U.S. patent application number 10/574671 was filed with the patent office on 2008-11-06 for intergrated ngl recovery and lng liquefaction.
This patent application is currently assigned to Fluor Technologies Corporation. Invention is credited to John Mak.
Application Number | 20080271480 10/574671 |
Document ID | / |
Family ID | 37215177 |
Filed Date | 2008-11-06 |
United States Patent
Application |
20080271480 |
Kind Code |
A1 |
Mak; John |
November 6, 2008 |
Intergrated Ngl Recovery and Lng Liquefaction
Abstract
Contemplated plants include a refluxed absorber and a
distillation column, wherein the absorber is operated at a higher
pressure than the distillation column to thereby produce a
cryogenic pressurized lean gas. The lean gas is further compressed
to a pressure suitable for liquefaction using energy from feed gas
vapor expansion. Desired separation of C2 products is ensured by
temperature control of the absorber and distillation column using
flow ratios of various streams within the plant, and by dividing
the separation process into two portions at different
pressures.
Inventors: |
Mak; John; (Santa Ana,
CA) |
Correspondence
Address: |
FISH & ASSOCIATES, PC;ROBERT D. FISH
2603 Main Street, Suite 1050
Irvine
CA
92614-6232
US
|
Assignee: |
Fluor Technologies
Corporation
Aliso Viejo
CA
|
Family ID: |
37215177 |
Appl. No.: |
10/574671 |
Filed: |
March 14, 2006 |
PCT Filed: |
March 14, 2006 |
PCT NO: |
PCT/US2006/009103 |
371 Date: |
July 15, 2008 |
Related U.S. Patent Documents
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Application
Number |
Filing Date |
Patent Number |
|
|
60673518 |
Apr 20, 2005 |
|
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|
Current U.S.
Class: |
62/626 ;
62/620 |
Current CPC
Class: |
F25J 1/0239 20130101;
F25J 2200/04 20130101; F25J 2230/08 20130101; F25J 2200/78
20130101; F25J 2270/90 20130101; F25J 2240/02 20130101; F25J
2230/60 20130101; F25J 2200/74 20130101; F25J 2280/02 20130101;
F25J 1/0035 20130101; F25J 2290/10 20130101; F25J 3/0233 20130101;
F25J 1/0022 20130101; F25J 2205/04 20130101; F25J 1/0255 20130101;
F25J 3/0209 20130101; F25J 3/0238 20130101 |
Class at
Publication: |
62/626 ;
62/620 |
International
Class: |
F25J 3/00 20060101
F25J003/00; F25J 1/00 20060101 F25J001/00 |
Claims
1. A plant for natural gas liquids recovery in which the plant is
coupled with an LNG liquefaction plant, comprising an absorber
configured to receive an absorber feed stream and a first and a
second reflux stream, and further configured to provide a bottom
product stream; a distillation column configured to receive a first
portion of the bottom product stream and a second portion of the
bottom product stream at different points, and wherein the
distillation column is further configured to operate a pressure
that is lower than an operating pressure in the absorber; and a
control unit that is configured to control a flow ratio of (a) the
feed stream to the second reflux stream and (b) the first portion
of the bottom product stream to the second portion of the bottom
product stream, wherein the flow ratio is a function of desired
ethane recovery in the distillation column bottom product
stream.
2. The plant of claim 1 further comprising at least one of a heat
exchanger and a reflux condenser that are configured to heat the
first portion of the bottom product stream, and still further
comprising an expansion device configured to cool the second
portion of the bottom product stream.
3. The plant of claim 1 wherein the distillation column is
configured to produce a distillation column overhead, and wherein
the plant further comprises a compressor that compresses the
distillation column overhead to at least absorber pressure.
4. The plant of claim 3 further comprising a cooling device
thermally coupled to the distillation column overhead and
configured to cool the compressed distillation column overhead.
5. The plant of claim 4 wherein the cooled compressed distillation
column overhead is the first reflux.
6. The plant of claim 1 wherein the absorber is configured to
produce an absorber overhead product that has a temperature of
equal or lower than -90.degree. F. and a pressure of between 500
psig and 700 psig.
7. The plant of claim 6 further comprising a compressor that is
configured to receive the absorber overhead product and to compress
the absorber overhead product to a pressure of at least 800
psig.
8. The plant of claim 7 wherein the compressor is operationally
coupled to an expander that expands the absorber feed stream.
9. A method of processing a gas for delivery to an LNG liquefaction
plant, comprising: providing an absorber that receives an absorber
feed stream and a first and a second reflux stream, and that
produces a bottom product stream; fluidly coupling the absorber to
a distillation column such that a first portion of the bottom
product stream and a second portion of the bottom product stream
are fed to the distillation column at different points; operating
the distillation column at a pressure that is lower than an
operating pressure of the absorber; and controlling a flow ratio of
(a) the feed stream to the second reflux stream and (b) the first
portion of the bottom product stream to the second portion of the
bottom product stream as a function of desired ethane recovery in
the distillation column bottom product stream.
10. The method of claim 9 further comprising a step of feeding a
distillation column overhead product to the absorber.
11. The method of claim 10 wherein the distillation column overhead
product is compressed, cooled, and fed to the absorber as the first
reflux stream.
