U.S. patent number 3,724,226 [Application Number 05/135,615] was granted by the patent office on 1973-04-03 for lng expander cycle process employing integrated cryogenic purification.
This patent grant is currently assigned to Gulf Research & Development Company. Invention is credited to Robert W. Pachaly.
United States Patent |
3,724,226 |
Pachaly |
April 3, 1973 |
LNG EXPANDER CYCLE PROCESS EMPLOYING INTEGRATED CRYOGENIC
PURIFICATION
Abstract
A process and apparatus for the liquefaction of natural gas
wherein raw feedstock is cryogenically fractionated to remove
essentially all of the carbon dioxide and C.sub.5.sub.+
hydrocarbons therefrom, and wherein the cryogenically purified
feedstock is cooled and liquefied under pressure in a cryogenic
heat exchanger. The pressurized cold liquid from the heat exchanger
is isenthalpically expanded to reduce the pressure and further cool
the liquid while at the same time flashing a minor gas fraction.
Refrigeration for the liquefaction of the natural gas is supplied
by a circulating refrigerant stream which is compressed and
work-expanded to obtain the necessary cooling. The minor flash gas
portion of the liquefaction step is commingled with the circulating
refrigerant stream so that the analysis of the refrigerant stream
is always rich in the lighter portions of the liquefaction stream,
thus aiding in maintaining refrigeration temperature differentials
to drive the liquefaction step. The work-expanded refrigerant
portion undergoes a compression cycle and is work-expanded in a
series of expansion turbines. The expansion turbines furnish at
least part of the power necessary to drive the compressor system in
the refrigerant gas cycle.
Inventors: |
Pachaly; Robert W. (Greenwich,
CT) |
Assignee: |
Gulf Research & Development
Company (Pittsburgh, PA)
|
Family
ID: |
22468861 |
Appl.
No.: |
05/135,615 |
Filed: |
April 20, 1971 |
Current U.S.
Class: |
62/613 |
Current CPC
Class: |
F25J
1/0282 (20130101); F25J 3/0266 (20130101); F25J
1/004 (20130101); F25J 1/0283 (20130101); F25J
1/0035 (20130101); F25J 1/0022 (20130101); F25J
1/0202 (20130101); F25J 1/0278 (20130101); F25J
3/0209 (20130101); F25J 1/0288 (20130101); F25J
3/0233 (20130101); F25J 3/0247 (20130101); F25J
1/0037 (20130101); F25J 2290/72 (20130101); F25J
2220/66 (20130101); F25J 2205/50 (20130101); F25J
2220/68 (20130101); F25J 2220/64 (20130101); F25J
2200/74 (20130101); Y02C 10/12 (20130101); F25J
2230/20 (20130101); Y02C 20/40 (20200801); F25J
2245/02 (20130101); F25J 2205/04 (20130101); F25J
2240/02 (20130101); F25J 2270/88 (20130101); F25J
2220/62 (20130101); F25J 2200/72 (20130101); F25J
2270/06 (20130101) |
Current International
Class: |
F25J
1/00 (20060101); F25J 3/02 (20060101); F25j
001/00 (); F25j 003/00 (); F25j 003/02 () |
Field of
Search: |
;62/23,24,27,28,26,29,38,39,40,54,52,53,30,40 |
References Cited
[Referenced By]
U.S. Patent Documents
Primary Examiner: Yudkoff; Norman
Assistant Examiner: Purcell; Arthur F.
Claims
What is claimed is:
1. A process for liquifying a natural gas mixture comprising
nitrogen, carbon dioxide, water and lower boiling hydrocarbons so
as to produce a substantially carbon dioxide- and water-free
liquified natural gas, which process comprises:
1. Condensing the natural gas mixture under elevated pressure to
form (a) a liquid water phase, (b) a heavy hydrocarbon liquid phase
containing nitrogen and carbon dioxide and (c) a light hydrocarbon
vapor phase containing nitrogen and carbon dioxide;
2. Removing the liquid water phase from the process;
3. Substantially isentropically expanding the light hydrocarbon
vapor phase in a work recovery engine to obtain mechanical energy
and to cool the light hydrocarbon vapor phase;
4. Fractionating the heavy hydrocarbon liquid phase from Step (1)
and the cooled light hydrocarbon vapor phase from Step (3) to form
(a) a methane-nitrogen fraction substantially free of carbon
dioxide and hydrocarbons heavier than methane and (b) a carbon
dioxide-hydrocarbon fraction substantially free of nitrogen and
containing C.sub.2 and heavier hydrocarbons;
5. Fractionating the carbon dioxide-hydrocarbon fraction from Step
(4) into (a) a substantially carbon dioxide fraction, (b) a C.sub.5
and heavier hydrocarbon fraction and, (c) C.sub.2 to C.sub.5
hydrocarbon fraction and removing the carbon dioxide and C.sub.5
and heavier hydrocarbon fractions from the process;
6. Compressing and cooling the C.sub.2 to C.sub.5 hydrocarbon
fraction from Step (5) to form a low temperature liquid stream;
7. Combining the low temperature liquid stream of Step (6) with at
least a portion of the methane-nitrogen fraction from Step (4) so
as to form a nitrogen-methane and C.sub.2 to C.sub.5 hydrocarbon
mixture;
8. Cooling the nitrogen-methane and C.sub.2 to C.sub.5 hydrocarbon
mixture under elevated pressure by heat exchange with a combined
gas mixture consisting essentially of methane and nitrogen produced
subsequently whereby the combined gas mixture is heated and a
predominant portion of the nitrogen-methane and C.sub.2 to C.sub.5
hydrocarbon mixture is liquified so as to form a predominantly
liquid, liquid-vapor mixture;
9. Substantially isenthalpically reducing the pressure of the
cooled predominantly liquid, liquid-vapor nitrogenmethane and
C.sub.2 to C.sub.5 hydrocarbon mixture so as further to cool it and
to form a minor gas fraction consisting essentially of a mixture of
methane and nitrogen while liquifying completely the balance of the
cooled stream to form a major liquid fraction containing methane
and substantially all the hydrocarbons heavier than methane through
C.sub.5 ;
10. separating the minor, methane-nitrogen gas fraction from the
major liquid fraction and recovering the major liquid fraction as
liquified natural gas product;
11. Combining the minor, methane-nitrogen gas fraction with an
isentropically expanded, refrigerant gas mixture produced
subsequently and consisting essentially of methane and nitrogen to
form the combined gas mixture;
12. Employing the combined gas mixture formed in Step (11) at least
partially to cool the nitrogen-methane and C.sub.2 to C.sub.5
hydrocarbon mixture in Step (8);
13. Compressing a major portion of the heated, combined gas mixture
from Step (8) to provide a compressed refrigerant gas mixture;
14. Heat exchanging at least the major portion of the combined gas
mixture prior to the compression of Step (13) with the compressed
refrigerant gas mixture whereby the at least major portion of the
heated, combined gas mixture is further heated to a temperature
from about -50.degree. to about 300.degree.F. and the compressed
refrigerant gas mixture is cooled;
15. Additionally cooling the compressed refrigerant gas to a
temperature such that in expansion in a work recovery engine the
temperature and enthalpy are sufficiently low to provide the
refrigeration required for Step (8);
16. Substantially isentropically expanding the additionally cooled,
compressed refrigerant gas mixture in a work recovery engine to
obtain mechanical energy and to further cool the refrigerant gas,
whereby the isentropically expanded, refrigerant gas mixture is
formed;
17. Removing from the process system subsequent to Step (8) a minor
portion of the combined gas mixture substantially equal in
magnitude to the minor methane-nitrogen gas fraction formed in Step
(9); and
18. Employing the mechanical energy obtained from the work recovery
engines of Steps (3) and (16) to compress, at least partly, the
major portion of the combined refrigerant gas mixture in Step
(13).
