Title: LNG production in cryogenic natural gas processing plants
Abstract: A process for liquefying natural gas in conjunction with processing natural gas to recover natural gas liquids (NGL) is disclosed. In the process, the natural gas stream to be liquefied is taken from one of the streams in the NGL recovery plant and cooled under pressure to condense it. A distillation stream is withdrawn from the NGL recovery plant to provide some of the cooling required to condense the natural gas stream. A portion of the condensed stream is expanded to an intermediate pressure and then used to provide some of the cooling required to condense the natural gas stream, and thereafter routed to the NGL recovery plant so that any heavier hydrocarbons it contains can be recovered in the NGL product. The remaining portion of the condensed stream is expanded to low pressure to form the liquefied natural gas stream.
Patent Number: 6,889,523 Issued on 05/10/2005 to Wilkinson, et al.
| Inventors:
|
Wilkinson; John D. (Midland, TX);
Hudson; Hank M. (Midland, TX);
Cuellar; Kyle T. (Katy, TX)
|
| Assignee:
|
ElkCorp (Dallas, TX)
|
| Appl. No.:
|
384038 |
| Filed:
|
March 7, 2003 |
| Current U.S. Class: |
62/613; 62/621 |
| Intern'l Class: |
F25J 001/00; F25J003/00 |
| Field of Search: |
62/613,611,620,621,625
|
References Cited [Referenced By]
U.S. Patent Documents
| 2952984 | Sep., 1960 | Marshall.
| |
| 3292380 | Dec., 1966 | Bucklin.
| |
| 3837172 | Sep., 1974 | Markbreiter et al.
| |
| 4140504 | Feb., 1979 | Campbell et al.
| |
| 4157904 | Jun., 1979 | Campbell et al.
| |
| 4171964 | Oct., 1979 | Campbell et al.
| |
| 4185978 | Jan., 1980 | McGalliard et al.
| |
| 4251249 | Feb., 1981 | Gulsby.
| |
| 4278457 | Jul., 1981 | Campbell et al.
| |
| 4445917 | May., 1984 | Chiu.
| |
| 4519824 | May., 1985 | Huebel.
| |
| 4525185 | Jun., 1985 | Newton.
| |
| 4545795 | Oct., 1985 | Liu et al.
| |
| 4600421 | Jul., 1986 | Kummann.
| |
| 4617039 | Oct., 1986 | Buck.
| |
| 4687499 | Aug., 1987 | Aghili.
| |
| 4689063 | Aug., 1987 | Paradowski et al.
| |
| 4690702 | Sep., 1987 | Paradowski et al.
| |
| 4707170 | Nov., 1987 | Ayres et al.
| |
| 4710214 | Dec., 1987 | Sharma et al.
| |
| 4755200 | Jul., 1988 | Liu et al.
| |
| 4851020 | Jul., 1989 | Montgomery IV.
| |
| 4854955 | Aug., 1989 | Campbell et al.
| |
| 4869740 | Sep., 1989 | Campbell et al.
| |
| 4889545 | Dec., 1989 | Campbell et al.
| |
| 4895584 | Jan., 1990 | Buck et al.
| |
| RE33408 | Oct., 1990 | Khan et al.
| |
| 5114451 | May., 1992 | Rambo et al.
| |
| 5275005 | Jan., 1994 | Campbell et al.
| |
| 5291736 | Mar., 1994 | Paradowski.
| |
| 5363655 | Nov., 1994 | Kikkawa et al.
| |
| 5365740 | Nov., 1994 | Kikkawa et al.
| |
| 5555748 | Sep., 1996 | Campbell et al.
| |
| 5566554 | Oct., 1996 | Vijayaraghavan et al.
| |
| 5568737 | Oct., 1996 | Campbell et al.
| |
| 5600969 | Feb., 1997 | Low.
| |
| 5615561 | Apr., 1997 | Houshmand et al.
| |
| 5651269 | Jul., 1997 | Prevost et al.
| |
| 5755114 | May., 1998 | Foglietta.
| |
| 5755115 | May., 1998 | Manley.
| |
| 5771712 | Jun., 1998 | Campbell et al.
| |
| 5799507 | Sep., 1998 | Wilkinson et al.
| |
| 5881569 | Mar., 1999 | Campbell et al.
| |
| 5890378 | Apr., 1999 | Rambo et al.
| |
| 5893274 | Apr., 1999 | Nagelvoort et al.
| |
| 5983664 | Nov., 1999 | Campbell et al.
