U.S. patent application number 14/828093 was filed with the patent office on 2016-03-10 for hydrocarbon gas processing.
This patent application is currently assigned to ORTLOFF ENGINEERS, LTD.. The applicant listed for this patent is ORTLOFF ENGINEERS, LTD.. Invention is credited to J. Ascencion Anguiano, Hank M. Hudson, Joe T. Lynch, John D. Wilkinson.
Application Number | 20160069610 14/828093 |
Document ID | / |
Family ID | 55437189 |
Filed Date | 2016-03-10 |
United States Patent
Application |
20160069610 |
Kind Code |
A1 |
Anguiano; J. Ascencion ; et
al. |
March 10, 2016 |
HYDROCARBON GAS PROCESSING
Abstract
A process and an apparatus are disclosed for the recovery of
ethane, ethylene, propane, propylene, and heavier hydrocarbon
components from a hydrocarbon gas stream. The stream is divided
into first and second streams. The first stream is cooled to
condense substantially all of it, expanded to lower pressure, and
supplied to a fractionation tower at an upper mid-column feed
position. The second stream is cooled sufficiently to partially
condense it and separated into vapor and liquid streams. The vapor
stream is divided into first and second portions. The first portion
is cooled to condense substantially all of it, expanded to the
tower pressure, and supplied to the tower at the top feed position.
The second portion is expanded to the tower pressure and supplied
to the fractionation tower at an intermediate mid-column feed
position. The liquid stream is expanded to the tower pressure and
supplied to the column at a lower mid-column feed position. The
quantities and temperatures of the feeds to the fractionation tower
are effective to maintain the overhead temperature of the
fractionation tower at a temperature whereby the major portion of
the desired components is recovered.
Inventors: |
Anguiano; J. Ascencion;
(Midland, TX) ; Wilkinson; John D.; (Midland,
TX) ; Lynch; Joe T.; (Midland, TX) ; Hudson;
Hank M.; (Midland, TX) |
|
Applicant: |
Name |
City |
State |
Country |
Type |
ORTLOFF ENGINEERS, LTD. |
Midland |
TX |
US |
|
|
Assignee: |
ORTLOFF ENGINEERS, LTD.
Midland
TX
|
Family ID: |
55437189 |
Appl. No.: |
14/828093 |
Filed: |
August 17, 2015 |
Related U.S. Patent Documents
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Application
Number |
Filing Date |
Patent Number |
|
|
62045908 |
Sep 4, 2014 |
|
|
|
Current U.S.
Class: |
62/621 |
Current CPC
Class: |
F25J 3/0238 20130101;
F25J 2270/60 20130101; F25J 2205/04 20130101; F25J 2200/30
20130101; F25J 2240/02 20130101; F25J 3/0233 20130101; F25J 2290/40
20130101; F25J 2200/70 20130101; F25J 2270/12 20130101; F25J
2200/02 20130101; F25J 2210/06 20130101; F25J 2270/02 20130101;
F25J 3/0209 20130101 |
International
Class: |
F25J 3/02 20060101
F25J003/02 |
Claims
1. In a process for the separation of a gas stream containing
methane, C.sub.2 components, C.sub.3 components, and heavier
hydrocarbon components into a volatile residue gas fraction and a
relatively less volatile fraction containing a major portion of
said C.sub.2 components, C.sub.3 components, and heavier
hydrocarbon components or said C.sub.3 components and heavier
hydrocarbon components, in which process (a) said gas stream is
cooled under pressure to provide a cooled stream; (b) said cooled
stream is expanded to a lower pressure whereby it is further
cooled; and (c) said further cooled stream is directed into a
distillation column and fractionated at said lower pressure whereby
the components of said relatively less volatile fraction are
recovered; the improvement wherein prior to cooling, said gas
stream is divided into first and second streams; and (1) said first
stream is cooled to condense substantially all of it; (2) said
substantially condensed first stream is expanded to said lower
pressure whereby it is further cooled, and is thereafter supplied
to said distillation column at an upper mid-column feed position;
(3) said second stream is cooled under pressure sufficiently to
partially condense it; (4) said partially condensed second stream
is separated thereby to provide a vapor stream and at least one
liquid stream; (5) said vapor stream is divided into first and
second portions; (6) said first portion is cooled to condense
substantially all of it; (7) said substantially condensed first
portion is expanded to said lower pressure whereby it is further
cooled, and is thereafter supplied to said distillation column at a
top feed position; (8) said second portion is expanded to said
lower pressure and is supplied to said distillation column at a
mid-column feed position below said upper mid-column feed position;
(9) at least a portion of said at least one liquid stream is
expanded to said lower pressure and is supplied to said
distillation column at a lower mid-column feed position below said
mid-column feed position; and (10) the quantities and temperatures
of said feed streams to said distillation column are effective to
maintain the overhead temperature of said distillation column at a
temperature whereby the major portions of the components in said
relatively less volatile fraction are recovered.
2. The improvement according to claim 1 wherein (1) said gas stream
is cooled under pressure sufficiently to partially condense it; (2)
said partially condensed gas stream divided into said first stream
and said second stream; and (3) said second stream is separated
thereby to provide said vapor stream and said at least one liquid
stream.
3. The improvement according to claim 1 wherein (1) said first
stream is combined with at least a portion of said at least one
liquid stream to form a combined stream; (2) said combined stream
is cooled to condense substantially all of it; (3) said
substantially condensed combined stream is expanded to said lower
pressure whereby it is further cooled, and is thereafter supplied
to said distillation column at said upper mid-column feed position;
and (4) any remaining portion of said at least one liquid stream is
expanded to said lower pressure and is supplied to said
distillation column at said lower mid-column feed position below
said mid-column feed position.
4. The improvement according to claim 2 wherein (1) said first
stream is combined with at least a portion of said at least one
liquid stream to form a combined stream; (2) said combined stream
is cooled to condense substantially all of it; (3) said
substantially condensed combined stream is expanded to said lower
pressure whereby it is further cooled, and is thereafter supplied
to said distillation column at said upper mid-column feed position;
and (4) any remaining portion of said at least one liquid stream is
expanded to said lower pressure and is supplied to said
distillation column at said lower mid-column feed position below
said mid-column feed position.
5. The improvement according to claim 1 wherein (1) said expanded
substantially condensed first portion is supplied at a top feed
position to a contacting and separating device that produces said
volatile residue gas fraction and a bottom liquid stream, whereupon
said bottom liquid stream is supplied to said distillation column;
(2) an overhead vapor stream is withdrawn from an upper region of
said distillation column and is supplied to said contacting and
separating device at a lower column feed position; and (3) the
quantities and temperatures of said feed streams to said contacting
and separating device are effective to maintain the overhead
temperature of said contacting and separating device at a
temperature whereby the major portions of the components in said
relatively less volatile fraction are recovered.
6. The improvement according to claim 5 wherein (1) said gas stream
is cooled under pressure sufficiently to partially condense it; (2)
said partially condensed gas stream divided into said first stream
and said second stream; and (3) said second stream is separated
thereby to provide said vapor stream and said at least one liquid
stream.
7. The improvement according to claim 5 wherein (1) said first
stream is combined with at least a portion of said at least one
liquid stream to form a combined stream; (2) said combined stream
is cooled to condense substantially all of it; (3) said
substantially condensed combined stream is expanded to said lower
pressure whereby it is further cooled, and is thereafter supplied
to said distillation column at said upper mid-column feed position;
and (4) any remaining portion of said at least one liquid stream is
expanded to said lower pressure and is supplied to said
distillation column at said lower mid-column feed position.
8. The improvement according to claim 6 wherein (1) said first
stream is combined with at least a portion of said at least one
liquid stream to form a combined stream; (2) said combined stream
is cooled to condense substantially all of it; (3) said
substantially condensed combined stream is expanded to said lower
pressure whereby it is further cooled, and is thereafter supplied
to said distillation column at said upper mid-column feed position;
and (4) any remaining portion of said at least one liquid stream is
expanded to said lower pressure and is supplied to said
distillation column at said lower mid-column feed position.
9. The improvement according to claim 1 or 2 wherein said expanded
said at least one liquid stream is heated and thereafter supplied
to said distillation column at said lower mid-column feed position
below said mid-column feed position.
10. The improvement according to claim 3 or 4 wherein said expanded
said any remaining portion of said at least one liquid stream is
heated and thereafter supplied to said distillation column at said
lower mid-column feed position below said mid-column feed
position.
11. The improvement according to claim 5 or 6 wherein said expanded
said at least one liquid stream is heated and thereafter supplied
to said distillation column at said lower mid-column feed
position.
12. The improvement according to claim 7 or 8 wherein said expanded
said any remaining portion of said at least one liquid stream is
heated and thereafter supplied to said distillation column at said
lower mid-column feed position.