2. The process of claim 1 wherein the nitrogen-methane and C.sub.2
to C.sub.5 hydrocarbon mixture in Step (8) is at a pressure in the
range of from about 350 to about 750 psia.
3. The process of claim 1 wherein the nitrogen-methane and C.sub.2
to C.sub.5 hydrocarbon mixture in Step (8) is maintained at a
pressure relative to its analysis such that the plot of the
temperature vs. enthalpy of the mixture approaches as straight a
line as is reasonably possible.
4. The process of claim 1 wherein the flow rate of the combined
methane-nitrogen gas mixture in Step (8) is from about 1.5 to about
4 times the flow rate of the nitrogen-methane and C.sub.2 to
C.sub.5 hydrocarbon mixture in Step (8), expressed as Mols per
hour.
5. The process of claim 3 wherein the flow rate of the combined
methane-nitrogen gas in Step (8) is maintained at such a rate that
the temperature vs. enthalpy curve for the combined gas as closely
as possible parallels the temperature vs. enthalpy curve of the
nitrogen-methane and C.sub.2 to C.sub.5 hydrocarbon mixture;
whereby more effective heat exchange between the combined gas and
the nitrogen-methane and C.sub.2 to C.sub.5 hydrocarbon mixture is
effected.
6. The process of claim 1 wherein the compressed, methane nitrogen
refrigerant gas mixture from Step (13) is at a pressure from about
400 to about 600 psia.
7. The process of claim 1 wherein the compression of Step (13) is
effected in a plurality of stages and the gas is cooled
interstage.
8. The process of claim 1 wherein the minor portion of the combined
gas mixture removed in Step (17) is charged to a separate work
recovery engine whereby mechanical energy is obtained, and wherein
the mechanical energy as obtained is employed at least to compress
partly the major portion of the combined gas mixture.
9. A process for liquifying a natural gas mixture comprising
nitrogen, carbon dioxide, water and lower boiling hydrocarbons so
as to produce a substantially carbon dioxide- and water-free liquid
natural gas, which process comprises:
1. Condensing the natural gas mixture under elevated pressure to
form (a) a liquid water phase, (b) a heavy hydrocarbon liquid phase
containing nitrogen and carbon dioxide and (c) a light hydrocarbon
vapor phase containing nitrogen and carbon dioxide;
2. Removing the liquid water phase from the process;
3. Substantially isentropically expanding the light hydrocarbon
vapor phase in a work recovery engine to obtain mechanical energy
and to cool the light hydrocarbon vapor phase;
4. Fractionating the heavy hydrocarbon liquid phase from Step (1)
and the cooled light hydrocarbon vapor phase from Step (3) to form
(a) a methane-nitrogen fraction substantially free of carbon
dioxide and hydrocarbons heavier than methane and (b) a carbon
dioxide-hydrocarbon fraction substantially free of nitrogen and
containing C.sub.2 and heavier hydrocarbons;
5. Separating the methane-nitrogen fraction into a first portion
and a second portion;
6. Heat exchanging the first portion with a first isentropically
expanded methane-nitrogen refrigerant gas obtained subsequently so
as to remove heat from the first portion and to heat the first
isentropically expanded methane-nitrogen refrigerant gas;
7. Returning the heat exchanged first portion to the fractionation
of Step (4) as reflux;
8. Fractionating the carbon dioxide-hydrocarbon fraction from Step
(4) to separate carbon dioxide from the C.sub.2 and heavier
hydrocarbons and removing the carbon dioxide from the process;
9. Fractionating the C.sub.2 and heavier hydrocarbons from Step (8)
to form a C.sub.2 -C.sub.4 hydrocarbon fraction and a C.sub.4 and
heavier hydrocarbon fraction and removing the C.sub.4 and heavier
hydrocarbon fraction from the process;
10. Compressing and cooling the C.sub.2 -C.sub.4 hydrocarbon
fraction from Step (9) to form a low temperature liquid stream;
11. Combining the low temperature liquid stream of Step (10) with
the second portion of the methane-nitrogen fraction from Step (4)
so as to form a nitrogen-methane and C.sub.2 -C.sub.4 hydrocarbon
mixture;
12. Cooling the nitrogen-methane and C.sub.2 -C.sub.4 hydrocarbon
mixture under elevated pressure by heat exchange with a first
isentropically expanded methane-nitrogen refrigerant gas obtained
subsequently, thereby heating the first isentropically expanded
refrigerant gas;
13. Cooling further the nitrogen-methane and C.sub.2 -C.sub.4
hydrocarbon mixture under elevated pressure by heat exchange with a
combined gas mixture consisting essentially of methane and nitrogen
produced subsequently whereby the combined gas mixture is heated
and a predominant portion of the nitrogenmethane and C.sub.2
-C.sub.4 hydrocarbon mixture is liquified so as to form a
predominantly liquid, liquid-vapor mixture;
14. Substantially isenthalpically reducing the pressure of the
cooled predominantly liquid, liquid-vapor nitrogen-methane and
C.sub.2 -C.sub.4 hydrocarbon mixture so as further to cool it and
to form a minor gas fraction consisting essentially of a mixture of
methane and nitrogen while liquifying completely the balance of the
cooled stream to form a major liquid fraction containing methane
and the C.sub.2 -C.sub.4 hydrocarbons;
15. Separating the minor, methane-nitrogen gas fraction from the
major liquid fraction and recovering the major liquid fraction as
liquified natural gas product;
16. Combining the minor, methane-nitrogen gas fraction with a
second isentropically expanded, refrigerant gas mixture produced
subsequently and consisting essentially of methane and nitrogen to
form the combined gas mixture;
17. Employing the combined gas mixture formed in Step (16) to cool
the nitrogen-methane and C.sub.2 -C.sub.4 hydrocarbon mixture in
Step (13);
18. Compressing at least a major portion of the heated, combined
gas mixture from Step (13) to provide a compressed refrigerant gas
mixture;
19. Heat exchanging at least a major portion of the combined gas
mixture prior to the compression of Step (18) with the compressed
refrigerant gas mixture whereby the at least major portion of the
heated, combined gas mixture is further heated to a temperature
from about -50.degree. to about 300.degree.F. and the compressed
refrigerant gas mixture is cooled;
20. Additionally cooling the compressed refrigerant gas to a
temperature such that in expansion in a work recovery engine the
temperature and enthalpy are sufficiently low to provide the
refrigeration required for Steps (12) and (13);
21. Substantially isentropically expanding the additionally cooled,
compressed refrigerant gas mixture in a work recovery engine to
obtain mechanical energy and to further cool the refrigerant gas,
whereby the first isentropically expanded, methane-nitrogen
refrigerant gas is formed;
22. Substantially isentropically expanding the heated first
isentropically expanded refrigerant gas from Step (12) in a work
recovery engine whereby (a) mechanical energy is obtained, (b) the
heated refrigerant gas is cooled, and (c) the second isentropically
expanded methane-nitrogen refrigerant gas is formed;
23. Removing from the process system subsequent to Step (13) a
minor portion of the combined gas mixture substantially equal in
magnitude to the minor methane-nitrogen gas fraction formed in Step
(14); and
24. Employing the mechanical energy obtained from the work recovery
engine of Steps (3), (21) and (22) at least to compress partly the
major portion of the combined refrigerant gas mixture in Step
(18).
