Note: Descriptions are shown in the official language in which they were submitted.
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PROCESS FOR IMPROVING FLOW PROPERTIES OF CRUDE PETROLEUM
BACKGROUND OF THE INVENTION
[0001] The field of the invention is improvement of the flow properties of
crude
petroleum.
RELATED PRIOR ART
[0002] When drilling for oil in remote places, considerable expense is
associated with
transporting the crude oil from the wellhead to a receiving facility. One
difficulty of
transporting crude oil is that certain crude oils may contain a significant
quantity of wax,
which has a high boiling point. The temperature at which the wax gels is the
pour point. The
temperature at which the wax solidifies is the cloud point. In instances where
the cloud point
or the pour point of a waxy crude oil is higher than the ambient temperature,
the likelihood of
wax solidification and buildup is a serious threat to a continuous
transportation of crude oil.
Clearing a pipeline that has become clogged with wax or gelled crude is very
expensive and
time-consuming.
[0003] Another specification for pipeline pumpability is the viscosity of the
oil. The
viscosity of the oil is proportional to the duty required to pump it. Hence,
each pipeline has a
viscosity, API and pour point specification. For example, to be accepted for
shipment in the
Enbridge Pipeline system in Canada and the U.S., the viscosity specification
is 350
Centistokes (cSt) at the pipeline operating temperature, which varies
seasonally.
[0004] Still another specification for pipeline pumpability is American
Petroleum
Institute (API) gravity index. Crude oil is often described in terms of
"lightness" or
"heaviness" by the API gravity index. A high number denotes a "light" crude,
and a low
number denotes a "heavy" crude.
[0005] A petroleum product with good flow properties such as low pour point,
high API
gravity, and low viscosity is desired by refiners.
[0006] FCC is a catalytic process for converting heavy hydrocarbons into
lighter
hydrocarbons by contacting the heavy hydrocarbons in a fluidized reaction zone
with a
catalyst absent substantial added hydrogen. Most FCC units now use zeolite-
containing
catalyst having high activity and selectivity. As the cracking reaction
proceeds, substantial
amounts of highly carbonaceous material referred to as coke are deposited on
the catalyst,
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forming spent catalyst. High temperature regeneration bums coke from the spent
catalyst.
Spent catalyst is continually removed from the reaction zone and replaced by
essentially
coke-free regenerated catalyst from the regeneration zone.
[0007] US 2007/0034550 Al teaches subjecting a portion of a crude stream to
FCC and
mixing a portion of the cracked stream with a second crude stream to
facilitate pipeline
transport. We have found that crudes subjected to FCC can produce olefins and
diolefins in
the gasoline and lighter portions of the product which are believed to cause
fouling of heat
exchangers and other equipment at the refinery end of the pipeline. A system
for extracting
and transporting crude oil from a remote field while maintaining a
sufficiently low
concentration of olefins and diolefins would be desirable.
[0008] Hydrotreating is a process in which hydrocarbon feeds are contacted
with catalyst
in the presence of added hydrogen to saturate olefins and diolefins and/or
desulfurize organic
sulfur. Hydrotreating is performed at elevated temperature and pressure.
Hydrotreating
cannot be performed without a source of hydrogen.
SUMMARY OF THE INVENTION
[0009] We have discovered a process and apparatus for preparing crude streams
for pipe
transport with sufficiently low olefin concentration. Hydrogen in the cracked
product is used
to hydrotreat a portion of the cracked hydrocarbon product to saturate
problematic olefins
over a hydrotreating catalyst. In one aspect of the invention, dry gas is
separated from the
FCC products and used for hydrotreating another portion of the FCC product to
saturate
problematic olefins. In a further aspect, the dry gas is purified to provide a
sufficiently
hydrogen-rich stream for hydrotreating the other portion of the FCC product.
In an even
further aspect, olefins in an LPG portion of the FCC product are oligomerized
and
hydrotreated. The hydrotreated streams may be blended with an unprocessed
crude stream to
prepare the unprocessed crude stream for pipeline transportation.
BRIEF DESCRIPTION OF THE SEVERAL VIEWS OF THE DRAWINGS
[0010] FIG. 1 is a flow scheme showing a process and apparatus of the present
invention.
[0011] FIG. 2 is a flow scheme showing an alternative the process and
apparatus of the
present invention.
[0012] FIG. 3 is a flow scheme of a hydrogen purification unit.
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DETAILED DESCRIPTION OF THE INVENTION
[0013] This invention may improve the flow properties of a crude petroleum
stream. The
process makes cutter stock from a portion of a crude oil using modularly
designed
components. Crude oil may comprise the crude feed to be catalytically cracked
by a fluidized
catalytic cracking (FCC) process and the product may be mixed with unprocessed
crude oil to
create a blend of processed and unprocessed crude to improve the flow
properties of the
crude by lowering the crude pour point, raising the API and/or reducing the
viscosity for
easing transportation of the blended product through a pipeline to a location
remote from the
oil field for further processing.
[0014] Residual fluidized catalytic cracking (RFCC) may be used to process
Conradson
carbon residue and metals-contaminated feedstocks such as atmospheric residues
or mixtures
of vacuum residue and gas oils. Depending on the level of carbon residue and
nickel and
vanadium contaminants, it is contemplated that these feedstocks may be
hydrotreated or
deasphalted before being fed to an RFCC unit.
[0015] Crude oil from a source may comprise all or part of a crude feed stream
to be
processed by FCC. Crude feed processed by this invention may be heavy
hydrocarbon
comprising heavy oil or bitumen. Whole bitumen may include resins and
asphaltenes, which
are complex polynuclear hydrocarbons, which add to the viscosity of the crude
oil and
increase the pour point. Crude feed may also include conventional crude oil,
atmospheric
tower bottom products, vacuum tower bottoms, coal oils, residual oils, tar
sands, shale oil and
asphaltic fractions.
[0016] Heavy crude oil is typically very viscous, having a API gravity of
between 8 and
13 API. Waxy crudes typically have a higher API in excess of 25, but a pour
point of between
20 and 50 C. Viscosity of crude oil may be between 10,000 and 15,000 cSt at
40 C. Crude
oil may be characterized as a hydrocarbon stream having properties in at least
one of the
following ranges: pour point of greater than 20 C, viscosity greater than
10,000 cSt at 38 C
(100 F) and an API gravity typically greater than 18 API.