13. In an apparatus for the separation of a gas stream containing
methane, C.sub.2 components, C.sub.3 components, and heavier
hydrocarbon components into a volatile residue gas fraction and a
relatively less volatile fraction containing a major portion of
said C.sub.2 components, C.sub.3 components, and heavier
hydrocarbon components or said C.sub.3 components and heavier
hydrocarbon components, in said apparatus there being (a) a first
cooling means to cool said gas stream under pressure connected to
provide a cooled stream under pressure; (b) a first expansion means
connected to receive at least a portion of said cooled stream under
pressure and expand it to a lower pressure, whereby said stream is
further cooled; and (c) a distillation column connected to receive
said further cooled stream, said distillation column being adapted
to separate said further cooled stream into said volatile residue
gas fraction and said relatively less volatile fraction; the
improvement wherein said apparatus includes (1) first dividing
means prior to said first cooling means to divide said gas stream
into first and second streams; (2) second cooling means connected
to said first dividing means to receive said first stream and cool
it sufficiently to substantially condense it; (3) said first
expansion means connected to said second cooling means, said first
expansion means being adapted to receive said substantially
condensed first stream and expand it to said lower pressure, said
first expansion means being further connected to said distillation
column to supply said expanded substantially condensed first stream
to said distillation column at an upper mid-column feed position;
(4) said first cooling means connected to said first dividing means
to receive said second stream and cool it under pressure
sufficiently to partially condense it; (5) separating means
connected to said first cooling means to receive said partially
condensed second stream and separate it into a vapor stream and at
least one liquid stream; (6) second dividing means connected to
said separating means to receive said vapor strewn and divide it
into first and second portions; (7) third cooling means connected
to said second dividing means to receive said first portion and
cool it sufficiently to substantially condense it; (8) second
expansion means connected to said third cooling means to receive
said substantially condensed first portion and expand it to said
lower pressure, said second expansion means being further connected
to said distillation column to supply said expanded substantially
condensed first portion to said distillation column at a top feed
position; (9) third expansion means being connected to said second
dividing means to receive said second portion and expand it to said
lower pressure, said third expansion means being further connected
to said distillation column to supply said expanded second portion
to said distillation column at a mid-column feed position below
said upper mid-column feed position; (10) fourth expansion means
connected to said separating means to receive at least a portion of
said at least one liquid stream and expand it to said lower
pressure, said fourth expansion means being further connected to
said distillation column to supply said expanded liquid stream to
said distillation column at a lower mid-column feed position below
said mid-column feed position; and (11) control means adapted to
regulate the quantities and temperatures of said feed streams to
said distillation column to maintain the overhead temperature of
said distillation column at a temperature whereby the major
portions of the components in said relatively less volatile
fraction are recovered.
14. The improvement according to claim 13 wherein (1) said first
cooling means is connected to receive said gas stream and cool it
under pressure sufficiently to partially condense it; (2) said
first dividing means is connected to said first cooling means to
receive said partially condensed gas stream and divide it into said
first stream and said second stream; and (3) said separating means
is connected to said first dividing means to receive said second
stream and separate it into said vapor stream and said at least one
liquid stream.
15. The improvement according to claim 13 wherein (1) a combining
means is connected to said first dividing means and said separating
means to receive said first stream and at least a portion of said
at least one liquid stream and form a combined stream; (2) said
second cooling means is connected to said combining means to
receive said combined stream and cool it sufficiently to
substantially condense it; (3) said first expansion means is
connected to said second cooling means to receive said
substantially condensed combined stream and expand it to said lower
pressure, said first expansion means being further connected to
said distillation column to supply said expanded substantially
condensed combined stream to said distillation column at said upper
mid-column feed position; and (4) said fourth expansion means is
connected to said separating means to receive any remaining portion
of said at least one liquid stream and expand it to said lower
pressure, said fourth expansion means being further connected to
said distillation column to supply said expanded liquid stream to
said distillation column at said lower mid-column feed position
below said mid-column feed position.
16. The improvement according to claim 14 wherein (1) a combining
means is connected to said first dividing means and said separating
means to receive said first stream and at least a portion of said
at least one liquid stream and form a combined stream; (2) said
second cooling means is connected to said combining means to
receive said combined stream and cool it sufficiently to
substantially condense it; (3) said first expansion means is
connected to said second cooling means to receive said
substantially condensed combined stream and expand it to said lower
pressure, said first expansion means being further connected to
said distillation column to supply said expanded substantially
condensed combined stream to said distillation column at said upper
mid-column feed position; and (4) said fourth expansion means is
connected to said separating means to receive any remaining portion
of said at least one liquid stream and expand it to said lower
pressure, said fourth expansion means being further connected to
said distillation column to supply said expanded liquid stream to
said distillation column at said lower mid-column feed position
below said mid-column feed position.
17. The improvement according to claim 13 wherein (1) said second
expansion means is connected to a contacting and separating means
to supply said expanded substantially condensed first portion to
said contacting and separating means at a top feed position, said
contacting and separating means being adapted to produce said
relatively less volatile fraction and a bottom liquid stream; (2)
said distillation column is connected to receive said bottom liquid
stream, said distillation column being adapted to separate said
bottom liquid stream into an overhead vapor stream and said
relatively less volatile fraction; (3) said distillation column is
further connected to said contacting and separating means to supply
said overhead vapor stream to said contacting and separating means
at a lower column feed position; and (4) said control means is
adapted to regulate the quantities and temperatures of said feed
streams to said contacting and separating means to maintain the
overhead temperature of said contacting and separating means at a
temperature whereby the major portions of the components in said
relatively less volatile fraction are recovered.
18. The improvement according to claim 17 wherein (1) said first
cooling means is connected to receive said gas stream and cool it
under pressure sufficiently to partially condense it; (2) said
first dividing means is connected to said first cooling means to
receive said partially condensed gas stream and divide it into said
first stream and said second stream; and (3) said separating means
is connected to said first dividing means to receive said second
stream and separate it into said vapor stream and said at least one
liquid stream.
19. The improvement according to claim 17 wherein (1) a combining
means is connected to said first dividing means and said separating
means to receive said first stream and at least a portion of said
at least one liquid stream and form a combined stream; (2) said
second cooling means is connected to said combining means to
receive said combined stream and cool it sufficiently to
substantially condense it; (3) said first expansion means is
connected to said second cooling means to receive said
substantially condensed combined stream and expand it to said lower
pressure, said first expansion means being further connected to
said distillation column to supply said expanded substantially
condensed combined stream to said distillation column at said upper
mid-column feed position; and (4) said fourth expansion means is
connected to said separating means to receive any remaining portion
of said at least one liquid stream and expand it to said lower
pressure, said fourth expansion means being further connected to
said distillation column to supply said expanded liquid stream to
said distillation column at said lower mid-column feed
position.
20. The improvement according to claim 18 wherein (1) a combining
means is connected to said first dividing means and said separating
means to receive said first stream and at least a portion of said
at least one liquid stream and form a combined stream; (2) said
second cooling means is connected to said combining means to
receive said combined stream and cool it sufficiently to
substantially condense it; (3) said first expansion means is
connected to said second cooling means to receive said
substantially condensed combined stream and expand it to said lower
pressure, said first expansion means being further connected to
said distillation column to supply said expanded substantially
condensed combined stream to said distillation column at said upper
mid-column feed position; and (4) said fourth expansion means is
connected to said separating means to receive any remaining portion
of said at least one liquid stream and expand it to said lower
pressure, said fourth expansion means being further connected to
said distillation column to supply said expanded liquid stream to
said distillation column at said lower mid-column feed
position.
21. The improvement according to claim 13 or 14 wherein a heating
means is connected to said fourth expansion means to receive said
expanded said at least one liquid stream and heat it, said heating
means being further connected to said distillation column to supply
said heated expanded said at least one liquid stream to said
distillation column at said lower mid-column feed point.
22. The improvement according to claim 15 or 16 wherein a heating
means is connected to said fourth expansion means to receive said
expanded said at least a portion of said at least one liquid stream
and heat it, said heating means being further connected to said
distillation column to supply said heated expanded said at least a
portion of said at least one liquid stream to said distillation
column at said lower mid-column feed point.
23. The improvement according to claim 17 or 18 wherein a heating
means is connected to said fourth expansion means to receive said
expanded said at least one liquid stream and heat it, said heating
means being further connected to said distillation column to supply
said heated expanded said at least one liquid stream to said
distillation column at said lower mid-column feed point.
24. The improvement according to claim 19 or 20 wherein a heating
means is connected to said fourth expansion means to receive said
expanded said at least a portion of said at least one liquid stream
and heat it, said heating means being further connected to said
distillation column to supply said heated expanded said at least a
portion of said at least one liquid stream to said distillation
column at said lower mid-column feed point.
Description
[0001] This invention relates to a process and an apparatus for the
separation of a gas containing hydrocarbons. The applicants claim
the benefits under Title 35, United States Code, Section 119(e) of
prior U.S. Provisional Application No. 62/045,908 which was filed
on Sep. 4, 2014.
BACKGROUND OF THE INVENTION
[0002] Ethylene, ethane, propylene, propane, and/or heavier
hydrocarbons can be recovered from a variety of gases, such as
natural gas, refinery gas, and synthetic gas streams obtained from
other hydrocarbon materials such as coal, crude oil, naphtha, oil
shale, tar sands, and lignite. Natural gas usually has a major
proportion of methane and ethane, i.e., methane and ethane together
comprise at least 50 mole percent of the gas. The gas also contains
relatively lesser amounts of heavier hydrocarbons such as propane,
butanes, pentanes, and the like, as well as hydrogen, nitrogen,
carbon dioxide, and other gases.
