U.S. patent application number 09/944511 was filed with the patent office on 2004-07-22 for process for upgrading fcc product with additional reactor.
Invention is credited to Lomas, David A..
Application Number | 20040140246 09/944511 |
Document ID | / |
Family ID | 32713942 |
Filed Date | 2004-07-22 |
United States Patent
Application |
20040140246 |
Kind Code |
A1 |
Lomas, David A. |
July 22, 2004 |
PROCESS FOR UPGRADING FCC PRODUCT WITH ADDITIONAL REACTOR
Abstract
A process is disclosed for taking a cut from an FCC reactor
product and reacting it in a separate reactor to upgrade the
product quality Cracking or reformulating reactions in the separate
reactor give reductions in olefins and reformulating
hydrogen-transfer reactions convert undesirable olefins to
isoparaffins and aromatics without reducing octane value Catalyst
particles from the FCC reactor may be cycled to the separate
reactor This process has also been found to substantially diminish
concentrations of nitrogen and sulfur compounds fed to the separate
reactor.
Inventors: |
Lomas, David A.;
(Barrington, IL) |
Correspondence
Address: |
JOHN G TOLOMEI, PATENT DEPARTMENT
UOP LLC
25 EAST ALGONQUIN ROAD
P O BOX 5017
DES PLAINES
IL
60017-5017
US
|
Family ID: |
32713942 |
Appl. No.: |
09/944511 |
Filed: |
August 31, 2001 |
Current U.S.
Class: |
208/134 ;
208/69 |
Current CPC
Class: |
C10G 35/04 20130101 |
Class at
Publication: |
208/134 ;
208/069 |
International
Class: |
C10G 035/04; C10G
057/00 |
Claims
What is claimed is:
1. A process for converting a hydrocarbon feed stream comprising:
passing a reformulation feed stream including saturated and
olefinic hydrocarbons with carbon numbers of 5-8 to a reformulating
reactor containing catalyst particles having a composition;
reformulating said reformulation feed stream in said reformulating
reactor to produce a reformulated product stream, said
reformulating proceeding at conditions that promote at least a 5%
net yield increase in aromatics on a fresh reformulation feed basis
indicating the occurrence of hydrogen transfer reactions; and
recovering said reformulated product stream.
2. The process of claim 1 wherein said reformulation feed stream is
prepared by: cracking a preliminary cracking feed stream with
catalyst particles in a cracking reactor to produce a cracked
product, said catalyst particles in said cracking reactor having a
same composition as the catalyst particles in said reformulating
reactor; separating said cracked product from said catalyst
particles in a separator vessel to obtain a cracked product stream;
and recovering at least a portion of said cracked product stream to
be said reformulation feed stream.
3. The process of claim 2 further including isolating said
reformulated product stream from said cracked product stream.
4. The process of claim 2 further comprising the step of cycling
catalyst particles that had previously resided in said cracking
reactor to said reformulating reactor.
5. The process of claim 1 wherein a greater proportion of
hydrocarbons with carbon numbers of 5-8 undergo hydrogen transfer
reaction than cracking reaction.
6. The process of claim 1 wherein olefins in said reformulation
feed stream convert to isoparaffins in the reformulating
reactor.
7. The process of claim 1 wherein the concentration of sulfur
compounds in the reformulated product stream is less than its
concentration in the reformulation feed stream.
8. The process of claim 1 wherein the concentration of nitrogen
compounds in the reformulated product stream is less than its
concentration in the reformulation feed stream.
9. The process of claim 1 wherein the reformulation feed stream has
an initial boiling point below about 200.degree. C. (392.degree.
F.).
10. A process for converting a hydrocarbon feed stream comprising:
contacting said hydrocarbon feed stream with catalyst particles
having a composition in a first reactor to produce a cracked
product; separating said cracked product from said catalyst
particles in a vessel to obtain a cracked product stream;
recovering a naphtha stream from said cracked product stream, said
naphtha stream having an initial boiling point below 127.degree. C.
(260.degree. F.); contacting said naphtha stream with catalyst
particles having said composition in a second reactor to produce an
upgraded product stream; and recovering said upgraded product
stream and isolating said upgraded product stream from said cracked
product stream.
11. The process of claim 10 wherein hydrogen transfer reactions
predominate over cracking reactions in the second reactor
12. The process of claim 10 wherein olefins convert to aromatics in
the second reactor.
13. The process of claim 10 wherein olefins convert to isoparaffins
in the secondary reactor.
14. The process of claim 10 wherein the concentration of sulfur
compounds in the upgraded product stream is 50% less than its
concentration in the naphtha stream.
15. The process of claim 10 wherein said naphtha stream has an end
point below 230.degree. C. (446.degree. F.).
16. The process of claim 10 wherein said catalyst particles in said
second reactor previously resided in the first reactor.
17. A process for converting a hydrocarbon feed stream comprising:
contacting said hydrocarbon feed stream with catalyst particles
having a composition in a first reactor to produce a cracked
product; separating said cracked product from said catalyst
particles in a vessel to obtain a cracked product stream;
recovering an oil stream from said cracked product stream having an
initial boiling point above about 200.degree. C. (392.degree. F.);
cycling catalyst particles that had resided in said first reactor
to a second reactor, said second reactor being discrete from said
vessel; contacting said oil stream with catalyst particles in a
second reactor to produce an upgraded product stream; and
recovering said upgraded product stream and isolating said upgraded
product stream from said cracked product stream.
18. The process of claim 17 further comprising the step of
hydrotreating said oil stream.
19. The process of claim 17 wherein no hydrogen is added to the
second reactor.
20. The process of claim 17 wherein the end point of said oil
stream is below about 288.degree. C. (550.degree. F.).
Description
BACKGROUND OF THE INVENTION
[0001] This invention relates generally to processes for the
fluidized catalytic cracking (FCC) of heavy hydrocarbon streams.
More specifically, this invention relates generally to processes
for upgrading catalytically cracked hydrocarbon feeds in a discrete
reactor vessel.
DESCRIPTION OF THE PRIOR ART
[0002] The FCC process is carried out by contacting the starting
material whether it be vacuum gas oil, reduced crude, or another
source of relatively high boiling hydrocarbons with a catalyst made
up of finely divided or particulate solid material. The catalyst is
transported in a fluid-like manner by passing gas or vapor through
it at sufficient velocity to produce a desired regime of fluid
transport. Contact of the oil with the fluidized material catalyzes
the cracking reaction. The cracking reaction deposits coke on the
catalyst. Coke is comprised of hydrogen and carbon and can include
other materials in trace quantities such as sulfur and metals that
enter the process with the starting material. Coke interferes with
the catalytic activity of the catalyst by blocking active sites on
the catalyst surface where the cracking reactions take place.
