. In these common reactor
embodiments, no intermediate vapor removal is required. However,
the introduction of a hydrogen-containing stream between the beds
may be optionally employed. In alternative embodiments, the
hydroisomerization catalyst and the hydrotreating catalyst which
are located in the same reactor do not need to be in separate beds
but rather the bed(s) can be comprised of a mixture of the
hydroisomerization and hydrotreating catalysts. In these single
reactor embodiments, it is preferred that the metal hydrogenation
component of the hydroisomerization catalyst is a non-noble Group
VIIIA metal optionally coupled with at least one Group VIA metal.
Suitable combinations of metals in this embodiment of the
hydroisomerization catalyst can include Ni, Co, or Fe with Mo or W,
preferably Ni with Mo or W.
[0068] In preferred embodiments, a vapor separation drum or a
fractionation stage is employed between the hydroisomerization or
combination hydrotreating/hydroisomerization steps described and
the processing of the hydrocarbons in the FCC unit, As described
prior, a vapor separation drum may be utilized in this step to
remove hydrogen, H.sub.2S, NH.sub.3 and/or light gas products.
However, in a preferred embodiment of the processes herein, a
distillation column (or "fractionator") is utilized to separate
some of the products from the hydroisomerization or combination
hydrotreating/hydroisomerization steps prior to sending the
remaining hydrocarbon products to the FCC unit for further
processing. In a particularly preferred embodiment, the product
from the hydroisomerization or combination
hydrotreating/hydroisomerization steps is sent to a distillation
column for further separation of components. Here at least one
overhead vapor stream is removed, at least one distillate product
stream is removed, and at least one other distillation product
stream is removed from the distillation column. This at least one
distillation product stream is then sent to the FCC unit for
further processes as noted in the invention herein. In a, preferred
embodiment, this at least one distillation product stream comprises
higher boiling point hydrocarbon fractions than the distillate
product stream. In another preferred embodiment, this at least one
distillation product stream comprises naphtha range hydrocarbon
fractions. In these embodiments employing this "pre-FCC"
distillation column, in a preferred embodiment, the products from
the FCC process are further fractionated in an FCC fractionator
column from which at least one distillate range product stream is
drawn from the FCC fractionator and combined with the distillate
product stream from the pre-FCC distillation column, In this
manner, overall distillate production can be increased.
[0069] Additionally or alternatively, the present invention can be
described according to one or more of the following
embodiments.
[0070] Embodiment 1. A process for increasing Fluid Catalytic
Cracking ("FCC") gasoline production comprising:
[0071] a) contacting a hydrocarbon-containing hydroisomerization
feedstream with a hydroisomerization catalyst under
hydroisomerization conditions to produce at least one
hydroisomerized liquid product stream that has a higher
iso-paraffin content than the hydroisomerization feedstream;
[0072] b) contacting in the reaction zone of an FCC reactor riser
an FCC feedstream comprising at least a portion of the
hydroisomerized liquid product stream of step a) with a fluid
catalytic cracking catalyst thereby catalytically cracking the FCC
feedstream into an FCC product that has an average lower boiling
point than the FCC feedstream, and producing a spent catalyst;
[0073] c) separating the FCC product from the spent catalyst;
[0074] d) cooling the FCC product; and
[0075] e) fractionating the FCC product into multiple FCC product
streams, is wherein at least one of the FCC product streams is a
naphtha boiling-range product stream; and
[0076] f) utilizing at least a portion of the naphtha boiling-range
product stream for gasoline production.
[0077] Embodiment 2. The process of embodiment 1, further
comprising:
[0078] contacting a hydrocarbon-containing hydrotreater feedstream
containing at least 250 ppmw of sulfur with a hydrotreating
catalyst under hydrotreating conditions to produce at least one
hydrotreated liquid product stream and one hydrotreated vapor
stream, wherein the hydrotreated liquid product stream has a lower
sulfur content than the sulfur-containing hydrocarbon
feedstream;
[0079] separating the hydrotreated liquid product stream from the
hydrotreated vapor stream; and
[0080] utilizing at least a portion of the hydrotreated liquid
product stream as the hydroisomerization feedstream in step a).
