U.S. patent application number 11/898675 was filed with the patent office on 2008-04-24 for process for selective sulfur removal from fcc naphthas using zeolite catalysts.
This patent application is currently assigned to ExxonMobil Research and Engineering Company. Invention is credited to Michael C. Clark, Peter W. Jacobs, Shifang L. Luo, John H. Thurtell, Benjamin S. Umansky.
Application Number | 20080093265 11/898675 |
Document ID | / |
Family ID | 39316917 |
Filed Date | 2008-04-24 |
United States Patent
Application |
20080093265 |
Kind Code |
A1 |
Umansky; Benjamin S. ; et
al. |
April 24, 2008 |
Process for selective sulfur removal from FCC naphthas using
zeolite catalysts
Abstract
A process for the removal of sulfur compounds from FCC catalytic
naphtha which is simple to implement and is economical in operation
operates by selectively hydrofining the cracked naphtha to remove
sulfur without sacrificing octane (hydrodesulfurization typically
around 85%) followed by a downstream alkylation which will remove
residual sulfur from the gasoline boiling range. The alkylation is
carried out using a solid, acidic molecular sieve catalyst under
mild conditions to shift sulfur species from the lighter,
olefin-rich portions of the naphtha to the heavy, olefin-poor
gasoline. Separation of the heavy portion of the treated product
followed by catalytic hydrodesulfurization or some other means
(e.g. sorption) removes the sulfur from the heavy portion without
losing significant octane value. Some of the heavy compounds formed
in this process can drop out of the gasoline range, into the
kero/light diesel range; this stream can be hydrotreated without
penalty since octane is not a requirement for the middle distillate
fuels.
Inventors: |
Umansky; Benjamin S.;
(Fairfax, VA) ; Luo; Shifang L.; (Clinton, NJ)
; Jacobs; Peter W.; (Bound Brook, NJ) ; Thurtell;
John H.; (Centreville, VA) ; Clark; Michael C.;
(Chantilly, VA) |
Correspondence
Address: |
ExxonMobil Research & Engineering Company
P.O. Box 900
1545 Route 22 East
Annandale
NJ
08801-0900
US
|
Assignee: |
ExxonMobil Research and Engineering
Company
Annandale
NJ
|
Family ID: |
39316917 |
Appl. No.: |
11/898675 |
Filed: |
September 14, 2007 |
Related U.S. Patent Documents
|
|
|
|
|
|
Application
Number |
Filing Date |
Patent Number |
|
|
60852404 |
Oct 18, 2006 |
|
|
|
Current U.S.
Class: |
208/216R ;
208/213 |
Current CPC
Class: |
C10G 2400/02 20130101;
C10G 29/205 20130101 |
Class at
Publication: |
208/216.00R ;
208/213 |
International
Class: |
C10G 45/60 20060101
C10G045/60 |
Claims
1. A process for the removal of sulfur compounds from a
catalytically cracked petroleum naphtha feed comprising organic
sulfur compounds and olefins, which comprises: Selectively
hydrodesulfurising the cracked naphtha under conditions and with a
catalyst which is effective to convert organic sulfur in the
naphtha feed to inorganic form with a selective, limited degree of
hydrogenation of olefinic compounds, Subjecting the
hydrodesulfurized naphtha to a fixed bed alkylation step in the
presence of a solid molecular sieve alkylation catalyst to alkylate
organic sulfur compounds in the hydrodesulfurized naphtha with the
olefins contained in the naphtha; Fractionating the
hydrodesulfurized, alkylated naphtha to remove alkylated sulfur
compounds which boil above the gasoline boiling range.
2. A process according to Paragraph 1 in which the
hydrodesulfurized naphtha is subjected to alkylation over a zeolite
alkylation catalyst comprising an intermediate pore size
zeolite.
3. A process according to Paragraph 2 in which the
hydrodesulfurized naphtha is subjected to alkylation over a zeolite
alkylation catalyst comprising ZSM-5 or ZSM-12.
4. A process according to Paragraph 1 in which the
hydrodesulfurized naphtha is subjected to alkylation over a zeolite
alkylation catalyst comprising a zeolite of the MWW family.
5. A process according to Paragraph 4 in which the zeolite of the
MWW family comprises MCM-22 or MCM-49.
6. A process according to Paragraph 1 in which additional benzene
from outside sources is added to the alkylation step.
7. A process according to Paragraph 1 in which the
hydrodesulfurized, alkylated naphtha is fractionated into a light
fraction and a heavy fraction with the heavy fraction being
subjected to a hydrodesulfurized step to remove sulfur
compounds.
8. A process according to Paragraph 7 in which the heavy fraction,
after being hydrodesulfurized to remove sulfur compounds, is
recombined with the light fraction.
9. A process according to claim 1 in which the desulfurization
during the selective hydrodesulfurization process is not greater
than 90%.
10. A process according to claim 1 in which the desulfurization
during the selective hydrodesulfurization process is not greater
than 80%.
11. A process according to claim 1 in which the sulfur removal from
the 190.degree. C.- fraction of the hydrodesulfurized naphtha is at
least 50%.
12. A process according to claim 1 in which the sulfur removal from
the 190.degree. C.- fraction of the hydrotreated naphtha is at
least 75%.
13. A process according to claim 1 in which the sulfur removal from
the 190.degree. C.- fraction of the hydrodesulfurized naphtha is at
least 85%.
14. A process according to claim 1 in which the boiling range
conversion of the 190 C- (375 F-) portion of the hydrodesulfurized
naphtha to the 190.degree. C.+ (375.degree. F.+) portion is less
than 20%.
15. A process according to claim 1 in which the selective naphtha
hydrodesulfurization process is carried out in the presence of a
supported cobalt-molybdenum catalyst which comprises a porous
inorganic refractory support, about 1 to 10 wt. % MoO.sub.3 and
about 0.1 to 5 wt. % CoO; and having a Co/Mo atomic ratio of about
0.1 to 1.0; and a median pore diameter of about 60 to 200 .ANG.;
and a MoO3 surface concentration of about 0.5.times.10.sup.-4 to
3.times.10.sup.-4 gm.sup.-2 MoO3; and an average particle size
diameter of less than about 2.0 mm.
16. A process according to claim 15 in which the
hydrodesulfurization catalyst has a metals sulfide edge plane area
from about 7600 to 2800 .mu.mol oxygen/g MoO3 as measured by oxygen
chemisorption.
17. A process according to claim 15 in which the
hydrodesulfurization is such that the inlet temperature of the
feedstock to the reaction unit is below the dew point of the
feedstock and 100% of the feedstock becomes vaporized in the
catalyst.
18. A process according to claim 15 in which the liquid hourly
space velocity of the cracked naphtha over the hydrodesulfurization
catalyst is from about 0.5 hr-1 to about 15 hr.sup.-1.
19. A process according to claim 15 in which the
hydrodesulfurization catalyst has a median pore diameter of from 80
to 150 .ANG..