PROCESSING APPARATUS
[0017] Referring to FIG. 1, apparatus 10 delivers a crude oil stream from the
oil field
ground 1 in line 3. The crude oil stream in line 3 is typically subjected to
heating and
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separation of an oil phase from a water phase to dewater the crude oil stream
in line 3. The
crude oil stream in line 3 is separated into two portions. A first crude
stream is carried in line
for processing while a second crude stream is carried in line 499 to bypass
the processing of
line 5. The first and second crude streams in lines 5 and 499, respectively,
will have the
5 characteristics of crude oil given above. The crude oil may be sent to a
fired heater 20 where
the crude oil may be preheated. Optionally, the crude oil in line 5 may also
be heated in heat
exchanger 18 by indirect heat exchange with bottoms recycle in line 22. After
leaving heater
20, the heated crude oil may be introduced into lower portion 31 of
fractionator 30. In some
FCC processes, the first crude stream in line 5 is not directed to
fractionator 30 but may
instead be introduced directly to riser 40 for catalytic cracking.
[0018] The recovery of resids, or bottom fractions, involves selective
vaporization or
fractional distillation of the crude oil with minimal or no chemical change in
the crude oil.
The fractionating process may provide a feed stock more suitable for FCC
processing. The
selective vaporization of the crude oil takes place under non-cracking
conditions, without any
reduction in the viscosity of the feedstock components. Light hydrocarbons,
those boiling
below 700 F (371 C), preferably those boiling below 675 F (357 C), and most
preferably
those boiling below 650 F (343 C), are flashed off of the crude oil in feed
zone 36. The light
hydrocarbons typically are not catalytically cracked. Hence, the feed zone 36
serves as a
stripper in which light hydrocarbons are stripped from the crude feed to
provide a stripped
first crude stream in FCC feed line 32.
[0019] The first crude feed stream 5 may be fed directly to a riser 40 of an
FCC unit
without the fractionating step, depending on the quantity of light ends,
gasoline, gas oils and
residuals. Direct feeding would be desirable if the quantity of hydrocarbons
boiling below
650 F (343 C) is relatively low and their segregation therefore unnecessary.
The bottoms
product of fractionator 30, in feed zone 36 comprising a stripped first crude
stream is
withdrawn via FCC feed line 32 and directed by pump 33 to a bottom of the
riser 40.
[0020] The feed rate to apparatus 10 may be between 5,000 and 200,000 barrels
per day,
preferably between 25,000 and 150,000 barrels per day, and more preferably
100,000 barrels
per day although the feed rate could vary from these ranges. Feed to the FCC
may be between
10 LV-% and 60 LV-% of the complex charge in line 3 from the oil field 1 with
lower rates
being preferable to higher rates unless utility balances require higher charge
rates. The
stripped first crude stream in line 32 is contacted with catalyst in the riser
40 perhaps in the
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presence of an inert fluidization gas such as steam. The first crude stream is
cracked into
lighter hydrocarbon products which are carried out of the riser 40 as a
cracked stream. The
catalyst becomes spent as carbon residue builds up on the catalyst surface.
The spent catalyst
and the cracked stream exit from the top of riser 40 and into a reactor vessel
50 in
downstream communication with the top of the riser 40 optionally through a
rough cut
separator 51 to separate cracked stream vapors from the spent catalyst. One or
more stages of
cyclones 52 further separate the spent catalyst from the cracked stream by
inducing the
mixture of catalyst and cracked stream gases to swirl so that the heavier
spent catalyst travels
downwardly and the lighter gaseous cracked stream travel upwardly.
[0021] Approximate operating conditions include heating the crude feed for
catalytic
cracking to between 300 and 500 F (between 149 and 260 C), preferably
between 350 and
450 F (between 177 and 232 C), and more preferably 400 F (204 C). The
temperature in
reactor vessel 50 may be between 850 and 1100 F (between 454 and 593 C),
preferably
between 900 and 1050 F (between 482 and 566 C), and more preferably between
950 and
1000 F (between 510 and 538 C). The FCC conversion may be between 40 and 80
LV-% to
gasoline and lighter products, between 65 LV-% and 75 % LV-% to gasoline and
lighter
products, or 70 LV-% to gasoline and lighter products.
[0022] Continuing with FIG. 1, the vapor products exit the top of reactor
vessel 50 and
may be directed via line 53 to product zone 37 in lower portion 31 of the
fractionator 30 in
downstream communication with the reactor vessel 50. Heat from product vapors
may be
absorbed within fractionator 30, so that the vapors are desuperheated and the
primary product
separation takes place. The heat required for the separation of the products
in fractionator 30
is primarily provided by the cracked product stream. Thus, in the case that
the crude feed is
sent directly to riser 40, no other heat is input to fractionator 30. The
fractionation of product
fed to product zone 37 may be by heat removal, rather than heat input. The
heat may be
removed from the fractionator by a series of pump-around exchanger flows
coupled with
fractionator bottoms steam generation and overhead cooling in the form of an
air/water
cooled condenser.
FCC PRODUCTS
[0023] Catalysts most appropriate for use in riser 40 are zeolitic molecular
sieves having
a large average pore size. Typically, molecular sieves with a large pore size
have pores with
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openings of greater than 0.7 nm in effective diameter defined by greater than
10 and typically
12 membered rings. Pore Size Indices of large pores are above 31. Suitable
large pore zeolite
components include synthetic zeolites such as X-type and Y-type zeolites,
mordenite and
faujasite. Y zeolites with low rare earth content may be the preferred
catalyst. Low rare earth
content denotes less than or equal to 1.0 wt-% rare earth oxide on the zeolite
portion of the
catalyst. The catalyst may be dispersed on a matrix comprising a binder
material such as
silica or alumina and/or an inert filer material such as kaolin. It is
envisioned that equilibrium
catalyst which has been used as catalyst in an FCC riser previously or other
types of cracking
catalyst may be suitable for use in the riser of the present invention.
[0024] In order to increase hydrogen production in the FCC for saturating
olefins, the
nickel activity of the catalyst may be optimized by adjusting the
concentration of nickel
passivation agent, such as antimony, injected with the feed. The nickel serves
as a
dehydrogenation catalyst under the conditions in the FCC riser 40. Although
vanadium is also
a dehydrogenation metal, it should be controlled by metal trapping agents,
such as a rare
earth metal to control vanadic attack on the zeolitic framework.
[0025] The FCC system cracks most of the crude feed into material in the C5+
range
boiling at 400 F. These products may have an API gravity of between 30 and 60,
between 35
and 55, or between 40 and 50, and therefore contribute significantly to the
increase in the net
API of the blended stream in line 502. Catalytic cracking of the crude oil
maximizes the API
gravity increase while processing a minimum amount of crude oil.
[0026] The combined liquid product from the FCC processing of crude oil may
contain
converted products from the crude stream and may be transported in line 500.