| |
| 6014869 | Jan., 2000 | Elion et al.
| |
| 6023942 | Feb., 2000 | Thomas et al.
| |
| 6053007 | Apr., 2000 | Victory et al.
| |
| 6062041 | May., 2000 | Kikkawa et al.
| |
| 6116050 | Sep., 2000 | Yao et al.
| |
| 6119479 | Sep., 2000 | Roberts et al.
| |
| 6125653 | Oct., 2000 | Shu et al.
| |
| 6182469 | Feb., 2001 | Campbell et al.
| |
| 6250105 | Jun., 2001 | Kimble.
| |
| 6269655 | Aug., 2001 | Roberts et al.
| |
| 6272882 | Aug., 2001 | Hodges et al.
| |
| 6308531 | Oct., 2001 | Roberts et al.
| |
| 6324867 | Dec., 2001 | Fanning et al.
| |
| 6336344 | Jan., 2002 | O'Brien.
| |
| 6347532 | Feb., 2002 | Agrawal et al.
| |
| 6363744 | Apr., 2002 | Finn et al.
| |
| 6367286 | Apr., 2002 | Price.
| |
| 6526777 | Mar., 2003 | Campbell et al.
| |
| 2003/0158458 | Aug., 2003 | Prim.
| |
| Foreign Patent Documents |
| WO 0188447 | Nov., 2001 | WO.
| |
Other References
U.S. Appl. No. 09/677,220, filed Oct. 2, 2000, Wilkinson et al.
U.S. Appl. No. 10/161,780, filed Jun. 4, 2002, Wilkinson et al.
U.S. Appl. No. 10/278,610, filed Oct. 23, 2002, Wilkinson et al.
Finn, Adrain J., Grant L. Johnson, and Terry R. Iomlinson, "LNG Technology for
Offshore and Mid-Scale Plants", Proceedings of the Seventy-Ninth Annual Convention
of the Gas Processors Association, pp. 429-450, Atlanta, Georgia, Mar. 13-15, 2000.
Kikkawa, Yoshitsugi, Masaaki Ohishi, and Noriyoshi Nozawa, "Optimize the Power
System of Baseload LNG Plant", Proceedings of the Eightieth Annual Convention of
the Gas Processors Association, San Antonino, Texas, Mar. 12-14, 2001.
Price, Brian C., "LNG Production for Peak Shaving Operations", Proceedings of
the Seventy-Eighth Annual Convention of the Gas Processors Association, pp. 273-280,
Nashville, Tennessee, Mar. 1-3, 1999.
|
Primary Examiner: Doerrler; William C.
Attorney, Agent or Firm: Fitzpatrick, Cella, Harper & Scinto
Claims
1. A process for liquefying a natural gas stream containing methane and heavier
hydrocarbon components wherein
(a) said natural gas stream is withdrawn from a cryogenic natural gas processing
plant recovering natural gas liquids;
(b) said natural gas stream is cooled under pressure to condense at least a portion
of it and form a condensed stream;
(c) a distillation stream is withdrawn from said plant to supply at least a portion
of said cooling of said natural gas stream;
(d) a first portion of said condensed stream is withdrawn, expanded to an intermediate
pressure, and directed in heat exchange relation with said natural gas stream to
supply at least a portion of said cooling, whereupon said first portion is directed
to said plant; and
(e) the remaining portion of said condensed stream is expanded to lower pressure
to form said liquefied natural gas stream.
2. A process for liquefying a natural gas stream containing methane and heavier
hydrocarbon components wherein
(a) said natural gas stream is withdrawn from a cryogenic natural gas processing
plant recovering natural gas liquids;
(b) said natural gas stream is cooled under pressure sufficiently to partially
condense it;
(c) a distillation stream is withdrawn from said plant to supply at least a portion
of said cooling of said natural gas stream;
(d) said partially condensed natural gas stream is separated into a liquid stream
and a vapor stream, whereupon said liquid stream is directed to said plant;
(e) said vapor stream is further cooled at pressure to condense at least a portion
of it and form a condensed stream;
(f) a first portion of said condensed stream is withdrawn, expanded to an intermediate
pressure, and directed in heat exchange relation with said expanded vapor stream
to supply at least a portion of said cooling, whereupon said first portion is directed
to said plant; and
(g) the remaining portion of said condensed stream is expanded to lower pressure
to form said liquefied natural gas stream.