20. A process according to claim 15 in which the
hydrodesulfurization catalyst has a MoO.sub.3 surface concentration
of about 1.times.10.sup.-4 to 2.times.10.sup.-4 g/m.sup.2
MoO.sub.3.
Description
CROSS REFERENCE TO RELATED APPLICATIONS
[0001] This application claims the priority of U.S. Patent
Application No. 60/852,404, filed 18 Oct. 2006.
FIELD OF THE INVENTION
[0002] This invention relates to a process for the upgrading of
hydrocarbon streams. It more particularly refers to a process for
upgrading gasoline boiling range petroleum fractions containing
substantial proportions of sulfur impurities by reducing the sulfur
content and also by removing benzene.
BACKGROUND OF THE INVENTION
[0003] Catalytically cracked gasoline currently forms a major part
of the gasoline product pool in the United States and it provides a
large proportion of the sulfur in the gasoline. The sulfur
impurities may require removal, usually by hydrotreating, in order
to comply with product specifications or to ensure compliance with
environmental regulations, both of which are expected to become
more stringent in the future, possibly permitting no more than
about 300 ppmw sulfur in motor gasolines; low sulfur levels result
in reduced emissions of CO, NOx and hydrocarbons.
[0004] Cracked naphtha, as it comes from the catalytic cracker and
without any further treatments, such as purifying operations, has a
relatively high octane number as a result of the presence of
olefinic components. In some cases, this fraction may contribute as
much as up to half the gasoline in the refinery pool, together with
a significant contribution to product octane, making it desirable
to remove the sulfur while retaining the more desirable olefins.
Hydrotreating of any of the sulfur containing fractions which boil
in the gasoline boiling range causes a reduction in the olefin
content, and consequently a reduction in the octane number and as
the degree of desulfurization increases, the octane number of the
normally liquid gasoline boiling range product decreases. Some of
the hydrogen may also cause some hydrocracking as well as olefin
saturation, depending on the conditions of the hydrotreating
operation. In any case, regardless of the mechanism by which it
happens, the decrease in octane which takes place as a consequence
of sulfur removal by hydrotreating creates a tension between the
growing need to produce gasoline fuels with higher octane number
and--because of current ecological considerations--the need to
produce cleaner burning, less polluting fuels, especially low
sulfur fuels. This inherent tension is yet more marked in the
current supply situation for low sulfur, sweet crudes.
[0005] A number of different techniques for reducing gasoline
sulfur have been recognized as promising. Basically, these
techniques have been categorized into four classes:
[0006] Hydrofining of FCC feed
[0007] Desulfurization of FCC gasoline
[0008] Use of catalysts and additives in the FCC process.
[0009] Sorption
[0010] All of these techniques have been applied with greater or
lesser success depending upon product requirements, economics and
the relative demand levels for gasoline and other fuel products. As
a generalization, the hydrofining of the FCC feed is considered an
expensive option in view of its hydrogen consumption even though
yields benefits result for the LPG and gasoline fractions at the
expense of cycle and slurry oils. Post-treatment of the FCC naphtha
is regarded as a lower cost option and has been implemented in a
number of competitive processes, each with their own distinguishing
characteristics. One group, exemplified by the UOP ISAL Process and
the Mobil OCTGAIN Process, suppress octane loss by isomerizing
saturated hydrocarbons following the non-selective
hydrodesulfurization. Selective hydrodesulfurization in processes
such as ExxonMobil SCANfining, CDHydro, CDHDS, Axens Prime G+ and
the BP-Axens OATS Process function by suppressing olefin
saturation, mainly in the naphtha front end, by various techniques
including catalytic selectivity as in the SCANfining Process or by
distillative separation as in the other processes of this type.
Sorption removal of sulfur is exemplified in the S Zorb Process
(Phillips) and the IRVAD Process.
[0011] In the non-selective naphtha hydrodesulfurization processes,
the FCC naphthas and other light fractions such as heavy cracked
gasoline are hydrotreated by passing the feed over a hydrotreating
catalyst at elevated temperature and somewhat elevated pressure in
a hydrogen atmosphere. After the hydrotreating operation is
complete, the product may be fractionated, or simply flashed, to
release the hydrogen sulfide and collect the now sweetened
gasoline.
[0012] The sulfur impurities tend to concentrate in the heavy
fraction of the gasoline, as noted in U.S. Pat. No. 3,957,625
(Orkin) which proposes a method of removing the sulfur by
hydrodesulfurization of the heavy fraction of the catalytically
cracked gasoline so as to retain the octane contribution from the
olefins which are found mainly in the lighter fraction. In one type
of conventional, commercial operation, the heavy gasoline fraction
is treated in this way. As an alternative, the selectivity for
hydrodesulfurization relative to olefin saturation may be shifted
by suitable catalyst selection, for example, by the use of a
magnesium oxide support instead of the more conventional
alumina.
[0013] U.S. Pat. No. 4,049,542 (Gibson) discloses a process in
which a copper catalyst is used to desulfurize an olefinic
hydrocarbon feed such as catalytically cracked light naphtha. This
catalyst is stated to promote desulfurization while retaining the
olefins and their contribution to product octane.
[0014] Processes for improving the octane rating of catalytically
cracked gasolines have been proposed. U.S. Pat. No. 3,759,821
(Brennan) discloses a process for upgrading catalytically cracked
gasoline by fractionating it into a heavier and a lighter fraction
and treating the heavier fraction over a ZSM-5 catalyst, after
which the treated fraction is blended back into the lighter
fraction. Another process in which the cracked gasoline is
fractionated prior to treatment is described in U.S. Pat. No.
4,062,762 (Howard) which discloses a process for desulfurizing
naphtha by fractionating the naphtha into three fractions each of
which is desulfurized by a different procedure, after which the
fractions are recombined.
[0015] The octane rating of the gasoline pool may be increased by
other methods, of which reforming is one of the most common. Light
and full range naphthas can contribute substantial volume to the
gasoline pool, but they do not generally contribute significantly
to higher octane values without reforming. They may, however, be
subjected to catalytically reforming so as to increase their octane
numbers by converting at least a portion of the paraffins and
cycloparaffins in them to aromatics. Fractions to be fed to
catalytic reforming, for example, with a platinum type catalyst,
need to be desulfurized before reforming because reforming
catalysts are generally not sulfur tolerant; they are usually
pretreated by hydrotreating to reduce their sulfur content before
reforming. The octane rating of reformate may be increased further
by processes such as those described in U.S. Pat. No. 3,767,568 and
U.S. Pat. No. 3,729,409 (Chen) in which the reformate octane is
increased by treatment of the reformate with ZSM-5.
[0016] Aromatics are generally the source of high octane number,
particularly very high research octane numbers and are therefore
desirable components of the gasoline pool. They have, however, been
the subject of severe limitations as a gasoline component because
of possible adverse effects on the ecology, particularly with
reference to benzene. It has therefore become desirable, as far as
is feasible, to create a gasoline pool in which the higher octanes
are contributed by the olefinic and branched chain paraffinic
components, rather than the aromatic components. Environmental
regulations related to motor fuels have produced substantial
changes in refinery operations. To comply with these regulations,
some refineries have excluded the C6 compounds from the reformer
feed to satisfy the low-level benzene requirement. A new refinery
process able to alkylate benzene and sulfur compounds with the
olefins contained in the same gasoline stream would be beneficial
not only to meet benzene specification but also to comply with
sulfur regulations.