Catalyst is traditionally transferred from a stripper, that removes
adsorbed hydrocarbons and gases from catalyst, to a regenerator for
purposes of removing the coke by oxidation with an
oxygen-containing gas. An inventory of catalyst having a reduced
coke content, relative to the catalyst in the stripper, hereinafter
referred to as regenerated catalyst, is collected for return to the
reaction zone. Oxidizing the coke from the catalyst surface
releases a large amount of heat, a portion of which escapes the
regenerator with gaseous products of coke oxidation generally
referred to as flue gas. The balance of the heat leaves the
regenerator with the regenerated catalyst. The fluidized catalyst
is continuously circulated from the reaction zone to the
regeneration zone and then again to the reaction zone. The
fluidized catalyst, as well as providing a catalytic function, acts
as a vehicle for the transfer of heat from zone to zone. Catalyst
exiting the reaction zone is spoken of as being spent, i.e.,
partially deactivated by the deposition of coke upon the catalyst.
The FCC processes, as well as separation devices used therein are
fully described in U.S. Pat. No. 5,584,985 B1 and U.S. Pat. No.
4,792,437 B1, the contents of which are hereby incorporated by
reference. Specific details of the various contact zones,
regeneration zones, and stripping zones along with arrangements for
conveying the catalyst between the various zones are well known to
those skilled in the art.
[0003] The FCC reactor cracks gas oil or heavier feeds into a broad
range of products. Cracked vapors from the FCC unit enter a
separation zone, typically in the form of a main column, that
provides a gas stream, a gasoline cut, light cycle oil (LCO) and
clarified oil (CO) which includes heavy cycle oil (HCO) components.
The gasoline cut may include light, medium and heavy gasoline
components. A major component of the heavy gasoline fraction
comprises condensed single ring aromatics. A major component of LCO
is condensed bicyclic ring aromatics.
[0004] Subjecting product fractions to additional reactions is
useful for upgrading product quality. The recracking of heavy
product fractions from the initially cracked FCC product is one
example. Typically, in recracking, uncracked effluent from a first
riser of an FCC reactor is recontacted with catalyst at a second
location to cleave larger molecules down into more useful smaller
molecules. For example, U.S. Pat. No. 4,051,013 B1 discloses
cracking both gasoline-range feed and gas oil feed in the same
riser at different elevations. U.S. Pat. No. 3,161,582 B1, U.S.
Pat. No. 5,176,815 B1 and U.S. Pat. No. 5,310,477 B1 all disclose
cracking a primary hydrocarbon feed in a riser of an FCC unit and
cracking a secondary hydrocarbon feed in a reactor into which the
riser exits. As a result, both cracked products mix in the reactor,
to some extent, which could negate the incremental upgrade
resulting from cracking the secondary hydrocarbon feed,
particularly if it is a fraction of the cracked primary hydrocarbon
feed.
[0005] FCC units employing two risers are known. U.S. Pat. No.
5,198,590 B1, U.S. Pat. No. 4,402,913 B1, U.S. Pat. No. 4,310,489
B1, U.S. Pat. No. 4,297,203 B1, U.S. Pat. No. 3,799,864 B1, U.S.
Pat. No. 3,748,251 B1, U.S. Pat. No. 3,714,024 B1 and WO 00/40672
disclose two riser FCC units in which feeds are predominantly
cracked in both risers In these patents, both risers communicate
with the same recovery conduit and/or reactor permitting
commingling of gaseous products. In U.S. Pat. No. 5,730,859 B1, all
of the effluent from one riser is fed to the other riser, without
first undergoing a product separation. U.S. Pat. No. 4,172,812 B1
teaches recracking all or a part of cracked product from a riser of
an FCC unit over a catalyst having a composition that is different
from the catalyst composition in the riser.
[0006] In U.S. Pat. No. 5,944,982 B1, although both risers
terminate in the same reactor vessel, gaseous products from each
riser are isolated from the other. This patent also discusses the
method of cracking a LCO fraction of a cracked product containing
very refractory, bicyclic aromatic components. Bicyclic aromatics
are very difficult to crack and boil only at high temperatures. The
presence of high boiling point bicyclic aromatics can cause a
gasoline pool to exceed maximum volatility standards. By a process
called J-cracking, LCO is hydrotreated to partially saturate the
bicyclic aromatic hydrocarbons such as naphthalene to produce
tetralin. The tetralin is then cracked to make benzene, toluene,
xylene and isoparaffins along with some naphthalene. The "J" in
J-cracking is a measure of unsaturation of the hydrocarbons having
the general formula:
C.sup.NH.sub.2N-J.
[0007] U.S. Pat. No. 3,928,172 B1 teaches an FCC unit with a
secondary dense fluidized catalyst bed in a separate reactor. Gas
oil is cracked in a riser of the FCC unit with unregenerated
catalyst from the separate dense fluidized catalyst bed. A heavy
naphtha fraction of the cracked gas oil, boiling between
127.degree. and 232.degree. C. (2600 and 450.degree. F.), from the
riser is recracked in the separate reactor over regenerated
catalyst. Apparently, the benefit of cracking lower boiling
fractions was not explored, presumably because the octane rating of
the lower boiling fraction was sufficiently high or because it was
predicted not to be effective. The data in the patent indicates
that nominal, if any, reformulation reactions occur in the separate
reactor because little, if any, new aromatics are produced.
[0008] It is also known to subject cracked product from a riser of
an FCC unit to a subsequent oligomerization reaction in a separate
reactor. In oligomerization, smaller olefins are bonded together to
make larger olefins of greater molecular weight. U.S. Pat. No.
5,009,851 B1 and U.S. Pat. No. 4,865,718 B1 disclose oligomerizing
a fraction of cracked product from an FCC unit in a separate
reactor.
[0009] Cracking and oligomerization differ from reformulation in
that the former involve decreases or increases in carbon numbers,
respectively. Whereas, in reformulation, carbon numbers are not
changed but hydrogen atoms are exchanged to alter the structure of
the molecule and make it more valuable.
[0010] In gasoline production, many governmental entities are
restricting the concentration of olefins allowed in the gasoline
pool. Reducing olefin concentration without also reducing value is
difficult because higher olefin concentrations typically promote
higher Research Octane Numbers (RON) and Motor Octane Numbers
(MON), but the latter to a lesser extent. Octane value or Road
Octane Number is the average of RON and MON. Merely saturating
olefins typically yields normal paraffins which typically have low
octane value. Additionally, saturation requires the addition of
hydrogen, which is expensive and in some regions, difficult to
obtain.
[0011] Feedstocks for FCC units typically include sulfur and
nitrogen. During FCC operation, the sulfur and nitrogen are
converted primarily to hydrogen sulfide and ammonia, which are
easily removed, but are also converted to organic sulfurs,
mercaptans and nitrogen oxides. Stricter environmental limits on
sulfur and nitrogen compound emissions along with lower sulfur
specifications for fuel products have raised interest in the need
to remove nitrogen and sulfur compounds from FCC gasoline. As
demands for cleaner fuels and use of high sulfur and high nitrogen
feedstocks increase, the need for sulfur and nitrogen removal from
FCC gasoline will become even greater.