12. The method of claim 9 wherein the distillation column is
operated at a pressure between 300 psig and 500 psig, and wherein
the absorber is operated at a pressure of between 500 psig and 800
psig.
13. The method of claim 9 further comprising a step of separating a
cooled feed gas into a liquid portion and a vapor portion, and
feeding the liquid portion after at least partial depressurization
and warming into the distillation column.
14. The method of claim 13 wherein the vapor portion is split into
a first and second stream to thereby form the second reflux stream
and the absorber feed stream.
15. The method of claim 9 wherein the absorber produces a cryogenic
absorber overhead stream, and further comprising a step of
compressing the cryogenic absorber overhead stream to a pressure
suitable for liquefaction.
16. The method of claim 15 wherein the step of compressing is
driven by expansion of the absorber feed stream.
17. A method of variably recovering C2 from a feed gas to a LNG
liquefaction plant comprising: feeding an expanded and heated
liquid portion of a feed gas to a distillation column and feeding a
vapor portion of the feed gas to an absorber; adjusting a flow
ratio of an absorber feed to a second reflux to the absorber, and
using a first reflux that is provided by a distillation column
overhead-product to thereby control an absorber overhead
temperature; adjusting a temperature of an absorber bottom product
that is fed to the distillation column to thereby control a
distillation column overhead temperature; and operating the
absorber at a higher pressure than the distillation column.
18. The method of claim 17 wherein the step of adjusting the
absorber bottom product temperature is performed by heating at
least one portion of the absorber bottom product in a heat
exchanger.
19. The method of claim 17 wherein the step of adjusting the
absorber bottom product temperature is performed by cooling at
least another portion of the absorber bottom product using a JT
valve.
20. The method of claim 17 wherein the step of adjusting the flow
ratio of the absorber feed to the second reflux to the absorber is
a function of desired C2 recovery.
Description
[0001] This application claims priority to our copending U.S.
provisional patent application with the Ser. No. 60/673,518, which
was filed Apr. 25, 2005.
FIELD OF THE INVENTION
[0002] The field of the invention is natural gas liquids (NGL)
recovery and LNG liquefaction, and particularly integrated plant
configurations for same.
BACKGROUND OF THE INVENTION
[0003] While the crude oil supply in the world is diminishing, the
supply of natural gas is still relatively abundant in many parts of
the world. Natural gas is typically recovered from oil and gas
production wells located onshore and offshore. Most typically,
natural gas predominantly comprises C1 (methane). Depending on the
particular formations and reservoirs, natural gas also contains
relatively low quantities of non-methane hydrocarbons, including
C2, C3, C4, C5, and heavier components. Still further components of
natural gas include water, nitrogen, carbon dioxide, hydrogen
sulfide, mercaptans, and other gases.
[0004] Natural gas from wellheads is commonly treated and
processed, and transported to gas processing plants in high
pressure transmission pipelines. However, and especially in remote
locations without the necessary pipeline infrastructure, natural
gas is commonly transported by liquefying the natural gas and
moving the liquefied gas (e.g., using LNG cargo carriers).
Unfortunately, direct liquefaction of natural gas is often
problematic as natural gas often contains C5, aromatics, and
heavier hydrocarbons, which solidify when cooled to cryogenic
temperatures. Consequently, such heavier components must be removed
to a relatively low level (typically less than 1 ppmv) to avoid
solidification and ultimately plugging of the cryogenic heat
exchange equipment. Additionally, lighter hydrocarbons such as C2,
C3, and C4 must also be removed at least to some degree for the
north American market, which typically requires the natural gas to
meet a heating value of between 1050 to 1070 Btu/SCF. There are
also economic incentives to extract components as they can be sold
at a premium price over natural gas. For example, C2 can often be
used as a feedstock for petrochemical manufacture, while C3 and C4
can be sold as LPG fuels and C5+ hydrocarbons can be used for
gasoline blending.
[0005] There are numerous configurations and methods known in the
art for C2 and C3+ NGL recovery from a natural gas feed. However,
past efforts have focused on removal of the NGL hydrocarbons from
natural gas using standalone NGL recovery plants, which operate
independently from LNG liquefaction plants. These processes
generally produce a relatively low pressure residue gas at ambient
temperature, which would necessitate re-compression and re-cooling
of the residue gas in an LNG liquefaction plant. Typical examples
include various expander processes described in U.S. Pat. Nos.
4,157,904 and 5,275,005 to Campbell et al., U.S. Pat. No. 4,251,249
to Gulsby, U.S. Pat. No. 4,617,039 to Buck, U.S. Pat. No. 4,690,702
to Paradowski et al., U.S. Pat. No. 5,799,507 to Wilkinson et al.,
and U.S. Pat. No. 5,890,378 to Rambo et al.
[0006] For high C2 recovery, some configurations as described in
U.S. Pat. No. 6,116,050, require letting down a portion of the
residue gas compressor discharge to the NGL recovery column as a
methane rich reflux using the Joule-Thomson (JT) valve. While these
processes can improve C2 recovery to some extent, additional
residue gas compressor horsepower is required, which may render the
process costly to operate. There are also more recent advances in
the C2 and C3 recovery area (see e.g., commonly owned U.S. Pat. No.