The liquid
product from the processing of the crude oil is characterized as having an API
gravity of at
least 30, preferably greater than 35, and more preferably greater than 37. The
liquid products
may also have a viscosity of less than 2 cSt, preferably less than 1.5 cSt and
more preferably
less than 1 cSt at 122 F (50 C). The liquid products formed may have a pour
point less than
40 F (4 C), preferably less than 30 F (-1 C), and more preferably less than 25
F (-3.8 C).
The combined liquid conversion products from the processing of the heavy oil
by FCC are
lighter and less viscous by virtue of the reduction in molecular weight. More
cracking in the
FCC may result in lower viscosity and density of the product.
[0027] The exact quantity of feed which is necessary to be processed depends
on the
specific acceptance requirements of the pipeline for pumpability. These may be
specified as
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maximum density or minimum API gravity, maximum viscosity at a certain
temperature,
maximum pour point or any combination of these specifications. Any of the
aforementioned
specifications could be the limiting factor for the amount of processing
needed, depending on
the crude type or the specification. In addition, the specifications may be
different for
different times of the year due to changing pipeline operation temperatures.
Adjustment of
the conversion level of the FCC or of amount processed can be exercised as a
convenient way
to meet the specifications at minimum operating cost.
[0028] The liquid products from the FCC reaction are mixed with a minimally or
unprocessed second crude stream in line 499 to form a mixed crude stream
suitable for
transport in line 502. Between 5 LV-% and 60 LV-% of the crude stream in line
3 may be
FCC processed and added to the second crude stream in line 499, preferably
between 10
LV-% and 40 LV-% of crude feed may be processed and added to the crude stream
in line
499, more preferably 30 LV-% of crude feed may be processed and added to the
crude stream
in line 499. A ratio of the unprocessed crude oil to the liquid products added
may be between
0.5:1 and 9:1, between 1:1 and 4:1, or between 2:1 and 3:1. Liquid streams
from fractionator
30, may be combined with the unprocessed second crude stream in line 499.
Depending on
the site requirements or crude grade desired, it may be desirable to bum all
or part of the
clarified oil in bottoms line 32, to balance the site energy needs or to
upgrade the quality of
the crude stream in line 500 and/or 502.
FRACTIONATOR
[0029] Continuing with FIG. 1, the fractionator column 30 may be a divided-
wall
fractionator with a partition 35 positioned vertically to isolate a feed zone
36 from a product
zone 37 at the bottom of the fractionator 30. Partition 35 may be formed of at
least one baffle
that is generally imperforate (at least 80 % imperforate, preferably 90 %
imperforate).
Multiple baffles may be used. The crude oil is directed to feed zone 36 and
heated to a
temperature between 600 and 800 F (between 315 and 427 C), preferably
between 650
and 750 F (between 343 and 399 C), and most preferably a temperature of 700 F
(371 C) at
a pressure of between 5 and 15 psig (between 35 and 103 kPa), preferably
between 7 and 13
psig (between 48 and 90 kPa), and most preferably 10 psig (69 kPa).
[0030] Fractionator 30 may condense superheated reaction products from the FCC
reaction to produce liquid hydrocarbon products. Fractionator 30 may also
provide some
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fractionation (or stripping) between liquid side stream products. After the
vapor products are
cooled from temperatures of between 900 and 1050 F (between 482 and 966 C),
preferably
between 950 and 1000 F (between 510 and 537 C), and more preferably 970 F
(521 C) to
temperatures of between 50 and 150 F (between 10 and 66 C), preferably
between 70 and
120 F (between 21 and 49 C), and more preferably 100 F (38 C), the vapor
products are
typically condensed into liquid products and the liquid products are
transported out of
fractionator 30 and directed to mix with the minimally processed or unreacted
second crude
stream from line 499 in line 500. In the embodiment of FIG. 1, the liquid
products taken as
cuts from fractionator 30 typically may comprise light cycle oil (LCO) and
fractionator
bottoms or clarified oil, also known as heavy cycle oil (HCO). In FIG. 1, HCO
does not have
a separate cut but is collected in the bottoms. The LCO stream in line 46 is
withdrawn from
the fractionator column 30 by a pump 48 and cooled in steam generator 49. A
reflux portion
is returned to the column 30 at a higher location via line 46a. LCO line 202
takes the
remainder to line 500. Lastly, clarified oil is removed in bottoms line 34
from the fractionator
column 30 by a pump 21 and a return portion is cooled in a feed heat exchanger
18 and
returned to the product zone 37 of the column 30 isolated from the feed side
36 by partition
35. Net bottoms line 203 may take a remainder of the clarified oil to line 500
for blending or
be diverted to the CO boiler 90 through lines 205 and 96.
[0031] FIG.1 shows a further embodiment in phantom in which the fractionator
30 makes
a cut between heavy naphtha and light naphtha. The heavy naphtha stream in
line 44 may be
withdrawn from the fractionator column 30 by a pump 45 and cooled in a boiler
feed water
preheater 47. A reflux portion may be returned to the column at a higher
location via line 44a.
Heavy naphtha line 201 takes the remainder to line 500 for blending. In this
embodiment,
only the light naphtha is taken in line 42 for hydrotreating.
HYDROTREATING
[0032] We have found that the naphtha and lighter FCC product hydrocarbons
boiling at
or below 135 to 177 C (275 to 350 F) contain a large concentration of
olefins and diolefins
which may cause fouling of heat exchanger tubes in a refinery, making the
upgraded crude
stream less problematic. This naphtha cut captures at least 80-90 wt-% of the
olefins and
rejects much of the organic sulfur that would cause hydrodesulfurization which
undesirably
consumes hydrogen in a hydrotreating reactor. We propose to hydrotreat the
naphtha and/or
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lighter FCC product hydrocarbons with hydrogen which may be in the dry gas
stream to
saturate olefins and diolefins. In a simplest embodiment, the entire naphtha
and lighter cut of
the cracked stream in line 42 is fed to a hydrotreating reactor 60 in
downstream
communication with the reactor vessel 50 and riser 40. In this embodiment, no
heavy naphtha
cut is separately taken from the fractionator 30, cooled and pumped around
back to the
fractionator 30. In this embodiment, the heavy naphtha and lighter materials
are removed in
the overhead line 42 from the fractionator 30 and cooled in a condenser 41 and
perhaps a
boiler feed water heater 43. In the alternative embodiment, only the light
naphtha cut boiling
at or below 135 to 177 C (275 to 350 F) is taken in line 42. Under both
alternative
embodiments, the naphtha cut in line 42 is flashed in a receiver 300 from
which water may be
removed from a boot in line 302. A wet gas stream is taken from the receiver
300 in line 306
and compressed in a compressor 310. The compressor 310 can pressurize the wet
gas to 862
to 2068 kPa (125 to 300 psia), and preferably 1448 to 2000 kPa (210 to 290
psia). The
compressed wet gas stream in line 324 may be cooled in heat exchanger 326 and
flashed in a
flash drum 328. A liquid stream from the flash drum 328 in line 330 is fed to
the receiver 300
while a vapor stream in line 332 is fed to an amine absorber 334. A lean
aqueous amine
scrubbing solution is introduced into absorber 334 via line 336 and scrubs
hydrogen sulfide
from the compressed vaporized light ends stream. A rich aqueous amine
scrubbing solution
containing hydrogen sulfide is removed from absorption zone 334 via line 337
and is
recovered and perhaps regenerated for recycle. A compressed vaporized light
ends stream
having a reduced concentration of hydrogen sulfide and carbon dioxide is
removed from
absorber 334 via line 338. A condensed naphtha stream is taken from the
receiver 300 in line
304 and pumped by pump 320 in line 312 into line 338. A reflux portion may be
split from
line 312 and be returned to the fractionator 30 in line 42a. Naphtha line 340
carries the
naphtha stream to the hydrotreating reactor 60.