[0012] It is an object of the present invention to provide a method
for enhancing the quality of product from an FCC unit. It is a
further object of the present invention to reduce the olefinicity
of product from an FCC unit without substantially reducing the
octane rating of the product and without the addition of hydrogen.
It is an even further object of the present invention to reduce the
concentration of sulfur and nitrogen compounds in an FCC
product.
BRIEF SUMMARY OF THE INVENTION
[0013] It has now been discovered that a separate reactor can be
used to either reformulate or crack a product fraction from an FCC
unit to reduce its olefinicity and maintain or boost its octane
rating without the separate addition of hydrogen. If the separate
reactor is incorporated into an FCC unit, catalyst can be
circulated between the FCC reactor and the separate reactor.
Additionally, it has been further found that higher boiling point
fractions from an FCC unit can be hydrotreated and sent to a
separate reactor, if incorporated in the FCC unit using catalyst
cycled through the FCC unit, to crack FCC product fractions down to
lower boiling point useful hydrocarbon components. Furthermore,
reacting fractions of FCC product in a separate reactor has been
found effective in substantially reducing sulfur and nitrogen
compounds in the fraction.
[0014] Accordingly, in one embodiment, the present invention
relates to a process for converting a hydrocarbon feed stream
comprising passing a reformulation feed stream including saturated
and olefinic hydrocarbons with carbon numbers of 5-8 to a
reformulating reactor. The reformulating reactor contains catalyst
particles having a composition. The reformulation feed stream is
reformulated in the reformulating reactor to produce a reformulated
product stream. The reformulating proceeds at conditions that
promote at least a 5% net yield increase in aromatics on a fresh
reformulation feed basis indicating the occurrence of hydrogen
transfer reactions. The reformulated product stream is then
recovered.
[0015] In another embodiment, the present invention relates to a
process for converting a hydrocarbon feed stream comprising
contacting the hydrocarbon feed stream with catalyst particles
having a composition in a first reactor to produce a cracked
product. The cracked product is separated from the catalyst
particles in a vessel to obtain a cracked product stream. A naphtha
stream is recovered from the cracked product stream. The naphtha
stream has an initial boiling point below 127.degree. C.
(260.degree. F.). The naphtha stream is contacted with catalyst
particles having the composition in a second reactor to produce an
upgraded product stream. The upgraded product stream is recovered
and isolated from the cracked product stream.
[0016] In a further embodiment, the present invention relates to a
process for converting a hydrocarbon feed stream comprising
contacting the hydrocarbon feed stream with catalyst particles
having a composition in a first reactor to produce a cracked
product. The cracked product is separated from the catalyst
particles in a vessel to obtain a cracked product stream. An oil
stream is recovered from the cracked product stream having an
initial boiling point above about 200.degree. C. (392.degree. F.).
Catalyst particles that had resided in the first reactor are cycled
to a second reactor that is discrete from the vessel. The oil
stream is contacted with catalyst particles in a second reactor to
produce an upgraded product stream. The upgraded product stream is
recovered and isolated from the cracked product stream.
[0017] Additional objects, embodiment and details of this invention
can be obtained from the following detailed description of the
invention.
BRIEF DESCRIPTION OF THE DRAWINGS
[0018] FIG. 1 is a sectional, elevational, schematical view of an
FCC unit incorporating a main column and a secondary reactor in
accordance with the present invention.
[0019] FIG. 2 is a sectional, elevational, schematical view of an
alternative embodiment of the present invention
[0020] FIG. 3 is a sectional, elevational, schematical view of a
further embodiment of the present invention.
DETAILED DESCRIPTION OF THE INVENTION
[0021] The present invention may be described with reference to
four components: an FCC reactor 10, a regenerator 50, a secondary
reactor 80, 80', 80" and a main column 100. Although many
configurations of the present invention are possible, three
specific embodiments are presented herdein by way of example. All
other possible embodiments for carrying out the present invention
are considered within the scope of the present invention. For
example, the secondary reactor 80, 80', 80" and/or the main column
100 need not be incorporated into an FCC unit as illustrated in
FIGS. 1-3 but may stand alone.
[0022] In the embodiment of the present invention in FIG. 1, the
FCC reactor 10 comprises a conduit in the form of a reactor riser
12 that extends upwardly through a lower portion of a reactor
vessel 14 as in a typical FCC arrangement. The central conduit or
reactor riser 12 preferably has a vertical orientation within the
reactor vessel 14 and may extend upwardly through the bottom of the
reactor vessel 14 or downwardly from the top of the reactor vessel
14. The reactor riser 12 terminates in a separation vessel 16 at
swirl arms 18. A hydrocarbon feed stream is fed to the riser at a
nozzle 20 which is contacted and vaporized by hot regenerated
catalyst fluidized by a gas such as steam from a nozzle 22. The
catalyst cracks the hydrocarbon feed stream and a mixture of
catalyst particles and gaseous cracked hydrocarbons exit the swirl
arms 18 into the separation vessel 16. Tangential discharge of
gases and catalyst from the swirl arms 18 produces a swirling
helical motion about the interior of the separation vessel 16,
causing heavier catalyst particles to fall into a dense catalyst
bed 24 and a mixture of gaseous cracked hydrocarbons and entrained
catalyst particles to travel up a gas recovery conduit 26 and enter
into cyclones 28. In the cyclones 28, centripetal force imparted to
the mixture induces the heavier entrained catalyst particles to
fall through diplegs 30 of the cyclone 28 and to the bottom of the
separation vessel 16 into a dense catalyst bed 32. The gases in the
cyclones 28 more easily change direction and begin an upward spiral
with the gases ultimately exiting the cyclones 28 through outlet
pipes 34. Cracked gases leave the reactor vessel 14 though an
outlet conduit 36. The cracked gases are optionally subjected to a
further separation (not shown) to further remove any light loading
of catalyst particles and are sent via a line 98 to fractionation
in the main column 100 which will be described later with reference
to all of FIGS. 1-3. Catalyst particles in the dense catalyst bed
32 enter the separation vessel 16 through windows 38 where they
join catalyst particles in the dense catalyst bed 24 in a stripping
section 40 of the separation vessel 16. The catalyst particles are
stripped of entrained cracked vapors over baffles 42 with a
stripping medium such as steam entering from at least one nozzle
44. The stripped cracked vapors travel up to the gas recovery
conduit 26 where they are processed with other cracked product
vapors.