6,837,070) in which a high pressure absorber is coupled with a
lower pressure distillation column to improve NGL recovery
efficiency. However, these NGL processes are designed for either
high C2 or C3 recovery, and generally not designed for varying
levels of C2 recovery without lowering C3 recovery. Thus, in most
cases, standalone NGL recovery plants are used to produce a low
pressure and ambient temperature residue gas that requires
re-compression and re-cooling in the LNG liquefaction plant,
thereby duplicating many of the refrigeration and heat exchange
equipment in the NGL recovery plant.
[0007] In other known approaches, attempts were made to include the
NGL recovery process as part of the LNG liquefaction plant, as
disclosed in the U.S. Pat. No. 6,401,486 to Lee et al. and U.S.
Pat. No. 6,662,589 to Roberts et al. Lee et al. teaches that a
methane rich stream can be used as reflux to the NGL recovery
column that is coupled with an overhead condenser in another NGL
column to achieve an propane recovery of 95%. However this
configuration requires the NGL column to operate at a pressure of
450 psig or even lower as the separation of NGL becomes
increasingly difficult at the higher pressures due to the
correspondingly reduced relative volatility. Consequently, these
processes require significant recompression from the column
overhead to the required LNG liquefaction pressure, typically from
450 psig to about 800 psig to 900 psig.
[0008] In still further known examples, Roberts et al in U.S. Pat.
No. 6,662,589 teach a C2 rich liquid being recycled from the NGL
fractionation unit that is used for C3 absorption in a high
pressure absorption column. While this process attempts to operate
the NGL column at a high pressure (e.g., 600 psig), NGL separation
efficiency suffers as the relative volatility of the NGL components
is reduced, which results in recovering significantly less NGL
components, especially C2 components. Without removal of a high
level of C2 and C3 components, the currently known processes cannot
produce a lean natural gas with sufficiently lower heating value
content to meet the North American pipeline specifications in an
economically manner. Additionally, the lean gas pressure from such
known processes would require significant refrigeration in the LNG
liquefaction plant due to the relatively low feed gas pressure (LNG
liquefaction generally requires significantly less refrigeration
duty when operating at a higher pressure, between 800 psig and 900
psig or higher). Additional configurations with similar problems
are described in U.S. Pat. No. 5,685,170 to Sorensen and U.S. Pat.
App. No. 2005/0247078 to Wilkinson et al.
[0009] Thus, while numerous compositions and methods for NGL
recovery are known in the art, all or almost all of them, suffer
from one or more disadvantages. Therefore, there is still a need
for improved NGL recovery, and especially where the NGL plant is
integrated with or coupled to a LNG liquefaction unit.
SUMMARY OF THE INVENTION
[0010] The present invention is directed to configurations and
methods of NGL recovery, preferably coupled to an LNG liquefaction
process, in which recovery of C2 components can be adjusted using
flow ratios of selected process streams. Most preferably, the
absorber in such configurations and methods is operated at
significantly higher pressure than the distillation column to
provide a cryogenic pressurized gas, while the absorber and
distillation column temperatures are adjusted to such that desired
quantities of C2 and C3+ products are recovered in the NGL.
Cryogenic absorber overhead product is then compressed to a
pressure suitable for liquefaction using energy derived from
expansion of a vapor portion of the feed gas.
[0011] In one aspect of the inventive subject matter, a plant
includes an absorber configured to receive an absorber feed stream
and a first and a second reflux stream, and that is further
configured to provide a bottom product stream. A distillation
column is configured to receive a first portion of the bottom
product stream and a second portion of the bottom product stream at
different points, wherein the distillation column is further
configured to operate at a pressure that is lower than an operating
pressure in the absorber. Contemplated plants will further include
a control unit having one or more control valves that are
configured to control flow ratios of (a) the feed stream to the
second reflux stream and (b) the first portion of the bottom
product stream to the second portion of the bottom product stream,
wherein the flow ratios are a function of desired ethane recovery
in the distillation column bottom product stream.
[0012] It is generally preferred that in such plants a heat
exchanger and/or a reflux condenser are configured to heat the
first portion of the bottom product stream, and that an expansion
device cools the second portion of the bottom product stream. It is
also generally preferred that the distillation column is configured
to produce a distillation column overhead, and that the plant
further comprises a compressor that compresses the distillation
column overhead to at least absorber pressure. Where desired, the
so compressed overhead is cooled and used as first reflux.
Typically, the absorber produces an absorber overhead product that
has a temperature of equal or lower than -90.degree. F. and a
pressure of between 500 psig and 700 psig, which may be further
compressed by a compressor to a pressure suitable for liquefaction
of the absorber overhead product (e.g., at least 800 psig). Most
preferably, the compressor is operationally coupled to an expander
that expands the absorber feed stream.
[0013] In another aspect of the inventive subject matter, a method
of processing a gas includes a step of providing an absorber that
receives an absorber feed stream and a first and a second reflux
stream, and that produces a bottom product stream. In another step,
the absorber is fluidly coupled to a distillation column such that
a first portion of the bottom product stream and a second portion
of the bottom product stream are fed to the distillation column at
different points, and in another step the distillation column is
operated at a pressure that is lower than an operating pressure of
the absorber. In yet another step, flow ratios of (a) the feed
stream to the second reflux stream and (b) the first portion of the
bottom product stream to the second portion of the bottom product
stream is controlled as a function of the desired ethane recovery
in the distillation column bottom product stream.