3. A process for liquefying a natural gas stream containing methane and heavier
hydrocarbon components wherein
(a) said natural gas stream is withdrawn from a cryogenic natural gas processing
plant recovering natural gas liquids;
(b) said natural gas stream is cooled under pressure sufficiently to partially
condense it;
(c) a distillation stream is withdrawn from said plant to supply at least a portion
of said cooling of said natural gas stream;
(d) said partially condensed natural gas stream is separated into a liquid stream
and a vapor stream, whereupon said liquid stream is directed to said plant;
(e) said vapor stream is expanded to an intermediate pressure and further cooled
at said intermediate pressure to condense at least a portion of it and form a condensed
stream;
(f) a first portion of said condensed stream is withdrawn, expanded to an intermediate
pressure, and directed in heat exchange relation with said expanded vapor stream
to supply at least a portion of said cooling, whereupon said first portion is directed
to said plant; and
(g) the remaining portion of said condensed stream is expanded to lower pressure
to form said liquefied natural gas stream.
4. A process for liquefying a natural gas stream containing methane and heavier
hydrocarbon components wherein
(a) said natural gas stream is withdrawn from a cryogenic natural gas processing
plant recovering natural gas liquids;
(b) said natural gas stream is cooled under pressure;
(c) a distillation stream is withdrawn from said plant to supply at least a portion
of said cooling of said natural gas stream;
(d) said cooled natural gas stream is expanded to an intermediate pressure and
further cooled at said intermediate pressure to condense at least a portion of
it and form a condensed stream;
(e) a first portion of said condensed stream is withdrawn, expanded to an intermediate
pressure, and directed in heat exchange relation with said expanded natural gas
stream to supply at least a portion of said cooling, whereupon said first portion
is directed to said plant; and
(f) the remaining portion of said condensed stream is expanded to lower pressure
to form said liquefied natural gas stream.
5. An apparatus for liquefying a natural gas stream containing methane and heavier
hydrocarbon components comprising
(a) first withdrawing means connected to a cryogenic natural gas processing plant
recovering natural gas liquids to withdraw said natural gas stream;
(b) heat exchange means connected to said first withdrawing means to receive
said natural gas stream and cool it under pressure to condense at least a portion
of it and form a condensed stream;
(c) second withdrawing means connected to said plant to withdraw a distillation
stream, said second withdrawing means being further connected to said heat exchange
means to heat said distillation stream and thereby supply at least a portion of
said cooling of said natural gas stream;
(d) third withdrawing means connected to said heat exchange means to withdraw
a first portion of said condensed stream;
(e) first expansion means connected to said third withdrawing means to receive
said first portion and expand it to an intermediate pressure, said first expansion
means being further connected to supply said expanded first portion to said heat
exchange means to heat said expanded first portion and thereby supply at least
a portion of said cooling, whereupon said heated expanded first portion is directed
to said plant; and
(f) second expansion means connected to said heat exchange means to receive the
remaining portion of said condensed stream and expand it to lower pressure to form
said liquefied natural gas stream.
6. An apparatus for liquefying a natural gas stream containing methane and heavier
hydrocarbon components comprising
(a) first withdrawing means connected to a cryogenic natural gas processing plant
recovering natural gas liquids to withdraw said natural gas stream;
(b) heat exchange means connected to said first withdrawing means to receive
said natural gas stream and cool it under pressure sufficiently to partially condense
it;
(c) second withdrawing means connected to said plant to withdraw a distillation
stream, said second withdrawing means being further connected to said heat exchange
means to heat said distillation stream and thereby supply at least a portion of
said cooling of said natural gas stream;
(d) separation means connected to said heat exchange means to receive said partially
condensed natural gas stream and to separate it into a vapor stream and a liquid
stream, whereupon said liquid stream is directed to said plant;
(e) said separation means being further connected to supply said vapor stream
to said heat exchange means, with said heat exchange means being adapted to further
cool said vapor stream at pressure to condense at least a portion of it and form
a condensed stream;
(f) third withdrawing means connected to said heat exchange means to withdraw
a first portion of said condensed stream;
(g) first expansion means connected to said third withdrawing means to receive
said first portion and expand it to an intermediate pressure, said first expansion
means being further connected to supply said expanded first portion to said heat
exchange means to heat said expanded first portion and thereby supply at least
a portion of said cooling, whereupon said heated expanded first portion is directed
to said plant; and
(h) second expansion means connected to said heat exchange means to receive the
remaining portion of said condensed stream and expand it to lower pressure to form
said liquefied natural gas stream.