[0081] Embodiment 3. The process of embodiment 1, further
comprising:
[0082] contacting a hydrocarbon-containing hydrotreater feedstream
containing at least 250 ppmw of sulfur with a hydrotreating
catalyst under hydrotreating conditions to produce the
hydroisomerization feedstream.
[0083] Embodiment 4. A process for increasing Fluid Catalytic
Cracking ("FCC") gasoline production comprising:
[0084] a) contacting a hydrocarbon-containing hydroisomerization
feedstream with a hydroisomerization catalyst under
hydroisomerization conditions to produce at least one
hydroisomerized product stream that has a higher iso-paraffin
content than the hydroisomerization feedstream;
[0085] b) contacting at least a portion of the hydroisomerized
product stream with a hydrotreating catalyst under hydrotreating
conditions to produce at least one hydrotreated liquid product
stream and one hydrotreated vapor stream, is wherein the
hydrotreated liquid product stream has a lower sulfur content than
the sulfur-containing hydrocarbon feedstream;
[0086] c) separating the hydrotreated liquid product stream from
the hydrotreated vapor stream;
[0087] d) contacting in the reaction zone of an FCC reactor riser
an FCC feedstream comprising at least a portion of the hydrotreated
liquid product stream of step c) with a fluid catalytic cracking
catalyst thereby catalytically cracking the FCC feedstream into an
FCC product that has an average lower boiling point than the FCC
feedstream, and producing a spent catalyst;
[0088] e) separating the FCC product from the spent catalyst;
[0089] f) cooling the FCC product; and
[0090] e) fractionating the FCC product into multiple FCC product
streams, wherein at least one of the FCC product streams is a
naphtha boiling-range product stream; and
[0091] g) utilizing at least a portion of the naphtha boiling-range
product stream for gasoline production.
[0092] Embodiment 5. The process of any of embodiments 1-4, wherein
at least 50 wt % of the normal paraffins in the hydroisomerization
feedstream are converted to iso-paraffins in the hydroisomerized
liquid product stream in step a).
[0093] Embodiment 6. The process of any of embodiments 1-5, wherein
the hydroisomerization catalyst comprises at least one Group VIIIA
metal, and further comprises a zeolite selected from EU-1, ZSM-35,
ZSM-11, ZSM-57, NU-87, SAPO-11, ZSM-22, and ZSM-48.
[0094] Embodiment 7. The process of embodiment 6, wherein the Group
VIIIA metal of the hydroisomerization catalyst is selected from Pt
and Pd.
[0095] Embodiment 8. The process of any of embodiments 6-7, wherein
the hydroisomerization catalyst further comprises at least one
Group VIA metal, wherein the Group VIA of the hydroisomerization
catalyst is selected from Mo and W, and the Group VIIIA metal of
the hydroisomerization catalyst is selected from Ni and Co.
[0096] Embodiment 9. The process of embodiment 8, wherein the Group
VIA of the hydroisomerization catalyst is W, the Group VIIIA metal
of the hydroisomerization catalyst is Ni.
[0097] Embodiment 10. The process of any of embodiments 6-9,
wherein the zeolite in the hydroisomerization catalyst is
ZSM-48.
[0098] Embodiment 11. The process of any of embodiments 1-10,
wherein the hydroisomerization feedstream contains over 300 ppmw of
sulfur.
[0099] Embodiment 12. The process of any of embodiments 1-11,
wherein the hydroisomerization conditions include a temperature of
from 400 to 850.degree. F. (204 to 454.degree. C.), a hydrogen
partial pressure of from 1.8 to 34.6 mPa (250 to 5000 psi), a
liquid hourly space velocity of from 0.2 to 10 v/v/hr, and a
hydrogen circulation rate of from 35.6 to 1781 m.sup.3/m.sup.3 (200
to 10,000 scf/B),
[0100] Embodiment 13. The process of any of embodiments 1-12,
wherein the conditions in the reaction zone of the FCC reactor
include a temperature from about 900 to about 1060.degree. F. (482
to 571.degree. C.), a hydrocarbon partial pressure from about 10 to
50 psia (70-345 kPa), and a catalyst to feed (wt/wt) ratio from
about 3 to 8, where the catalyst weight is total weight of the
fluid catalytic cracking catalyst.