[0014] In especially preferred methods, the distillation column
overhead product is fed to the absorber, most preferably as
compressed and cooled first reflux stream. In further preferred
methods, the distillation column is operated at a pressure between
300 psig and 500 psig, and the absorber is operated at a pressure
of between 500 psig and 800 psig. Typically, a cooled feed gas is
separated into a liquid portion and a vapor portion, and a portion
of the liquid portion is after at least partially depressurized and
heated and fed into the distillation column. The vapor portion in
such methods is preferably split into a first and second stream to
thereby form the second reflux stream and the absorber feed stream.
It is further preferred that the cryogenic absorber overhead stream
is compressed to a pressure suitable for liquefaction using energy
from the expansion of the absorber feed stream.
[0015] Therefore, and viewed from a different perspective, a method
of variably recovering C2 from a feed gas includes a step of
feeding an expanded and heated liquid portion of a feed gas to a
distillation column and feeding a vapor portion of the feed gas to
an absorber. In another step, a flow ratio of an absorber feed to a
second reflux to the absorber is controlled, and a first reflux
that is provided by a distillation column overhead product is used
to thereby control an absorber overhead temperature, and the degree
of C2 recovery. In yet another step, the temperature of the
absorber bottom product that is fed to the distillation column is
adjusted to thereby control distillation column overhead
temperature, controlling the desirable NGL recovery levels. The
absorber is typically operated at a higher pressure than the
distillation column.
[0016] Most typically, the step of adjusting the absorber bottom
product is performed by heating at least one portion of the
absorber bottom product in a heat exchanger, while the step of
adjusting the absorber bottom product is performed by cooling at
least another portion of the absorber bottom product using a JT
valve. In further preferred aspects, the step of adjusting the flow
ratio of the absorber feed to the second reflux to the absorber is
a function of desired C2 recovery.
[0017] Various objects, features, aspects and advantages of the
present invention will become more apparent from the following
detailed description of preferred embodiments of the invention.
BRIEF DESCRIPTION OF THE DRAWINGS
[0018] Prior Art FIG. 1 is a schematic of an exemplary known
plant.
[0019] FIG. 2 is a schematic of an exemplary plant configuration
according to the inventive subject matter.
[0020] FIG. 3 is a graph depicting composite heat curves for heat
exchanger 51 and 54 in a plant according to FIG. 2 for C3
recovery.
[0021] FIG. 4 is a graph depicting composite heat curves for heat
exchanger 51 and 54 in a plant according to FIG. 2 for C2
recovery.
[0022] FIG. 5 is a graph comparing the relative volatilities of the
NGL components at 600 psig between the prior art and a plant
according to FIG. 2 in C2 recovery.
DETAILED DESCRIPTION
[0023] The inventor has discovered that C2 and C3+ components can
be effectively and economically separated from natural gas in a
plant (which is preferably coupled to an LNG liquefaction plant)
using an absorber that operates at high pressure and produces a
cryogenic pressurized lean gas, while a distillation column located
downstream of an absorber operates at low pressure and produces the
NGL as a bottom product and a reflux stream for the absorber. In
especially preferred configurations and methods, recovery of C2 in
the NGL can be adjusted by controlling process streams within the
plant. Moreover, it should be appreciated that the cryogenic
absorber overhead product stream is already at relatively high
pressure, and energy for recompression to a pressure suitable for
liquefaction is typically provided by expansion of a vapor portion
of the natural gas feed stream.
[0024] In contrast, as depicted in Prior Art FIG. 1, a C2 NGL
recovery plant has a single column operating at lower pressure
typically 450 psig or lower, which necessitates substantial
recompression for delivery of the natural gas to the liquefaction
plant. Here, contaminants-free and dried feed gas stream 1,
typically supplied at about 1200 psig, is cooled in exchanger 51
using column overhead vapor side reboiler stream 22 and external
refrigerant 32. Liquid is removed from separator 52 and sent to the
NGL column 58 that acts as a demethanizer. The flashed vapor from
the separator 52 is split into two portions, one portion is cooled
in the exchanger 54 to provide reflux to the column, while the
other portion is expanded in turbo-expander 64, cooled and sent to
the lower section for rectification. It should be especially noted
that the above gas subcooled process produces a residue gas at
ambient temperature and about 550 psig which must be recompressed
using re-compressor 100 prior to the standalone LNG plant. Thus, it
should be recognized that such non-integrated NGL recovery plant is
inefficient as the residue gas pressure is not adequate to meet the
LNG liquefaction plant requirement, requiring additional
compression and that the residue gas must be re-chilled in the LNG
liquefaction plant.