[0003] The present invention is generally concerned with the
recovery of ethylene, ethane, propylene, propane and heavier
hydrocarbons from such gas streams. A typical analysis of a gas
stream to be processed in accordance with this invention would be,
in approximate mole percent, 69.4% methane, 11.8% ethane and other
C.sub.2 components, 5.6% propane and other C.sub.3 components, 0.9%
iso-butane, 1.8% normal butane, and 1.1% pentanes plus, with the
balance made up of nitrogen and carbon dioxide. Sulfur containing
gases are also sometimes present.
[0004] The historically cyclic fluctuations in the prices of both
natural gas and its natural gas liquid (NGL) constituents have at
times reduced the incremental value of ethane, ethylene, propane,
propylene, and heavier components as liquid products. This has
resulted in a demand for processes that can provide more efficient
recoveries of these products, for processes that can provide
efficient recoveries with lower capital investment, and for
processes that can be easily adapted or adjusted to vary the
recovery of a specific component over a broad range. Available
processes for separating these materials include those based upon
cooling and refrigeration of gas, oil absorption, and refrigerated
oil absorption. Additionally, cryogenic processes have become
popular because of the availability of economical equipment that
produces power while simultaneously expanding and extracting heat
from the gas being processed. Depending upon the pressure of the
gas source, the richness (ethane, ethylene, and heavier
hydrocarbons content) of the gas, and the desired end products,
each of these processes or a combination thereof may be
employed.
[0005] The cryogenic expansion process is now generally preferred
for natural gas liquids recovery because it provides maximum
simplicity with ease of startup, operating flexibility, good
efficiency, safety, and good reliability. U.S. Pat. Nos. 3,292,380;
4,061,481; 4,140,504; 4,157,904; 4,171,964; 4,185,978; 4,251,249;
4,278,457; 4,519,824; 4,617,039; 4,687,449; 4,689,063; 4,690,702;
4,854,955; 4,869,740; 4,889,545; 5,275,005; 5,555,748; 5,566,554;
5,568,737; 5,771,712; 5,799,507; 5,881,569; 5,890,378; 5,983,664;
6,182,469; 6,578,379; 6,712,880; 6,915,662; 7,191,617; 7,219,513;
8,590,340; 8,881,549; 8,919,148; 9,021,831; 9,021,832; 9,052,136;
9,052,137; 9,057,158; 9,068,774; 9,074,814; 9,080,810; and
9,080,811; reissue U.S. Pat. No. 33,408; and co-pending application
Ser. Nos. 11/839,693; 12/750,862; 12/772,472; 12/781,259;
12/868,993; 12/849,007; 12/869,139; 14/462,056; and Ser. No.
14/462,083 describe relevant processes (although the description of
the present invention in some cases is based on different
processing conditions than those described in the listed U.S.
Patents and applications).
[0006] In a typical cryogenic expansion recovery process, a feed
gas stream under pressure is cooled by heat exchange with other
streams of the process and/or external sources of refrigeration
such as a propane compression-refrigeration system. As the gas is
cooled, liquids may be condensed and collected in one or more
separators as high-pressure liquids containing some of the desired
C.sub.2+ components. Depending on the richness of the gas and the
amount of liquids formed, the high-pressure liquids may be expanded
to a lower pressure and fractionated. The vaporization occurring
during expansion of the liquids results in further cooling of the
stream. Under some conditions, pre-cooling the high pressure
liquids prior to the expansion may be desirable in order to further
lower the temperature resulting from the expansion. The expanded
stream, comprising a mixture of liquid and vapor, is fractionated
in a distillation (demethanizer or deethanizer) column. In the
column, the expansion cooled stream(s) is (are) distilled to
separate residual methane, nitrogen, and other volatile gases as
overhead vapor from the desired C.sub.2 components, C.sub.3
components, and heavier hydrocarbon components as bottom liquid
product, or to separate residual methane, C.sub.2 components,
nitrogen, and other volatile gases as overhead vapor from the
desired C.sub.3 components and heavier hydrocarbon components as
bottom liquid product.
[0007] If the feed gas is not totally condensed (typically it is
not), the vapor remaining from the partial condensation can be
split into two streams. One portion of the vapor is passed through
a work expansion machine or engine, or an expansion valve, to a
lower pressure at which additional liquids are condensed as a
result of further cooling of the stream. The pressure after
expansion is essentially the same as the pressure at which the
distillation column is operated. The combined vapor-liquid phases
resulting from the expansion are supplied as feed to the
column.
[0008] The remaining portion of the vapor is cooled to substantial
condensation by heat exchange with other process streams, e.g., the
cold fractionation tower overhead. Some or all of the high-pressure
liquid may be combined with this vapor portion prior to cooling.
The resulting cooled stream is then expanded through an appropriate
expansion device, such as an expansion valve, to the pressure at
which the demethanizer is operated. During expansion, a portion of
the liquid will vaporize, resulting in cooling of the total stream.
The flash expanded stream is then supplied as top feed to the
demethanizer. Typically, the vapor portion of the flash expanded
stream and the demethanizer overhead vapor combine in an upper
separator section in the fractionation tower as residual methane
product gas. Alternatively, the cooled and expanded stream may be
supplied to a separator to provide vapor and liquid streams. The
vapor is combined with the tower overhead and the liquid is
supplied to the column as a top column feed.
[0009] In the ideal operation of such a separation process, the
residue gas leaving the process will contain substantially all of
the methane in the feed gas with essentially none of the heavier
hydrocarbon components, and the bottoms fraction leaving the
demethanizer will contain substantially all of the heavier
hydrocarbon components with essentially no methane or more volatile
components. In practice, however, this ideal situation is not
obtained because the conventional demethanizer is operated largely
as a stripping column. The methane product of the process,
therefore, typically comprises vapors leaving the top fractionation
stage of the column, together with vapors not subjected to any
rectification step. Considerable losses of C.sub.2, C.sub.3, and
C.sub.4+ components occur because the top liquid feed contains
substantial quantities of these components and heavier hydrocarbon
components, resulting in corresponding equilibrium quantities of
C.sub.2 components, C.sub.3 components, C.sub.4 components, and
heavier hydrocarbon components in the vapors leaving the top
fractionation stage of the demethanizer. The loss of these
desirable components could be significantly reduced if the rising
vapors could be brought into contact with a significant quantity of
liquid (reflux) capable of absorbing the C.sub.2 components,
C.sub.3 components, C.sub.4 components, and heavier hydrocarbon
components from the vapors.
[0010] In recent years, the preferred processes for hydrocarbon
separation use an upper absorber section to provide additional
rectification of the rising vapors. The source of the reflux stream
for the upper rectification section is typically a recycled stream
of residue gas supplied under pressure. The recycled residue gas
stream is usually cooled to substantial condensation by heat
exchange with other process streams, e.g., the cold fractionation
tower overhead. The resulting substantially condensed stream is
then expanded through an appropriate expansion device, such as an
expansion valve, to the pressure at which the demethanizer is
operated. During expansion, a portion of the liquid will usually
vaporize, resulting in cooling of the total stream. The flash
expanded stream is then supplied as top feed to the demethanizer.
Typically, the vapor portion of the expanded stream and the
demethanizer overhead vapor combine in an upper separator section
in the fractionation tower as residual methane product gas.
Alternatively, the cooled and expanded stream may be supplied to a
separator to provide vapor and liquid streams, so that thereafter
the vapor is combined with the tower overhead and the liquid is
supplied to the column as a top column feed. Typical process
schemes of this type are disclosed in U.S. Pat. Nos. 4,889,545;
5,568,737; 5,881,569; and 9,052,137, assignee's co-pending
application Ser. No. 12/717,394; and in Mowrey, E. Ross.
"Efficient, High Recovery of Liquids from Natural Gas Utilizing a
High Pressure Absorber", Proceedings of the Eighty-First Annual
Convention of the Gas Processors Association, Dallas, Tex., Mar.
11-13, 2002. Unfortunately, these processes require the use of a
large amount of compression power to provide the motive force for
recycling the reflux stream to the demethanizer, adding to both the
capital cost and the operating cost of facilities using these
processes.
[0011] The present invention also employs an upper rectification
section (or a separate rectification column in some embodiments).
However, the reflux for this rectification section is provided by
cooling two streams derived from the feed gas to substantial
condensation and then expanding both streams to the operating
pressure of the fractionation tower. During expansion, a portion of
each stream is vaporized, resulting in cooling of each total
stream. One cooled, expanded stream is then supplied to the
fractionation tower at a top column feed point and the other
cooled, expanded stream is supplied to the tower at an upper
mid-column feed point. The condensed liquid in the top column feed,
which is predominantly liquid methane, can then be used to absorb
C.sub.2 components, C.sub.3 components, C.sub.4 components, and
heavier hydrocarbon components from the vapors rising through the
upper rectification section and thereby capture these valuable
components in the bottom liquid product from the demethanizer.
[0012] In accordance with the present invention, it has been found
that C.sub.2 recovery in excess of 93% and C.sub.3 and C.sub.4+
recoveries in excess of 99% can be obtained. In addition, the
present invention makes possible essentially 100% separation of
methane and lighter components from the C.sub.2 components and
heavier components at higher recovery levels compared to the prior
art while maintaining the same energy requirements. The present
invention, although applicable at lower pressures and warmer
temperatures, is particularly advantageous when processing feed
gases in the range of 400 to 1500 psia [2,758 to 10,342 kPa(a)] or
higher under conditions requiring NGL recovery column overhead
temperatures of -50.degree. F. [-46.degree. C.] or colder.