[0101] Embodiment 14. The process of any of embodiments 1-13,
wherein the fluid catalytic cracking catalyst comprises at least
one large-pore size faujasite zeolite and at least one medium-pore
size zeolite selected from ZSM-5, ZSM-12, ZSM-22, ZSM-23, ZSM-34,
ZSM-35, ZSM-38, ZSM-48, ZSM-50, silicalite, and silicalite
[0102] Embodiment 15. The process of any of embodiments 1-14,
wherein the FCC feedstream further comprises a heavy hydrocarbon
feedstream boiling in the range of about 430.degree. F. to about
1050.degree. F. (221 to 566.degree. C.).
[0103] Embodiment 16. The process of embodiment 15, wherein the
heavy hydrocarbon feed stream is comprised of a hydrocarbon stream
selected from gas oil, heavy and reduced petroleum crude oil;
petroleum atmospheric distillation bottoms, petroleum vacuum
distillation bottoms, pitch, asphalt, bitumen, heavy hydrocarbon
residues, tar sand oils, shale oil, and liquid products derived
from coal liquefaction processes.
[0104] Embodiment 17. The process of any of embodiments 1-16,
wherein feed residence time in the reaction zone of the FCC reactor
riser is less than about 5 seconds.
[0105] Embodiment 18. The process of any of embodiments 1-3 as
further limited by any of embodiments 5-17, wherein the at least
one hydroisomerized liquid product stream of step a) is sent to a
distillation column of to produce the at least a portion of the
hydroisomerized liquid product stream of step b), as well as
producing a distillation column overhead vapor stream and at least
a first distillate product stream from the distillation column,
wherein the distillation column overhead vapor stream and the first
distillate product stream are not sent to the reaction zone of the
FCC reactor riser.
[0106] Embodiment 19. The process of embodiment 4 as further
limited by any of embodiments 5-17, wherein the separation of step
c) is performed in a distillation column to produce the
hydrotreated liquid product stream and the hydrotreated vapor
stream, as well as producing a distillation column overhead. vapor
stream and at least a first distillate product stream from the
distillation column, wherein the distillation column overhead vapor
stream and the distillate product stream are not sent to the
reaction zone of the FCC reactor riser.
[0107] Embodiment 20. The process of any of embodiments 18-19,
wherein at least one of the FCC product streams is an FCC
distillate boiling-range product stream and at least a portion of
the first distillate product stream is combined with at least a
portion of the FCC distillate boiling-range product stream to form
a combined distillate product stream.
[0108] Embodiment 21. The process of embodiment 20, wherein at
least a portion of the combined distillate product stream is
utilized for diesel product blending.
[0109] Embodiment 22. The process of any of embodiments 2 or 4 as
further limited by any of embodiments 5-21, wherein the
hydrotreated liquid product stream contains less than 30 ppmw of
sulfur.
[0110] Embodiment 23. The process of any of embodiments 2-4 as
further limited by any of embodiments 5-21, wherein the
hydrotreating catalyst comprises at least one Group VIA metal and
at least one Group VIIIA metal on a refractory oxide support,
wherein the refractory oxide support comprises silica, alumina, or
silica-alumina; and the hydroisomerization catalyst is comprised of
at least one Group VIIIA metal, and a zeolite selected from EU-1,
ZSM-35, ZSM-11, ZSM-57, NU-87, SAPO-11, ZSM-22, and ZSM-48.
[0111] Embodiment 24. The process of embodiment 23, wherein the
hydrotreating catalyst has a has an average pore size of from about
100 .ANG. to about 1000 .ANG., and a surface area of from about 100
to 350 m.sup.2/g.
[0112] Embodiment 25. The process of any of embodiments 2-24,
wherein the hydrotreating conditions include a temperature in the
range 450.degree. F. to 750.degree. F. is (232.degree. C. to 399T),
pressure in the range of 1480 to 20786 kPa (200 to 3000 psig), a
space velocity of from 0.1 to 10 LHSV, and a hydrogen treat gas
rate of from 18 to 890 m.sup.3/m.sup.3 (100 to 5000 scf/B).