[0025] Such difficulties can be overcome in configurations
according to the inventive subject matter in which a high pressure
cryogenic vapor stream predominantly comprising methane is produced
from a NGL recovery plant that advantageously reduces, or even
eliminates the duplication of refrigeration and heat exchange steps
and equipment in the LNG liquefaction plant. Such configurations
and methods greatly reduce the refrigeration requirement in the LNG
liquefaction plant, while advantageously allowing recovery of 99%
propane and of up to 85% ethane from the feed gas. Contemplated NGL
recovery plants will typically produce a lean gas predominantly
comprising methane with a predetermined heating value (e.g., to
meet the North America natural gas pipeline requirement). Moreover,
contemplated plants and configurations can be integrated with an
LNG liquefaction plant, which will in turn increase the throughput
of the LNG liquefaction train for the same energy input. Still
further, it is noted that contemplated configurations can be
changed from C3 recovery to C2 recovery by adjusting a flow ratio
between the top reflux and the expansion flow while diverting at
least a portion of the absorber bottoms product flow to a
distillation column.
[0026] In especially preferred configurations and methods, the gas
processing portion of an integrated plant comprises a refluxed
absorber producing a bottom stream and receiving a feed gas and an
absorber reflux stream that is produced from the overhead vapor
from a distillation column (preferably after the overhead vapor is
compressed and cooled). Most typically, the distillation column is
fluidly coupled to the absorber, receives a column feed stream and
operates at a pressure that is at least 50 to 100 psi lower, and
more preferably 100 psi to 300 psi lower than the operating
pressure of the absorber. Seamless changeover of C3 recovery
operation to C2 recovery operation (or vice versa) while
maintaining 95% or higher C3 recovery for any level of C2 recovery
is achieved by increasing the second reflux to the absorber while
reducing the flow to the expander, with simultaneously diverting at
least a portion of the absorber bottoms product flow to the
distillation column.
[0027] One exemplary configuration for integrated NGL recovery with
LNG liquefaction is depicted in FIG. 2 in which two columns operate
at a pressure differential of about 300 psi, and in which the
absorber has an operating pressure about 600-700 psig. It should be
pointed out that very high recovery of C2 plus components is
possible in such plants and that the high pressure column operation
is particularly beneficial in reducing the refrigeration duty in
the LNG liquefaction plant. It should be further noted that such
configurations and methods can be used for flexible C3 and C2
recovery. Exemplary compositions, temperatures, pressures, and flow
rates of feed gas, product gas, and liquid product for typical
operation are shown in the tables below. Table 1 is a table
depicting the overall mass balance for C2 recovery when the plant
is operated on C2 recovery mode, while Table 2 is a table depicting
the overall mass balance for C3 recovery when the plant is operated
on C3 recovery mode.
TABLE-US-00001 TABLE 1 C2 Recovery Mode LIQUID FROM RESIDUE GAS NGL
TO STREAM FEED RECOVERY LIQUEFACTION CO2 0.000 0.000 0.000 N2 4.569
0.000 5.007 C1 86.161 1.020 94.320 C2 5.046 51.017 0.641 C3 1.854
21.093 0.011 iC4 0.395 4.514 0.000 nC4 0.590 6.751 0.000 iC5 0.248
2.833 0.000 nC5 0.205 2.342 0.000 C6 0.224 2.565 0.000 C7 0.662
7.570 0.000 MMscfd 1,227 107 1,119 Barrel per Day 75,743
Temperature, .degree. F. 120 115 -75 Pressure, psig 1,200 470
900
TABLE-US-00002 TABLE 2 C3 Recovery Mode LIQUID FROM RESIDUE GAS NGL
TO STREAM FEED RECOVERY LIQUEFACTION CO2 0.000 0.000 0.000 N2 4.569
0.000 4.768 C1 86.161 0.000 89.906 C2 5.046 0.867 5.228 C3 1.854
43.345 0.051 iC4 0.395 9.478 0.000 nC4 0.590 14.173 0.000 iC5 0.248
5.946 0.000 nC5 0.205 4.917 0.000 C6 0.224 5.384 0.000 C7 0.662
15.889 0.000 MMscfd 1,227 51 1,175 Barrel per Day 40,371
Temperature, F. 120 258 -60 Pressure, psig 1,200 440 886
[0028] In FIG. 2, feed gas stream 1 enters the plant at about 1200
psig and 120.degree. F., and is cooled in heat exchanger 51 to
typically -10.degree. F. to -40.degree. F., forming stream 2 using
letdown absorber bottoms stream 15, liquid stream 5 from separator
52, side reboiler stream 22 from the distillation column, and
external refrigeration stream 32. It should be noted that any type
of refrigeration system is suitable, including pure component
cascade refrigeration cycles, mixed refrigerant cycles, or a
combination of both systems. Further refrigeration is provided to
the system with expansion using the turbo-expander 64 and various
Joule-Thomson (JT) valves. The particularly high energy efficiency
of contemplated processes and configurations is illustrated by the
close temperature approaches of the heating and cooling curves in
the minimization of loss work (or close temperature approaches) as
shown in FIG. 3 and FIG. 4. Here, the combined hot composite curve
and the combined cold composite curve of the feed gas exchanger 51
and the reflux exchanger 54 are shown in FIG. 3 and FIG. 4 for C3
and C2 recovery, respectively.