[0033] In one embodiment of the present invention, the naphtha and lighter
stream in line
340 which may include light or full range naphtha, liquefied petroleum gas
(LPG) and dry
gas containing hydrogen is introduced into the hydrotreating reactor 60 to
saturate the olefins
and diolefins present therein. The hydrogen present in the dry gas drives the
hydrotreating
reaction. Preferred hydrotreating reaction conditions include a temperature
from 260 C
(500 F) to 426 C (800 F), a pressure of 862 to 2068 kPa (125 to 300 psia), and
preferably
1448 to 2000 kPa (210 to 290 psia) substantially as provided by compressor
310, and a liquid
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hourly space velocity from 0.1 hr-1 to 10 hr-1. The mild pressure in the
hydrotreating reactor
is chosen to just saturate olefins and to avoid hydrodesulfurization of the
organic sulfur in the
naphtha to conserve hydrogen. If, however, sufficient hydrogen is in the dry
gas component
of the naphtha stream, pressure can be increased to hydrodesulfurize the
naphtha stream if
desired.
[0034] Suitable hydrotreating catalyst for use in the present invention are
any known
conventional hydrotreating catalysts and include those which are comprised of
at least one
Group VIII metal, preferably iron, cobalt and nickel, more preferably cobalt
and/or nickel and
at least one Group VI metal, preferably molybdenum and tungsten, on a high
surface area
support material, preferably alumina. Other suitable hydrotreating catalysts
include zeolitic
catalysts, as well as noble metal catalysts where the noble metal is selected
from palladium
and platinum. It is within the scope of the present invention that more than
one type of
hydrotreating catalyst be used in the same reaction vessel. Two or more
catalyst beds and one
or more quench points may be utilized in the reaction vessel or vessels. The
Group VIII metal
is typically present in an amount ranging from 2 to 20 weight percent,
preferably from 4 to 12
weight percent. The Group VI metal will typically be present in an amount
ranging from 1 to
weight percent, preferably from 2 to 25 weight percent.
[0035] The resulting effluent from the hydrotreating reactor 60 in line 350
with a lower
concentration of olefins than in the stream in line 340 is preferably
contacted with an aqueous
20 stream from line 352 to dissolve any ammonium salts and partially condense
the
hydrotreating effluent. The hydrotreated effluent in line 350 is then
introduced into a high
pressure vapor-liquid separator 62 operated at a pressure substantially equal
to the
hydrotreating reactor and a temperature in the range from 38 C (100 F) to 71 C
(160 F). An
aqueous hydrotreated naphtha stream is recovered from the vapor-liquid
separator 62 in line
25 200 and delivered to line 500 for blending, preferably after dewatering
(not shown). Line 500
may be a conduit that carries the second crude stream from line 499 that has
been minimally
processed or not processed. A hydrogen-rich dry gas stream is removed from the
vapor-liquid
separator in line 354. The dry gas in line 354 may be delivered to the fired
heater 20 via line
210 and/or by line 96 to the CO boiler 90.
[0036] In an additional embodiment, shown in FIG. 1, a portion of the dry gas
stream
may be optionally split off of line 354 in line 356 regulated by a control
valve, compressed in
compressor 358 and recycled in line 360 to line 340 feeding the naphtha to the
hydrotreating
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reactor 60. The recycle gas compressor may increase the average hydrogen
purity in the
hydrotreating reactor 60 and further increase catalyst life.
BLENDED PRODUCT
[0037] As shown in FIG. 1, the separate conversion products; hydrotreated
naphtha and
lighter products in line 200, LCO in line 202 and optionally heavy naphtha in
line 201 are
combined in line 500 where they combine with minimally processed or
unprocessed second
crude stream from line 499, thus forming a blended stream 502, or a synthetic
product. The
second crude stream may be supplied directly from the oilfield, but more
preferably may be
stripped to remove light hydrocarbons and dewatered. In an alternate
embodiment, a portion
of one or more of the conversion products is taken off as a side-product and
further treated or
processed as a saleable commodity. If this option is desired, a greater
portion of the feed will
need to be processed in the FCC riser 40 to make up for a loss of low
viscosity material for
blending.
[0038] Liquid products may include bottoms, light cycle oil, hydrotreated
naphtha, and
perhaps unhydrotreated heavy naphtha and the portions of each one may be
selected to
combine with the unprocessed crude to achieve desired flow properties. The
minimally or
unprocessed second crude stream may be a portion of the crude source that was
not FCC
processed. Specifically, all liquid streams may be combined with the second
crude stream.
The blended stream in line 502 may have the following characteristics, 18 API
or greater,
preferably at least 19 API, more preferably greater than 19.5 API. The blended
stream may
have a viscosity at 100 F (38 C) of no more than 10,000 cSt, preferably no
more than 5000
cSt, and more preferably no more than 25 cSt. The blended stream may also have
a pour
point of no more than 20 C, preferably no more than 15 C, and more preferably
no more than
0 C. The blended stream may then be pumped in a pipeline 502 to a remote
location for
further processing such as in a refinery or a distribution station. A remote
location is typically
greater than 20 miles away from the well in the oil field 1.