10. The process of claim 9 wherein the nitrogen-methane and C.sub.2
-C.sub.4 hydrocarbon mixture in Step (13) is at a pressure in the
range from about 350 to about 750 psia.
11. The process of claim 9 wherein the nitrogen-methane and C.sub.2
-C.sub.4 hydrocarbon mixture in Step (13) is maintained at a
pressure relative to its analysis such that the plot of the
temperature vs. enthalpy of the mixture approaches as straight a
line as is reasonably possible.
12. The process of claim 9 wherein the flow rate of the combined
methane-nitrogen gas mixture in Step (13) is from about 1.5 to
about four times the flow rate of the nitrogen-methane and C.sub.2
-C.sub.4 hydrocarbon mixture in Step (13), expressed as moles per
hour.
13. The process of claim 12 wherein the flow rate of the combined
methane-nitrogen gas is maintained at such a rate that the
temperature vs. enthalpy curve for the combined gas as closely as
possible parallels the temperature vs. enthalpy curve of the
nitrogen-methane and C.sub.2 -C.sub.4 hydrocarbon gas mixture,
whereby more effective heat exchange between the combined gas and
the nitrogen-methane and C.sub.2 -C.sub.4 hydrocarbon gas mixture
is effected.
14. The process of claim 9 wherein the compressed, methane-nitrogen
refrigerant gas mixture from Step (18) is at a pressure from about
400 to about 600 psia.
15. The process of claim 9 wherein the compression of Step (18) is
effected in a plurality of stages and the gas is cooled
interstage.
16. A continuous process for the liquefaction of natural gas which
comprises the steps of:
Condensing the natural gas mixture under elevated pressure to form
(a) a liquid water phase, (b) a heavy hydrocarbon liquid phase
containing nitrogen and carbon dioxide and (c) a light hydrocarbon
vapor phase containing nitrogen and carbon dioxide;
removing the liquid water phase from the process;
adjusting the pressure of the heavy hydrocarbon liquid phase and
the light hydrocarbon vapor phase to a range of from about 350 to
about 750 psia;
fractionating the pressure adjusted heavy and light hydrocarbon
phases to form (a) a methane-nitrogen fraction substantially free
of carbon dioxide and hydrocarbons heavier than methane and (b) a
carbon dioxide-hydrocarbon fraction substantially free of nitrogen
and containing C.sub.2 and heavier hydrocarbons;
fractionating further the carbon dioxide-hydrocarbon fraction to
form a C.sub.2 -C.sub.4 hydrocarbon fraction and a C.sub.4 and
heavier hydrocarbon fraction and removing the C.sub.4 and heavier
hydrocarbon fraction from the process;
combining the C.sub.2 -C.sub.4 hydrocarbon fraction with the carbon
dioxide free methane-nitrogen fraction to form a substantially
purified natural gas mixture;
cooling and substantially completely liquifying the purified
natural gas mixture in a countercurrent heat
exchanger-liquifier;
flash evaporating a minor gas fraction from the cooled effluent
from the heat exchanger-liquifier by substantially isenthalpically
reducing the pressure thereof, said flash evaporation causing the
temperature of both the minor gas fraction and the remaining major
liquid fraction to decrease;
collecting the major liquid fraction as product and returning the
flashed fraction to a point where it is combined with an
isentropically expanded refrigeration gas fraction produced
subsequently in the process;
passing said combined gases countercurrently to the purified
natural gas mixture in said heat exchanger-liquifier at a flow rate
of two to four times that of the purified natural gas stream;
separating the warmed effluent from said heat exchanger-liquifier
into a major refrigeration gas fraction and a minor fuel gas
fraction;
passing the refrigeration gas fraction through a counter-current
heat exchanger precooler to a compressor system to be compressed to
a pressure in the range of about 400 to 600 psia;
passing the pressurized gas from said compressor system through
said heat exchanger precooler to cool the refrigeration gas;
isentropically expanding the cooled effluent from said heat
exchanger precooler with an expansion engine at a temperature and
pressure sufficient to maintain the gas in the gaseous state in
said expansion turbine to form the isentropically expanded
refrigeration gas, whereby mechanical energy is obtained and the
pressure of the refrigeration gas fraction is reduced, thereby
further cooling such fraction;
combining the further cooled, reduced pressure, isentropically
expanded refrigeration gas effluent from said expansion engine with
said flash gas to a point prior to entry into said heat
exchanger-liquifier; and
utilizing (a) the fuel gas fraction to furnish fuel for a gas
driven turbine which provides a portion of the mechanical energy
necessary to drive said compressor system and (b) the mechanical
energy obtained from the isentropic expansion to provide an
additional portion of the mechanical energy necessary to drive the
said compressor system.
17. In an integrated apparatus for purifying and liquifying natural
gas, the combination comprising:
gas-liquid separating means for separating condensed water and a
heavy hydrocarbon condensate from a pressurized natural gas mixture
comprising nitrogen, carbon dioxide, water and lower boiling
hydrocarbons;
means for adjusting the pressure of the heavy hydrocarbon
condensate and the water and heavy hydrocarbon condensate-free
remainder of the natural gas mixture to about 350 to about 750
psia;
means for fractionating the combined, pressure adjusted hydrocarbon
condensate and water-free natural gas mixture to form a
methane-nitrogen fraction substantially free of carbon dioxide and
a carbon dioxide-hydrocarbon fraction substantially free of
nitrogen and containing predominantly C.sub.2.sub.+
hydrocarbons;
means for fractionating the carbon dioxide-hydrocarbon fraction to
form a carbon dioxide fraction and a C.sub.2.sub.+ hydrocarbon
fraction, said carbon dioxide fraction including all of the methane
present in the carbon dioxide-hydrocarbon and a portion of the
C.sub.2 hydrocarbons contained therein;
means for combining the C.sub.2.sub.+ hydrocarbon fraction with the
methane-nitrogen fraction, thus forming a substantially water-free
and carbon dioxide-free natural gas mixture;
a heat exchanger-liquifier for cooling and liquifying the
water-free and carbon dioxide-free natural gas by countercurrent
heat exchange with a combined flash gas and work expanded
refrigerant gas from an expansion engine;
means for isenthalpically expanding and partially flashing the
natural gas effluent from said heat exchanger-liquifier;
gas-liquid separating means for separating said expanded effluent
into a major liquid fraction and a minor flash gas fraction;
means for recovering the major liquid fraction as product;
means for passing the flash gas to the heat
exchanger-liquifier;
compression means for compressing the combined flash gas and work
expanded refrigerant gas from the heat exchanger-liquifier;
a heat exchanger precooler for precooling the gas coming from the
compression means by countercurrent heat exchange with the combined
flash gas and work expanded gas from said heat
exchanger-liquifier;
an expansion engine for work expanding said precooled gas, said
engine further cooling the precooled gas and reducing the pressure
thereof while generating mechanical energy;
means for feeding the work expanded and cooled gas from said
expansion engine to said heat exchanger-liquifier to join said
flash gas; and
means for bleeding from the apparatus a minor portion of the
combined flash gas and work expanded gas prior to compression
thereof, said expansion engine and said compression means being
operatively cooperative whereby the mechanical energy obtained from
said engine supplies at least a portion of the work required to
operate said compression means.