[0029] The chilled feed gas stream 2 is separated in separator 52,
forming a gaseous portion 3 and a liquid portion 4. The liquid
portion 4 is letdown in pressure via JT valve 53 forming stream 5
typically at about -40.degree. F. During C3 recovery, stream 5 is
heated in exchanger 51 to about 80.degree. F. forming stream 6,
using the heat content in the feed gas stream 1. Stream 6 enters
the stripping section of the distillation column 61 for removal of
the C2 and lighter components. The gaseous portion 3 from separator
52 is split into two portions. One portion (stream 7) is routed to
the exchanger 54 to provide reflux 12 to the absorber via stream 9
and JT valve 55, while the other portion (stream 8) is expanded in
turbo-expander 64 which generates power to operate compressor 65 to
thereby form a chilled vapor stream 10, typically at -80.degree. F.
to -100.degree. F. or lower. The chilled vapor is letdown in
pressure to the absorber 58, which operates at 500 psig to 700
psig, typically at 600 psig.
[0030] For the different C2 recovery levels, the flow ratio (i.e.,
ratio of stream 8 to stream 3), can be adjusted to maintain a high
C3 recovery. Table 3 shows examples of different flow ratios and
the results of C3 and C2 recovery. Most notably, high C3 recoveries
(typically over 98%) are maintained for all C2 recovery levels.
TABLE-US-00003 TABLE 3 SPLIT RATIO (RATIO OF STREAM 8 C3 C2 TO
STREAM 3) RECOVERY, % RECOVERY, % 0.7 98 85 0.8 98 62 0.9 99 31 1.0
99 25
[0031] Absorber 58 is refluxed with two cold streams, with a first
reflux stream 11 (top reflux) supplied by stream 27 from the
distillation column 61, and a second reflux stream 12 from
exchanger 54. Using twin reflux streams and suitable flow ratios,
high C3 recovery can be maintained for the various levels of ethane
recovery. During C3 recovery, the C2 content in the NGL product
from the distillation column is lowered by increasing the column
bottom temperature using heat supplied from the side reboiler and
the bottom reboiler 63.
[0032] The absorber produces an overhead vapor stream 28 at about
-100.degree. F. to -110.degree. F. and a bottoms stream 14 at about
-90.degree. F. to -100.degree. F. The overhead vapor is compressed
by residue gas compressor 65 using power generated by
turbo-expander 64 forming a discharge stream 29, typically at about
900 psig and -70.degree. F. to -80.degree. F. It should be
especially appreciated that compression of a cold vapor is energy
efficient as the achievable compression ratio across the compressor
is significantly higher than that using a warm vapor of heretofore
known plants. Thus, contemplated processes produce a high pressure
and cryogenic temperature vapor that can be fed to the LNG
liquefaction plant 67 for LNG production forming stream 30 at about
-255.degree. F. to -260.degree. F.
[0033] The absorber bottoms stream 14 is letdown in pressure in JT
valve 59 to about 460 psig, and is chilled to about -100.degree. F.
forming stream 15. During C3 recovery, this cold stream is used to
provide a least a portion of the cooling duty of feed exchanger 51
and the reflux duty in condenser 62 to form streams 17 and 18,
respectively. The overhead condenser 62 typically includes a heat
exchange coil that is integral to the distillation column,
generating internal reflux stream 19 to the rectification section
of the distillation column. Alternatively, the integral condenser
62 system can also be replaced by an external system, which would
include an external heat exchanger, a separator and a reflux pump
(not shown).
[0034] During C2 recovery, at least a portion of absorber bottom
stream 14 is routed directly to the top of the distillation column
for absorption of the C2 and heavier components in the distillation
column. In this operation, JT valve 59 is partially, and more
typically entirely closed and JT valve 60 is partially, and more
typically entirely open forming stream 20, routing the cold
absorber bottoms to the distillation column for C2 recovery. Thus,
using various process streams within the NGL recovery plant, it
should be recognized that the temperature profile of distillation
column 61 can (gradually) vary between C3 recovery and C2 recovery
as exemplified in Table 4 below.
TABLE-US-00004 TABLE 4 TEMPERATURE PROFILE OF 61 C2 RECOVERY C3
RECOVERY Top (Stream 24) -100 to -120.degree. F. -35 to -45.degree.
F. Bottom (Stream 25) -90 to -120.degree. F. 240 to 270.degree.
F.
[0035] The NGL product composition is controlled with a side
reboiler integral to exchanger 51, and a bottom reboiler 63 using
an external heat source 34. The distillation column 61 produces an
NGL bottoms products 25 (C2 plus and/or C3 plus) and an overhead
vapor stream 24 that is compressed in compressor 66 to about 500
psig to 700 psig, or as needed to enter the absorber as a top
reflux to form stream 26, further cooled in heat exchanger 54, and
then used after pressure reduction in JT valve 56 as the first
(top) reflux 11 to the absorber. Refrigerant stream 31 is supplied
from an external refrigeration unit to the exchanger 54 for cooling
and partially or totally condensing this recycle stream. Heated
refrigerant stream 33 is later returned to the refrigeration
unit.