[0017] A series of patents originating from Mobil Oil Corp.
describe a process for the upgrading of gasoline by sequential
hydrotreating and selective cracking steps. In the first step of
the process, the naphtha is desulfurized by hydrotreating and
during this step some loss of octane results from the saturation of
olefins. The octane loss is restored in the second step by a
shape-selective cracking, preferably carried out in the presence of
an intermediate pore size zeolite such as ZSM-5. The product is a
low-sulfur gasoline of good octane rating. The patents in this
series are typified by the first, U.S. Pat. No. 5,346,609.
Developments of the basic process with alternative methods of
sulfur removal intended to further minimize octane loss are U.S.
Pat. No. 5,318,690, in which the naphtha is split into two
fractions with the light fraction subjected to an extractive sulfur
removal operation which preserves olefin content the heavy fraction
which contains relatively fewer of the desirable high-octane
olefins is desulfurized by hydrodesulfurization and any resulting
octane loss is restored by a selective cracking over a zeolite. A
further development described in U.S. Pat. No. 5,360,532 uses a
final extraction step to remove recombinant mercaptans.
[0018] When this process is applied to heavy catalytic naphtha
(HCN), the yield loss can be between 6 to 15% while maintaining a
similar octane value. The yield loss is highly dependent on
octane-recovered value and the type of feed. Heavy catalytic
naphtha is the ideal stream for this technology as higher losses
are obtained with full range or intermediate catalytic
naphthas.
[0019] A different approach was taken to sulfur removal by Exxon
Research and Engineering Company in the selective naphtha
hydrofining process described in various patents including: U.S.
Pat. No. 5,985,136; U.S. Pat. No. 6,126,184; U.S. Pat. Nos.
6,231,753; 6,409,913; U.S. Pat. No. 6,231,754; US6013598; U.S. Pat.
No. 6,387,249; U.S. Pat. No. 6,596,157. The ExxonMobil selective
naphtha naptha hydrofining process, SCANfining.TM., which
incorporates aspects of the processes described in these patents,
which is commercially available under license from ExxonMobil
Research and Engineering Company has demonstrated itself to be a
very effective naphtha desulfurization process. This selective
naphtha hydrofining process was developed for deep
hydrodesulfurization with maximum preservation of the olefins
(octane). The single stage version of the process can be used with
a full range catalytic naphtha or with an intermediate catalytic
naphtha (ICN), for example a nominal 65-175 C (150-350.degree. F.)
or a heavy catalytic naphtha (HCN), for example, a nominal 175 C+
(350.degree. F.+) naphtha, or both. The two-stage version of the
process, as described in U.S. Pat. No. 6,231,753, WO 03/048273 and
WO 03/099963, adds a second reactor and inter-stage removal of H2S
allowing very deep HDS with very good olefin retention. The
operation of this process relies on a combination of a highly
selective catalyst with process conditions designed to achieve
hydrodesulfurization with minimum olefin saturation.
[0020] In cases where the sulfur content of the naphtha is modest
and very deep HDS is not required, the SCANfining process can be an
attractive option. If severe HDS conditions are needed, it is
generally better to treat the LCN stream separately to preserve the
maximum amount of olefins. Co-pending application USSN, claiming
priority from U.S. Ser. No. 60/852,405, filed 18 Oct. 2006,
entitled "Process for benzene reduction and sulfur removal from FCC
naphthas", describes a process in which the cracked naphtha feed is
split into a light fraction and a heavy fraction so as to permit
optimized processing of each fraction for sulfur removal while
retaining octane values.
SUMMARY OF THE INVENTION
[0021] We have now devised a process scheme to allow the removal of
sulfur compounds and reduce benzene content of the FCC catalytic
naphthas. This scheme is simple to implement and is economical in
operation. According to the present process scheme, the selective
naphtha hydrofining process is run to remove sulfur without
sacrificing octane (hydrodesulfurization typically around 85%)
followed by a downstream alkylation which will remove residual
sulfur from the gasoline boiling range. In the alkylation step, the
partly desulfurized naphtha is contacted with a solid, acidic
molecular sieve alkylation catalyst under mild conditions to shift
sulfur species from the lighter, olefin-rich portions of the
naphtha to the heavy, olefin-poor gasoline. Separation of the heavy
portion of the treated product followed by catalytic
hydrodesulfurization or some other means (e.g. sorption) removes
the sulfur without losing significant octane value. Some of the
heavy compounds formed in this process can drop out of the gasoline
range, into the kero/light diesel range; this stream can be
hydrotreated without penalty since octane is not a requirement for
the middle distillate fuels. In addition, other reactions may occur
in the presence of the catalyst to increase product octane,
including isomerization of olefins to iso-olefins of higher octane
and alkylation of aromatics to alkylaromatics, a reaction which is
of particular utility in converting benzene in the naphtha to
toluene and other alkylaromatics of reduced toxicity and higher
octane. In this step of the process, it is also possible to add
additional benzene from other refinery processes.
[0022] By carrying out an initial sulfur removal under conditions
which do not result in excessive olefin saturation, the octane
value of the olefins is preserved while, at the same time, enabling
a substantial amount of the sulfur compounds to be removed from the
cracked naphtha. The residual sulfur is then removed
non-hydrogenatively by alkylation in a process step in which again,
octane is largely retained. The alkylated sulfur compounds can then
be removed from the gasoline boiling range fraction by
fractionation, conveniently during product recovery. Thus, in this
way, the objective of sulfur removal is satisfied by the balancing
of conditions in the two selected process steps so that olefin
octane is retained to the extent possible; very high levels of
desulfurization can be achieved with little impact on gasoline
yield and octane.
Typical reaction pathway for the alkylation step are:
[0023] Olefin+Sulfur compound.fwdarw.Alkylated Sulfur Compounds
[0024] OlefinIso-olefins
[0025] Olefin+Aromatic.fwdarw.Alkylaromatic
[0026] The alkylation catalyst can be combined in one reactor with
the selective hydrofining catalyst beds or can be loaded in a
separate reactor.
[0027] The use of the solid molecular sieve alkylation catalyst is
particularly favorable. Unlike catalysts such as solid phosphoric
acid (sPA) catalyst, the molecular sieves are readily regenerable,
either in situ or ex situ, and can be readily handled. They are
also tolerant to the presence of water in the feed and do not tend
to disintegrate or form lumps during operation. Suitable alkylation
catalysts can be the intermediate pore size aluminosilicate
zeolites such as ZSM-5, ZSM-11, ZSM-12 or members of the MWW family
of zeolites such as MCM-22 and MCM-49. Large pore size zeolites
such as zeolite Y (usually used as USY) or zeolite beta may also be
utilized. Zeolite beta may be especially useful in view of its
capability to isomerize paraffins in the presence of aromatics,
thus, giving a further potential octane boost.