7. An apparatus for liquefying a natural gas stream containing methane and heavier
hydrocarbon components comprising
(a) first withdrawing means connected to a cryogenic natural gas processing plant
recovering natural gas liquids to withdraw said natural gas stream;
(b) heat exchange means connected to said first withdrawing means to receive
said natural gas stream and cool it under pressure sufficiently to partially condense
it;
(c) second withdrawing means connected to said plant to withdraw a distillation
stream, said second withdrawing means being further connected to said heat exchange
means to heat said distillation stream and thereby supply at least a portion of
said cooling of said natural gas stream;
(d) separation means connected to said heat exchange means to receive said partially
condensed natural gas stream and to separate it into a vapor stream and a liquid
stream, whereupon said liquid stream is directed to said plant;
(e) first expansion means connected to said separation means to receive said
vapor stream and expand it to an intermediate pressure, said first expansion means
being further connected to supply said expanded vapor stream to said heat exchange
means, with said heat exchange means being adapted to further cool said expanded
vapor stream at said intermediate pressure to condense at least a portion of it
and form a condensed stream;
(f) third withdrawing means connected to said heat exchange means to withdraw
a first portion of said condensed stream;
(g) second expansion means connected to said third withdrawing means to receive
said first portion and expand it to an intermediate pressure, said second expansion
means being further connected to supply said expanded first portion to said heat
exchange means to heat said expanded first portion and thereby supply at least
a portion of said cooling, whereupon said heated expanded first portion is directed
to said plant; and
(h) third expansion means connected to said heat exchange means to receive the
remaining portion of said condensed stream and expand it to lower pressure to form
said liquefied natural gas stream.
8. An apparatus for liquefying a natural gas stream containing methane and heavier
hydrocarbon components comprising
(a) first withdrawing means connected to a cryogenic natural gas processing plant
recovering natural gas liquids to withdraw said natural gas stream;
(b) heat exchange means connected to said first withdrawing means to receive
said natural gas stream and cool it under pressure;
(c) second withdrawing means connected to said plant to withdraw a distillation
stream, said second withdrawing means being further connected to said heat exchange
means to heat said distillation stream and thereby supply at least a portion of
said cooling of said natural gas stream;
(d) first expansion means connected to said heat exchange means to receive said
cooled natural gas stream and expand it to an intermediate pressure, said first
expansion means being further connected to supply said expanded natural gas stream
to said heat exchange means, with said heat exchange means being adapted to further
cool said expanded natural gas stream at said intermediate pressure to condense
at least a portion of it and form a condensed stream;
(e) third withdrawing means connected to said heat exchange means to withdraw
a first portion of said condensed stream;
(f) second expansion means connected to said third withdrawing means to receive
said first portion and expand it to an intermediate pressure, said second expansion
means being further connected to supply said expanded first portion to said heat
exchange means to heat said expanded first portion and thereby supply at least
a portion of said cooling, whereupon said heated expanded first portion is directed
to said plant; and
(g) third expansion means connected to said heat exchange means to receive the
remaining portion of said condensed stream and expand it to lower pressure to form
said liquefied natural gas stream.
Description
BACKGROUND OF THE INVENTION
This invention relates to a process for processing natural gas to produce liquefied
natural gas (LNG) that has a high methane purity. In particular, this invention
is well suited to co-production of LNG by integration into natural gas processing
plants that recover natural gas liquids (NGL) and/or liquefied petroleum gas (LPG)
using a cryogenic process.
Natural gas is typically recovered from wells drilled into underground reservoirs.
It usually has a major proportion of methane, i.e., methane comprises at least
50 mole percent of the gas. Depending on the particular underground reservoir,
the natural gas also contains relatively lesser amounts of heavier hydrocarbons
such as ethane, propane, butanes, pentanes and the like, as well as water, hydrogen,
nitrogen, carbon dioxide, and other gases.
Most natural gas is handled in gaseous form. The most common means for transporting
natural gas from the wellhead to gas processing plants and thence to the natural
gas consumers is in high pressure gas transmission pipelines. In a number of circumstances,
however, it has been found necessary and/or desirable to liquefy the natural gas
either for transport or for use. In remote locations, for instance, there is often
no pipeline infrastructure that would allow for convenient transportation of the
natural gas to market. In such cases, the much lower specific volume of LNG relative
to natural gas in the gaseous state can greatly reduce transportation costs by
allowing delivery of the LNG using cargo ships and transport trucks.