[0013] For a better understanding of the present invention,
reference is made to the following examples and drawings. Referring
to the drawings:
[0014] FIG. 1 is a flow diagram of a prior art natural gas
processing plant in accordance with U.S. Pat. No. 4,278,457;
[0015] FIG. 2 is a flow diagram of a prior art natural gas
processing plant in accordance with U.S. Pat. No. 5,568,737;
[0016] FIG. 3 is a flow diagram of a prior art natural gas
processing plant in accordance with U.S. Pat. No. 7,191,617;
[0017] FIG. 4 is a flow diagram of a prior art natural gas
processing plant in accordance with assignee's co-pending
application Ser. No. 11/839,693;
[0018] FIG. 5 is a flow diagram of a natural gas processing plant
in accordance with the present invention; and
[0019] FIGS. 6 through 9 are flow diagrams illustrating alternative
means of application of the present invention to a natural gas
stream.
[0020] In the following explanation of the above figures, tables
are provided summarizing flow rates calculated for representative
process conditions. In the tables appearing herein, the values for
flow rates (in moles per hour) have been rounded to the nearest
whole number for convenience. The total stream rates shown in the
tables include all non-hydrocarbon components and hence are
generally larger than the sum of the stream flow rates for the
hydrocarbon components. Temperatures indicated are approximate
values rounded to the nearest degree. It should also be noted that
the process design calculations performed for the purpose of
comparing the processes depicted in the figures are based on the
assumption of no heat leak from (or to) the surroundings to (or
from) the process. The quality of commercially available insulating
materials makes this a very reasonable assumption and one that is
typically made by those skilled in the art.
[0021] For convenience, process parameters are reported in both the
traditional British units and in the units of the Systeme
International d'Unites (SI). The molar flow rates given in the
tables may be interpreted as either pound moles per hour or
kilogram moles per hour. The energy consumptions reported as
horsepower (HP) and/or thousand British Thermal Units per hour
(MBTU/Hr) correspond to the stated molar flow rates in pound moles
per hour. The energy consumptions reported as kilowatts (kW)
correspond to the stated molar flow rates in kilogram moles per
hour.
DESCRIPTION OF THE PRIOR ART
[0022] FIG. 1 is a process flow diagram showing the design of a
processing plant to recover C.sub.2+ components from natural gas
using prior art according to U.S. Pat. No. 4,278,457. In this
simulation of the process, inlet gas enters the plant at
104.degree. F. [40.degree. C.] and 896 psia [6,181 kPa(a)] as
stream 31. If the inlet gas contains a concentration of sulfur
compounds which would prevent the product streams from meeting
specifications, the sulfur compounds are removed by appropriate
pretreatment of the feed gas (not illustrated). In addition, the
feed stream is usually dehydrated to prevent hydrate (ice)
formation under cryogenic conditions. Solid desiccant has typically
been used for this purpose.
[0023] The feed stream 31 is divided into two portions, streams 32
and 33. Stream 32 is cooled to substantial condensation in heat
exchanger 13 by heat exchange with cold residue gas (steam 41),
flashed separator liquids (stream 35a), and propane refrigerant.
The resulting substantially condensed stream 32a at -132.degree. F.
[-91.degree. C.] is then flash expanded through expansion valve 17
to the operating pressure (approximately 316 psia [2,181 kPa(a)])
of fractionation tower 19. During expansion a portion of the stream
is vaporized, resulting in cooling of the total stream. In the
process illustrated in FIG. 1, the expanded stream 32b leaving
expansion valve 17 reaches a temperature of -159.degree. F.
[-106.degree. C.] before it is supplied to fractionation tower 19
as the top column feed.
[0024] The remaining portion of feed stream 31, stream 33, is
cooled in heat exchanger 10 by heat exchange with cool residue gas
(stream 41a) and propane refrigerant. Note that in all cases
exchangers 10 and 13 are representative of either a multitude of
individual heat exchangers or a single multi-pass heat exchanger,
or any combination thereof. (The decision as to whether to use more
than one heat exchanger for the indicated cooling services will
depend on a number of factors including, but not limited to, inlet
gas flow rate, heat exchanger size, stream temperatures, etc.) The
cooled stream 33a enters separator 12 at -30.degree. F.
[-35.degree. C.] and 870 psia [6,002 kPa(a)] where the vapor
(stream 34) is separated from the condensed liquid (stream 35). The
separator liquid (stream 35) is expanded to slightly above the
operating pressure of fractionation tower 19 by expansion valve 18.
Expanded stream 35a is heated from -63.degree. F. [-53.degree. C.]
to 23.degree. F. [-5.degree. C.] in heat exchanger 13 as described
earlier before heated stream 35b is supplied to fractionation tower
19 at a lower mid-column feed point.
[0025] The vapor (stream 34) from separator 12 enters a work
expansion machine 15 in which mechanical energy is extracted from
this portion of the high pressure feed. The machine 15 expands the
vapor substantially isentropically to the tower operating pressure,
with the work expansion cooling the expanded stream 34a to a
temperature of approximately -99.degree. F. [-73.degree. C.]. The
typical commercially available expanders are capable of recovering
on the order of 80-85% of the work theoretically available in an
ideal isentropic expansion. The work recovered is often used to
drive a centrifugal compressor (such as item 16) that can be used
to re-compress the residue gas (stream 41b), for example. The
partially condensed expanded stream 34a is thereafter supplied as
feed to fractionation tower 19 at an tipper mid-column feed
point.
[0026] The demethanizer in tower 19 is a conventional distillation
column containing a plurality of vertically spaced trays, one or
more packed beds, or some combination of trays and packing. As is
often the case in natural gas processing plants, the fractionation
tower may consist of three sections. The upper section 19a is a
separator wherein the partially vaporized top feed is separated
into its respective vapor and liquid portions, and wherein the
vapor rising from the rectifying section 19b below is combined with
the vapor portion of the top feed to form the cold demethanizer
overhead vapor (stream 41) which exits the top of the tower at
-137.degree. F. [-94.degree. C.]. The middle absorbing (rectifying)
section 19b contains the trays and/or packing to provide the
necessary contact between the vapor portions of the expanded
streams 34a and 35b rising upward and cold liquid falling downward.
The lower stripping (demethanizing) section 19c contains additional
trays and/or packing to provide the necessary contact between the
liquids falling downward and the vapors rising upward.
Demethanizing section 19c also includes one or more reboilers (such
as the reboiler 21 and side reboiler 20 shown in FIG. 1) which heat
and vaporize a portion of the liquids flowing down the column to
provide the stripping vapors which flow up the column to strip the
liquid product, stream 44, of methane and lighter components.
Stream 34a enters demethanizer 19 at an intermediate feed position
located below rectifying section 19h and above demethanizing
section 19c. The liquid portion of the expanded stream 34a
commingles with liquids falling downward from rectifying section
19b and the combined liquid continues downward into demethanizing
section 19c of column 19. The vapor portion of the expanded stream
34a commingles with vapors rising upward from demethanizing section
19c and the combined vapor rises upward through rectifying section
19b and is contacted with cold liquid falling downward to condense
and absorb the C.sub.2 components, C.sub.3 components, and heavier
components from the vapor.
[0027] The liquid product (stream 44) exits the bottom of tower 19
at 79.degree. F. [26.degree. C.], based on a typical specification
of a methane concentration of 0.07% on a volume basis in the bottom
product. The cold residue gas stream 41 passes countercurrently to
one portion of the feed gas in heat exchanger 13 where it is heated
to -31.degree. F. [-35.degree. C.](stream 41a), and
countercurrently to the other portion of the feed gas in heat
exchanger 10 where it is heated to 88.degree. F. [31.degree.
C.](stream 41b). The residue gas is then re-compressed in two
stages. The first stage is compressor 16 driven by expansion
machine 15. The second stage is compressor 22 driven by a
supplemental power source which compresses the residue gas (stream
41d) to sales line pressure. After cooling to 110.degree. F.
[43.degree. C.] in discharge cooler 23, the residue gas product
(stream 41e) flows to the sales gas pipeline at 918 psia [6,330
kPa(a)], sufficient to meet line requirements (usually on the order
of the inlet pressure).
[0028] A summary of stream flow rates and energy consumption for
the process illustrated in FIG. 1 is set forth in the following
table:
TABLE-US-00001 TABLE I (FIG. 1) Stream Flow Summary-Lb. Moles/Hr
[kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31
22,875 3,898 1,843 1,240 32,972 32 6,428 1,095 518 348 9,265 33
16,447 2,803 1,325 892 23,707 34 14,127 1,566 382 94 18,283 35
2,320 1,237 943 798 5,424 41 22,867 504 20 2 26,507 44 8 3,394
1,823 1,238 6,465 Recoveries* Ethane 87.06% Propane 98.92% Butanes+
99.88% Power Residue Gas Compression 16,044 HP [26,376 kW]
Refrigerant Compression 7,492 HP [12,317 kW] Total Compression
23,536 HP [38,693 kW] *(Based on un-rounded flow rates)
[0029] FIG. 2 represents an alternative prior art process according
to U.S. Pat. No. 5,568,737. The process of FIG. 2 has been applied
to the same feed gas composition and conditions as described above
for FIG. 1. In the simulation of this process, as in the simulation
for the process of FIG. 1, operating conditions were selected to
maximize the recovery level for a given energy consumption.