Description
RELATED APPLICATIONS
This application is related to Ser. No. 36,277, filed May 11, 1970
and entitled Apparatus and Process for Liquefaction of Natural
Gases.
BACKGROUND OF THE INVENTION
This invention relates to the liquefaction of a gas; and, more
particularly, to a method and apparatus for the liquefaction of
natural gas. Even more particularly, this invention relates to a
LNG expander cycle process employing an integrated cryogenic
purification, and to apparatus therefor.
Natural gas is mostly composed of methane, but usually contains
small amounts of heavier hydrocarbons such as ethane, propane,
butane and the like as well as small amounts of aromatic
hydrocarbons. Natural gas also contains minor amounts of
non-hydrocarbons, such as water, nitrogen, carbon dioxide and the
like. Thus, the present invention is directed to a cryogenic
purification of raw natural gas to remove substantially all of the
water, carbon dioxide and heavy hydrocarbons therefrom, and to the
subsequent liquefaction of the purified natural gas to form a
liquid product stream comprised mainly of methane and substantially
free from carbon dioxide and heavy hydrocarbons.
Recently, there has been an increasing demand for a simple yet
economic process for reducing natural gas to a liquified state. One
of the reasons is that natural gas wells are being discovered in
remote parts of the world where transporting the gas to the point
of consumption by pipeline is difficult or impossible.
Transportation of natural gas in the gaseous state by marine
vessels would be uneconomical, unless the gaseous materials were
compressed, and this in turn would be impractical since the
containers would have to be exceptionally strong and extremely
large to hold the gas to be transported. The cost of such
containers and the industrial hazards that attend their use are so
great that the transportation of compressed gas is impractical.
Accordingly, an efficient process for the liquefaction of natural
gas is of great importance, particularly, where the supply thereof
is in a remote area and there is a demand at a distant market
place. Such a process is particularly important where the gas must
be transported by marine vessels since, by liquefaction of natural
gas, the volume thereof can be reduced to nearly one six-hundredth
its volume and the containers need not be of the thickness,
strength, and capacity necessary for the shipment of compressed
gas.
DESCRIPTION OF THE PRIOR ART
Many attempts have been made to find an economical process for
liquifying natural gas, but numerous of the prior art processes
have major disadvantages or limitations. For example, the power
consumption of certain prior art liquifier systems is prohibitively
high, while in others the need for the use of an inordinately large
number of expensive heat exchangers and compressors exists. In
addition, some liquifying systems which are acceptable in more
industralized areas are often less suited for use in some of the
more underdeveloped and remote areas of the world. For example, a
system that requires molecular sieves and/or specific adsorbents
for the initial purification of raw natural gas feedstock is less
economically operated in areas where such sieves and adsorbents are
not readily available and/or where spent sieves and adsorbants have
essentially no salvage value. Still other prior art system require
complicated high pressure equipment which is difficult to maintain
and to control automatically. And yet others require the use of
expensive refrigerants which must be shipped to the liquifying
plant.
At present, there are three known basic cycles for liquefaction of
natural gases. These are generally referred to as the "Cascade
Cycle", "Multicomponent Refrigerant Cycle" and "Expander Cycle".
Many minor variations can be effected in the design of each type of
cycle to adapt it to the specific process requirements.
Briefly, the "Cascade Cycle" consists of a series of heat exchanges
with the feed gas, each exchange being at successively lower
temperature until the desired liquefaction is accomplished. The
levels of refrigeration are obtained with different refrigerants or
with the same refrigerant at different evaporating pressures.
Frequently, a combination of both approaches is used. The high
efficiency of the "Cascade Cycle" is offset by rather high
investment cost in the extensive heat exchange and compression
systems. Piping costs are high and considerable area is necessary
for the large amount of equipment. This type of plant is more
easily justified when fuel costs are high such as at the delivery
end of a pipeline system. This system would be impractical,
however, where there is limited space or poor soil conditions as
exists in many of the remotely located gas wells discovered in the
world.
A "Multicomponent Refrigerant Cycle" has been operated on a pilot
scale but no large commercial operation has been achieved to date,
although the basic design is known to those skilled in the art.
This process was designed in an effort to eliminate some of the
complexity of the "Cascade Cycle" and still retain the low power
requirement. The process uses a carefully controlled analysis of
the hydrocarbon refrigerant stream so that at successive
temperature levels of cooling specific liquid fractions are
obtained. Each of these liquid fractions, condensed at high
pressure, is then evaporated to obtain the required refrigeration.
The composition changes of the successively condensed fractions
when evaporated at a constant pressure furnish distinct temperature
levels of heat exchange. In this respect, the system operates as an
autocascade cycle.
The compression system of the foregoing plant is simple since only
one refrigerant is used. However, the heat exchange system and its
controls are extensive and costly. The nature of the equipment
requires series flow of all refrigerant gas over the cascade type
of exchangers. This can be accomplished by vertically stacking the
exchangers to a height of over 200 feet. Unless good soil bearing
is available, the support of such a unit is difficult and the
erection of such unit on the deck of a floating vessel presents
serious problems of stability. To segmentize the heat exchangers
would require very large vapor lines in the range of 5 to 6 feet in
diameter. Therefore, this type of cycle would be impractical in a
liquefaction plant near off-shore wells, where the soil bearing is
poor or where the wells are in a nonindustralized part of the
world. A large Multicomponent Refrigerant cycle also requires a
complete hydrocarbon fractionating unit to prepare the pure
components required to maintain the carefully controlled
refrigerant analysis.
The "Expander Cycle" is similar to that used on most of the large
air separation plants today and it does have the advantage of
simplicity over a "Cascade Cycle". In this cycle, gas is compressed
to a selected pressure, cooled, then allowed to expand through an
expansion turbine, thereby performing work and reducing the
temperature of the gas. It is quite possible to liquify a portion
of the gas in such an expansion. The low temperature gas is then
heat exchanged to effect liquefaction of the feed. The power
obtained from the expansion is usually used to supply part of the
main compression power utilized in the refrigeration cycle. If the
expander cycle is a closed cycle, any suitable refrigerant gas can
be used. If it is an open cycle liquid natural gas plant, the
refrigerant would have to be methane or a methane-nitrogen mixture
as this would be flashed from the gas-liquid separator in the
process.
An expander cycle plant is compact, has minimum items of equipment,
simple control and utilizes all standard machinery and heat
exchangers. This type of plant has an important added advantage of
mechanical simplicity that is particularly significant when
considering operations in remote areas of the world.