[0113] Embodiment 26. The process of any of embodiments 2 or 3, as
further limited by any of embodiments 5-25, wherein the
hydrotreater feedstream has a T5 boiling point of at least
400.degree. F. and a T95 boiling point of less than about
1150.degree. F.
[0114] Embodiment 27. The process of embodiment 26, wherein the
hydrotreater feedstream is comprised of at least 75 wt % of a
hydrocarbon feedstream derived from a fossil-based oil material,
and is further comprised of from 5 to 25 wt % of oil derived from
renewable biofuel sources.
[0115] Embodiment 28. The process of any of embodiments 3 or 4, as
further limited by any of embodiments 5-27, wherein the
hydrotreating catalyst and the hydroisomerization catalyst are in a
single reactor.
[0116] The principles and modes of operation of this invention have
been described above with reference to various exemplary and
preferred embodiments. As understood by those of skill in the art,
the overall invention, as defined by the claims, encompasses other
preferred embodiments not specifically enumerated herein.
EXAMPLE
[0117] In the Example herein, an FCC kinetic research model was
utilized to test the effects of isomerizing the normal paraffins in
a typical heavy hydrocarbon FCC feedstream composition. This model
represents and models the effects of converting all of the normal
paraffins to isoparaffins via, a hydroisomerization catalystic
process and then catalytically cracking the resulting
hydroisomerized hydrocarbon material in a fluid catalytic cracking
is (FCC) process.
[0118] The feed compositions for the non-hydroisomerized FCC feed
("Base Case") and the hydroisomerized. FCC feed ("Isomerized Case")
of the present invention are shown in Table 1.
TABLE-US-00001 TABLE 1 FCC Feed Compositions FCC Feed Composition
Base Case Isomerized Case Specific Gravity 0.905 0.906 IBP,
(.degree. C.) 247 247 FBP, (.degree. C.) 608 608 Molecular Weight
384 387 Sulfur, (ppm) 380 380 Nitrogen, (ppm) 954 654 N-paraffins,
(vol %) 5.8 0 I-paraffins, (vol %) 12.6 18.4 Naphthenes, (vol %)
31.8 31.8 Aromatics, (vol %) 49.8 49.8
[0119] Table 2 shows a comparison of the predicted FCC product
compositions utilizing the Base Case and the Isomerized Case FCC
feed compositions shown in Table 1.
TABLE-US-00002 TABLE 2 FCC Product Compositions FCC Product
Composition, (vol %) Base Case Isomerized Case C2 and lighter 1.4
1.3 C3 1.5 1.6 C3= 6.8 6.9 C4 6.8 7.0 C4= 9.1 9.4 LCN (light
gasoline) (123.degree. C. FBP) 31.0 31.8 HCN (heavy gasoline)
(167.degree. C. FBP) 16.6 16.8 LCO (255.degree. C. FBP) 19.7 19.8
HCO (397.degree. C. FBP) 16.2 15.4 Cat Bottoms 6.7 6.0 TABLE 2
NOTES: 1) "C3=" represents C3 olefins; "C4=" represents C4 olefins;
"LCN" = light cat naphtha, "HCN" = heavy cat naphtha, "LCO" = light
cycle oil, and "HCO" = heavy cycle oil. 2) Error range in
calculated values +/- 0.05% (absolute). 3) Total values reflect an
increase in volumetric yield of the product from the FCC.
[0120] As can be seen in the product data above, the process of
invention unexpectedly resulted in a 2.5% increase in desired FCC
light gasoline production (31.0 vol % to 31.8 vol %) or a 2.1%
increase in FCC overall gasoline production (i.e., LCN+HCN, 47.6
vol % to 48.6 vol %). As noted prior, in a typical refinery, this
can translate into an FCC gasoline production increase of over
40,000 gallons per day in a refinery with a typical size FCC unit.
Just as unexpected, and very economically beneficial, is the fact
that almost all of the increase in the light gasoline production
(most desired product) is offset by an almost corresponding
decrease in the cat bottoms production (least desired product).
This increases the overall refinery economics of practicing this
process configuration.
* * * * *