Another circumstance that favors the liquefaction of natural gas is for its
use as a motor vehicle fuel. In large metropolitan areas, there are fleets of buses,
taxi cabs, and trucks that could be powered by LNG if there was an economic source
of LNG available. Such LNG-fueled vehicles produce considerably less air pollution
due to the clean-burning nature of natural gas when compared to similar vehicles
powered by gasoline and diesel engines which combust higher molecular weight hydrocarbons.
In addition, if the LNG is of high purity (i.e., with a methane purity of 95 mole
percent or higher), the amount of carbon dioxide (a "greenhouse gas") produced
is considerably less due to the lower carbon:hydrogen ratio for methane compared
to all other hydrocarbon fuels.
The present invention is generally concerned with the liquefaction of natural
gas as a co-product in a cryogenic gas processing plant that also produces natural
gas liquids (NGL) such as ethane, propane, butanes, and heavier hydrocarbon components.
A typical analysis of a natural gas stream to be processed in accordance with this
invention would be, in approximate mole percent, 92.3% methane, 4.4% ethane and
other C
2 components, 1.5% propane and other C
3 components,
0.3% iso-butane, 0.3% normal butane, 0.3% pentanes plus, with the balance made
up of nitrogen and carbon dioxide. Sulfur containing gases are also sometimes present.
There are a number of methods known for liquefying natural gas. For instance,
see Finn, Adrian J., Grant L. Johnson, and Terry R. Tomlinson, "LNG Technology
for Offshore and Mid-Scale Plants", Proceedings of the Seventy-Ninth Annual Convention
of the Gas Processors Association, pp. 429-450, Atlanta, Ga., Mar. 13-15, 2000
and Kikkawa, Yoshitsugi, Masaaki Ohishi, and Noriyoshi Nozawa, "Optimize the Power
System of Baseload LNG Plant", Proceedings of the Eightieth Annual Convention of
the Gas Processors Association, San Antonio, Tex., Mar. 12-14, 2001 for surveys
of a number of such processes. U.S. Pat. Nos. 4,445,917; 4,525,185; 4,545,795;
4,755,200; 5,291,736; 5,363,655; 5,365,740; 5,600,969; 5,615,561; 5,651,269; 5,755,114;
5,893,274; 6,014,869; 6,053,007; 6,062,041; 6,119,479; 6,125,653; 6,250,105 B1;
6,269,655 B1; 6,272,882 B1; 6,308,531 B1; 6,324,867 B1; 6,347,532 B1; International
Publication Number WO 01/88447 A1 published Nov. 22, 2001; our co-pending U.S.
patent application Ser. No. 09/839,907 filed Apr. 20, 2001; our co-pending U.S.
patent application Ser. No. 10/161,780 filed Jun. 4, 2002; and our co-pending U.S.
patent application Ser. No. 10/278,610 filed Oct. 23, 2002 also describe relevant
processes. These methods generally include steps in which the natural gas is purified
(by removing water and troublesome compounds such as carbon dioxide and sulfur
compounds), cooled, condensed, and expanded. Cooling and condensation of the natural
gas can be accomplished in many different manners. "Cascade refrigeration" employs
heat exchange of the natural gas with several refrigerants having successively
lower boiling points, such as propane, ethane, and methane. As an alternative,
this heat exchange can be accomplished using a single refrigerant by evaporating
the refrigerant at several different pressure levels. "Multi-component refrigeration"
employs heat exchange of the natural gas with one or more refrigerant fluids composed
of several refrigerant components in lieu of multiple single-component refrigerants.
Expansion of the natural gas can be accomplished both isenthalpically (using Joule-Thomson
expansion, for instance) and isentropically (using a work-expansion turbine, for instance).
While any of these methods could be employed to produce vehicular grade LNG,
the capital and operating costs associated with these methods have generally made
the installation of such facilities uneconomical. For instance, the purification
steps required to remove water, carbon dioxide, sulfur compounds, etc. from the
natural gas prior to liquefaction represent considerable capital and operating
costs in such facilities, as do the drivers for the refrigeration cycles employed.
This has led the inventors to investigate the feasibility of integrating LNG production
into cryogenic gas processing plants used to recover NGL from natural gas. Such
an integrated LNG production method would eliminate the need for separate gas purification
facilities and gas compression drivers. Further, the potential for integrating
the cooling/condensation for the LNG liquefaction with the process cooling required
for NGL recovery could lead to significant efficiency improvements in the LNG liquefaction method.