An efficient "Expander Cycle" method for liquefaction of low
boiling gases such as oxygen and nitrogen is presently known. The
heat exchanger cycle of this process is operated under 400 to 1000
psia pressure and cooling is made more efficient by causing
components of the warming stream to undergo a plurality of work
expansion steps with intervening reheating. In this process, two or
more heat exchangers are employed in series with an intermittent
refrigeration of the incoming gas. A portion of the feed stream
which had been previously cooled by a warm-leg heat exchanger and a
refrigeration unit is work expanded and thereafter used to absorb
heat from the remaining portion of the feed stream in a
countercurrent heat exchanger. The warmed effluent gas from the
heat exchanger is work-expanded a second time and this cooled and
expanded gas is combined with the flash gas to be used to absorb
heat from the incoming feed stream in the second heat exchanger.
The cold effluent from the second heat exchanger is isenthalpically
expanded and passed to a gas-liquid separator to remove the liquid
for storage and the flash gas is combined with the work-expanded
gas as discussed above. The combined warmed effluent gas from the
heat exchangers is recycled to the feed stream to be recompressed
and undergoes the foregoing liquefaction.
This type of process suffers the drawback of expensive refrigerants
and separate compression and expansion systems driven by an outside
source of power to maintain its operation. Furthermore, such a
process is not practicable in the liquefaction of natural gas since
there is no provision for handling the heavy gases which freeze at
the temperatures encountered in the heat exchangers. In addition,
if the flash gas was recycled back to the feed stream, the lower
boiling ends, i.e., nitrogen, etc. of the feed mixture would
increase in the heat exchanger liquifier causing an imbalance in
the system requiring additional energy to liquify and cause
thermodynamic inefficiency. This latter problem is not apparent
when there is only one pure material to be liquified such as
nitrogen and oxygen, but when dealing with the liquefaction of
natural gases which contain a plurality of gases having boiling
points lower than methane, the problem is paramount.
In the above-mentioned related application, Ser. No. 36,277, filed
May 11, 1970, it is pointed out that the "Expander Cycle" process
can be successfully utilized in liquifying natural gas by
separating a combined residual flash gas and refrigerant gas
effluent from the heat exchanger-liquifier into a bleed portion and
a refrigerant gas portion, in stead of recycling this combined
mixture to the feed stream as described in the heretofore mentioned
prior art. While this latter process offers significant advantages
over the previously described prior art processes, it requires the
use of molecular sieves and adsorbents to purify the raw feed gas
before subjecting it to liquefaction. Accordingly, the use of this
process in a remote part of the world would require periodic
shipments of fresh sieves and adsorbents to the plant site, as well
as shipments of the spent sieves and adsorbents to an industralized
area for salvage.
In light of the inherent problems associated with the production of
liquified natural gas, the foregoing prior art processes do not
offer a sufficiently economic and practicable system for liquifying
natural gas in remote non-industrial parts of the world where the
soil bearing is poor and where expended materials have little or no
salvage value.
SUMMARY OF THE INVENTION
It has now been discovered that in an "Expansion Cycle" process, a
greater proportion of the natural gas feed is liquified efficiently
and a substantially self-balancing process is obtained when a raw
natural gas feedstock is purified by an integrated cryogenic
fractionation prior to liquefaction and when the liquefaction
comprises separating the combined residual flash gas and
refrigerant gas used to liquify and subcool the natural gas into a
bleed gas portion and a recycle refrigerant gas portion instead of
recycling the entire of the flash and refrigerant gases.
Since the liquefaction apparatus and process of the present
invention is compact, efficient and self-sustaining, it is capable
of being prefabricated on a mobile platform which may be of a
marine type that can be shipped to distant and remote parts of the
world. The mobile marine means contemplated by the present
invention can be of the type disclosed in the U.S. Pat. No.
3,161,492 to Keith et al., which disclosure is incorporated herein
by reference. Additionally, the mobile marine platform can be of
the floating type wherein the platform is not supported by
mechanical means, such as caissons, engaging the marine floor but
rather is supported by the buoyancy of the platform itself. Such a
floating platform can take the form of a barge or conventional,
self-powered ocean-going vessel such as a tanker. Such an apparatus
and process prefabricated on a marine means is so designed as to
permit direct loading of the liquid into a transport means such as
a ship or tanker. Furthermore, by employing the apparatus of the
present invention it is possible to bring the prefabricated plant
to off-shore gas wells while it would be impossible using the
cumbersome processes of the prior art.
More specifically, the invention contemplates the preassembly of a
simplified expansion type of liquefaction plant on a marine
platform which is only a portion of the size of the conventional
cascade liquefaction plant disclosed in the above-cited patent, and
which avoids the need for purification systems which employ
molecular sieves and adsorbents that must be transported to a plant
site and periodically regenerated, replaced and salvaged.
Accordingly, the present invention provides, in its preferred
aspects, a process for producing a large supply of liquified
natural gas in areas remotely located in the world where the source
of power is limited or non-existent and where the use of molecular
sieves and adsorbents is uneconomic.
An object of the present invention is to provide a process which
will produce liquified natural gas employing moderate pressures
(200-1000 psia) and still maintain thermodynamic efficiency.
Accordingly, this efficiency is accomplished by controlling the
condensing gas pressure with relation to its analysis so that the
temperature-enthalpy curve is essentially a straight line. This
enhances the possibility of refrigeration by a cold gaseous stream
eliminating many of the cryogenic liquid handling problems of other
processes and which would be extremely difficult and hazardous in
floating platform construction as described above.
Another object of the present invention is to accomplish
liquefaction by employing a countercurrent heat exchanger-liquifier
or subcooler whereby the cooling or refrigerant stream is
maintained in the form of a gas instead of a boiling liquid
refrigerant stream. This object can be accomplished since it has
been unexpectedly found that the nitrogen content of the feed gas
stabilizes in the refrigerant system and the actual refrigerant
becomes an equilibrium mixture of methane and nitrogen gases. This
analysis change assists greatly in providing the required
temperature driving force to promote the liquefaction of the main
stream of natural gas which because of its heavier constituents
condenses at a higher temperature.
Still another object is to employ as the refrigerant material a
combination of the flash gas obtained in the pressure letdown from
the heat exchanger-liquifier and the cold expansion turbine
effluent gas thereby causing the lower boiling point gases in the
feed gas stream to reach an equilibrium in the expander refrigerant
cycle. The refrigerant system is thus allowed to operate
effectively at lower temperatures and pressures than processes
heretofore employed. By being able to operate at lower temperatures
and pressures, the equipment attending to the process of the
present invention is far less expensive than processes heretofore
employed.
It is another object of the present invention to form a bleed
stream of the warmed effluent refrigerant gas from the heat
exchanger-liquifier equivalent to the flow rate of the flash gas
coming from the LNG gas-liquid separator. This bleed stream can be
employed as fuel for a gas turbine system which in turn furnishes
part of the power required to drive the compressors in the
refrigeration cycle.
An additional object of the present invention is to control the
temperature and pressure of the refrigerant gas so as continually
to maintain a gaseous state in the heat exchangers and expansion
turbines. It has been unexpectedly found that if the
temperature-enthalpy curve of the feed gas is sufficiently straight
as controlled by its pressure it is not necessary to cool the
refrigerant gas in the expansion engine to such an extent that some
of the gas liquifies in order to liquify the incoming gas.