[0030] After feed stream 31 is divided into two portions, streams
32 and 33, stream 32 is cooled to substantial condensation in heat
exchanger 13 by heat exchange with cold residue gas (stream 41),
flashed separator liquids (stream 35a), and propane refrigerant.
The resulting substantially condensed stream 32a at -140.degree. F.
[-96.degree. C.] is then flash expanded through expansion valve 17
to the operating pressure (approximately 340 psia [2,346 kPa(a)])
of fractionation tower 19. During expansion a portion of the stream
is vaporized, resulting in cooling of expanded stream 32b to
-159.degree. F. [106.degree. C.] before it is supplied to
fractionation tower 19 at an upper mid-column feed point.
[0031] The remaining portion of feed stream 31, stream 33, is
cooled in heat exchanger 10 by heat exchange with a portion (stream
47) of the cool demethanizer overhead vapor (stream 45a) and
propane refrigerant, and cooled stream 33a enters separator 12 at
-25.degree. F. [-32.degree. C.] and 870 psia [6,002 kPa(a)] where
the vapor (stream 34) is separated from the condensed liquid
(stream 35). The separator liquid (stream 35) is expanded to
slightly above the operating pressure of fractionation lower 19 by
expansion valve 18. Expanded stream 35a is heated from -54.degree.
F. [-48.degree. C.] to 23.degree. F. [-5.degree. C.] in heat
exchanger 13 as described earlier before heated stream 35b is
supplied to fractionation tower 19 at a lower mid-column feed
point.
[0032] The vapor (stream 34) from separator 12 enters a work
expansion machine 15 in which mechanical energy is extracted from
this portion of the high pressure feed. The machine 15 expands the
vapor substantially isentropically to the tower operating pressure,
with the work expansion cooling the expanded stream 34a to a
temperature of approximately -89.degree. F. [-67.degree. C.]. The
partially condensed expanded stream 34a is thereafter supplied as
feed to fractionation tower 19 at an intermediate mid-column feed
point.
[0033] The recompressed and cooled distillation vapor stream 45e is
divided into two streams. One portion, stream 41, is the volatile
residue gas product. The other portion, recycle stream 48, flows to
heat exchanger 11 where it is cooled to 0.degree. F. [-18.degree.
C.] by heat exchange with the remaining portion (stream 46) of cool
demethanizer overhead vapor stream 45a. The cooled recycle stream
48a then flows to exchanger 13 where it is further cooled to
-140.degree. F. [-96.degree. C.] and substantially condensed by
heat exchange with cold distillation vapor stream 45 and propane
refrigeration. The substantially condensed stream 48b is then
expanded through an appropriate expansion device, such as expansion
valve 14, to the demethanizer operating pressure, resulting in
cooling of the total stream to -167.degree. F. [-111.degree. C.].
The expanded stream 48c then supplied to fractionation tower 19 as
the top column feed. The vapor portion of stream 48c combines with
the vapors rising from the top fractionation stage of the column to
form distillation vapor stream 45, which is withdrawn from an upper
region of the tower.
[0034] The liquid product (stream 44) exits the bottom of tower 19
at 88.degree. F. [31.degree. C.], based on a methane concentration
of 0.07% on a volume basis in the bottom product. The demethanizer
overhead vapor stream 45 passes countercurrently to one portion of
the feed gas (stream 32) and the partially cooled recycle stream
(stream 48a) in heat exchanger 13 where it is heated to -27.degree.
F. [-33.degree. C.](stream 45a) and then divided into two portions,
stream 46 and stream 47. Stream 46 passes countercurrently to
recycle stream 48 in heat exchanger 11 and is heated to 105.degree.
F. [41.degree. C.](stream 46a), while stream 47 passes
countercurrently to the other portion of the feed gas in heat
exchanger 10 where it is heated to 91.degree. F. [33.degree. C.]
(stream 47a). Streams 46a and 47a recombine as stream 45b at
92.degree. F. [33.degree. C.], which is then re-compressed in two
stages, compressor 16 driven by expansion machine 15 and compressor
22 driven by a supplemental power source. After cooling to
110.degree. F. [43.degree. C.] in discharge cooler 23, stream 45e
is split into the residue gas product (stream 41) and the recycle
stream 48 as described earlier. Residue gas stream 41 then flows to
the sales gas pipeline 918 psia [6,330 kPa(a)].
[0035] A summary of stream flow rates and energy consumption for
the process illustrated in FIG. 2 is set forth in the following
table:
TABLE-US-00002 TABLE II (FIG. 2) Stream Flow Summary-Lb. Moles/Hr
[kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31
22,875 3,898 1,843 1,240 32,972 32 5,513 939 444 299 7,946 33
17,362 2,959 1,399 941 25,026 34 15,187 1,761 450 114 19,761 35
2,175 1,198 949 827 5,265 48 3,035 69 0 0 3,518 45 25,902 592 0 0
30,022 41 22,867 523 0 0 26,504 44 8 3,375 1,843 1,240 6,468
Recoveries* Ethane 86.57% Propane 100.00% Butanes+ 100.00% Power
Residue Gas Compression 17,363 HP [28,544 kW] Refrigerant
Compression 6,214 HP [10,216 kW] Total Compression 23,577 HP
[38,760 kW] *(Based on un-rounded flow rates)
[0036] A comparison of Tables I and II shows that, compared to the
FIG. 1 process, the FIG. 2 process has slightly lower ethane
recovery (86.57% versus 87.06%), but improves propane recovery from
98.92% to 100.00% and butanes+recovery from 99.88% to 100.00%.
Comparison of Tables I and II further shows that the improvement in
yields for the FIG. 2 process was achieved using essentially the
same power requirements.
[0037] FIG. 3 represents an alternative prior art process according
to U.S. Pat. No. 7,191,617. The process of FIG. 3 has been applied
to the same feed gas composition and conditions as described above
for FIGS. 1 and 2. In the simulation of this process, as in the
simulation for the processes of FIGS. 1 and 2, operating conditions
were selected to maximize the recovery level for a given energy
consumption.
[0038] After feed stream 31 is divided into two portions, streams
32 and 33, stream 32 is cooled to substantial condensation in heat
exchanger 13 by heat exchange with cold residue gas (stream 41),
flashed separator liquids (stream 35a), and propane refrigerant.
The resulting substantially condensed stream 32a at -120.degree. F.
[-84.degree. C.] is then flash expanded through expansion valve 17
to the operating pressure (approximately 324 psia [2,235 kPa(a)])
of fractionation tower 19. During expansion a portion of the stream
is vaporized, resulting in cooling of expanded stream 32b to
-153.degree. F. [-103.degree. C.] before it is supplied to
fractionation tower 19 at an upper mid-column feed point.
[0039] The remaining portion of feed stream 31, stream 33, is
cooled in heat exchanger 10 by heat exchange with cool residue gas
(stream 41a) and propane refrigerant, and cooled stream 33a enters
separator 12 at -34.degree. F. [-36.degree. C.] and 870 psia [6,002
kPa(a)] where the vapor (stream 34) is separated from the condensed
liquid (stream 35). The separator liquid (stream 35) is expanded to
slightly above the operating pressure of fractionation tower 19 by
expansion valve 18. Expanded stream 35a is heated from -66.degree.
F. [-54.degree. C.] to 21.degree. F. [-6.degree. C.] in heat
exchanger 13 as described earlier before heated stream 35b is
supplied to fractionation tower 19 at a lower mid-column feed
point.
[0040] The vapor (stream 34) from separator 12 enters a work
expansion machine 15 in which mechanical energy is extracted from
this portion of the high pressure feed. The machine 15 expands the
vapor substantially isentropically to the tower operating pressure,
with the work expansion cooling the expanded stream 34a to a
temperature of approximately -100.degree. F. [-74.degree. C.]. The
partially condensed expanded stream 34a is thereafter supplied as
feed to fractionation tower 19 at an intermediate mid-column teed
point.
[0041] A portion of the distillation vapor (stream 49) is withdrawn
from an upper region of the demethanizing section in fractionation
column 19, below the feed position of expanded stream 34a.
Distillation vapor stream 49 is then cooled from -100.degree. F.
[-73.degree. C.] to -146.degree. F. [-99.degree. C.] and partially
condensed (stream 49a) in heat exchanger 24 by heat exchange with
the cold demethanizer overhead stream 48 exiting the top of
demethanizer 19 at -150.degree. F. [-101.degree. C.]. The cold
demethanizer overhead stream is warmed to -118.degree. F.
[-84.degree. C.] (stream 48a) as it cools and condenses a portion
of stream 49.
[0042] The operating pressure in reflux separator 25 is maintained
slightly below the operating pressure of demethanizer 19. This
provides the driving force which causes distillation vapor stream
49 to flow through heat exchanger 24 and thence into the reflux
separator 25 where the condensed liquid (stream 51) is separated
from the uncondensed vapor (stream 50). Stream 51 then combines
with the warmed demethanizer overhead stream 48a from heat
exchanger 24 to form cold residue gas stream 41 at -123.degree. F.