[0028] The front end of the cracked feed, which is relatively rich
in olefins is spared the saturating effect of severe
hydrodesulfurization but is nevertheless subjected to a controlled
extent of desulfurization. The back end, by contrast, is relatively
olefin-poor but high in sulfur compounds such as thiophenes and
substituted thiophenes which are not amenable to extraction by
conventional extractive processes. This higher-boiling, sulfur-rich
fraction is effectively desulfurized by the alkylation followed by
fractionation, optionally with the initial selective hydrofining
step. The sulfur from thiophenes, substituted thiophenes and other
higher boiling sulfur compounds initially present in the higher
boiling fraction, if not removed by the initial hydrofining, are
subjected to a similar type of alkylation usually in the presence
of a zeolite catalyst to convert them into sulfur compounds boiling
above the gasoline boiling range.
DRAWINGS
[0029] FIG. 1 of the drawings is a schematic of a process unit for
carrying out the sulfur reduction process.
[0030] FIG. 2 is a graph showing the sulfur and conversion
hydrocarbon yields from treatment of a naphtha with a zeolite
catalyst as described in Example 8.
[0031] FIG. 3 is a graph showing the hydrocarbon and boiling point
shifts resulting from the treatment described in Example 8.
DETAILED DESCRIPTION
[0032] Depending on the refinery operation, there are typically
three components in the FCC naphtha: Light Cat Naphtha (LCN),
Intermediate Cat Naphtha (ICN) and Heavy Cat Naphtha (HCN). The
typical LCN (nominally C5-65.degree. C., C5-150.degree. F.)
fraction contains mostly light (C1-C4) mercaptan sulfur species
with lesser amounts of carbon disulfide, methylethylsulfide (MES),
and dimethylsulfide (DMS). The LCN end point is generally set to
ensure that only minimal amounts of thiophene (b.p. 84.degree. C.,
183.degree. F.) enter the LCN stream. The typical ICN nominally in
the boiling range of 65-175.degree. C. (150-345.degree. F.)
contains most of the olefins of the FCC naphtha. Normal
hydrotreating process of this stream will drastically affect its
octane value. The typical HCN nominally boiling above 177.degree.
C. (350.degree. F.) usually contains a higher concentration of
sulfur and has the most difficult sulfur species to remove
including benzothiophenes and substituted benzothiophenes. Thus,
the problems in treating the FCC naphtha can be summarized
according to the naphtha fraction: to remove mercaptans from the
light naphtha while retaining olefins, to remove benzene, thiophene
and other sulfur compounds from the intermediate naphtha again
while retaining olefin content and with the heavy naphtha, to
remove sulfur including the refractory substituted thiophenes.
[0033] According to the present invention, the sulfur in the
cracked naphtha is effectively removed by the combination of the
selective hydrofining (hydrodesulfurization) and the alkylation
which uses the olefins contained in the naphtha (mainly in the
front end) as the source of alkylating agent. Although this results
in some olefin loss, the initial removal of sulfur in the selective
hydrofining leaves only the residual sulfur to be removed and the
olefin loss which takes place by using the olefins to alkylate the
residual sulfur is less than the loss which would take place if the
selective hydrofining were operated at the higher severity required
to convert the sulfur compounds to inorganic form. In addition, if
the alkylation conditions allow conversion of benzene (octane
.about.100) to alkylbenzenes such as toluene an octane boost may be
achieved since the alkylaromatics are of higher blending octane
(toluene octane=120), thereby compensating for the loss of
olefins.
[0034] FIG. 1 shows an illustrative unit configuration for carrying
out the present processing scheme. The FCC naphtha from the main
column enters the unit by way of line 10, coming into the selective
hydrofining reactor 11 in which the sulfur compounds are removed by
the selective hydrogenative process with its minimal tendency to
saturate the olefins. The hydrofined effluent passes in line 12 to
alkylation reactor 13 in which the sulfur compounds are converted
to higher boiling alkyl analogs. The effluent from the alkylation
reactor passes out through line 14 and the alkylated sulfur
compounds which boil above the gasoline boiling range are separated
from in fractionator 20 in the product recovery section. If the
alkylation reactions result in alkylated sulfur products which
remain in the back end (higher boiling portion) of the naphtha, the
effluent from the alkylation reactor may be fractionated to remove
the olefin-rich front end with the back end being passed to a
hydrogenative sulfur removal step, e.g. hydrotreating, interposed
between the alkylation reactor and the final product fractionator.
In this case, the effluent from the alkylation reactor will pass
through the portions of the unit indicated in feint, passing
through line 15 to fractionator 16 which splits out the front end
and sends it directly to the product recovery section through line
14. The back end passes to hydrotreater 18 in which it is
hydrotreated to convert the sulfur compounds to inorganic form. The
hydrotreated naphtha passes to the product recovery section through
line 19 and thence to fractionator 20. The sulfur is recovered from
the separator gases in the normal way (separators and scrubbers not
shown). When the secondary hydrotreater is used, the alkylated
effluent can be effectively hydrotreated to remove the sulfur since
the olefin content of this fraction is relatively limited, so that
the resulting octane loss is minimal. Light ends resulting from the
processing leave fractionator 20 through line 21 and the
desulfurized gasoline leaves through line 22. Middle distillate
either from the feed (if the gasoline cut point on fractionator 20
is lower than the feed cut point) or from the higher boiling
materials produced from the alkylation is taken off the bottom of
the fractionator through line 23.
Selective Sulfur Removal
[0035] A desirable preliminary step is the removal of basic
nitrogenous compounds from the naphtha since these may have an
inhibiting effect on the acidic molecular sieve catalyst used in
the subsequent alkylation step. An acidic wash treatment using
dilute sulfuric acid or another acid is usually effective but a
preferred treatment is with a heterogeneous acidic sorbent,
preferably an acidic cation exchange resin such as the
macroreticular cross-linked polymer sulfonate exchange resins sold
under the Amberlyst.TM. trademark, such as Amberlyst.TM. 15. Resins
of this type may be readily regenerated by acid treatment in the
conventional manner.
[0036] The sulfur in the cracked naphtha is removed by a process
which has minimal effect on the olefin content of the naphtha. To
this end, a mercaptan extraction process such as the extractive
Merox process may be used. The Merox sweetening type processes such
as the Minalk process and the oxidative Merox process may also be
used provided that the fractionation step which follows in the
process sequence is effective to remove the disulfides and enable
the gasoline fraction to meet applicable product sulfur
specifications. The mercaptan extraction processes, however, are
less effective with thiophene and other forms of sulfur and if a
significant amount of non-mercaptan sulfur is present in the
cracked naphtha, too much will be carried through to the alkylation
step with the result that olefin consumption in this step will be
increased. The selection of one of these sulfur removal processes
is therefore dependent very much upon the sulfur species present in
the feed.