[0023] Stripped catalyst from the stripping section 40 of the FCC
reactor 10 travels through a first stripped catalyst pipe 46
regulated by a control valve 48 and into the regenerator 50 at a
lower chamber 52. In the lower chamber 52, stripped catalyst is
subjected to hot oxygen-containing gas such as air from a
distributor 54. Coke is burned from the catalyst and as the
catalyst is heated, it ascends upwardly in the lower chamber 52 and
is distributed into an upper chamber 55 of the regenerator through
a distributor 56. Regenerated catalyst collects in a dense catalyst
bed 58 whereas entrained catalyst is removed from regenerator
effluent gases in cyclones 60 and 62. Flue gas exits the cyclone 62
through an outlet pipe 64 to exit the regenerator through an outlet
66. Regenerated catalyst from the dense catalyst bed 58 travels
through a regenerated catalyst pipe 68 regulated by a control valve
70 into the reactor riser 12 where it is fluidized and contacted
with fresh feed. Stripped catalyst also exits the stripping section
40 through a second stripped catalyst pipe 72 regulated by a
control valve 74 into a dense catalyst bed 82 in the secondary
reactor 80. The degree to which the control valve 74 is opened can
be automatically controlled to obtain the temperature desired in
the secondary reactor 80. For example, if higher temperature is
desired in the secondary reactor 80, more of the relatively hot
catalyst can be permitted to pass through the control valve 74 to
add heat to the secondary reactor 80. The secondary reactor 80 is
preferably a fluidized bed. However, a riser reactor or other
reactor configuration may be suitable. A partition defines a hopper
section 81 of the secondary reactor 80. Catalyst in the dense
catalyst bed 82 that falls into the hopper section 81 is fluidized
by steam or some other fluidizing media through a distributor 84
and is stripped of entrained gases over baffles 83. A desired cut
of hydrocarbon feed from the FCC reactor 10 and fractionated in the
main column 100 is fed to a secondary reactor 80. The feed to the
secondary reactor 80 from the main column 100 is fed through a
distributor 86 where it is contacted with catalyst in the dense
catalyst bed 82. The distributor 86 distributes feed in such a way
as to fluidize the dense catalyst bed 82. Cyclones 88 and 90 remove
entrained catalyst from a gaseous product which leaves the
secondary reactor 80 through a conduit 92. Catalyst leaves the
secondary reactor 80 after being stripped in the hopper section 81
through a pipe 76 regulated by a control valve 78. The degree to
which the control valve 78 is opened can be automatically
controlled to obtain the level desired in the secondary reactor 80.
The level of the catalyst in the secondary reactor 80 determines
the weight hourly space velocity (WHSV) of reactants through the
secondary reactor 80. For example, if a greater WHSV is desired,
the control valve 78 would be opened relatively more to reduce the
level of catalyst in the dense catalyst bed 82.
[0024] FIG. 2 is an alternative embodiment of the present invention
in which regenerated catalyst is fed to the secondary reactor 80'.
In FIG. 2, the elements of the FCC reactor 10 and the regenerator
50 have generally the same configuration as in FIG. 1. Elements in
FIG. 2 with different configurations from FIG. 1, such as in the
secondary reactor 80', will be distinguished by adding a "'" symbol
to the reference numeral. Hydrocarbon feed processed in the FCC
reactor 10 is recovered at the outlet conduit 36 and is carried by
the line 98 to be fractionated in the main column 100, perhaps
after interim processing, to obtain a desired cut to be fed to the
secondary reactor 80'. The feed to the secondary reactor 80' is fed
by a fluidizing nozzle 85 to be contacted in a riser 86' with
regenerated catalyst from a regenerated catalyst pipe 68' regulated
by a control valve 70'. Both feed and catalyst are distributed by
the riser 86' into a dense catalyst bed 82' which is fluidized by
the feed from the riser 86'. Products exit the secondary reactor
80' out a conduit 92' after entrained catalyst is removed in
cyclones 88' and 90'. A partition defines a hopper section 81' of
the secondary reactor 80'. Catalyst from the dense catalyst bed 82'
in the secondary reactor 80' that falls into the hopper section 81'
is fluidized with a medium such as steam from a distributor 84' and
is stripped of entrained product gases over baffles 83'. Stripped
catalyst passes through a pipe 76' regulated by a control valve 78'
to the reactor riser 12 where it contacts the primary hydrocarbon
feed stream injected by the nozzle 20. Stripped catalyst from the
stripping section 40 of the FCC reactor 10 passes through a
stripped catalyst pipe 46' regulated by a control valve 48' into
the lower chamber 52 of the regenerator 50 where coke deposits are
burned from catalyst by means of a hot oxygen-containing gas such
as air. Regenerated catalyst from the upper chamber 55 passes
through the regenerated catalyst pipe 68' and is regulated by the
control valve 70' before it enters the riser 86' of the secondary
reactor 80'. All other elements in FIG. 2 have generally the same
function as in FIG. 1.
[0025] FIG. 3 shows another embodiment of an FCC unit utilizing a
secondary reactor 80" which receives catalyst from and returns
catalyst to the regenerator 50. Again, because the FCC reactor 10
and the regenerator 50 are both very similar to those depicted in
FIG. 1, all of their elements in both drawings will retain the same
reference numerals. However, those elements in FIG. 3 that differ
from the corresponding elements in FIG. 1 will be distinguished by
adding a """ symbol to the reference numeral. Primary hydrocarbon
feed is fed to the reactor riser 12 by means of the nozzle 20. The
primary feed is contacted with regenerated catalyst and cracked to
yield product that is withdrawn from the FCC reactor 10 via the
outlet conduit 36. Catalyst separated from the cracked product is
stripped in the stripping section 40 and passed through a stripped
catalyst pipe 46" regulated by a control valve 48" into the lower
chamber 52 of the regenerator 50. Regenerated catalyst from the
upper chamber 55 of the regenerator 50 is distributed to the
reactor riser 12 through a first regenerated catalyst pipe 68"
regulated by a control valve 70" where it contacts fresh primary
feed and is also distributed through a second regenerated catalyst
pipe 72" regulated by a control valve 74" to the secondary reactor
80". The gaseous vapor effluent from the FCC reactor 10 is carried
from the outlet conduit 36 through the line 98, perhaps to further
processing and then to the main column 100 to be fractionated. A
desired fraction is fed to the secondary reactor 80" through a
distributor 86" which fluidizes a dense catalyst bed 82" with a
medium such as steam. The feed contacts regenerated catalyst in the
dense catalyst bed 82". A partition defines a hopper section 81" in
the secondary reactor 80". Catalyst from the dense catalyst bed 82"
of the secondary reactor 80" that falls into the hopper section 81"
is fluidized by steam of some other fluidizing media through a
distributor 84" and is stripped of entrained gases over baffles
83". Stripped catalyst passes through a pipe 76" regulated by a
control valve 78" to the regenerator 50. The product from the
secondary reaction is recovered through cyclones 88" and 90" which
remove entrained catalyst and send the catalyst back to the dense
catalyst bed 82". A conduit 92" carries gaseous product to further
processing which could consist of heating and fractionating.
[0026] The secondary reactor 80, 80', 80" may stand alone instead
of being incorporated into an FCC unit. If the secondary reactor
80, 80', 80" stands alone, the preferred feed will be a cut of
product from an FCC unit.