Therefore, the present invention has a decided advantage over the
prior art processes which partially liquify the gas in the
expansion turbines, thus eliminating the problem of erosion in the
vanes of the turbine, which, in turn, becomes costly in the overall
operating expense of the liquifaction plant.
In accordance with the present invention there is provided an
integrated process for separating substantially all of the carbon
dioxide and water and most of the heavy hydrocarbons from raw
natural gas feedstock, and then liquifying the resulting purified
natural gas mixture to form a product comprising nitrogen and
hydrocarbons containing up to about 6 carbon atoms. The raw natural
gas mixture is purified by first producing the crude mixture, under
pressure, into an inlet separator vessel wherein condensed water
and a hydrocarbon condensate are separated from the gas mixture as
two substantially immiscible liquid layers. The water layer is
drained from the bottom of the separator and discarded. The
condensed hydrocarbon layer and the gaseous portion of the raw feed
are transferred from the separator to a cryogenic LNG fractionator
distilling all of the nitrogen, substantially all of the methane
and essentially none of the carbon dioxide contained in the raw
feed as overhead. The bottoms from the LNG fractionator, which
contain a small amount of methane, as well as all of the
C.sub.2.sub.+ hydrocarbons and essentially all of the carbon
dioxide contained in the raw feed, is further fractionated to
separate (a) the carbon dioxide, methane and part of the ethane at
one end of the spectrum and (b) the C.sub.5.sub.+ hydrocarbons at
the other end of the spectrum from (c) the remainder of the
hydrocarbons (predominantly ethane with but minor amounts of
C.sub.3 and C.sub.4 hydrocarbons); the latter being combined with
the portion of the LNG fractionator overhead not used for reflux.
This combined stream is then liquified to form the LNG product. In
a preferred embodiment, after removal of the carbon dioxide
fraction, the remaining C.sub.2.sub.+ hydrocarbons are fractionated
in an LPG fractionator to separate a C.sub.2 -C.sub.4 fraction from
a C.sub.4.sub.+ gas liquid fraction, whereupon only the C.sub.2
-C.sub.4 fraction is combined with the overhead from the LNG
fractionator, thereby forming a substantially purified natural gas
mixture that is essentially free from carbon dioxide and completely
devoid of C.sub.5.sub.+ hydrocarbons. In any event, the carbon
dioxide fraction split from the LNG fractionator bottoms may be
used as fuel gas to supply part of the power requirements of the
process equipment.
In the event of the temperature and pressure conditions of the raw
gas mixture leaving the inlet separator are such that hydrates may
form during subsequent processing, a quantity of an inert
dessicant, such as a solution of triethylene glycol, may be
injected into the gas mixture to reduce its dew point. This may be
conveniently carried out, for example, by injecting the triethylene
glycol solution into the gas mixture as it leaves the inlet
separator, whereafter the water-enriched glycol solution is removed
from the LNG fractionator bottoms by passing the bottoms to a heavy
hydrocarbon surge drum wherein the glycol solution and hydrocarbons
separate, the latter then being decanted and transferred to the
carbon dioxide fractionator.
The substantially purified natural gas comprising the LNG
fractionator overhead and either the carbon dioxide free fraction
from the carbon dioxide fractionator or the C.sub.2 -C.sub.4
fraction from the LPG fractionator is then cooled under elevated
pressure by heat exchange with a combined, methanenitrogen
refrigerant gas mixture obtained elsewhere in the process so as to
effect liquefaction of a predominant portion of the purified
natural gas mixture while heating the combined refrigerant gas
mixture. The pressure of the cooled natural gas is then
substantially isenthalpically reduced, thereby further cooling it,
forming (by flash evaporation) a minor gas fraction consisting
essentially of methane and nitrogen and liquifying completely the
balance of the natural gas mixture to form a major liquid fraction
containing methane and substantially all of the C.sub.2.sub.+
hydrocarbons present in the purified natural gas feed. The minor,
methane-nitrogen gas fraction is then separated from the major
liquid fraction and the major liquid fraction is recovered as
product.
The separated minor, methane-nitrogen gas fraction is combined with
an isentropically expanded methane-nitrogen refrigerant gas mixture
circulated in the process. This gas mixture can be produced
elsewhere in the process. This combination of methane-nitrogen gas
constitutes the combined refrigerant gas mixture mentioned
previously and, as also mentioned previously, this combined
refrigerant gas mixture is employed to cool the purified natural
gas mixture in the liquefaction heat exchange step. A major portion
of the heated combined refrigerant gas mixture leaving the
liquefaction heat exchange step is compressed to provide a
compressed, methane-nitrogen refrigerant gas mixture. This gas
mixture, after suitable precooling, is substantially isentropically
expanded in one or more work recovery engines whereby mechanical
energy is obtained, the gas mixture is further cooled, and the
isentropically expanded, methane-nitrogen refrigerant gas mixture
mentioned previously is obtained. The compressed methane-nitrogen
gas mixture, isentropically expanded methane-nitrogen gas mixture
and the corresponding portion of the combined gas mixture
constitute a circulating refrigerant system.
At least a major portion of the heated, combined refrigerant gas
mixture leaving the liquefaction heat-exchange step and prior to
compression is heat exchanged with the compressed, precooled
methane-nitrogen refrigerant gas mixture whereby such combined gas
mixture is further heated to a temperature suitable for entering
the compressors. Concurrently, the compressed, methane-nitrogen
refrigerant gas mixture is cooled to a temperature level such that
upon work expansion as previously described appropriate temperature
levels and refrigeration capacity are developed for liquefaction of
the natural gas mixture.
A minor portion of the combined refrigerant gas mixture after
leaving the liquefaction heat exchange step is removed from the
circulating refrigerant system and is employed as fuel gas to
supply a portion of the power needed to drive the above-mentioned
compressors. Although this minor portion can be removed at any
point in the circulating refrigerant system, it is preferred to
remove it prior to compression of the combined refrigerant gas
mixture. Advantageously, this minor portion is removed subsequent
to the heat exchange of the combined refrigerant gas mixture with
the compressed, precooled methane-nitrogen refrigerant gas mixture
and prior to compression of the combined refrigerant gas mixture.
The mechanical energy obtained from the work recovery engines is
also employed to provide at least part of the power needed to
compress the major portion of the combined gas mixture. The
quantity of the minor portion removed from the refrigerant system
is substantially equal in magnitude to the minor methane-nitrogen
gas fraction formed by the isenthalpic pressure reduction (flash
evaporation) previously described.
The raw natural gas feedstock suitable for charging to the initial
separating vessel of this invention may comprise natural gas
obtained from a crude oil well (associated gas) or from a gas well
(non-associated gas). However, various natural gas streams may not
be suitable for immediate employment in the present invention due
to the source, both ultimate and immediate, from which they are
obtained. Thus, for example, both associated and non-associated gas
obtained directly from the well head may be at an undesirably high
temperature, e.g. substantially in excess of 100.degree.F. Further,
associated gas or gas obtained from certain storage facilities may
be of inadequate pressure for employment in the present process,
e.g. only about 100 or 200 psia. Thus, certain natural gas streams,
depending upon their source, may require certain pretreatment or
preparation prior to employment in the present invention.