[0036] It should be appreciated that recycling the overhead stream
from the distillation permits recovery of the desirable NGL
components allowing the distillation column to operate at the most
efficient pressure for fractionation of separation of the desirable
NGL components. Furthermore, and as illustrated in FIG. 5,
contemplated configurations take advantage of the relative
volatilities of the NGL components (that is C1 to C2) in the
columns using the pressure differential as compared to known
configurations. Here, both the present configurations and the known
configurations are compared for a column operating at 600 psig
pressure for C2 recovery. The relative volatilities of the NGL
components of the known configurations are represented by curve A,
which drops to a very low value of 2 in the mid section of the
column. These low relative volatilities are the primary reason of
lower separation efficiency and lower NGL recoveries, even with a
large number of fractionation trays. In contrast, employing two
columns operating at different pressures, with the first column
(absorber) operating at a high pressure of 600 psig (curve B) and
the second column operating at 450 psig (curve C), dramatically
increases relative volatilities of the NGL components (e.g., to
values of over 10), resulting in higher separation efficiency and
higher NGL recoveries.
[0037] Thus, it should be recognized that the inventor discovered
an efficient and flexible configuration and process that produces a
high pressure cryogenic vapor stream (containing predominately
methane) suitable for feeding to an LNG liquefaction plant, and
that further produces a liquid stream containing predominantly
ethane and heavier hydrocarbons. Most typically, contemplated
configurations and processes can achieve 95% or higher propane
recovery when operating in a propane recovery mode, and can also
achieve up to 50 to 85% ethane recovery when operating in an ethane
recovery mode without substantial reduction (i.e., less than 5%
absolute in reduction) in propane recovery. Contemplated
configurations and processes can also provide a seamless and
gradual changeover from a C3 recovery operation to a C2 recovery
operation (or vice versa) by only adjusting flow ratios of streams
to the absorber and the distillation column.
[0038] Viewed from a different perspective, the inventor discovered
that high C2 and C3 recovery from a feed gas with relatively high
pressure (e.g., between about 800 psig to 1600 psig) can be
realized by operating an absorber in a gas processing plant at a
higher pressure than a distillation column (e.g., a demethanizer or
deethanizer), and in which a compressor is used to recycle the
distillation column overhead to the absorber. In such
configurations, the absorber bottoms product is preferably expanded
to provide cooling for the feed gas and the reflux stream. The
overhead vapor from the absorber is then further compressed with
power provided by the turboexpander, forming a feed gas to an LNG
liquefaction plant without further recompression. Further
components and considerations related to some aspects of the
inventive subject matter are disclosed in our copending U.S. patent
application with the Ser. No. 10/478,705, which is incorporated by
reference herein. It is still further contemplated that the
configurations according to the inventive subject matter may find
wide applicability in plants where high propane and ethane recovery
are desirable, and where feed gas is available at a pressure
greater than about 800 psig.
[0039] With respect to the feed gas, it is contemplated that
numerous natural gas sources are suitable, including non-associated
gas field production or gas fields associated with oil production,
whether they are located on-shore or offshore. Consequently, the
pressure of contemplated feed gas streams may also vary
considerably, and where desirable, lower pressures may therefore be
increased using boosters or compressors. However, it is generally
preferred that appropriate feed gas pressures for plant
configurations according to FIG. 2 will generally be in the range
between about 800 psig and about 1600 psig, and that at least a
portion of the feed gas is expanded in a turboexpander to provide
cooling and/or power for the residue gas recompression.
[0040] In especially preferred aspects, the absorber is configured
to separately receive a first and a second portion of a feed gas
vapor and a distillation column overhead, wherein the first portion
of the feed gas vapor and the distillation column overhead provide
the reflux to the absorber. In such configurations, a flow control
adjusts the ratio of at least one of the first and second portions
of the feed gas vapor to produce the desired recovery levels of
ethane. Among other advantages, it should be recognized that an
optimum flow ratio of the first and second flow of the feed gas is
employed for the variable C2 recovery while maintaining a high C3
(95% or above) recovery. Most typically, the flow control will
include one or more manual and/or automatic valves, which are most
preferably operated using microprocessor-based control equipment
well known in the art. Such control units may operate entirely
automatic without user intervention and may be configured to
receive compositional information by sensors. Alternatively, or
additionally, an operator may also provide compositional
information based on a known or determined feed gas composition,
and/or flow control may also be regulated at least in part by
product volume flow in one or more products.
[0041] The absorber preferably produces an overhead vapor product
that is predominantly methane (e.g., at least 85%, more typically
at least 90%, and most typically at least 93%) at cryogenic
temperature (-80.degree. F. or lower), which is further compressed
in a compressor using power generated by the turbo-expansion of the
feed gas. Such configuration produces a high pressure cryogenic
vapor at 800 psig to 900 psig or higher that can advantageously be
directly fed to the LNG liquefaction plant. It should be recognized
that compressing a cryogenic vapor from the high pressure absorber
is more energy efficient, and therefore allows operation of an LNG
liquefaction plant with relatively low, and more typically no
additional compression of the feed gas.
[0042] It should also be recognized that at least a portion of the
absorber bottoms product is used as a lean oil for C2 absorption in
the distillation column during C2 recovery. In such configurations,
the flow to the feed cooler is reduced or stopped, thereby
directing most or all of the absorber bottom product to the
distillation column. C2 recovery in such configurations increases
when the first portion of the feed gas vapor increases relative to
the second portion of the feed gas vapor. Thus, preferred
configurations permit a seamless changeover of the C3 recovery
operation to the C2 recovery operation (or vise versa) while
maintaining 95% or higher C3 recovery. The bottom product of the
absorber is preferably expanded in a range of 50 psi to 350 psi,
thereby chilled by Joule-Thomson effect to -90.degree. F. to
-130.degree. F. It is also contemplated that the cooled and
expanded bottoms product stream is fed as the distillation column
feed stream into the distillation column, and it is further
contemplated that the expanded bottoms product stream may further
provide cooling for the feed gas and reflux to the column, and at
least of a portion of the expanded absorber bottoms product may
also be routed directly to the distillation column for C2
absorption during the C2 recovery operation.