[0037] A preferred sulfur removal technique which is effective
against sulfur species present in naphtha streams together with
other impurities is the selective hydrofining process referred to
above which is capable of converting organic sulfur in the naphtha
feed to inorganic form with a selective limited degree of
hydrogenation of olefinic and aromatic compounds. As noted above,
the selective naphtha hydrofining process is described in U.S. Pat.
No. 5,985,136; U.S. Pat. Nos. 6,013,598; 6,126,184; U.S. Pat. No.
6,197,718; U.S. Pat. No. 6,231,753; U.S. Pat. No. 6,231,754; U.S.
Pat. Nos. 6,387,249; 6,409,913; U.S. Pat. No. 6,596,157; U.S. Pat.
No. 6,589,418; US 20030106839A1; US 20030127362A1; US
20030188992A1; US 20030220186A1; US20030217952A1; US20030183556A1;
US 20040026298A1, US 20030221994A1; WO 2004/062796; WO 2005/037959,
to which reference is made for a description and definition of the
process and of variations in it. The ExxonMobil selective naphtha
naptha hydrofining process, SCANfining.TM., which incorporates
aspects of the processes described in these patents, is
commercially available under license from ExxonMobil Research and
Engineering Company and is highly suitable for this purpose.
[0038] The selective naphtha hydrofining process can be defined as
one using a supported cobalt-molybdenum catalyst which is
characterized by a combination of parameters including: [0039]
about 1 to 10 wt. % MoO.sub.3; and [0040] about 0.1 to 5 wt. % CoO;
and [0041] a Co/Mo atomic ratio of about 0.1 to 1.0, preferably 0.2
to 0.8 and most preferably from 0.25 to 0.72; and [0042] a median
pore diameter of about 60 to 200 .ANG., preferably 75 to 175 .ANG.
and most preferably 80 to 150 .ANG.; and [0043] a MoO.sub.3 surface
concentration of about 0.5.times.10.sup.-4 to 3.times.10.sup.-4
g/m.sup.2 MoO.sub.3, preferably 0.75.times.10.sup.-4 to
2.5.times.10.sup.-4 g/m.sup.2 MoO.sub.3 and most preferably from
1.times.10.sup.-4 to 2.times.10.sup.-4 g/m.sup.2 MoO.sub.3 [0044]
an average particle size diameter of less than about 2.0 mm,
preferably less than 1.6 mm and most preferably less than 1.4 mm.
Preferably, the catalyst has a metals sulfide edge plane area from
about 7600 to 2800 .mu.mol oxygen/g MoO3 as measured by oxygen
chemisorption. It is also preferable for the MoO.sub.3 surface
concentration to be within the range of about 0.75.times.10.sup.-4
to about 2.5.times.10.sup.-4 g/m.sup.2 MoO, with the Co/Mo atomic
ratio from about 0.20 to 0.85.
[0045] In a preferred manner of operation, the process conditions
of the hydrofining are such that the inlet temperature of the
feedstock to the reaction unit is below the dew point of the
feedstock and 100% of the feedstock becomes vaporized in the
catalyst bed. The process generally operates at a liquid hourly
space velocity of from about 0.5 hr-1 to about 15 hr.sup.-1,
preferably from about 0.5 hr.sup.-1 to about 10 hr.sup.-1, and most
preferably from about 1 hr.sup.-1 to about 5 hr.sup.-1.
[0046] The metal sulfide edge plane area as measured by the Oxygen
Chemisorption Test is described in "Structure and Properties of
Molybdenum Sulfide: Correlation of O.sub.2 Chemisorption with
Hydrodesulfurization Activity", S. J. Tauster et al., Journal of
Catalysis 63, pp 515-519 (1980). The Oxygen Chemisorption Test
involves edge-plane area measurements made wherein pulses of oxygen
are added to a carrier gas stream and thus rapidly traverse the
catalyst bed. For example, the oxygen chemisorption will be from
about 800 to 2,800, preferably from about 1,000 to 2,200, and more
preferably from about 1,200 to 2,000 .mu.mol oxygen/gram
MoO.sub.3.
[0047] Further features of the preferred selective
hydrodesulfurization process and the selective catalyst used in the
process are found in the US patents referred to above in connection
with the SCANfining.TM. process, to which reference is made for a
description of them.
Alkylation-Process Parameters
[0048] Operation may take place under vapor phase, liquid phase or
supercritical phase conditions (reactor inlet). Frequently, mixed
phase conditions will prevail, depending on the feed composition
and the conditions used. At the reactor outlet, liquid phase will
prevail under normal conditions with the product including
significant proportions of C.sub.8, C.sub.10 and higher
hydrocarbons. Pressure need not be above autogenous, sufficient to
maintain the reactants in the liquid phase if liquid phase
operation is desired. The pressure will therefore normally be
dependent on unit constraints but usually will not exceed about
10,000 kPag (about 1450 psig) with low to moderate pressures,
normally not above 7,000 kPag (about 1,000 psig) being favored from
equipment and operating considerations although higher pressures
are not excluded. In most cases, the pressure will be in the range
of 1000 to 5500 kPag (about 145 to 800 psig) in order to make use
of existing equipment. Space velocities can be quite high, giving
good catalyst utilization. Space velocities are normally in the
range of 0.5 to 5 hr.sup.-1 WHSV for the olefin feed, in most
cases, 1 to 2 hr.sup.-1 WHSV. Optimum conditions may be determined
empirically, depending on feed composition, catalyst aging and unit
constraints.
[0049] Two factors affecting choice of temperature will be the feed
composition and the level of sulfur and other impurities. The
sulfur acts as a catalyst poison at relatively low reaction
temperatures, typically about 120.degree. C., but has relatively
little effect at higher temperatures about 180.degree. C. or
higher, e.g. 200.degree. C., 220.degree. C., so that the preferred
temperature regime is above about 150.degree. C., with temperatures
above 180.degree. C. or higher being preferred, e.g. 200.degree. or
220.degree. C. or higher. In general terms, the temperature will be
from about 120.degree. to 350.degree. C. (about 250 to 660.degree.
F.) and in most cases between 150.degree. and 250.degree. C. (about
300 to 480.degree. F.).
[0050] If the naphtha fraction contains large proportions of
benzene, it may be desirable to add olefins to the alkylation step
from external sources in order to promote alkylation of the benzene
to alkylaromatics. In such cases, lighter olefins may be used if
available. FCC off-gas may be added with the advantage that the
resulting alkylaromatic products have the lower carbon numbers
characteristic of the preferred gasoline aromatics. The ratio
between the total olefin and aromatic feed components is normally
chosen to achieve the desired degree of benzene reduction
consistent with the use of this process step to reduce sulfur.
Optimal conditions may be determined empirically depending on feed
composition, available feed rates, product objectives and unit
type.