[0027] In reference to all of FIGS. 1-3, the cracked product stream
in the line 98 from the FCC reactor 10, relatively free of catalyst
particles and including the stripping fluid, exits the reactor
vessel 14 through the outlet conduit 36. The cracked product stream
in the line 98 may be subjected to additional treatment to remove
fine catalyst particles or to further prepare the stream prior to
fractionation. The line 98 transfers the product stream containing
the cracked product to a fractionator in the form of the main
column 100. A variety of products are withdrawn from the main
column 100. In this case, the main column 100 recovers an overhead
stream of light products comprising unstabilized gasoline and
lighter gases. A line 102 transfers the overhead stream through a
condenser 104 and a cooler 106 before it enters a receiver 108. A
line 110 withdraws a light off-gas stream from the receiver 108. A
bottom liquid stream of light gasoline leaves the receiver 108 via
a line 112 which may have to undergo further treatment to stabilize
the light gasoline. The main column 100 also provides a heavy
gasoline stream, an LCO stream and an HCO stream through lines 120,
122 and 124, respectively. Parts of the streams in the lines 120,
122 and 124 are all circulated through heat exchangers 126, 128 and
130 and reflux loops 132, 134 and 136, respectively, to remove heat
from the main column 100. Streams of heavy gasoline, LCO and HCO
are transported from the main column 100 through respective lines
140, 142 and 144. A CO fraction may be recovered from the bottom of
the main column 100 via a line 146. Part of the CO fraction is
recycled through a reboiler 148 and returned to the main column 100
through a line 150. The CO stream is removed from the main column
100 via a line 152.
[0028] The light gasoline or light naphtha fraction preferably has
an initial boiling point (IBP) below 127.degree. C. (260.degree.
F.) in the C.sub.5 range; i.e., about 35.degree. C. (95.degree.
F.), and an end point (EP) at a temperature greater than or equal
to 127.degree. C. (260.degree. F.). The boiling points for these
fractions are determined using the procedure known as ASTM D86-82.
The heavy gasoline or heavy naphtha fraction has an IBP at or above
127.degree. C. (260.degree. F.) and an EP at a temperature above
200.degree. C. (392.degree. F.), preferably between 204.degree. and
221.degree. C. (4000 and 430.degree. F.), particularly at
216.degree. C. (420.degree. F.). The LCO stream has an IBP at about
the EP temperature of the heavy gasoline and an EP in a range of
260.degree. to 371.degree. C. (5000 to 700.degree. F.) and
preferably 288.degree. C. (550.degree. F.). The HCO stream has an
IBP of the EP temperature of the LCO stream and an EP in a range of
3710 to 427.degree. C. (7000 to 800.degree. F.), and preferably
about 399.degree. C. (750.degree. F.). The CO stream has an IBP of
the EP temperature of the HCO stream and includes everything
boiling at a higher temperature. One or more of each of these
streams or other cuts from the main column 100 are sent to the
secondary reactor 80, 80', 80" to be contacted with the catalyst
therein. In one embodiment, a stream such as the line 142 which
carries LCO may be hydrotreated in a hydrotreating reactor 154
before it is sent to the secondary reactor 80, 80', 80" for
cracking. Other streams from the main column 100 could be
hydrotreated before entering the secondary reactor 80, 80',
80".
[0029] In the secondary reactor 80, 80', 80", the predominant
reaction may be cracking in which a hydrocarbon molecule is broken
into two smaller hydrocarbon molecules, so that the number of
carbon atoms in each molecule diminishes. Alternatively, the
predominant reaction in the secondary reactor 80, 80', 80" may be a
hydrogen-transfer reaction such as reformulation or isomerization
in which the structures of the molecules are changed but the number
of carbon atoms in each molecule does not change. In determining
which type of reaction, cracking or hydrogen transfer, predominates
over the other, reactions involving compounds with 5 to 8 carbons
may be the most relevant because they include most of the olefins
which can either crack or reform.
[0030] Olefins, naphthenes and cyclo-olefins are reformulated into
paraffins, aromatics and some naphthenes as shown in formulas (1),
(2), (3) and (4).
3C.sub.nH.sub.2n+C.sub.mH.sub.2m.fwdarw.3C.sub.nH.sub.2n+2+C.sub.mH.sub.2m-
-6olefins+naphthene.fwdarw.paraffins+aromatic (1)
4C.sub.nH.sub.2n.fwdarw.3C.sub.nH.sub.2n+2+C.sub.nH.sub.2n-6olefins+paraff-
ins.fwdarw.aromatic (2)
C.sub.mH.sub.2m-2+2C.sub.nH.sub.2n.fwdarw.C.sub.mH.sub.2m-6+2
C.sub.nH.sub.2n+2cyclo-olefins+olefins.fwdarw.aromatic+paraffins
(3)
C.sub.nH.sub.2n+H.sub.2.fwdarw.CnH.sub.2n+2olefins+hydrogen.fwdarw.paraffi-
ns (4)
[0031] Olefins have a higher octane value than their paraffinic
counterpart. Hence, the conversion of olefins to paraffins
typically degrades octane value. When the olefins cyclitize to
become aromatics as shown in formulas (1) and (2) and when
cyclo-olefins aromaticize to yield aromatics as in formula (3),
they donate much hydrogen. Other olefins pick up the hydrogen to
become paraffins as shown in formula (4). In the present invention
using the secondary reactor 80, 80', 80", normal olefins and
iso-olefins predominantly reformulate to isoparaffins which carry a
higher octane rating than normal paraffins. Additionally, aromatics
also boost the octane rating of the product. Because the
isoparaffins and aromatics have a high octane rating, the hydrogen
transfer reformulation in the secondary reactor 80, 80', 80"
maintains the high octane ratings despite the typical octane rating
decline that accompanies conversion of olefins to paraffins
Accordingly, the hydrogen-transfer reactions in the secondary
reactor 80, 80', 80" which yield more isoparaffins and aromatics
are superior to a process which saturates the olefins into normal
paraffins. Advantageously, the hydrogen transfer reactions are
performed without the addition of hydrogen, which can be expensive
and difficult to obtain
[0032] Production of aromatics is a gauge for the degree of
hydrogen transfer that occurs in the reaction When conditions are
set to promote hydrogen transfer reactions in the secondary reactor
80, 80', 80", a net yield increase in aromatics of 5% on a fresh
feed basis is typical and at least a 40% increase is easily
attainable.