CATALYST REGENERATION
[0039] As shown in FIG. 1, the spent catalyst separated from products by
cyclones 52 fall
downwardly into a bed and are stripped of hydrocarbons by steam in stripper 54
and
delivered via spent catalyst conduit 55 regulated by a valve to a regenerator
70 in
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downstream communication with the reactor vessel 50. In the regenerator, 70
coke is burned
off of the surface of the spent catalyst to produce a fresh or regenerated
catalyst. Air is
pumped from line 72 by blower 73 and enters the bottom of regenerator 70 to
burn the coke
at a temperature of between 900 and 1600 F (between 482 and 871 C),
preferably between
1000 and 1400 F (between 538 and 760 C), more preferably between 1200 and
1300 F
(between 649 and 704 C). Regenerator 70 may regenerate catalyst at between
1100 and
1500 F (between 593 and 896 C), preferably between 1200 and 1400 F
(preferably
between 649 and 760 C), more preferably between 1220 and 1350 F (between 660
and
732 C).
[0040] After the coke has been substantially burned off, the spent catalyst
becomes
regenerated catalyst again. The carbon that has been burned off makes up
regeneration flue
gas containing H2, CO, C02, and light hydrocarbons. Cyclones 75 separate
regenerated
catalyst from the regeneration flue gas. Regenerated catalyst may be returned
to riser 40 in
downstream communication with the regenerator vessel 70 via regenerated
catalyst conduit
74 to contact incoming crude feed in line 32.
[0041] The regeneration flue gas may be carried out of regenerator 70 by flue
line 56 and
into CO boiler 90. The CO/CO2 ratio in the regeneration flue gas in stream 56
may be
between 0.3:1 and 1.5:1, preferably between 0.7:1 and 01.25:1, more preferably
1:1. Running
regenerator 70 in partial burn is most appropriate for use with heavy
residuals where
regenerator heat release and air consumption are high due to high coke yield.
In addition,
oxygen-lean regeneration offers improved catalyst activity maintenance at high
catalyst
vanadium levels, due to reduced vanadium mobility at lower oxygen levels. By
running
regenerator 70 in deep partial burn to maximize the CO yield the unit will
limit the amount of
heat that could be released if the carbon were allowed to completely burn to
C02. This will
lower the regenerator temperature and permit a higher catalyst to oil ratio.
[0042] The heating value of the CO-containing gas may be low due to dilution
with much
nitrogen, therefore for efficient burning an auxiliary fuel such as dry gas is
optionally injected
in line 96 with air in line 95 to promote combustion and heat the burning zone
to a
temperature at which substantially all CO is oxidized to CO2 in CO boiler 90.
In the CO
boiler 90 the regeneration flue gas reaches temperatures of at least 1500 F
(815 C),
preferably at least 1700 F (926 C), and more preferably at least 1800 F (982
C). The
combustion in the CO boiler 90 heats and vaporizes water fed by water line 99
to generate
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high pressure superheated steam which leaves CO boiler through steam line 101
for use in the
FCC complex. The regeneration flue gas containing CO2 leaves the CO boiler 90
and is
released to the stack 102. An alternative auxiliary fuel may comprise
clarified oil diverted
from line 203 in line 205.
[0043] In addition to running the regenerator 70 in deep partial burn,
additional heat may
be removed from the regenerator 70 through the operation of a catalyst cooler
on the
regenerator 70. The regenerator may be equipped with between 1 and 5 catalyst
coolers, more
preferably 2 and 4 catalyst coolers 71, and more preferably 3 catalyst
coolers. Catalyst
coolers may remove heat through steam generation. The steam from the catalyst
coolers 71
may be delivered via line 94 to the CO boiler 90 to be superheated in the CO
boiler.
DEBUTANIZER
[0044] In a further embodiment, the naphtha stream may be directed to a
debutanizer to
form liquefied petroleum gas (LPG) and gasoline. The LPG and the gasoline may
be added to
the unprocessed crude, in selected amounts to achieve desired flow properties.
The ability to
modify the relative amounts of light hydrocarbons (propane through pentane) in
the blended
pipeline crude is advantageous because it may be held in tankage and therefore
subjected to a
still further specification of Reid vapor pressure (RVP) to minimize the boil-
off of material at
ambient conditions which may violate environmental regulations, cause material
loss to
flaring or require expensive vapor recovery systems. LPG addition to the
unprocessed crude
must be gauged to balance vapor pressure and flow properties.
[0045] The embodiment of FIG. 2 includes a debutanizer 600 in downstream
communication with the fractionator 30 to separate naphtha from the LPG and
lighter
material stream. Many of the elements in FIG. 2 have the same configuration as
in FIG. 1 and
bear the same reference number. Elements in FIG. 2 that correspond to elements
in FIG. 1 but
have a different configuration bear the same reference numeral as in FIG. 1
but are marked
with a prime symbol ('). Every element upstream of the compressor 310 and pump
320 is the
same as in FIG. 1 and the foregoing description is applicable in FIG. 2. FIG.
2 does not show
the heavy naphtha cut and pump around with line 44, pump 45, steam generator
47 and line
44a in phantom as in FIG. 1. The naphtha and lighter hydrocarbons stripped
from the crude
oil may leave upper portion 39 of fractionator 30 in line 42. It is also
contemplated in this
embodiment that full range naphtha be withdrawn in line 42 without heavy
naphtha being
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separately withdrawn in line 44. However, the embodiment will be described
with only the
light naphtha product boiling at or below 135 to 177 C (275 to 350 F) being
withdrawn in
line 42 and heavy naphtha withdrawn in line 44. The naphtha portion of the
cracked stream in
line 42 may be condensed by a condenser 41 and an optional boiler feed water
heater 43
before it is directed to overhead receiver 300. Water is decanted from the
receiver 300 in line
302 while vaporous wet gas is separated in line 306 from unstabilized naphtha
liquid in line
304. The wet gas is pressurized in compressor 310 to the hydrotreating
pressure previously
mentioned. In an embodiment, the compressed wet gas stream in line 324' is
diverted in line
325 because valve 327 is closed while the valve on line 325 is open. The
compressed wet gas
stream is then cooled in a heat exchanger 326 and flashed in a flash drum 328.
A liquid
stream from the flash drum 328 in line 330' is fed to the receiver 300 while a
vapor stream is
removed in line 338'. Unstabilized naphtha in line 304 is pumped by pump 320
in line 312'.
Because valve 313 is closed and the valve on line 311 is open, line 311
diverts the stream into
line 338' to provide a mixed stream in line 314. The stream in line 314 is
split between line
402 which transports the mixed stream to a debutanizer column 600 and line 220
which may
send naphtha to line 500 for blending. A portion of the unstabilized naphtha
is refluxed to the
fractionator column 30 via line 42a.