Alkylation Catalyst
[0051] The catalysts used in the alkylation contain a solid
molecular sieve with acidic functionality as their essential
catalytic component, preferably an intermediate or large pore size
sieve. The intermediate pore size molecular sieves are a well
established class and may comprise zeolites such as the
aluminosilicate zeolites or other metallosilicate zeolites such as
the aluminophosphosilicates and the aluminophosphates. The
aluminosilicate zeolites are, however, preferred from the viewpoint
of their catalytic activity and stability. Examples of intermediate
pore size aluminosilicate zeolites which may be used are ZSM-5,
ZSM-11 and ZSM-12. The more highly constrained intermediate pore
size zeolites such as ZSM-22, ZSM-23 and ZSM-35 will not normally
be preferred since their constrained pore structure does not allow
the reactants and reaction products to access or to leave the
internal pore structure of the zeolite. The large pore size
zeolites such as zeolite Y (usually as USY), ZSM-4, ZSM-20 may also
be used. Zeolite beta with its characteristic properties similar to
both the large and intermediate pore size zeolites is also an
appropriate selection.
[0052] A highly favored class of intermediate pore size zeolites
are those of the MWW type. The MWW family of zeolite materials has
achieved recognition as having a characteristic framework structure
which presents unique and interesting catalytic properties. The MWW
topology consists of two independent pore systems: a sinusoidal
ten-member ring [10 MR] two dimensional channel separated from each
other by a second, two dimensional pore system comprised of 12 MR
super cages connected to each other through 10 MR windows. The
crystal system of the MWW framework is hexagonal and the molecules
diffuse along the [100] directions in the zeolite, i.e., there is
no communication along the c direction between the pores. In the
hexagonal plate-like crystals of the MWW type zeolites, the
crystals are formed of relatively small number of units along the c
direction as a result of which, much of the catalytic activity is
due to active sites located on the external surface of the crystals
in the form of the cup-shaped cavities. In the interior structure
of certain members of the family such as MCM-22, the cup-shaped
cavities combine together to form a supercage. The MCM-22 family of
zeolites has attracted significant scientific attention since its
initial announcement by Leonovicz et al. in Science 264, 1910-1913
[1994] and the later recognition that the family is currently known
to include a number of zeolitic materials such as PSH 3, MCM-22,
MCM 49, MCM 56, SSZ 25, ERB-1, ITQ-1, and others: Lobo et al. AlChE
Annual Meeting 1999, Paper 292J.
[0053] The relationship between the various members of the MCM-22
family have been described in a number of publications. Three
significant members of the family are MCM-22, MCM-36, MCM-49, and
MCM-56. When initially synthesized from a mixture including sources
of silica, alumina, sodium, and hexamethylene imine as an organic
template, the initial product will be MCM-22 precursor or MCM-56,
depending upon the silica:alumina ratio of the initial synthesis
mixture. At silica:alumina ratios greater than 20, MCM-22 precursor
comprising H-bonded vertically aligned layers is produced whereas
randomly oriented, non-bonded layers of MC-56 are produced at lower
silica:alumina ratios. Both these materials may be converted to a
swollen material by the use of a pillaring agent and on
calcination, this leads to the laminar, pillared structure of
MCM-36. The as-synthesized MCM-22 precursor can be converted
directly by calcination to MCM-22 which is identical to calcined
MCM-49, an intermediate product obtained by the crystallization of
the randomly oriented, as-synthesized MCM-56. In MCM-49, the layers
are covalently bonded with an interlaminar spacing slightly greater
than that found in the calcined MCM-22/MCM 49 materials. The
unsynthesized MCM-56 may be calcined itself to form calcined MCM 56
which is distinct from calcined MCM-22/MCM-49 in having a randomly
oriented rather than a laminar structure. In the patent literature
MCM-22 is described in U.S. Pat. No. 4,954,325 as well as in U.S.
Pat. Nos. 5,250,777; 5,284,643 and 5,382,742. MCM-49 is described
in U.S. Pat. No. 5,236,575; MCM-36 in U.S. Pat. No. 5,229,341 and
MCM-56 in U.S. Pat. No. 5,362,697.
[0054] The preferred zeolitic material for use in the catalyst of
the present process is MCM-22 although zeolite MCM-49 may be found
to have certain advantages relative to MCM-22. It has been found
that the MCM-22 may be either used fresh, that is, not having been
previously used as a catalyst or alternatively, regenerated MCM-22
may be used. Regenerated MCM-22 may be used after it has been used
in any of the catalytic processes for which it is suitable,
including the present process in which the catalyst has shown
itself remain active after even multiple regenerations. It may also
be possible to use MWW catalysts which have previously been used in
other commercial processes and for which they are no longer
acceptable, for example, MCM-22 catalyst which has previously been
used for the production of aromatics such as ethylbenzene or
cumene, normally using reactions such as alkylation and
transalkylation. The cumene production (alkylation) process is
described in U.S. Pat. No. 4,992,606 (Kushnerick et al).
Ethylbenzene production processes are described in U.S. Pat. No.
3,751,504 (Keown); U.S. Pat. No. 4,547,605 (Kresge); and U.S. Pat.
No. 4,016,218 (Haag); U.S. Pat. Nos. 4,962,256; 4,992,606;
4,954,663; 5,001,295; and 5,043,501 describe alkylation of aromatic
compounds with various alkylating agents over catalysts comprising
MWW zeolites such as PSH-3 or MCM-22. U.S. Pat. No. 5,334,795
describes the liquid phase synthesis of ethylbenzene with MCM-22.
As noted above, MCM-22 catalysts may be regenerated after catalytic
use in these processes and other aromatics production processes by
conventional air oxidation techniques similar to those used with
other zeolite catalysts. Conventional air oxidation techniques are
also suitable when regenerating the catalysts after use in the
present process.
[0055] In addition to the MWW active component, the catalysts for
use in the present process will often contain a matrix material or
binder in order to give adequate strength to the catalyst as well
as to provide the desired porosity characteristics in the catalyst.
High activity catalysts may, however, be formulated in the
binder-free form by the use of suitable extrusion techniques, for
example, as described in U.S. Pat. No. 4,908,120. When used, matrix
materials suitably include alumina, silica, silica alumina,
titania, zirconia, and other inorganic oxide materials commonly
used in the formulation of molecular sieve catalysts. For use in
the present process, the level of MCM-22 in a finished matrixed
catalyst will be typically from 20 to 70% by weight, and in most
cases from 25 to 65% by weight. In manufacture of a matrixed
catalyst, the active ingredient will typically be mulled with the
matrix material using an aqueous suspension of the catalyst and
matrix, after which the active component and the matrix are
extruded into the desired shape, for example, cylinders, hollow
cylinders, trilobe, quadlobe, etc. A binder material such as clay
may be added during the mulling in order to facilitate extrusion,
increase the strength of the final catalytic material and to confer
other desirable solid state properties. The amount of clay will not
normally exceed 10% by weight of the total finished catalyst.
Self-bound catalysts (alternatively referred to as unbound or
binder-free catalysts), that is, catalysts which do not contain a
separately added matrix or binder material, are useful and may be
produced by the extrusion method described in U.S. Pat. No.