[0033] The reaction in the secondary reactor 80, 80', 80" is
preferably conducted with the same catalyst circulated through the
regenerator 50 and the FCC reactor 10. Of course, if a secondary
reactor 80, 80', 80" stands alone without incorporation into an FCC
unit, the catalyst in the secondary reactor need not be circulated
through an FCC unit. If hydrogen-transfer reactions are intended to
predominate over cracking reactions in the secondary reactor, the
WHSV will typically range from 0.1 to 5 hr.sup.-1. If cracking
reactions are to predominate over hydrogen-transfer reactions, the
WHSV will typically range from 5 to 50 hr.sup.-1. Additionally, the
conditions in a hydrogen-transfer reaction are less severe, with
temperatures in the range of 3990 to 510.degree. C. (750.degree. to
950.degree. F.) than in a cracking reaction with temperatures in
the range of 482.degree. to 649.degree. C. (9000 to 1200.degree.
F.).
[0034] An additional advantage of the hydrogen transfer reaction in
the secondary reactor 80, 80', 80" is that it is endothermic.
Hence, the spent catalyst which contacts the hydrocarbon stream in
the dense catalyst bed 82, 182, 282 is cooled before it is sent
back to the reactor riser 12 of the FCC reactor 10 or the
regenerator 50. Consequently, heat will be removed from the whole
system which permits use of a greater catalyst-to-oil ratio in the
reactor riser 12, resulting in higher conversion in the FCC reactor
10.
[0035] The reformulation of the fraction from the main column 100
by hydrogen transfer in the secondary reactor 80, 80', 80" reduces
the concentrations of organic sulfur and nitrogen compounds in the
products. The reaction of the gasoline fraction in the secondary
reactor 80, 80', 80" can lower sulfur concentration in the reactor
products by as much as 80 wt- % and nitrogen concentration in the
products by as much as 98 wt-%. Hence, the products from the
secondary reactor 80, 80', 80" will contain low concentrations of
sulfur and nitrogen compounds. Leftover sulfur and nitrogen
compounds can be removed from the product by hydrotreating and
taken off in the overhead of a finishing distillation column if
necessary to meet specifications.
[0036] Typically, the catalyst circulation rate through the reactor
riser 12 and the input of feed and any lift gas that enters the
riser will produce a flowing density of between 48 and 320
kg/m.sup.3 (3 and 20 lbs/ft.sup.3) and an average velocity of about
3 to 31 m/sec (10 to 100 ft/sec) for the catalyst and gaseous
mixture. In the FCC reactor 10, catalyst will usually contact the
hydrocarbons in a catalyst to oil ratio in a range of from 3 to 8,
and more preferably in a range of from 4 to 6. The length of the
reactor riser 12 will usually be set to provide a residence time of
between 0.5 to 10 seconds at these average flow velocity
conditions. Other reaction conditions in the reactor riser 12
usually include a temperature of from 468.degree. to 566.degree. C.
(8750 to 105.degree. F.).
[0037] This invention can employ a wide range of commonly used FCC
catalysts. These catalyst compositions include high activity
crystalline alumina silicate or zeolite containing catalysts.
Zeolite catalysts are preferred because of their higher intrinsic
activity and their higher resistance to the deactivating effects of
high temperature exposure to steam and exposure to the metals
contained in most feedstocks. Zeolites are usually dispersed in a
porous inorganic carrier material such as silica, aluminum, or
zirconium. These catalyst compositions may have a zeolite content
of 30% or more. Zeolites including high silica-to-alumina
compositions such as LZ-210 and ZSM-5 type materials are preferred
when lighter products are desired. Another particularly useful type
of FCC catalysts comprises silicon substituted aluminas. As
disclosed in U.S. Pat. No. 5,080,778 B1, the zeolite or silicon
enhanced alumina catalysts compositions may include intercalated
clays, also generally known as pillared clays The preferred
catalysts for the present invention include USY zeolites. When
hydrogen-transfer reactions are desired to predominate over
cracking reactions in the secondary reactor 80, 80', 80", high rare
earth content Y zeolites are preferred. The term "high rare earth
content" denotes greater than about 2.0 wt-% rare earth oxide on
the zeolite portion of the catalyst. High rare earth content Y
zeolites such as USY zeolite may have as much as 4 wt-% rare earth.
The high rare earth content promotes hydrogen transfer by
increasing adjacent acid site density on the catalyst. Strongly
acidic catalyst sites on the catalyst promote cracking. Y zeolites
with low rare earth content can still effectively promote hydrogen
transfer but with longer reactor residence times. When cracking
reactions are desired to predominate over hydrogen transfer
reactions in the secondary reactor 80, 80', 80", low rare earth Y
zeolite catalysts are preferred which have a rare earth oxide
content of 2.0 wt-% or less. Additives, such as sulfur-reducing
additives, may be added to the catalyst. It is anticipated that
such additives may experience enhanced effectiveness in the
secondary reactor for longer residence times.
[0038] Feeds suitable for processing by this invention include
conventional FCC feedstocks or higher boiling hydrocarbon feeds.
The most common of the conventional feedstocks is a vacuum gas oil
which is typically a hydrocarbon material having a boiling range of
from 343.degree. to 552.degree. C. (650.degree. to 1025.degree. F.)
and is prepared by vacuum fractionation of atmospheric residue Such
fractions are generally low in coke precursors and heavy metals
which can deactivate the catalyst.
[0039] When LCO is the feed to the secondary reactor 80, 80', 80",
a portion of the LCO fraction will typically pass through the
hydrotreating reactor 154 and be transported through a line 156 to
the secondary reactor 80, 80', 80" in which J-cracking occurs. When
operating in the LCO mode of this invention, the LCO cut carries
bicyclic aromatic compounds into the secondary reactor 80, 80', 80"
which cannot be cracked unless they are pretreated. These bicyclic
compounds include indenes, biphenyls and naphthalenes which are
refractory to cracking under the conditions in the reactor riser
12. In the J-cracking process, one of the rings of the bicyclic
hydrocarbons are saturated. The saturated ring is then cracked in
the secondary reactor 80, 80', 80" and cleaved from the aromatic
ring as shown in exemplary formulas (5) and (6). 1
[0040] In formula (5), one of the rings of dimethyl naphthalene is
saturated to make dimethyl tetrahydronaphthalenes. In formula (6),
the saturated ring of two dimethyl tetrahydronaphthalenes are
cracked and accept hydrogen donated from a ring of another dimethyl
tetrahydronaphthalene that aromaticizes. The cracked rings yield
toluene and isobutane.
[0041] Suitable methods for carrying out J-cracking are further
described in U.S. Pat. No. 3,479,279 B1 and U.S. Pat. No. 3,356,609
B1 which are incorporated herein by reference. The J-cracking
process eliminates about two-hirds of the high boiling aromatics
from an LCO cut bringing the effluent from the secondary reactor
80, 80', 80" into the gasoline boiling range. The LCO fraction can
pass through the hydrotreating reactor 154 as a separate stream or
together with another fraction from the main column 100.