[-86.degree. C.].
[0043] The liquid stream 51 from reflux separator 25 is pumped by
pump 26 to a pressure slightly above the operating pressure of
demethanizer 19, and stream 51a is then supplied as cold top column
feed (reflux) to demethanizer 19 at -145.degree. F. [-98.degree.
C.]. This cold liquid reflux absorbs and condenses the C.sub.2
components, C.sub.3 components, and heavier components rising in
the upper region of the rectifying section of demethanizer 19.
[0044] The liquid product (stream 44) exits the bottom of tower 19
at 81.degree. F. [27.degree. C.], based on a methane concentration
of 0.07% on a volume basis in the bottom product. The cold residue
gas stream 41 passes countercurrently to one portion of the feed
gas in heat exchanger 13 where it is heated to -35.degree. F.
[-37.degree. C.](stream 41a), and countercurrently to the other
portion of the feed gas in heat exchanger 10 where it is heated to
88.degree. F. [31.degree. C.](stream 41b). The residue gas is then
re-compressed in two stages, compressor 16 driven by expansion
machine 15 and compressor 22 driven by a supplemental power source.
After cooling to 110.degree. F. [43.degree. C.] in discharge cooler
23, the residue gas product (stream 41e) flows to the sales gas
pipeline at 918 psia [6,330 kPa(a)].
[0045] A summary of stream flow rates and energy consumption for
the process illustrated in FIG. 3 is set forth in the following
table:
TABLE-US-00003 TABLE III (FIG. 3) Stream Flow Summary-Lb. Moles/Hr
[kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31
22,875 3,898 1,843 1,240 32,972 32 6,062 1,033 488 329 8,737 33
16,813 2,865 1,355 911 24,235 34 14,234 1,525 360 86 18,356 35
2,579 1,340 995 825 5,879 49 4,853 364 18 1 5,443 50 3,384 49 0 0
3,622 51 1,469 315 18 1 1,821 48 19,483 288 2 0 22,698 41 22,867
337 0 0 26,320 44 8 3,561 1,841 1,240 6,652 Recoveries* Ethane
91.35% Propane 99.88% Butanes+ 100.00% Power Residue Gas
Compression 15,932 HP [26,192 kW] Refrigerant Compression 7,640 HP
[12,560 kW] Total Compression 23,572 HP [38,752 kW] *(Based on
un-rounded flow rates)
[0046] A comparison of Tables I, II, and III shows that the FIG. 3
process improves the ethane recovery from 87.06% (for FIG. 1) and
86.57% (for FIG. 2) to 91.35%. The propane recovery for the FIG. 3
process (99.88%) is significantly higher than that of the FIG. 1
process (98.92%) but slightly lower than that of the FIG. 2 process
(100.00%). The butanes+recovery for the FIG. 3 process (100.00%) is
slightly higher than that of the FIG. 1 process (99.88%) and the
same as that of the FIG. 2 process (100.00%). Comparison of Tables
I, II, and III further shows that the improvement in yields for the
FIG. 3 process was achieved using essentially the same power
requirements.
[0047] FIG. 4 represents an alternative prior art process according
to co-pending application Ser. No. 11/839,693. The process of FIG.
4 has been applied to the same feed gas composition and conditions
as described above for FIGS. 1 through 3. In the simulation of this
process, as in the simulation for the process of FIGS. 1 through 3,
operating conditions were selected to maximize the recovery level
for a given energy consumption.
[0048] After feed stream 31 is divided into two portions, streams
32 and 33, stream 32 is cooled to substantial condensation in heat
exchanger 13 by heat exchange with cold residue gas (stream 41),
flashed separator liquids (stream 35a), and propane refrigerant.
The resulting substantially condensed stream 32a at -151.degree. F.
[-101.degree. C.] is then flash expanded through expansion valve 17
to the operating pressure (approximately 319 psia [2,202 kPa(a)])
of fractionation tower 19. During expansion a portion of the stream
is vaporized, resulting in cooling of expanded stream 32b to
-165.degree. F. [-109.degree. C.] before it is supplied to
fractionation tower 19 at an upper mid-column feed point.
[0049] The remaining portion of feed stream 31, stream 33, is
cooled in heat exchanger 10 by heat exchange with cool residue gas
(stream 41a) and propane refrigerant, and cooled stream 33a enters
separator 12 at -40.degree. F. [-40.degree. C.] and 870 psia [6,002
kPa(a)] where the vapor (stream 34) is separated from the condensed
liquid (stream 35). The separator liquid (stream 35) is expanded to
slightly above the operating pressure of fraction tower 19 by
expansion valve 18. Expanded stream 35a is heated from -73.degree.
F. [-59.degree. C.] to 4.degree. F. [-16.degree. C.] in heat
exchanger 13 as described earlier before heated stream 35b is
supplied to fractionation tower 19 at a lower mid-column feed
point.
[0050] The vapor (stream 34) from separator 12 enters a work
expansion machine 15 in which mechanical energy is extracted from
this portion of the high pressure feed. The machine 15 expands the
vapor substantially isentropically to the tower operating pressure,
with the work expansion cooling the expanded stream 34a to a
temperature of approximately -107.degree. F. [-77.degree. C.]. The
partially condensed expanded stream 34a is thereafter supplied as
feed to fractionation tower 19 at an intermediate mid-column feed
point.
[0051] A portion of the distillation vapor (stream 49) is withdrawn
from an upper region of the demethanizing section in fractionation
column 19 below expanded stream 34a at -102.degree. F. [-75.degree.
C.] and is compressed to approximately 486 psia [3,353 kPa(a)] by
vapor compressor 27. The compressed stream 49a is then cooled from
-52.degree. F. [-47.degree. C.] to -151.degree. F. [-101.degree.
C.] and substantially condensed (stream 49b) in heat exchanger 13
as described earlier.
[0052] Since substantially condensed stream 49b is at a pressure
greater than the operating pressure of demethanizer 19, it is flash
expanded through expansion valve 14 to the operating pressure of
fractionation tower 19. During expansion a portion of the stream is
vaporized, resulting in cooling of the total stream to -159.degree.
F. [-106.degree. C.]. The expanded stream 49e is then supplied as
cold top column feed (reflux) to demethanizer 19. The vapor portion
of stream 49c combines with the distillation vapor rising from the
upper fractionation stage to form residue gas stream 41 exiting the
top of demethanizer 19 at -154.degree. F. [-103.degree. C.], while
the cold liquid reflux portion absorbs and condenses the C.sub.2
components, C.sub.3 components, and heavier components rising in
the upper region of the rectifying section of demethanizer 19.
[0053] The liquid product (stream 44) exits the bottom of tower 19
at 79.degree. F. [26.degree. C.], based on a methane concentration
of 0.07% on a volume basis in the bottom product. The cold residue
gas stream 41 passes countercurrently to one portion of the feed
gas in heat exchanger 13 where it is heated to -56.degree. F.
[-49.degree. C.] (stream 41a), and countercurrently to the other
portion of the feed gas in heat exchanger 10 where it is heated to
92.degree. F. [33.degree. C.] (stream 41b). The residue gas is then
re-compressed in two stages, compressor 16 driven by expansion
machine 15 and compressor 22 driven by a supplemental power source.
After cooling to 110.degree. F. [43.degree. C.] in discharge cooler
23, the residue gas product (stream 41e) flows to the sales gas
pipeline at 918 psia [6,330 kPa(a)].
[0054] A summary of stream flow rates and energy consumption for
the process illustrated in FIG. 4 is set forth in the following
table:
TABLE-US-00004 TABLE IV (FIG. 4) Stream Flow Summary-Lb. Moles/Hr
[kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31
22,875 3,898 1,843 1,240 32,972 32 4,529 772 365 245 6,529 33
18,346 3,126 1,478 995 26,443 34 15,105 1,522 341 79 19,366 35
3,241 1,604 1,137 916 7,077 49 2,429 179 8 0 2,722 41 22,867 263 1
0 26,244 44 8 3,635 1,842 1,240 6,728 Recoveries* Ethane 93.26%
Propane 99.95% Butanes+ 100.00% Power Residue Gas Compression
15,859 HP [26,072 kW] Vapor Compression 351 HP [577 kW] Refrigerant
Compression 7,366 HP [12,110 kW] Total Compression 23,572 HP
[38,759 kW] *(Based on un-rounded flow rates)
[0055] A comparison of Tables I, II, III, and IV shows that the
FIG. 4 process improves the ethane recovery from 87.06% (for FIG.
1), 86.57% (for FIG. 2), and 91.35% (for FIG. 3) to 93.26%. The
propane recovery for the FIG. 4 process (99.95%) is significantly
higher than that of the FIG. 1 process (98.92%) and higher than
that of the FIG. 3 process (99.88%), but slightly lower than that
of the FIG. 2 process (100.00%) The butanes+recovery for the FIG. 4
process (100.00%) is slightly higher than that of the FIG. 1
process (99.88%) and the same as that of the FIG. 2 process and the
FIG. 3 process (100.00% for both). Comparison of Tables I, II, III,
and IV further shows that the improvement in yields for the FIG. 4
process was achieved using essentially the same power
requirements.