4,582,815, to which reference is made for a description of the
method and of the extruded products obtained by its use. The method
described there enables extrudates having high constraining
strength to be produced on conventional extrusion equipment and
accordingly, the method is eminently suitable for producing the
high activity self-bound catalysts. The catalysts are produced by
mulling the zeolite, as described in U.S. Pat. No. 4,582,815, with
water to a solids level of 25 to 75 wt % in the presence of 0.25 to
10 wt % of basic material such as sodium hydroxide. Further details
are to be found in U.S. Pat. No. 4,582,815. Generally, the
self-bound catalysts can be characterized as particulate catalysts
in the form, for instance, of extrudates or pellets, containing at
least 90 wt. pct., usually at least 95 wt. pct., of the active
zeolite component with no separately added binder material e.g.
alumina, silica-alumina, silica, titania, zirconia etc. Extrudates
may be in the conventional shapes such as cylinders, hollow
cylinders, trilobe, quadlobe, flat platelets etc.
[0056] As noted above, MCM-22 and other catalysts of this family
may be regenerated after catalytic use for example, in the present
process or in the cumene, ethylbenzene and other aromatics
production processes, with the regeneration carried out by
conventional air oxidation techniques similar to those used with
other zeolite catalysts. Regeneration of the catalyst after use in
the present process results in only a modest activity loss, with
the catalyst maintaining more than 95% of fresh activity after the
first regeneration. Even after multiple regenerations, a reasonable
and acceptable level of activity is retained. The catalyst has been
found to maintain more than 80% of fresh activity after 6
regenerations. Following the air oxidation, the catalyst may be
reconditioned by aqueous reconditioning treatment using water or a
mildly alkaline solution, for example, a dilute solution of ammonia
or sodium carbonate. Treatment with water alone at ambient
temperatures has been found to be effective: the air-regenerated
catalyst is cooled and then immersed in a water bath after which it
is dried and returned to service. The reconditioning treatment may
be continued for the empirically determined time which results in
an improvement in catalyst properties. It is theorized that the
reconditioning treatment enables the silanol groups on the surface
of the zeolite to be re-formed after the regeneration treatment
with a consequent restoration of catalytic properties which, in
favorable cases, may provide a catalyst almost comparable to a
fresh catalyst.
[0057] A guard bed may be used ahead of the selective hydrofining
reactor and, if used, may conveniently be the same catalyst used in
the alkylation reactor as a matter of operating convenience but
this is not required: if desired another catalyst or sorbent to
remove contaminants from the feed may used, typically a cheaper
guard bed sorbent, e.g. a used catalyst from another process or
alumina. The objective of the guard bed is to remove the
contaminants from the feed before the feed comes to the hydrofining
catalyst and provided that this is achieved, there is wide variety
of choice as to guard bed catalysts and conditions useful to this
end. The volume of the guard bed will normally not exceed about 20%
of the total catalyst bed volume of the unit.
Product Recovery and Treatment
[0058] After treatment in the alkylation step, the naphtha is
treated for recovery of the product gasoline. Stabilization to
remove light ends formed in the processing is typical as well as
fractionation to separate the gasoline from heavier fractions
formed in the alkylation step and any higher boiling sulfur
compounds formed in the sweetening step (if used). At this time,
the product specifications need to be observed in order to obtain
proper flash point, boiling point and other specifications.
Experimental Work
[0059] Two types of FCC gasoline feeds were used in an
investigation of the potential of zeolites for alkylating sulfur
species, a C.sub.5+ full range FCC gasoline and C.sub.5 to
190.degree. C. (375.degree. F.) gasoline. Different samples of each
type of gasoline feed were employed in the tests but differences
between the samples for a given feed type were minimal and of no
consequence. In general, 85 to 90 wt. % of the full range FCC
gasoline boils below 190.degree. C. (375.degree. F.); the sulfur
content was 2000 to 2500 ppm, 40-50% of which is in the 190.degree.
C.- (375.degree. F.-) portion of the feed. For the C.sub.5 to
190.degree. C. (375.degree. F.) FCC gasoline, more than 95 wt %
boils below 190.degree. C. (375.degree. F.); these gasolines
contained from 900-1300 ppm of sulfur, 85-95% of which is in the
190.degree. C.- (375.degree. F.-) portion of the feed.
[0060] The examples below illustrate the effectiveness of the
catalytic treatments for shifting sulfur species from lighter,
olefin-rich portions of the FCC gasoline to the heavy, olefin-lean
gasoline or to the kero/light diesel range. Separation of the
portion containing the sulfur compounds followed by catalytic
hydrodesulfurization or some other means (e.g. sorption) can
therefore remove the sulfur species without octane loss. Thus, very
high levels of desulfurization can be achieved with little impact
on overall gasoline yield and octane.
EXAMPLE 1
[0061] 1 part by weight of 100% H-MCM-22 (preparation below) was
placed in a vial and contacted at room temperature with 3 parts by
weight C.sub.5+ full range gasoline. The contents of the vial were
allowed to equilibrate over a 2-3 days, with periodic manual
agitation. The supernatant liquid was then removed and subjected to
GC analyses for boiling range distribution and sulfur content and
distribution. Around 9% of the feed boiling below 190.degree. C.
(375.degree. F.) was converted to the 190.degree. C.+ (375.degree.
F.+) portion and 91% of the sulfur was removed from the 190.degree.
C.- (375.degree. F.-) portion of the feed.
EXAMPLE 2
[0062] The same experiment as described in Example 1 was performed,
except a C.sub.5- 190.degree. C. (C.sub.5- 375.degree. F.) FCC
gasoline cut was employed. The 910 ppm S in this feed is very
similar to the amount of 190.degree. C.- (375.degree. F.-) sulfur
in the full range gasoline used in Example 1. Once again a high
percentage of the light sulfur was shifted to heavy sulfur. Around
12% of the feed boiling below 190.degree. C. (375.degree. F.) was
moved to the 190.degree. C.+ (375.degree. F.+) portion.
EXAMPLE 3
[0063] An H-MCM-22 powder (10 g) was placed in a 600 ml autoclave
and dehydrated under a slow, continuous N.sub.2 flow at 200.degree.
C. After 16 hours, the nitrogen flow was stopped, all inlets and
outlets of the autoclave were closed, and the dried H-MCM-22 was
cooled to 120.degree. C. under the static N.sub.2 atmosphere. At
120.degree. C., a light (C.sub.5- 190 C) FCC gasoline feed (100 g)
was added via an ISCO.TM. pump to the autoclave over a 5 minute
interval. After the gasoline addition, the autoclave's pressure was
adjusted to 2070 kPag (300 psig) with N.sub.2. The autoclave was
then isolated from the gasoline source and held at 120.degree. C.
and 2070 kPag for 5 hr. Thereafter, the autoclave was cooled to
ambient temperature and vented to ambient pressure slowly. A
portion of the product was quickly sampled, passed through a
syringe filter to remove catalyst particles, and subjected to
sulfur and simulated distillation analysis. Boiling range
conversion of 190.degree. C.- (375.degree. F.-) to the 190.degree.
C.+ (375.degree. F.+) portion is less than 2% at 55-60% sulfur
removal from the 190.degree. C.- (375.degree. F.-) portion of the
feed.