[0042] The hydrotreatment of the fraction in the hydrotreating
reactor 154 takes place at low severity conditions to avoid the
saturation of the single ring aromatic compounds in the gasoline
fraction. In the method of this invention, up to 100% of the
fraction may be hydrotreated. Hydrotreating is carried out in the
presence of a nickel-molybdenum or cobalt-molybdenum catalyst and
relatively mild hydrotreating conditions including a temperature of
3160 to 371.degree. C. (6000 to 700.degree. F.), a liquid hourly
space velocity (LHSV) of from 0.2 to 2 hr.sup.-1 and a pressure of
3447 to 10342 kPa (500 to 1500 psig).
[0043] The present invention can be operated in several ways, four
of which are explained herein. In the first exemplary operation,
higher proportions of LCO and LPG are obtained. The FCC reactor 10
is run at relatively low severity with a temperature between
482.degree. and 521.degree. C. (900.degree. and 970.degree. F.) and
a short contact time of 1 to 3 seconds. The FCC reactor 10 will
thus operate at low conversion to yield a high proportion of LCO,
HCO and CO, some gasoline and some liquefied petroleum gas (LPG),
all withdrawn from the main column 100. If the feed to the FCC
reactor 10 is highly paraffinic, all of the CO can be fed to the
secondary reactor 80, 80', 80". However, if the feed is not highly
paraffinic, only the HCO fraction should be fed to the secondary
reactor. Fractions of LCO and LPG product can be recovered from the
main column 100. If gasoline is desired, it can be recovered from
the main column and sent to the gasoline pool. If gasoline is not
desired, it can be sent with CO and HCO or alone to the secondary
reactor 80, 80', 80" which cracks the CO and HCO mixture at high
severity temperatures such as 521.degree. to 560.degree. C. (9700
to 1040.degree. F.) and preferably 549.degree. C. (1020.degree. F.)
and at a space time of 1 to 10 hr.sup.-1. LPG and LCO are then
recovered from the secondary reactor which can be added to the
fractions of LCO and LPG recovered from the main column 100. A
medium or smaller pore, shape selective zeolite additive such as
ZSM-5 may be added to the catalyst to obtain greater yields of LPG
in this operation. Because the secondary reactor is operated at
high severity, FIG. 2 or 3 would be most appropriate for this
operation because the hotter catalyst from the regenerator 50 can
provide the necessary heat requirements.
[0044] A second operation in which the present invention can be
used to produce gasoline, LPG and benzene, toluene and xylene (BTX)
gasoline. The FCC reactor 10 is run at a high severity temperature
ranging from 521.degree. to 560.degree. C. (9700 to 1040.degree.
F.), preferably 549.degree. C. (1020.degree. F.) and a contact time
of over 3 seconds. The high severity cracking operation gives a
high conversion with gasoline, LPG, LCO and CO in the product
stream. Gasoline and LPG are recovered from the main column 100
while LCO is fed from the main column 100 to the hydrotreating
reactor 154 to saturate one of the bicyclic aromatic rings to
prepare it for cracking. The hydrotreated LCO is then sent to the
secondary reactor 80, 80', 80" operated at high severity
temperatures of 521.degree. to 560.degree. C. (970.degree. to
1040.degree. F.) sufficient to J-crack it to obtain BTX gasoline
which can be mixed with gasoline to upgrade gasoline product
quality. The embodiments in FIG. 2 or 3 can be used for this
exemplary operation.
[0045] In a third exemplary operation, the desired product yields
up to an 80% reduction in gasoline sulfur and nitrogen and
possesses an olefin concentration as low as 1 wt- %. The primary
reactor is run at a severity appropriate to obtain the desired
conversion. Either a full range cut of gasoline having an IBP below
127.degree. C. (260.degree. F.) and an EP at or below 200.degree.
C. (392.degree. F.) or a fraction thereof from the main column 100
is fed to the secondary reactor 80, 80', 80" which is run at
482.degree. to 521.degree. C. (900.degree. to 970.degree. F.). In
the secondary reactor, the olefins reformulate via hydrogen
transfer to isoparaffins and aromatics with minimal gasoline yield
loss and an octane gain and without need of additional hydrogen.
Moreover, sulfur levels are reduced by as much as 80 wt-% and
nitrogen levels are reduced by as much as 98 wt-%. If necessary,
the gasoline can then be hydrotreated to reduce sulfur and nitrogen
compounds to even lower levels to meet specifications by converting
them to hydrogen sulfide and ammonia, respectively, which can be
removed in the light ends of a downstream gasoline fractionation
unit (not shown) with minimal octane debit and consumption of
hydrogen. This operation can be performed with any of the three
embodiments in FIGS. 1-3 of the present invention.
[0046] When the desired products are LCO and low olefinicity,
moderate octane gasoline, a fourth exemplary operation may be used.
The FCC reactor 10 is run at low severity at a temperature of
482.degree. to 521.degree. C. (9000 to 970.degree. F.) and a
contact time of 1 to 3 seconds. The low conversion operation yields
high quantities of LCO, some gasoline and not much LPG. The LCO can
be recovered from the main column 100. The gasoline fraction can be
fed to the secondary reactor at low severity 4820 to 521.degree. C.
(900 to 970.degree. F.) and low WHSV, 0.1 to 5 hr.sup.-1, so the
gasoline reforms to convert olefins to aromatics and isoparaffins
to upgrade the gasoline quality.
EXAMPLES
Example 1
[0047] A fraction of gasoline from an FCC reactor effluent having
the properties in Table I was subjected to coked USY zeolite
catalyst with 1 to 1.5 wt-% rare earth in a reactor at the
conditions in Table I. The reaction yielded a product with the
properties in Table I.
1 TABLE I FEED PROPERTIES IBP, .degree. C. (.degree. F.) 121 (250)
Aromatics, wt-% 61.8 Olefins, wt-% 14.2 Paraffins/Naphthenes, wt-%
24 RON 93.3 MON 81.9 REACTOR CONDITIONS WHSV, hr.sup.-1 1 Reaction
Temperature, .degree. C. (.degree. F.) 454 (850) Catalyst-to-Oil
Ratio 6.0 Pressure, kPa (psig) 69 (10) PRODUCT PROPERTIES
C.sub.2.sup.-, wt-% 0.6 C.sub.3, wt-% 1.2 C.sub.4, wt-% 2.0
C.sub.5.sup.+/232.degree. C. (450.degree. F.), wt-% 89.4 LCO, wt-%
4.7 CO, wt-% 2.1 Gasoline RON 95.8 Gasoline MON 84 Aromatics, wt-%
70 Olefins, wt-% 1 Paraffins/Naphthenes, wt-% 29
[0048] In this example, the olefin concentration dropped from 14.2%
to 1 wt-% as a result of the secondary reaction. Whereas, the
aromatics concentration increased from 61.8 to 70 wt-%.
Additionally, both the RON and the MON increased. The relatively
small concentrations of C.sub.4 and smaller hydrocarbons reveal
that cracking reactions were minor compared to the reformulating,
hydrogen transfer reactions indicated by the increase in
aromatics.