DESCRIPTION OF THE INVENTION
[0056] FIG. 5 illustrates a flow diagram of a process in accordance
with the present invention. The feed gas composition and conditions
considered in the process presented in FIG. 5 are the same as those
in FIGS. 1 through 4. Accordingly, the FIG. 5 process can be
compared with that of the FIGS. 1 through 4 processes to illustrate
the advantages of the present invention.
[0057] In the simulation of the FIG. 5 process, inlet gas enters
the plant at 104.degree. F. [40.degree. C.] and 896 psia [6,181
kPa(a)] as stream 31. After feed stream 31 is divided into two
portions, streams 32 and 33, stream 32 is cooled to substantial
condensation in heat exchanger 13 by heat exchange with cold
residue gas (stream 41), flashed separator liquids (stream 35a),
and propane refrigerant. The resulting substantially condensed
stream 32a at -148.degree. F. [-100.degree. C.] is then flash
expanded through expansion valve 17 to the operating pressure
(approximately 324 psia [2,235 kPa(a)]) of fractionation tower 19.
During expansion a portion of the stream is vaporized, resulting in
cooling of expanded stream 32b to -164.degree. F. [-109.degree. C.]
before it is supplied to fractionation tower 19 at an upper
mid-column feed point.
[0058] The remaining portion of feed stream 31, stream 33, is
cooled in heat exchanger 10 by heat exchange with cool residue gas
(stream 41a) and propane refrigerant, and cooled stream 33a enters
separator 12 at -36.degree. F. [-38.degree. C.] and 870 psia [6,002
kPa(a)] where the vapor (stream 34) is separated from the condensed
liquid (stream 35). The separator liquid (stream 35) is expanded to
slightly above the operating pressure of fractionation tower 19 by
expansion valve 18. Expanded stream 35a is heated from -68.degree.
F. [-56.degree. C.] to 10.degree. F. [-12.degree. C.] in heat
exchanger 13 as described earlier before heated stream 35b is
supplied to fractionation tower 19 at a lower mid-column feed
point.
[0059] The vapor (stream 34) from separator 12 is divided into two
streams, 36 and 39. Stream 36, containing about 16% of the total
vapor, is cooled to -148.degree. F. [-100.degree. C.] and
substantially condensed (stream 36a) in heat exchanger 13 as
described earlier. Substantially condensed stream 36a is flash
expanded through expansion valve 14 to the operating pressure of
fractionation tower 19. During expansion a portion of the stream is
vaporized, resulting in cooling of the total stream. In the process
illustrated in FIG. 5, the expanded stream 36b leaving expansion
valve 14 reaches a temperature of -169.degree. F. [-112.degree. C.]
before it is supplied as cold top column feed (reflux) to
demethanizer 19. This cold liquid reflux absorbs and condenses the
C.sub.2 components, C.sub.3 components, and heavier components
rising in the upper region of the rectifying section of
demethanizer 19.
[0060] The remaining 84% of the vapor from separator 12 (stream 39)
enters a work expansion machine 15 in which mechanical energy is
extracted from this portion of the high pressure feed. The machine
15 expands the vapor substantially isentropically to the tower
operating pressure, with the work expansion cooling the expanded
stream 39a to a temperature of approximately -103.degree. F.
[-75.degree. C.]. The partially condensed expanded stream 39a is
thereafter supplied as feed to fractionation tower 19 at an
intermediate mid-column feed point.
[0061] The demethanizer in tower 19 is a conventional distillation
column containing a plurality of vertically spaced trays, one or
more packed beds, or some combination of trays and packing. The
demethanizer tower consists of three sections. The upper section
19a is a separator wherein the partially vaporized top feed (stream
36b) is separated into its respective vapor and liquid portions,
and wherein the vapor rising from the rectifying section 19b below
is combined with the vapor portion of the top feed to form the cold
demethanizer overhead vapor (stream 41) which exits the top of the
tower at -151.degree. F. [-102.degree. C.]. The middle absorbing
(rectifying) section 19b contains the trays and/or packing to
provide the necessary contact between the vapor portions of the
expanded streams 32b, 39a, and 35b rising upward and cold liquid
falling downward. The lower stripping (demethanizing) section 19c
contains additional trays and/or packing to provide the necessary
contact between the liquids falling downward and the vapors rising
upward. Demethanizing section 19c also includes one or more
reboilers (such as the reboiler 21 and side reboiler 20 shown in
FIG. 5) which heat and vaporize a portion of the liquids flowing
down the column to provide the stripping vapors which flow up the
column to strip the liquid product, stream 44, of methane and
lighter components. Stream 39a enters demethanizer 19 at an
intermediate feed position located below rectifying section 19b and
above demethanizing section 19c. The liquid portion of the expanded
stream 39a commingles with liquids falling downward from rectifying
section 19b and the combined liquid continues downward into
demethanizing section 19c of column 19. The vapor portion of the
expanded stream 39a commingles with vapors rising upward from
demethanizing section 19c and the combined vapor rises upward
through rectifying section 19b and is contacted with cold liquid
falling downward to condense and absorb the C.sub.2 components,
C.sub.3 components, and heavier components from the vapor.
[0062] In demethanizing section 19c of demethanizer 19, the feed
streams are stripped of their methane and lighter components. The
resulting liquid product (stream 44) exits the bottom of tower 19
at 80.degree. F. [27.degree. C.] based on a methane concentration
of 0.07% on a volume basis in the bottom product. The cold residue
gas stream 41 leaves demethanizer 19 and passes countercurrently to
one portion of the feed gas and a portion of the separator vapor in
heat exchanger 13 where it is heated to -44.degree. F. [-42.degree.
C.] (stream 41a) and countercurrently to the other portion of the
feed gas in heat exchanger 10 where it is heated to 91.degree. F.
[33.degree. C.] (stream 41b) as it provides cooling as previously
described. The residue gas is then re-compressed in two stages,
compressor 16 driven by expansion machine 15 and compressor 22
driven by a supplemental power source. After cooling to 110.degree.
F. [43.degree. C.] in discharge cooler 23, the residue gas product
(stream 41e) flows to the sales gas pipeline at 918 psia [6,330
kPa(a)].
[0063] A summary of stream flow rates and energy consumption for
the process illustrated in FIG. 5 is set forth in the following
table:
TABLE-US-00005 TABLE V (FIG. 5) Stream Flow Summary-Lb. Moles/Hr
[kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31
22,875 3,898 1,843 1,240 32,972 32 4,804 819 387 260 6,924 33
18,071 3,079 1,456 980 26,048 34 15,151 1,588 368 87 19,496 35
2,920 1,491 1,088 893 6,552 36 2,379 249 58 14 3,061 49 12,772
1,339 310 73 16,435 41 22,867 265 6 0 26,251 44 8 3,633 1,837 1,240
6,721 Recoveries* Ethane 93.22% Propane 99.70% Butanes+ 99.99%
Power Residue Gas Compression 15,881 HP [26,108 kW] Refrigerant
Compression 7,678 HP [12,623 kW] Total Compression 23,559 HP
[38,731 kW] *(Based on un-rounded flow rates)
[0064] A comparison of Tables I through V shows that, compared to
the prior art, the present invention improves the ethane recovery
from 87.06% (for FIG. 1), 86.57% (for FIG. 2), and 91.35% (for FIG.
3) to 93.22%, with essentially the same recovery as the FIG. 4
process (93.26%). The propane recovery for the present invention
(99.70%) is significantly higher than that of the FIG. 1 process
(98.92%) and somewhat lower than that of the FIG. 2 process
(100.00%), the FIG. 3 process (99.88%), and the FIG. 4 process
(99.95%). The butanes+recovery for the present invention (99.99%)
is slightly higher than that of the FIG. 1 process (99.88%) and
essentially the same as that of the FIG. 2 process, the FIG. 3
process, and the FIG. 4 process (100.00% for all three). Comparison
of Tables I through V further shows that the improvement in yields
for the present invention shown in FIG. 5 was achieved using
essentially the same power requirements as the FIG. 1 through 4
processes.
[0065] The improvement in the recovery efficiency of the present
invention over that of the prior art processes can be understood by
examining the improvement in the rectification that the present
invention provides for rectifying section 19b. Compared to the
prior art of the FIG. 1 process, the present invention produces a
better top reflux stream containing more methane and less C.sub.2+
components. Comparing reflux stream 32 in Table I for the FIG. 1
prior art process with first reflux stream 36 in Table V for the
present invention, it can be seen that the present invention
provides a top reflux stream with a significantly lower
concentration of C.sub.2+ components (10.5% for the present
invention versus 21.2% fir the FIG. 1 prior art process). Further,
the present invention uses a second reflux stream (stream 32) at an
intermediate feed point of rectifying section 19b to provide bulk
recovery of the C.sub.2 components, C.sub.3 components, and heavier
hydrocarbon components contained in expanded feed 39a and the
vapors rising from demethanizing section 19c. This means that less
rectification is required in the upper region of rectifying section
19b, minimizing the required flow rate for top reflux stream 36 so
that more of the separator vapor (stream 34) flows to expansion
machine 15 in stream 39, which produces more power to drive
compressor 16 and reduces the power required by compressor 22. Note
that the total reflux provided to rectification section 19b in
streams 36 and 32 for the present invention is nearly 8% higher
than that provided in the single reflux stream 32 of the FIG. 1
process.