[0064] The significant degree of sulfur removal from the
190.degree. C.- (375.degree. F.-) gasoline under markedly different
conditions with respect to the above examples (lower catalyst to
oil ratio, shorter contact time, higher temperature) illustrates
the effectiveness of the MCM-22 catalyst for accomplishing this
chemistry.
EXAMPLE 4
[0065] The same amount of the light FCC gasoline was treated with
10 g of USY zeolite (100% H-form, preparation below), using the
procedure and conditions described in Example 3. Although the
sulfur removal from the 190.degree. C.- (375.degree. F.-) portion
of the feed is smaller with respect to the MCM-22 catalyst under
the same conditions, the results showed that USY was nonetheless
able to shift the sulfur out of the 190.degree. C.- (375.degree.
F.)- gasoline. Boiling range conversion of 190.degree. C.-
(375.degree. F.-) gasoline to 190.degree. C.+ (375.degree. F.+)
gasoline is less than 1% with 20-25% of the sulfur removed from the
190.degree. C.- (375.degree. F.-) portion of the feed.
EXAMPLE 5
[0066] The same amount of the light FCC gasoline was treated with
10 g zeolite Beta (100% H-form, preparation below), using the
procedure and conditions described in the Example 3. Sulfur removal
from the 190.degree. C.- (375.degree. F.-) portion of the feed is
roughly similar to that observed with MCM-22 under the same
conditions. Boiling range conversion of 190.degree. C.-
(375.degree. F.-) to 190.degree. C.+ (375.degree. F.+) portion is
around 5% at 50-55% sulfur removal from the 190.degree. C.-
(375.degree. F.-) portion of the feed. In conjunction with the
above examples, this example illustrates that a wide variety of
aluminosilicate molecular sieves can be effective for shifting
sulfur into heavier portions of the gasoline.
EXAMPLE 6
[0067] 1 part by weight of 100% USY (previously reduced with
H.sub.2) was contacted at room temperature with 3 parts by weight
of the full range C.sub.5+ gasoline. The contents of the vial were
allowed to equilibrate over 16-24 hours, with periodic manual
agitation. The supernatant liquid was then removed and subjected to
GC analyses for boiling range distribution and sulfur content and
distribution. Almost 80% of the sulfur was removed from the
190.degree. C.- (375.degree. F.-) portion of the feed. This
example, together with example 4 above, illustrates that zeolite
performance can be enhanced by modifying operating conditions.
EXAMPLE 7
[0068] The product from Example 1 was subjected to a compositional
analysis by mass spectrometry and the results were used to
calculate the octane properties of the treated gasoline. Compared
to results for the feed based on the same type of analysis, the
product road octane (i.e. (R+M)/2) showed only a modest change
(86.0 for the treated gasoline vs. 87.4 for the feed). This example
illustrates that high conversions of 190.degree. C.- (375.degree.
F.-) sulfur to 190.degree. C.+ (375.degree. F.+) sulfur can be
achieved without greatly affecting the gasoline octane quality.
EXAMPLE 8
[0069] The effectiveness of the alkylation step for removing sulfur
from the gasoline fraction was demonstrated using a narrow cut
intermediate (C.sub.6-C.sub.8) catalytic naphtha fraction
containing 49% olefins, 12% aromatics, 360 wppm sulfur. The
nitrogen content was reduced to about 1 wppm by treatment with an
ion exchange resin (Amberlyst.TM.) and alumina. This fraction was
passed without treat gas over an MCM-49 catalyst diluted 4:1 with
inerts, in an upflow reactor. The pressure (total system, gauge)
was held at 6200 kPag (900 psig) and space velocity at 5 hr.sup.-1
v/v; the temperature was varied upwards during the course of the 11
day run during which 2 mass balances were taken each day. A
177.degree. C.- (350.degree. F.-) product fraction was taken and
the sulfur conversion from this fraction determined. FIG. 2 shows
that the sulfur conversion out of the product fraction increases
with temperature and that the total hydrocarbon yield decreases.
Similar results were obtained with a 204.degree. C.- (400.degree.
F.-) fraction. FIG. 3 shows the hydrocarbon and boiling point
shifts resulting from the treatment at a temperature of 185.degree.
C. (365.degree. F.) attained at end-of-run when Mass Balance 22 was
taken after 11 days of operation; it shows that while the shift in
boiling point of the hydrocarbons is relatively limited (compare
the boiling point shift between lines "Feed FID" and "Balance 22
FID"), a significantly greater shift in the boiling points of the
sulfur species is obtained (compare the boiling point shift between
lines "Feed SCD" and "Balance 22 SCD"), demonstrating that a
subsequent fractionation will be readily capable of separating the
alkylated sulfur compounds from the hydrocarbon components.
Preparation of H-MCM-22 Crystal Sample
[0070] A Na-form MCM-22 with a silica-to-alumina molar ratio of 24
was used for this catalyst preparation. The Na-form MCM-22 was
calcined at 482.degree. C. for 3 hours under N.sub.2 atmosphere.
Then the crystals were ammonium exchanged with 1 M ammonium nitrate
solution (5 cc/g zeolite) two times at the natural pH and washed
with deionized water. Then the wet MCM-22 zeolite was dried in an
oven at 120.degree. C. overnight. The material was air calcined (5
v/v/min) in a tray for 6 hours at 538.degree. C.
Preparation of Beta Crystal Sample
[0071] A commercial Na-form Beta with a silica-to-alumina molar
ratio of 35 was used for this catalyst preparation. The Na-form
Beta was ammonium exchanged with 1 M ammonium nitrate solution two
times at the natural pH and washed with deionized water. Then wet
Beta zeolite was dried in an oven at 120.degree. C. overnight. The
material was calcined in a tray under N.sub.2 at 482.degree. C. for
3 hours, then under air (5 v/v/min) for 6 hours at 538.degree.
C.
Preparation of H USY Crystal Sample (Y-30681)
[0072] A commercial Na-form USY with a silica-to-alumina ratio of
5.5 and a unit cell size of 24.54 .ANG. was used for this catalyst
preparation. The Na-form USY was made a slurry with deionized water
to target a 35 wt % solids level. A solution of 30 wt % ammonium
sulfate was prepared, and then the pH was adjusted to 3.0 using 20
wt % H.sub.2SO.sub.4 solution. The pH 3.0 ammonium sulfate solution
was added slowly to the USY slurry (1.3 g of 30% aluminum sulfate
solution per 1 g zeolite) while the overall solution pH was
adjusted to 3.0. The mixture slurry solution was stirred for
.about.30 minutes while the pH was adjusted to 3.0. The exchanged
USY zeolite was filtered and washed with deionized water (10 cc/g
zeolite). Then the USY zeolite was ammonium exchanged one more time
by pouring pH 3.0 ammonium sulfate solution (1.3 g solution per 1 g
USY) over the USY zeolite on the filter. The exchanged USY zeolite
was washed with deionized water (10 cc/g zeolite) and then dried in
an oven at 120.degree. C. overnight. The material was air calcined
(5 cc air/g zeolite/min) for 3 hours at 538.degree. C.
* * * * *