In accordance with the present invention, it has been found that LNG with a methane
purity in excess of 99 percent can be co-produced from a cryogenic NGL recovery
plant without reducing the NGL recovery level using less energy than prior art
processes. 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° F. [-46° C.] or colder.
For a better understanding of the present invention, reference is made to the
following examples and drawings. Referring to the drawings:
FIG. 1 is a flow diagram of a prior art cryogenic natural gas processing plant
in accordance with U.S. Pat. No. 4,278,457;
FIG. 2 is a flow diagram of said cryogenic natural gas processing plant when
adapted for co-production of LNG in accordance with a prior art process;
FIG. 3 is a flow diagram of said cryogenic natural gas processing plant when
adapted for co-production of LNG using a prior art process in accordance with U.S.
Pat. No. 5,615,561;
FIG. 4 is a flow diagram of said cryogenic natural gas processing plant when
adapted for co-production of LNG in accordance with an embodiment of our co-pending
U.S. patent application Ser. No. 09/839,907;
FIG. 5 is a flow diagram of said cryogenic natural gas processing plant when
adapted for co-production of LNG in accordance with the present invention;
FIG. 6 is a flow diagram illustrating an alternative means of application of
the present invention for co-production of LNG from said cryogenic natural gas
processing plant; and
FIG. 7 is a flow diagram illustrating another alternative means of application
of the present invention for co-production of LNG from said cryogenic natural gas
processing plant.
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.
For convenience, process parameters are reported in both the traditional British
units and in the units of the International System of Units (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. The
LNG production rates reported as gallons per day (gallons/D) and/or pounds per
hour (Lbs/hour) correspond to the stated molar flow rates in pound moles per hour.
The LNG production rates reported as cubic meters per day (m
3/D) and/or
kilograms per hour (kg/H) correspond to the stated molar flow rates in kilogram
moles per hour.
DESCRIPTION OF THE PRIOR ART
Referring now to FIG. 1, for comparison purposes we begin with an example
of an NGL recovery plant that does not co-produce LNG. In this simulation of a
prior art NGL recovery plant according to U.S. Pat. No. 4,278,457, inlet gas enters
the plant at 90° F. [32° C.] and 740 psia [5,102 kPa(a)] as stream 31.
If the inlet gas contains a concentration of carbon dioxide and/or sulfur compounds
which would prevent the product streams from meeting specifications, these 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.
The feed stream 31 is cooled in heat exchanger 10 by heat exchange
with cool demethanizer overhead vapor at -66° F. [-55° C.] (stream 36
a),
bottom liquid product at 56° F. [13° C.] (stream 41
a) from
demethanizer bottoms pump 18, demethanizer reboiler liquids at 36°
F. [2° C.] (stream 40), and demethanizer side reboiler liquids at -35°
F. [-37° C.] (stream 39). Note that in all cases heat exchanger 10
is 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 31
a enters
separator 11 at -43° F. [-42° C.] and 725 psia [4,999 kPa(a)]
where the vapor (stream 32) is separated from the condensed liquid (stream 35).
The vapor (stream 32) from separator 11 is divided into two streams,
33 and 34. Stream 33, containing about 27% of the total vapor,
passes through heat exchanger 12 in heat exchange relation with the demethanizer
overhead vapor stream 36, resulting in cooling and substantial condensation
of stream 33
a. The substantially condensed stream 33
a at
-142° F. [-97° C.] is then flash expanded through an appropriate expansion
device, such as expansion valve 13, to the operating pressure (approximately
320 psia [2,206 kPa(a)]) of fractionation tower 17. 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 33
b leaving expansion
valve 13 reaches a temperature of -153° F. [-103° C.], and is
supplied to separator section 17
a in the upper region of fractionation
tower 17. The liquids separated therein become the top feed to demethanizing
section 17
b.
The remaining 73% of the vapor from separator 11 (stream 34) enters
a work expansion machine 14 in which mechanical energy is extracted from
this portion of the high pressure feed. The machine 14 expands the vapor
substantially isentropically from a pressure of about 725 psia [4,999 kPa(a)] to
the tower operating pressure, with the work expansion cooling the expanded stream
34
a to a temperature of approximately -107° F. [-77° 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
15) that can be used to re-compress the residue gas (stream 38),
for example. The expanded and partially condensed stream 34
a is supplied
as a feed to the distillation column at an intermediate point. The separator liquid
(stream 35) is likewise expanded to the tower operating pressure by expansion
valve 16, cooling stream 35
a to -72° F. [-58° C.]
before it is supplied to the demethanizer in fractionation tower 17 at a
lower mid-column feed point.