[0046] In the debutanizer column 600, a portion of the cracked stream
comprising
naphtha is subjected to fractionation to separate LPG from naphtha.
Fractionation yields a C4-
overhead in overhead line 602 which is condensed in condenser 606 with the
production of
steam and dewatered in receiver 608. The condensed LPG is pumped and split
between reflux
line 610 which is returned to the debutanizer 600 and LPG line 612. The LPG
line 612 feeds
a blend line 614 which blends LPG with the processed products in line 500 and
an optional
product line 616 which recovers LPG as product which may be stored and/or sold
locally.
LPG is an excellent cutter component, but because of its high vapor pressure
can be blended
only up to the flash specification. Hence, the split between lines 610 and 612
and 614 and
616 should be set to maximize the LPG blended in line 500 up to the flash
specification. Any
excess can be captured and sold as LPG perhaps after further stripping of dry
gas therefrom
or used in the fired heater 20 or the CO boiler 90. A dry gas stream 618 from
the receiver 608
may then be fed to a hydrogen purification unit 700 in downstream
communication with the
fractionator 30 and an overhead line 602 of the debutanizer 600. The dry gas
stream in dry
gas line 618 contains hydrogen and may be considered a hydrogen stream. The
debutanizer
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column 600 also produces a bottoms stream in bottoms line 604 typically
comprising C5+
material. The bottoms stream 604 is split into several streams. A reboil line
620 is heated by
reboiler 622 and returned to the debutanizer column 600. A naphtha feed stream
in line 340'
transports naphtha to the hydrotreating reactor 60 which is in downstream
communication
with the bottoms line 604 of the debutanizer 600. A portion of the naphtha
stream may be
split off in line 626 and recovered as product in line 626 to be stored and/or
sold locally.
ABSORBER
[0047] An alternative embodiment, shown in phantom in FIG. 2, utilizes an
absorber 400
in downstream communication with the fractionator 30 to separate a naphtha
portion of the
cracked stream into a C3+ naphtha stream and a dry gas stream. The compressed
wet gas in
line 324' may continue on in line 324a through an open valve 327 and is fed to
the bottom of
the absorber 400 instead of proceeding in line 325 because the control valve
on line 325 is
closed. Similarly, the unstabilized liquid naphtha is pumped in line 312'
which may continue
on in line 312a through an open valve 313 to a top of the absorber 400 because
the valve on
line 311 is closed. In this embodiment which utilizes the absorber 400,
streams in lines 324'
and 312' are not combined and fed to the debutanizer via line 314 but are kept
separate. In
the absorber 400, the unstabilized liquid naphtha absorbs liquefied petroleum
gas (LPG) from
the wet gas and exits the absorber 400 in a bottoms line 401 comprising C3+
naphtha. The
absorbent line is split between product line 220 for delivering C3+ to line
500 for blending
and a debutanizer feed line 402. The debutanizer 600 is in downstream
communication with
the bottoms line 401 of the absorber 400 via line 402. Additionally, an
optional naphtha
recycle stream in line 624 from the bottoms of the debutanizer may be recycled
to the
absorber 400 to recover more LPG. In a further embodiment, a portion or all of
the heavy
naphtha in line 201 may be diverted via line 503 to the naphtha recycle line
624 to
supplement the naphtha feed to the absorber 400 and increase the recovery of
LPG in line
401. A dry gas stream with less LPG than in the wet gas in line 324 comprising
C2-, H2S and
H2 exit the absorber 400 in an overhead line 404. The dry gas stream in line
404 flows
through an open control valve 405 to join dry gas stream 618 and provide
combined dry gas
stream in line 406. The dry gas streams in overhead line 404 and dry gas lines
618 and 406
contain hydrogen and may be considered hydrogen streams. Dry gas stream
containing
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hydrogen is carried by dry gas line 406 to the hydrogen purification unit 700
in downstream
communication with the overhead line 404 of the absorber 400.
DRY GAS PURIFICATION
[0048] In an embodiment, the hydrogen in the dry gas stream may be purified
before it is
used for hydrotreating to increase the hydrotreating catalyst life. Several
types of hydrogen
purification units may be suitable.
[0049] The dry gas in line 406 may be fed to an amine absorber 334 to remove
hydrogen
sulfide and carbon dioxide. A lean aqueous amine scrubbing solution is
introduced into
absorber 334 via line 336 and scrubs hydrogen sulfide and carbon dioxide from
the dry gas
stream. A rich aqueous amine scrubbing solution containing hydrogen sulfide is
removed
from absorber 334 via line 337 and is recovered and perhaps regenerated for
recycle. A dry
gas stream with a smaller concentration of hydrogen sulfide and carbon dioxide
than in line
406 is removed from absorber 334 via line 408.
[0050] The dry gas stream in line 408 at a pressure determined by compressor
310 that
will be adequate for hydrogen purification while sufficiently above dew point
to maintain a
gaseous state is fed to the hydrogen purification unit 700. The hydrogen
purification unit 700
may be a pressure swing adsorption system 750 shown in FIG. 3. Other types of
hydrogen
purification units may be suitable. The pressure of the dry gas in line 408
may be between
862 and 2068 kPa (125 and 300 psia).