[0066] Compared to the prior art of the FIG. 2 process, the present
invention supplies its feed streams to column 19 at significantly
lower temperatures, reducing the quantity of vapor entering
rectification section 19b and reducing the quantity of reflux
required. Recycle stream 48 used in the FIG. 2 process to produce
top reflux stream 48c adds to the cooling load imposed on cold
demethanizer overhead stream 45, reducing the cooling available in
heat exchanger 10 such that the portion of the feed gas in stream
33a entering separator 12 for the FIG. 2 process is much warmer
than that of the present invention. As can be seen from Table II
for the FIG. 2 process, despite top reflux stream 48 having almost
15% more flow than the present invention and a much lower
concentration of C.sub.2+ components (2.0% for the FIG. 2 prior art
process versus 10.5% for the present invention), it cannot rectify
the vapors rising in rectification section 19b as well as top
reflux stream 36 of the present invention.
[0067] Compared to the prior art of the FIG. 3 process, the present
invention produces a much greater quantity of top reflux containing
more methane and less C.sub.2+ components. Comparing reflux stream
51 in Table III for the FIG. 3 prior art process with top reflux
stream 36 in Table V for the present invention, it can be seen that
the present invention provides a top reflux stream with 68% more
flow and with a significantly lower concentration of C.sub.2+
components (10.5% for the present invention versus 18.3% for the
FIG. 3 prior art process). Note that although the total reflux
provided to rectification section 19b in streams 51 and 32 of the
FIG. 3 process is nearly 6% higher than that of the present
invention, the much higher concentration of C.sub.2 components in
top reflux stream 51 of the FIG. 3 process (nearly twice that in
top reflux stream 36 of the present invention) prevents it from
providing efficient rectification.
[0068] Compared to the prior art of the FIG. 4 process, the present
invention achieves essentially the same rectification with its two
reflux streams. Comparing reflux stream 49 in Table IV for the FIG.
4 prior art process with top reflux stream 36 in Table V for the
present invention, it can be seen that the present invention
provides a top reflux stream with 12% more flow but with a higher
concentration of C.sub.2+ components (10.5% for the present
invention versus 6.9% for the FIG. 4 prior art process). Further,
the total reflux provided to rectification section 19b for the
present invention is nearly 8% higher than that in streams 49 and
32 of the FIG. 4 process. This higher reflux flow allows the
present invention to match the rectification of the FIG. 4 prior
art process despite the higher concentration of C.sub.2+ components
in its top reflux stream.
[0069] However, the present invention is able to match the recovery
of the FIG. 4 process without vapor compressor 27 required by the
prior art process. Vapor compressors in this service are expensive
to install and to operate, adding to both the capital cost and the
maintenance cost of the plant, reducing revenue and reducing the
return on investment. Rotating equipment like this also adds to the
complexity of operating the plant, making it more difficult to
optimize the process for maximum recovery with minimum energy
consumption. Thus, the present invention is less expensive to
build, less expensive to maintain, and easier to operate than the
prior art of the FIG. 4 process.
Other Embodiments
[0070] FIGS. 6 through 9 display other embodiments of the present
invention. FIGS. 5 through 7 depict fractionation towers
constructed in a single vessel. FIGS. 8 and 9 depict fractionation
towers constructed in two vessels, absorber (rectifying) column 28
(a contacting and separating device) and partial rectification
stripper (distillation) column 19. In such cases, the overhead
vapor stream 52 from partial rectification stripper column 19 flows
to the lower section of absorber column 28 to be contacted and
further rectified by substantially condensed stream 36b. Pump 29 is
used to route the liquids (stream 53) from the bottom of absorber
column 28 to the top of stripper column 19 so that the two towers
effectively function as one distillation system. The decision
whether to construct the fractionation tower as a single vessel
(such as demethanizer 19 in FIGS. 5 through 7) or multiple vessels
will depend on a number of factors such as plant size, the distance
to fabrication facilities, etc.
[0071] In accordance with this invention, it is generally
advantageous to design the absorbing (rectifying) section of the
demethanizer to contain multiple theoretical separation stages.
However, the benefits of the present invention can be achieved with
as few as two theoretical stages. For instance, all or a part of
expanded substantially condensed stream 36b leaving expansion valve
14 and all or a part of expanded substantially condensed stream 32b
(FIG. 5) or 38b (FIGS. 6 through 9) from expansion valve 17 can be
combined (such as in the piping joining the expansion valves to the
demethanizer) and if thoroughly intermingled, the vapors and
liquids will mix together and separate in accordance with the
relative volatilities of the various components of the total
combined streams. Such commingling of the two streams, combined
with contacting at least a portion of expanded stream 39a, shall be
considered for the purposes of this invention as constituting an
absorbing section.
[0072] In accordance with the present invention, the splitting of
the feed gas may be accomplished in several ways. In the processes
of FIGS. 5, 6, and 8, the splitting of the feed gas occurs before
any cooling of the feed gas. In such cases, cooling and substantial
condensation of one portion of the feed gas in multiple heat
exchangers may be favored in some circumstances, such as heat
exchangers 11 and 13 shown in FIGS. 6 and 8. The feed gas may also
be split, however, following cooling (but prior to separation of
any liquids which may have been formed) as shown in FIGS. 7 and
9.
[0073] The high pressure liquid (stream 35 in FIGS. 5 through 9)
need not be expanded, heated, and fed to a mid-column feed point on
the distillation column. Instead, all or a portion of it may be
combined with the portion of the cooled feed gas (stream 32a in
FIGS. 6 and 8 or stream 32 in FIGS. 7 and 9) flowing to heat
exchanger 13. (This is shown by the dashed stream 37 in FIGS. 6
through 9.) Any remaining portion of the liquid (stream 40 in FIGS.
6 through 9) may be expanded through an appropriate expansion
device, such as expansion valve 18 or an expansion machine, and fed
to a mid-column feed point on the distillation column (stream 40a).
Stream 40 may also be used for inlet gas cooling or other heat
exchange service before or after the expansion step prior to
flowing to the demethanizer.
[0074] As described earlier, a portion of the feed gas (stream 32)
and a portion of the separator vapor (stream 36) are substantially
condensed and the resulting condensate used to absorb valuable
C.sub.2 components, C.sub.3 components and heavier components from
the vapors rising through rectifying section 19b of demethanizer 19
(FIGS. 5 through 7), or through absorber column 28 and the upper
section of partial rectification stripper column 19 (FIGS. 8 and
9). However, the present invention is not limited to this
embodiment. It may be advantageous, for instance, to treat only a
portion of these vapors in this manner, or to use only a portion of
the condensate as an absorbent, in cases where other design
considerations indicate portions of the vapors or the condensate
should bypass rectifying section 19b of demethanizer 19 (FIGS. 5
through 7), or absorber column 28 and/or partial rectification
stripper column 19 (FIGS. 8 and 9).
[0075] Feed gas conditions, plant size, available equipment, or
other factors may indicate that elimination of work expansion
machine 15, or replacement with an alternate expansion device (such
as an expansion valve), is feasible. Although individual stream
expansion is depicted in particular expansion devices, alternative
expansion means may be employed where appropriate. For example,
conditions may warrant work expansion of the substantially
condensed portion of the separator vapor (stream 36a in FIGS. 5
through 9), the substantially condensed portion of the feed stream
(stream 32a in FIG. 5), or the substantially condensed combined
stream (stream 38a in FIGS. 6 through 9).
[0076] In accordance with the present invention, the use of
external refrigeration to supplement the cooling available to the
inlet gas and/or separator vapor from other process streams may be
employed, particularly in the case of a rich inlet gas. The use and
distribution of separator liquids and demethanizer side draw
liquids for process heat exchange, and the particular arrangement
of heat exchangers for inlet gas and separator vapor cooling must
be evaluated for each particular application, as well as the choice
of process streams for specific heat exchange services.
[0077] It will also be recognized that the relative amount of feed
found in each branch of the split vapor feeds will depend on
several factors, including gas pressure, feed gas composition, the
amount of heat which can economically be extracted from the feed,
and the quantity of horsepower available. More feed to the top of
the column may increase recovery while decreasing power recovered
from the expander thereby increasing the recompression horsepower
requirements. Increasing feed lower in the column reduces the
horsepower consumption but may also reduce product recovery. The
relative locations of the mid-column feeds may vary depending on
inlet composition or other factors such as desired recovery levels
and amount of liquid formed during inlet gas cooling. Moreover, two
or more of the feed streams, or portions thereof, may be combined
depending on the relative temperatures and quantities of individual
streams, and the combined stream then fed to a mid-column feed
position.
[0078] The present invention allows reduced capital expenditures
and/or provides improved recovery of C.sub.2 components, C.sub.3
components, and heavier hydrocarbon components or of C.sub.3
components and heavier hydrocarbon components per amount of utility
consumption required to operate the process. An improvement in
utility consumption required for operating the demethanizer or
deethanizer process may appear in the form of reduced power
requirements for compression or re-compression, reduced power
requirements for external refrigeration, reduced energy
requirements for lower reboilers, or a combination thereof.
[0079] While there have been described what are believed to be
preferred embodiments of the invention, those skilled in the art
will recognize that other and further modifications may be made
thereto, e.g. to adapt the invention to various conditions, types
of feed, or other requirements without departing from the spirit of
the present invention as defined by the following claims.
* * * * *