[0043] The distillation column in preferred configurations
typically comprises a demethanizer or deethanizer column and
produces an overhead stream, which is compressed, cooled, and fed
to the absorber as the first absorber reflux stream. It should be
noted that the compressor can be located in various locations
between the absorber and distillation column, and that this reflux
stream can be fed to various locations in the absorber for C3
recovery and/or C2 recovery. In yet further contemplated aspects,
an external refrigeration unit is fluidly coupled to feed gas
exchangers and reflux exchangers to supply feed gas chilling and
column reflux duties. Such external refrigeration unit can be
implemented using numerous configurations well known in the art
using pure components or mixed refrigerants. For example, a cascade
refrigeration process may employ heat exchange of the natural gas
with several pure component refrigerants having successively lower
boiling points. Alternatively, a single mixed refrigerant with
multiple pure components can be used by evaporating the refrigerant
at several different pressure levels. Also, where desirable,
natural gas cooling can be achieved by expansion of the natural gas
using either Joule-Thomson expansion or expansion turbine.
[0044] It is also contemplated that an external refrigeration unit
can be employed to cool at least one of the first and second reflux
streams, and may further cool at least one of the natural gas feed
and a vapor portion of the natural gas feed. Where C2 recovery is
particularly preferred, it is contemplated that the first lean
reflux stream from the feed gas may be fed into the absorber as a
subcooled liquid (i.e., a liquid cooled below its bubble point
temperature), and that the distillation column is a
demethanizer.
[0045] Consequently, in one aspect of the inventive subject matter,
a method of operating a plant includes a step of providing an
absorber and a distillation column, wherein the absorber receives a
plurality of absorber feed streams and provides a bottom product to
the distillation column, and an overhead absorber vapor that is
used as feed to the LNG liquefaction. In another step, at least one
of the feed streams is split into a first and second portion,
wherein the first and second portions are introduced into the
absorber at different locations, and in still another step, the
flow ratio between the first and second portions is used to control
the degree of recovery of C2 components in the bottom product of
the distillation column. In yet a further step, the absorber
bottoms product is used as a lean oil for absorption of the C2
component in the distillation column.
[0046] Viewed from a different perspective, contemplated
configurations may also be used in a method of increasing
throughput and/or energy consumption in an LNG liquefaction plant
that is coupled to a natural gas recovery plant having an absorber
and a distillation column. Such methods will typically include one
step in which a first reflux stream is provided to the absorber,
wherein the first reflux stream comprises an overhead product from
the distillation column. In another step, a bypass is provided
upstream of a turbo expander, wherein the bypass receives a vapor
portion of a cooled natural gas and provides the vapor portion to
the absorber, and in yet another step, the pressure of the vapor
portion is cooled, at least partially condensed and reduced in
pressure before the vapor portion enters the absorber as a second
reflux stream. In a still further step, a heat exchanger cools at
least one of the first and second reflux streams using an external
refrigerant. Therefore, a method of operating a plant may include
one step in which an absorber and a distillation column are
provided. In a further step, a cooled lean overhead product from
the distillation column is fed to the absorber as a first reflux
stream, and in another step, the pressure of a cooled vapor portion
of a natural gas feed is reduced via a device other than a turbo
expander, wherein the cooled vapor portion that is at least
partially condensed or subcooled, reduced in pressure and fed to
the absorber as a second reflux stream.
[0047] With respect to the remaining components of contemplated
configurations (e.g., heat exchangers, pumps, pipes, valves,
compressors, expanders, etc.) it should be appreciated that such
components are well known to the artisan and that all
known/commercially available components are deemed suitable for use
in conjunction with the teachings presented herein. Furthermore,
the term "about" where used in conjunction with a numeral refers to
a numeric range that encompasses +/-10% of the numeral, inclusive.
For example, the term about 10% refers to a range of 9% to 11%,
inclusive.
[0048] Thus, specific embodiments and applications for high propane
and ethane recovery processes and configurations have been
disclosed. It should be apparent, however, to those skilled in the
art that many more modifications besides those already described
are possible without departing from the inventive concepts herein.
The inventive subject matter, therefore, is not to be restricted
except in the spirit of the appended claims. Moreover, in
interpreting both the specification and the claims, all terms
should be interpreted in the broadest possible manner consistent
with the context. In particular, the terms "comprises" and
"comprising" should be interpreted as referring to elements,
components, or steps in a non-exclusive manner, indicating that the
referenced elements, components, or steps may be present, or
utilized, or combined with other elements, components, or steps
that are not expressly referenced. Furthermore, where a definition
or use of a term in a reference, which is incorporated by reference
herein is inconsistent or contrary to the definition of that term
provided herein, the definition of that term provided herein
applies and the definition of that term in the reference does not
apply.
* * * * *