[0051] In an embodiment, the hydrogen in the dry gas can be purified in a
pressure swing
adsorption (PSA) unit 750 shown in FIG. 3 to provide a hydrogen rich gaseous
stream having
a reduced concentration of carbon oxides, methane and ethane. The pressure
swing
adsorption process provides a well established means for separating and
purifying hydrogen
from a feed gas mixture of larger molecules. The process provides adsorption
of the
adsorbable species, such as carbon oxides, water and light hydrocarbon
molecules, on an
adsorbent at a high adsorption pressure with passage of the smaller hydrogen
molecules and
pressure reduction to a lower desorption pressure to desorb the adsorbed
species. It is
generally desirable to employ the PSA process in multiple bed systems such as
those
described in US 3,430,418, herein incorporated by reference, in which at least
four adsorption
beds are employed. The PSA process is carried out in such systems on a
cyclical basis,
employing a processing sequence. Referring to FIG. 3, the PSA unit 750 may
have four beds
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761-764 having inlet ends 761a-764a and outlet ends 761b-764b. Valving is
generally shown
in FIG. 4. In the first step, the dry gas in line 770 in downstream
communication with the dry
gas line 408 is fed to an inlet end 761 a of a first adsorbent bed 761 at high
adsorption
pressure to adsorb adsorbable species onto the adsorbent with passage of
product hydrogen
gas to a discharge end 761b of the bed 761. Purified hydrogen gas may pass
from the PSA
unit 750 through product line 772 with a greater hydrogen purity than in feed
line 770. Feed
flow is terminated to the first bed 761 before the carbon oxides, water and
hydrocarbons
break through to the discharge end 761b of the first bed. Second, the first
bed 761 is
cocurrently depressurized to an intermediate pressure by releasing void space
gas from the
discharge end 761b of the first bed to a discharge end 762b of a second bed
762 thereby
repressurizing the second bed which has just been purged of desorbed carbon
oxides, water
and hydrocarbons. Further cocurrent depressurization of the first bed 761 can
occur by
releasing remaining void space gas to a discharge end 763b of a third bed 763
to purge the
third bed of desorbed carbon oxides, water and hydrocarbons. In a third step,
the inlet 761 a to
the first bed 761 is opened in a countercurrent depressurization or blow down
step, in which
gas departs the first bed through the inlet end 761 a leaving the first bed
761 at sufficiently
low pressure to desorb adsorbed species from the adsorbent. Desorbed species
are released
through the inlet 761 a and recovered in desorbent line 774 with a greater
concentration of
adsorbable species than in the feed line 770. In a fourth step, void space gas
from a fourth bed
764 may be released through a discharge end 764b thereof and fed through the
discharge end
761 a of the first bed 761 to purge out the desorbed species. In a last step,
void space gas from
the second bed 762 is fed from its discharge end 762b into the discharge end
761b of the first
bed 761 to repressurize the first bed. Product gas from the discharge end 763b
of the third bed
763 is then fed into the discharge end 761b of the first bed 761 to achieve
adsorption pressure
in the first bed 761. Since the first bed 761 is now at adsorption pressure,
the cycle in the first
bed begins anew. The same process sequence is operated with the other beds 762-
764, with
differences relating to the position of the bed 762-764 in the order.
[0052] A suitable adsorbent may be activated calcium zeolite A with or without
activated
carbon. If this combination of adsorbents is used, the activated carbon will
adsorb the carbon
dioxide and water, while the zeolite A will adsorb the carbon monoxide and
hydrocarbons.
[0053] Purified hydrogen with a hydrogen concentration greater than in dry gas
line 408
is transported in line 772 which is in upstream communication with the
hydrotreating reactor
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60. The desorbent line 774 containing dry gas with a reduced concentration of
hydrogen
relative to the concentration in line 408 communicates with a waste dry gas
line 210 which
may be delivered to the fired heater 20 or to the CO boiler 90.
[0054] In the embodiment of FIG. 2, the C5+naphtha in line 340' is mixed with
purified
hydrogen in line 772 and introduced into the hydrotreating reactor 60 via line
780 to saturate
the olefins and diolefins present therein. The hydrotreating reactor 60 is in
downstream
communication with the hydrogen purification unit 700. The resulting effluent
from the
hydrotreating reactor 60 in line 350 with a lower concentration of olefins
than in the stream in
line 340' is preferably contacted with an aqueous stream from line 352 to
dissolve any
ammonium salts and partially condense the hydrotreated effluent. The
hydrotreated effluent
in line 350 is then introduced into a high pressure vapor-liquid separator 62
operated at a
pressure substantially equal to the hydrotreating reactor and a temperature in
the range from
38 C (100 F) to 71 C (160 F). An aqueous hydrotreated naphtha stream is
recovered from
the vapor-liquid separator 62 in line 200' and delivered to line 500 for
blending with the
minimally processed or unprocessed second crude stream from line 499. A
hydrogen-rich dry
gas stream is removed from the vapor-liquid separator in line 354'. The dry
gas in line 354'
may be delivered to the fired heater 20 by line 210' and by line 96 to the CO
boiler 90.
[0055] In an additional embodiment, shown in FIG. 2, at least a portion of the
dry gas
stream may be optionally split off of line 354' in line 356' regulated by a
control valve,
compressed in compressor 358 and recycled in line 360' to mix with purified
hydrogen from
line 772 and naphtha in line 340' to feed the hydrotreating reactor 60 via
line 780.
OLIGOMERIZATION
[0056] In an additional option, control valve 630 is opened to allow LPG in
recovery line
612 to flow through line 632 to an oligomerization reactor 80. The olefin
containing LPG
stream in recovery line 632 has C3 and C4 olefins that can be oligomerized
into heavier
naphtha molecules. The diolefins in the LPG stream in line 632 are first
reacted with a
selective hydrogenation catalyst in selective hydrogenation zone 78 to
selectively saturate
diolefins without completely saturating them to paraffins. Hydrogen may be
provided from
the hydrogen purification zone 700 by line 782 diverging from hydrogen stream
in line 772
regulated by a control valve. Suitable conditions for operation of a selective
hydrogenation
process are described, for example, in US 6,166,279 and US 6,075,173. Such
conditions
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include passing the LPG stream in the liquid phase in the presence of hydrogen
at molar ratio
0.5 to 5 moles hydrogen per mole of diolefin over a catalyst comprising at
least one metal
selected from the group formed by nickel, palladium and platinum, deposited on
a support
such as aluminum oxide, at a temperature of 20 to 200 C (68 to 392 F), a
pressure of 689
to 3447 kPa(g) (100 to 500 psig), and a space velocity of 0.5 to 10 hr-1. Two
or more reaction
zones may be used although only one is shown. Each reaction zone may employ a
recycle of
reactor effluent to the reactor inlet with a ratio of recycle to fresh
olefinic feed stream ranging
from 0 to 20. The residual diolefin content of such a process can be in the
range 1 to 100
wppm, depending on the severity of the operation.
[0057] The LPG effluent from the selective hydrogenation reactor in line 79
with a
diolefin concentration that is less than in line 632 may be mixed with none,
one, some or all
of a paraffinic diluent in line 230, a selectivity modifier that may enter
through process line
81, an effluent recycle stream in recycle line 82 and a LPG recycle stream in
line 83 to form
an oligomerization reactor feed in feed line 84 that is then fed to a
oligomerization reactor 80.
The feed line 84 is in downstream communication with the overhead line 602 of
the
debutanizer 600 via line 632. The paraffinic diluent in line 230 may be a
portion of
hydrotreated naphtha from line 200'. In the oligomerization reactor 80, LPG is
contacted with
an oligomerization catalyst at oligomerization conditions to oligomerize the
lighter olefins to
produce heavier olefins in the naphtha range.