While there is no theoretical maximum to the pressure which can be
employed in the various fractionators and heat exchangers of the
present invention, it would appear that there is little practical
advantage to be gained by employing pressures in excess of 1000
psia. This is particularly true in connection with the heat
exchanger or subcooler employed in the liquefaction step unless the
purified natural gas entering the subcooler contains a very high
methane content (95%+). Preferably a maximum pressure no greater
than about 750 psia is employed, thereby avoiding the necessity of
very high pressure vessels. Advantageously, a maximum pressure of
about 650 psia is maintained. On the other hand, it is found that a
minimum pressure of about 350 psia is required for this heat
exchange liquefaction and preferably a pressure of about 600 psia
is maintained. The pressure selected is such that a plot of the
temperature vs. enthalpy of the gas to be liquified approaches a
straight line, i.e. not have a large enthalpy increase over a
narrow temperature range. Thus the pressure selected will depend on
the gas composition. Accordingly, therefore, a low pressure natural
gas stream, such as an associated gas, can be compressed to achieve
the desired pressure range either prior to introduction into the
initial separator vessel or prior to introduction into the LNG
fractionator, whereas an extremely high pressure natural gas
stream, such as a nonassociated gas, can be reduced in pressure
either prior to introduction to the initial separator vessel or,
preferably, after leaving the separator vessel but prior to
introduction into the LNG fractionator. This reduction of pressure
can be accomplished by throttling, employment of a turbo-expander
or any other expansion or pressure reducing means whether capable
of recovering mechanical energy or not.
In the event that the initial temperature of the natural gas is too
high it must be cooled, such as, for example, by heat exchange or
other techniques well-known in the art. If the natural gas is at an
undesirably high temperature and an undesirably high pressure it is
possible to reduce both the temperature and pressure of the gas
simultaneously by the simple technique of adiabatic expansion. By
use of these well-known techniques it is possible to limit the
temperature of the natural gas charged to the LNG fractionator, for
example, to a maximum of about 150.degree.F., preferably, however,
the maximum temperature will not exceed about 125.degree.F. It is
particularly advantageous to employ temperatures below about
100.degree.F. Selection of the exact processing temperature can
depend on the processing sequence selected for gas purification as
discussed below.
It will be understood, of course, that the selection of the
particular operating pressure as well as the temperature employed
in the various fractionators and in the liquefaction heat exchanger
will be determined to a great extent by the particular composition
of the natural gas being treated as well as by the temperature of
the natural gas being charged. In any event, however, it is
necessary that the temperature and pressure conditions in the
liquefaction heat exchanger be maintained such that at least a
predominant portion of the natural gas effluent therefrom is in the
liquid state. In this latter connection, the temperature and
pressure conditions in the liquefaction heat exchanger are
generally such that about 85 percent or more of such natural gas is
in the liquid state. Preferably at least about 90 percent, more
preferably at least about 95 percent, is liquified. Usually the
temperature of the cooled natural gas from the liquefaction heat
exchange will be about -175.degree. to about -225.degree.F. when
employing pressures within the ranges indicated above.
After liquefaction heat exchange, the pressure of the natural gas
isenthalpically reduced to a pressure just slightly above
atmospheric pressure, generally falling in the range of from about
16 to about 35 psia, and preferably in the range of from about 20
to about 30 psia. This isenthalphic reduction in pressure results
in the flash evaporation of the minor gas fraction, liquefaction of
the balance of the natural gas and the overall reduction in
temperature of both the minor gas fraction and the remaining major
liquid fraction. The minor gas fraction generally comprises up to
about 15 mol percent of the total liquified natural gas mixture and
preferably comprises from about 5 to about 10 mol percent of the
liquified natural gas mixture. Again the exact extent of the
reduction in temperature effected by the isenthalpic pressure
reduction is dependent primarily upon the magnitude of reduction in
pressure effected and upon the composition of the natural gas
mixture. Accordingly, the magnitude of the reduction in pressure
must be sufficient to reduce the temperature to a level sufficient
to ensure liquefaction of a major portion of the particular
composition of natural gas at the lower pressure. Thus, by way of
illustration, the isenthalpic reduction in pressure of some
substantially completely liquified natural gas mixtures from the
range of about 600 psia down to about 20 psia will result in a
further reduction in temperature of the natural gas in the range of
about 40.degree. to about 70.degree.F. and the achievement of a
temperature sufficient to liquify completely the main portion of
the stream at the lower pressure, e.g., -240.degree. to
-260.degree.F.
The minor gas fraction which has been flashed comprises about 95
mole percent (dry) methane and about 5 mole percent (dry) nitrogen,
and contains a substantial portion of the nitrogen initially
contained in the raw feedstock. The liquid fraction remaining after
pressure let down comprises about 90 to 95 percent of the methane
and essentially all of the C.sub.2.sub.+ hydrocarbons fed to the
liquifier heat exchanger. The mole percent (dry) of nitrogen
contained in the remaining liquid fraction is substantially lower
than the mole percent of nitrogen in the liquified natural gas
prior to pressure let down.
After combining the above minor gas fraction with the
isentropically expanded methane-nitrogen refrigerant gas mixture,
such combined gas mixture will be at a temperature lower than the
temperature of the liquified natural gas from the liquefaction heat
exchange, but will be completely in the gaseous state.
As will be understood, the composition of the minor
methane-nitrogen gas fraction determines the composition of the
combined gas mixture, and while the quantity of the minor gas
fraction is grossly smaller than the quantity of the purified
natural gas mixture to be liquified, such minor gas fraction is
allowed to build up and is recirculated within the refrigerant
system such that during normal operation the flow rate of the
combined gas mixture is from about 1.5 to about 4 times the flow
rate of the purified natural gas mixture, expressed in moles per
hour (dry). Preferably, the flow rate of the combined gas mixture
is in the range from about 2 to about 3 times the flow rate of the
purified natural gas mixture. The amount of refrigerant and
combined gas employed is determined by the relative plots of
temperature vs. enthalpy for the combined gas and the natural gas
being liquified so that these curves do not cross, thereby ensuring
proper heat exchange in the gas liquifier.
After having achieved the desired flow rate of the combined gas
mixture, it is essential to normal operation of the process for a
quantity of such combined gas mixture to be bled from the system
after heat exchange with the purified natural gas mixture. The
quantity of the combined gas stream removed from the system is
substantially equal in magnitude to the minor methane-nitrogen gas
fraction obtained by flashing. This removal is necessary to
maintain the system in balance and to prevent unwanted build-up
within the system once normal operations have been achieved. This
removal of a minor portion of the heated combined gas mixture can
be effected immediately subsequent to removal of the combined gas
mixture from the liquefaction heat exchanger. Alternatively, such
bleed can be effected subsequent to heat exchanging the combined
gas mixture with the compressed methane-nitrogen refrigerant gas
mixture but prior to such compression. Removal of this later point
permits heat exchanging a somewhat greater quantity of uncompressed
gas against a somewhat smaller quantity of compressed gas and can
favor heat balances in certain design conditions. Again this minor
gas fraction can be removed after any portion of the compression
step if for any reason higher pressure gas is desired. Although
this bleeding can be effected at substantially any place subsequent
to removing the combined gas mixture from the liquefaction heat
exchange and prior to combining the isentropically expanded
methane-nitrogen refrigerant gas mixture with the minor gas
fraction obtained by flash vaporization, it is considered to be
advantageous to effect such bleed after heat exchange, but prior to
the compression so as to avoid needlessly compressing a quantity of
gas which is to be removed from the refrigerant system in any
event.