The demethanizer in fractionation tower 17 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 two sections. The upper section
17
a is a separator wherein the partially vaporized top feed is divided
into its respective vapor and liquid portions, and wherein the vapor rising from
the lower distillation or demethanizing section 17
b is combined with
the vapor portion of the top feed to form the cold demethanizer overhead vapor
(stream 36) which exits the top of the tower at -150° F. [-101°
C.]. The lower, demethanizing section 17
b contains the trays and/or
packing and provides the necessary contact between the liquids falling downward
and the vapors rising upward. The demethanizing section also includes reboilers
which heat and vaporize a portion of the liquids flowing down the column to provide
the stripping vapors which flow up the column.
The liquid product stream 41 exits the bottom of the tower at 51°
F. [10° C.], based on a typical specification of a methane to ethane ratio
of 0.028:1 on a molar basis in the bottom product. The stream is pumped to approximately
650 psia [4,482 kPa(a)] (stream 41
a) in pump 18. Stream 41
a,
now at about 56° F. [13° C.], is warmed to 85° F. [29°
C.] (stream 41
b) in heat exchanger 10 as it provides cooling
to stream 31. (The discharge pressure of the pump is usually set by the
ultimate destination of the liquid product. Generally the liquid product flows
to storage and the pump discharge pressure is set so as to prevent any vaporization
of stream 41
b as it is warmed in heat exchanger 10.)
The demethanizer overhead vapor (stream 36) passes countercurrently to
the incoming feed gas in heat exchanger 12 where it is heated to -66°
F. [-55° C.] (stream 36
a) and heat exchanger 10 where
it is heated to 68° F. [20° C.] (stream 36
b). A portion
of the warmed demethanizer overhead vapor is withdrawn to serve as fuel gas (stream
37) for the plant, with the remainder becoming the residue gas (stream 38).
(The amount of fuel gas that must be withdrawn is largely determined by the fuel
required for the engines and/or turbines driving the gas compressors in the plant,
such as compressor 19 in this example.) The residue gas is re-compressed
in two stages. The first stage is compressor 15 driven by expansion machine
14. The second stage is compressor 19 driven by a supplemental power
source which compresses the residue gas (stream 38
b) to sales line
pressure. After cooling to 120° F. [49° C.] in discharge cooler 20,
the residue gas product (stream 38
c) flows to the sales gas pipeline
at 740 psia [5,102 kPa(a)], sufficient to meet line requirements (usually on the
order of the inlet pressure).
A summary of stream flow rates and energy consumption for the process illustrated
in FIG. 1 is set forth in the following table:
| TABLE I |
| (FIG. 1) |
| Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] |
| |
| Stream |
Methane |
Ethane |
Propane |
Butanes+ |
Total |
| 31 |
35,473 |
1,689 |
585 |
331 |
38,432 |
| 32 |
35,210 |
1,614 |
498 |
180 |
37,851 |
| 35 |
263 |
75 |
87 |
151 |
581 |
| 33 |
9,507 |
436 |
134 |
49 |
10,220 |
| 34 |
25,703 |
1,178 |
364 |
131 |
27,631 |
| 36 |
35,432 |
211 |
6 |
0 |
35,951 |
| 37 |
531 |
3 |
0 |
0 |
539 |
| 38 |
34,901 |
208 |
6 |
0 |
35,412 |
| 41 |
41 |
1,478 |
579 |
331 |
2,481 |
| |
Recoveries* |
|
|
|
|
| |
Ethane |
87.52% |
| |
Propane |
98.92% |
| |
Butanes+ |
99.89% |
| |
Power |
| |
Residue Gas Compression |
14,517 |
HP |
[23,866 |
kW] |
| |
| |
*(Based on un-rounded flow rates) |
FIG. 2 shows one manner in which the NGL recovery plant in FIG. 1 can be adapted
for co-production of LNG, in this case by application of a prior art process for
LNG production similar to that described by Price (Brian C. Price, "LNG Production
for Peak Shaving Operations", Proceedings of the Seventy-Eighth Annual Convention
of the Gas Processors Association, pp. 273-280, Atlanta, Ga., Mar. 13-15, 2000).
The inlet gas composition and conditions considered in the process presented in
FIG. 2 are the same as those in FIG. 1. In this example and all that follow,
the simulation is based on co-production of a nominal 50,000 gallons/D [417 m
3/D]
of LNG, with the volume of LNG measured at flowing (not standard) conditions.