Example 2
[0049] A separate study was performed to determine the effect on
product properties of four sets of operating conditions on full
range FCC gasoline as shown in Table II.
2TABLE II FEED PROPERTIES IBP, .degree. C. (.degree. F.) 35 (95)
Paraffrns, wt-% 27 Olefins, wt-% 51 Naphthenes, wt-% 6 Aromatics,
wt-% 14 C.sub.4, wt-% 2.3 Feed Boiling Over 1.3 221.degree. C.
(430.degree. F.), wt-% PROCESS A B C D CONDITIONS Reaction 399
(750) 399 (750) 454 (850) 482 (900) Temperature, .degree. C.
(.degree. F.) Catalyst-to-Oil Ratio 3 5 5.1 5.1 PRODUCT YIELDS,
wt-% C.sub.2.sup.- 0.06 0.13 0.43 0.60 C.sub.3 0.82 1.22 2.85 4.16
C.sub.4 3.5 4.53 6.75 8.35 C.sub.5.sup.+/220.degree. C.
(429.degree. F.) 91.3 86.4 83.1 80.0 LCO 2.5 3.69 3.28 2.87 CO 0.2
1.5 1.4 1.9 Coke 1.6 2.5 2.2 2.1 Gasoline Recovery 94.9 90.0 86.7
83.6 Paraffins 42 47 48 44 Olefins 31 21 18 13 Naphthenes 8 8 7 7
Aromatics 21 23 27 36
[0050] As the temperature is increased, the gasoline recovery
diminished while the aromatics concentration increased and the
olefins concentration decreased. Additionally, cracking as
indicated by the amount of C.sub.4 and lower carbon number
concentration increases as the reaction temperature and/or
catalyst-to-oil ratio increases. Accordingly, the reaction
conditions can be tailored to obtain a desired product quality.
Example 3
[0051] The feed in the next set of experiments had the properties
given in Table III.
3 TABLE III Paraffins, wt-% 28.1 Olefins,wt-% 50.4 Naphthenes, wt-%
5.9 Aromatics, wt-% 14.4 C.sub.12 Non-Aromatics, wt-% 1.32 RON 91.0
MON 79.3 Road Octane Number 85.2 Sulfur, ppm 136 Nitrogen, ppm 46
C.sub.4, wt-% 2.3 221.degree. C. (430.degree. F.) plus, wt-% 1.3
IBP, .degree. C. (.degree. F.) 35 (95) T10 51 (123) T30 67 (153)
T50 88 (190) T70 118 (244) T90 152 (306) EP, .degree. C. (.degree.
F.) 179 (354)
[0052] The foregoing feed was reacted under three different sets of
conditions with corresponding product yields and quality given in
Table IV.
4 TABLE IV Run A B C PROCESS CONDITIONS Reactor Temperature,
.degree. C. (.degree. F.) 427 (800) 454 (850) 482 (900)
Catalyst-to-Oil Ratio 6.5 6.1 5.9 Hydrocarbon Partial Pressure, 117
(17.0) 114 (16.5) 122 (17.7) kPa (psia) System Pressure, kPa (psig)
278 (40.3) 276 (40.0) 273 (39.6) LHSV, hr.sup.-1 4.6 4.6 4.6
PRODUCT YIELDS, wt-% Dry Gas 0.4 0.7 1.1 C.sub.3's 1.6 2.4 3.4
C.sub.4's 6.1 7.8 9.4 C.sub.5.sup.+ Gasoline 85.5 83.0 80.0
Paraffins 53.3 54.7 52.3 Olefins 13.8 12.4 12.3 Naphthenes 8.1 5.5
6.2 Aromatics 24.8 27.4 29.2 Sulfur, ppm 69 62 68 Nitrogen, ppm 1 2
4 RON 87.4 88.4 90.4 MON 80.5 81.5 81.8 Road Octane Number 84.0
85.0 86.1
[0053] The foregoing qualities and yields pertaining to the
C.sub.5+gasoline have been adjusted to reflect the fact that
C.sub.4's were present in the feed which did not participate in the
reaction and would not be present in the feed to the secondary
reactor. Moreover, the data indicates that not much cracking
occurred in the reaction because relatively small quantities of
C.sub.4.sup.- material is generated. The process also reduces the
olefin concentration while increasing the paraffin and aromatics
concentration, all without substantial change in the Road Octane
Number.
[0054] Table V gives the breakdown of the product composition from
foregoing Run B by carbon number and compound type. The number that
is not in parentheses in Table V is the weight percentage of that
compound in the feed. Whereas, the number in parentheses is the
weight percentage of the compound in the product.
5TABLE V Gasoline Composition Full Range Feed vs. Product Carbon #
Total Naphthenes Isoparaffins n-Paraffins Cyclic-Olefins
Iso-Olefins n-Olefins Aromatics 5 24.93 (25.32) 0.1 (0.0) 7.33
(17.48) 1.53 (2.49) 0.61 (0.14) 8.24 (3.17) 7.12 (2.03) -- (--) 6
23.00 (23.92) 1.22 (1.79) 6.31 (15.28) 0.92 (1.63) 2.04 (0.32) 7.33
(3.12) 4.68 (1.11) 0.51 (0.67) 7 18.17 (16.43) 1.79 (1.94) 3.87
(7.59) 0.51 (0.90) 2.24 (0.29) 4.88 (1.33) 2.54 (0.29) 2.34 (4.10)
8 14.96 (14.41) 1.53 (0.87) 2.54 (3.88) 0.51 (0.63) 1.02 (--) 3.15
(0.46) 1.32 (--) 4.88 (8.57) 9 12.72 (17.58) 0.92 (0.64) 1.83
(2.31) 0.41 (0.43) 0.31 (--) 1.83 (0.16) 0.81 (--) A9 + 10 3.47
(1.79) 0.32 (0.24) 1.12 (1.10) 0.41 (0.46) 0.00 (--) 1.12 (--) 0.51
(--) 6.61 (14.04) 11 1.42 (0.52) -- (--) 0.51 (0.54) 0.31 (0.00)
0.00 (--) 0.41 (--) 0.2 (--) Total 98.68 (100) 5.87 (5.49) 23.51
(48.19) 4.58 (6.55) 6.21 (0.74) 26.97 (8.23) 17.2 (3.42) 14.35
(27.37) 12 C.sub.12.sup.+ Non-Aromatics: 1.3
[0055] With regard to Table V, aromatics with nine or more carbon
numbers are grouped together. Therefore, the numbers given for
carbon numbers 10 and 11 in the "Total" column include only
non-aromatic C.sub.10's and C.sub.11's. The minimal changes in
total concentration of each carbon number fraction, especially in
the C.sub.5-C.sub.8 range shows that reformulating hydrogen
transfers are predominant over cracking reactions under this set of
conditions. Moreover, the large increase in isoparaffins compared
to the moderate increase in paraffins greatly offsets the octane
value debit resulting from olefin reduction.
* * * * *