[0058] Conditions for the operation of a oligomerization process include
passing the LPG
liquid over a catalyst such as SPA or a sulfonic acid ion exchange resin such
as Amberlyst
A-15, A-35, A-16, A-36, Dowex 50 or the like. Several means can be used to
restrict the
formation of dodecene and higher oligomers. These include addition of a
paraffinic diluent to
the oligomerization reactor when SPA catalyst is used, recycle of a portion of
the
oligomerization reactor effluent to the oligomerization reactor feed stream
and addition of 0.1
to 3.0 wt-% oxygenated selectivity modifier to the oligomerization reactor
when resin
catalyst is used. Since, this oligomerization may occur in the field where
process streams are
less available and because the process is only making cutter stock that will
be refined at a
downstream refinery, avoiding heavy olefin production is not critical.
Additionally, if heavier
oligomers are desired to conserve hydrogen in the hydrotreating reactor 60,
none of the
measures to avoid heavy oligomerization need be taken.
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[0059] The preferred operating conditions applicable when an SPA catalyst is
used differ
from those when an ion exchange resin catalyst is used. Preferred temperatures
for operation
with an SPA catalyst are in the range 40 to 260 C, and more typically in the
range 75 to
230 C, while preferred temperatures for operation with an ion-exchange resin
catalyst are in
the temperature range 0 to 200 C, and more typically in the range 40 to 150
C. Preferred
pressures for operation with an SPA catalyst are in the range 689 to 8274
kPa(g) (100 to 1200
psig), and more typically in the range 1379 to 6895 kPa(g) (200 to 1000 psig),
while
preferred pressures for operation with an ionic resin catalyst are in the
range 345 to 3447
kPa(g) (50 to 500 psig), and more typically in the range 1379 to 2413 kPa(g)
(200 to 350
psig). These pressures may be kept in the lower end of the range, so an
additional compressor
is not required to boost the pressure above the system pressure needed for the
hydrotreating
reactor 70. A preferred space velocity range for operation with SPA catalyst
is 0.5 to 5 hr_'
and for operation with an ion-exchange resin catalyst is 0.3 to 20 hr_'
depending on the
properties of the oligomerization reactor feed such as olefin content and
type.
[0060] An oligomerization reactor product is withdrawn from oligomerization
reactor 80
through effluent line 85. A portion of the oligomerization reactor effluent
may be recycled to
the oligomerization reactor feed through recycle line 82 to control the
exotherm. A second
portion of the oligomerization reactor product is passed through process line
86 to a flash
drum 800, in which an unreacted LPG vapor stream and an oligomerization
product rich
liquid stream are formed. The LPG vapor stream leaves flash drum 800 in vapor
line 87 for
further processing. A portion of vapor stream in line 87 may be recycled by
line 83 to the
oligomerization reactor 80 after condensing and compression while the
remaining stream is
processed through line 634 to be mixed with crude via line 500. A portion of
the LPG may be
recovered in line 636 if desired. The oligomerization product-rich liquid
stream containing
naphtha range molecules is sent through process line 89 to join naphtha in
line 340' in route
to the hydrotreating reactor 60 in downstream communication with effluent line
85 of the
oligomerization reactor 80 to saturate the olefins.
[0061] It is also contemplated that if higher hydrogen requirements are
necessary, that a
steam reformer may be used to convert hydrocarbons in dry gas streams into
hydrogen gas.
All LPG and dry gas streams would be feed candidates to a steam reformer for
hydrogen
production.
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EXAMPLE
[0062] We simulated the operation of the process of the present invention on
the basis of
charging 2,385 m3/d (15,000 bbl/d) of crude to the FCC unit. The properties of
the feed
simulated are in Table 1.
Table 1
API 12.8
UOP K 11.4
Nickel, wppm 42.0
Vanadium, wppm 152.0
Sulfur, wt-% 1.28
Conradson Carbon, wt-% 12.88
[0063] The cracked stream from the FCC unit had the composition in Table 2
expressed
in weight percentages.
Table 2
Hydrogen Sulfide 0.41
Hydrogen 0.44
Methane 1.00
Ethylene 0.86
Ethane 0.85
Propylene 3.40
-Propane 0.96
Butylenes 4.44
Isobutane 1.72
Normal Butane 0.59
Light Naphtha C5-164 C 327 F 21.65
LCO and Heavy Naphtha 33.61
[0064] The properties of the light naphtha in the cracked stream are given in
Table 3.
Table 3
API 62.9
Sulfur, wt-% 0.04
Paraffins/Olefins/Naphthenes/Aromatics, wt-% 42/24/12/22
Bromine Number 39.8
IBP/EP, ASTM, C F 46/164 115/327
[0065] Case 1 is the embodiment of FIG. 2 with the valve 630 closed to the
oligomerization zone and valves 313 and 327 closed to the absorber 400. The
hydrogen
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production assumed that the PSA hydrogen purification unit would retain 86 wt-
% of the
hydrogen in dry gas feed. Additionally, all of the dry gas from the
debutanizer is fed to the
PSA unit. Table 4 gives the hydrogen balance for Case 1.
Table 4
Hydrogen from PSA unit, kg/hr (lbs/hr) 368 812
Naphtha hydrotreater demand, kg/hr (lbs/hr) 255 (561)
Excess hydrogen, kg/hr (lbs/hr) 114 251
[0066] In Case 1, a surplus of hydrogen exists to saturate the olefins in the
naphtha
stream.
[0067] In Case 2, valve 630 is opened, so all of the LPG in line 612 is fed to
the
oligomerization reactor 80. Additionally, absorber 400 was utilized and all of
the dry gas in
the absorber overhead and the debutanizer overhead was fed to the PSA unit.
Again, the
hydrogen production assumed that the PSA hydrogen purification unit would
retain 86 wt-%
of the hydrogen in dry gas feed. Table 5 gives the hydrogen balance for Case
2.
Table 5
Hydrogen from PSA unit, kg/hr (lbs/hr) 368 812
Naphtha hydrotreater demand, kg/hr (lbs/hr) 255 (561)
Additional naphtha hydrotreater demand for 82(181)
olefinic oligomers, kg/hr (lbs/hr)
Selective hydrotreater demand, kg/hr (lbs/hr) 3(6)
Excess hydrogen, kg/hr (lbs/hr) 29 64
[0068] Even when additional hydrogen is required to saturate diolefins in the
selective
hydrotreater and to saturate the olefinic oligomers from the oligomerization
reactor in the
naphtha hydrotreater, the dry gas in the cracked stream still provides
sufficient hydrogen to
saturate all the olefins in the naphtha stream.
[0069] The existence of excess hydrogen indicates the naphtha cut point can be
adjusted
to allow heavier naphtha into the cracked stream in line 42. In both cases,
the total olefinic
concentration in line 500 is less than 0.1 wt-%. After the second crude stream
is added to the
first processed crude stream the olefin concentration will be decreased even
further.
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