In the simulation of the FIG. 2 process, the inlet gas cooling, separation, and
expansion scheme for the NGL recovery plant is exactly the same as that used in
FIG. 1. In this case, the compressed and cooled demethanizer overhead vapor
(stream 45
c) produced by the NGL recovery plant is divided into two
portions. One portion (stream 38) is the residue gas for the plant and is
routed to the sales gas pipeline. The other portion (stream 71) becomes
the feed stream for the LNG production plant.
The inlet gas to the NGL recovery plant (stream 31) was not treated for
carbon dioxide removal prior to processing. Although the carbon dioxide concentration
in the inlet gas (about 0.5 mole percent) will not create any operating problems
for the NGL recovery plant, a significant fraction of this carbon dioxide will
leave the plant in the demethanizer overhead vapor (stream 36) and will
subsequently contaminate the feed stream for the LNG production section (stream
71). The carbon dioxide concentration in this stream is about 0.4 mole percent,
well in excess of the concentration that can be tolerated by this prior art process
(about 0.005 mole percent). Accordingly, the feed stream 71 must be processed
in carbon dioxide removal section 50 before entering the LNG production
section to avoid operating problems from carbon dioxide freezing. Although there
are many different processes that can be used for carbon dioxide removal, many
of them will cause the treated gas stream to become partially or completely saturated
with water. Since water in the feed stream would also lead to freezing problems
in the LNG production section, it is very likely that the carbon dioxide removal
section 50 must also include dehydration of the gas stream after treating.
The treated feed gas enters the LNG production section at 120° F. [49°
C.] and 730 psia [5,033 kPa(a)] as stream 72 and is cooled in heat exchanger
51 by heat exchange with a refrigerant mixture at -261° F. [-163°
C.] (stream 74
b). The purpose of heat exchanger 51 is to cool
the feed stream to substantial condensation and, preferably, to subcool the stream
so as to eliminate any flash vapor being generated in the subsequent expansion
step. For the conditions stated, however, the feed stream pressure is above the
cricondenbar, so no liquid will condense as the stream is cooled. Instead, the
cooled stream 72
a leaves heat exchanger 51 at -256° F.
[-160° C.] as a dense-phase fluid. (The cricondenbar is the maximum pressure
at which a vapor phase can exist in a multi-phase fluid. At pressures below the
cricondenbar, stream 72
a would typically exit heat exchanger 51
as a subcooled liquid stream.)
Stream 72
a enters a work expansion machine 52 in which
mechanical energy is extracted from this high pressure stream. The machine 52
expands the dense-phase fluid substantially isentropically from a pressure of about
728 psia [5,019 kPa(a)] to the LNG storage pressure (18 psia [124 kPa(a)]), slightly
above atmospheric pressure. The work expansion cools the expanded stream 72
b
to a temperature of approximately -257° F. [-160° C.], whereupon
it is then directed to the LNG storage tank 53 which holds the LNG product
(stream 73).
All of the cooling for stream 72 is provided by a closed cycle refrigeration
loop. The working fluid for this cycle is a mixture of hydrocarbons and nitrogen,
with the composition of the mixture adjusted as needed to provide the required
refrigerant temperature while condensing at a reasonable pressure using the available
cooling medium. In this case, condensing with ambient air has been assumed, so
a refrigerant mixture composed of nitrogen, methane, ethane, propane, and heavier
hydrocarbons is used in the simulation of the FIG. 2 process. The composition of
the stream, in approximate mole percent, is 5.2% nitrogen, 24.6% methane, 24.1%
ethane, and 18.0% propane, with the balance made up of heavier hydrocarbons.
The refrigerant stream 74 leaves partial condenser 56 at 120°
F. [49° C.] and 140 psia [965 kPa(a)]. It enters heat exchanger 51
and is condensed and then subcooled to -256° F. [-160° C.] by the flashed
refrigerant stream 74
b. The subcooled liquid stream 74
a
is flash expanded substantially isenthalpically in expansion valve 54
from about 138 psia [951 kPa(a)] to about 26 psia [179 kPa(a)]. During expansion
a portion of the stream is vaporized, resulting in cooling of the total stream
to -261° F. [-163° C.] (stream 74
b). The flash expanded
stream 74
b then reenters heat exchanger 51 where it provides
cooling to the feed gas (stream 72) and t