The compression of the combined methane-nitrogen refrigerant gas
mixture can be effected in either single or multistage compression.
Advantageously, multi-stage compression is employed utilizing, for
example, at least two stages and preferably at least three stages.
The multi-stage compression can be used together with interstage
cooling. Such cooling can be effected via heat exchange so as to
recover heat values for use elsewhere or such cooling can be
effected employing air or water as a cooling means simply for
dissipating the heat from the system. Additionally, the compressed
methane-nitrogen refrigerant gas mixture can be subjected to an
after cooling prior to heat exchange with the combined gas mixture.
Generally, the heat exchange between the combined gas mixture and
the compressed methane-nitrogen refrigerant gas mixture as well as
the interstage cooling is conducted so as to provide a gas
temperature at compressor inlet, either single or multi-stage,
above the cryogenic range, i.e. above about -50.degree.F.
Preferably, this compressor inlet temperature is maintained above
about -20.degree.F. and advantageously above about 60.degree.F. On
the other hand, however, the inlet temperature of the gas to be
compressed must not be excessively high or it will place a needless
burden upon the compression equipment. Accordingly, therefore, the
maximum temperature employed is no greater than about 300.degree.F.
and preferably is no greater than the temperature produced in a gas
line in tropical climates, i.e. about 120.degree.F. Advantageously,
however, the maximum temperature is no greater than about
90.degree. to 100.degree.F. Again, when employing multi-stage
compression, the maximum temperature to each compression stage can
be controlled by interstage cooling in the manner mentioned
above.
The combined gas mixture is heat exchanged against the compressed
methane-nitrogen refrigerant gas mixture so as to reduce the
temperature of the compressed methane-nitrogen refrigerant gas
mixture. This heat exchange is designed to control the temperature
of the compressed refrigerant gas to the work recovery engines so
as to assist in control of the final combined gas temperature to
the natural gas liquefaction. This latter heat exchange is also
designed to increase the temperature of the combined gas mixture
prior to compression since, from an equivalent viewpoint, the
temperature of the combined gas from the liquefaction heat exchange
is at a temperature which is undesirably low, e.g. -100.degree.F.
or lower, to be handled by typical commercial compressors.
Additionally, the increase in the temperature of the gas due to
compression must be be offset, and if the temperature of the gas
exiting from a compressor is too low, e.g. less than 32.degree.F.,
fresh water coolers are inoperable and air coolers are ineffective
in many, if not most, climates to accomplish the desired
temperature reduction.
The isentropic expansion of the compressed methane-nitrogen
refrigerant gas mixture is conducted so as to cool the
methane-nitrogen gas mixture to as low a temperature as can be
achieved but to a temperature about the liquefaction point of such
a mixture at the outlet pressure of the respective work recovery
engines, the outlet pressure of the final work recovery engine
being maintained slightly above atmospheric pressure. Generally,
this pressure is in the range from about 16 to about 35 psia and
preferably in the range from about 20 to about 30 psia. In any
event, the outlet pressure of the work recovery engine is
maintained at the same level as the flash evaporated minor gas
fraction. The expansion temperature is controlled by the inlet
temperature as previously mentioned.
In the process of this invention the operating conditions of
temperature and pressure are maintained such that at all times the
combined gas mixture, the compressed methane-nitrogen refrigerant
gas mixture and the isentropically expanded methane-nitrogen
refrigerant gas mixture are substantially completely in the gaseous
state. To phrase it in another manner, there is little, if any,
liquid in the refrigeration gas cycle, of this process. However,
liquid can be generated if desired but when designed for liquid
type of operation special expansion engine designs are necessary to
prevent erosion and special liquid distributors are required in the
heat exchange system.
The mechanical energy obtained from the respective work recovery
engines is employed to supply a portion of the horse-power required
by the compression operation. It will be understood, that since the
same quantity of material is being compressed as is being expanded,
and since neither the compressors nor work recovery engines, e.g.,
the turboexpanders, are 100 percent efficient, the mechanical
energy available from the work recovery engines will provide only a
portion of the horsepower required for compression. Accordingly,
therefore, the remaining portion of the mechanical energy required
for compression is supplied by utilizing the fuel gas streams bled
at various stages from the system and by employing one or more
external sources of energy.
In order to describe this invention in greater detail, reference is
made to the following specific example which will be described in
connection with the accompanying two sheets of drawings designated
FIGS. 1a and 1b and together illustrating an integral flow scheme.
All compositions and percentage in the example are given on a dry
basis, unless specified otherwise.
Referring to FIG. 1a, 30,408 Mols per hour of raw natural
non-associated gas is delivered from a suitable source (not shown)
through conduit 10 to an inlet separator 12 at a pressure of about
2350 psia and a temperature of about 85.degree.F. The composition
in mol percent of the feed entering the inlet separator 12 is
essentially as follows: 1.00% nitrogen, 2.00% carbon dioxide,
83.22% methane, 6.18% ethane, 2.76% propane, 1.50% butanes, and
3.34% C.sub.5 and heavier hydrocarbons. The raw feed also includes
a total of about 3430 pounds per hour of water. In the separator 12
about 3260 pounds per hour of condensed water and about 2598 Mols
per hour of a hydrocarbon condensate are separated from the
remainder of the water saturated feed gas. The condensed water is
removed as bottoms via conduit 14.
The hydrocarbon condensate is transferred through conduit 16 to a
demethanizer or LNG fractionator 18 at a rate of about 2598 Mols
per hour. The composition in mol percent of the hydrocarbon
condensate is essentially as follows: 0.35% nitrogen, 1.50% carbon
dioxide, 49.37% methane, 7.75% ethane, 5.17% propane, 4.12%
butanes, and 31.66% C.sub.5 and heavier hydrocarbons. The
condensate also comprises about 49 pounds of dissolved water.
A restriction orifice 20 is placed in the conduit 16 to control the
flow rate of the condensate and to prevent rapid build up of
pressure in the fractionator 18. The normal operating pressure of
the fractionator 18 is about 600 psia. The gas stream leaving the
separator 12 through conduit 22 contains about 121 pounds per hour
or about 230 ppm of water. The hydrate formation temperature of
this gas is about 73.degree.F. Therefore, to prevent hydrate
formation in either conduit 22 or in subsequent processing, about
3772 pounds per hour of a 62 percent triethylene glycol solution is
injected into conduit 22 near the outlet of the separator 12 by
conventional means (not shown). The glycol is diluted to about 60
percent in picking up water from the gas in conduit 22, whereby the
dew point of the gas is reduced to below -50.degree.F. The freezing
point of the glycol solution is well below the processing
temperatures.
A minor portion of the high pressure gas in conduit 22, i.e., about
465 Mols per hour (dry) is bled off through conduit 24 and is
employed as fuel gas in a later stage of the process. The remaining
major portion of the high pressure gas from conduit 22, i.e., about
27345 Mols per hour, is passed by means of conduit 26 (as shown in
FIG. 1a) to a first stage expansion turbine or turboexpander 28
(shown in FIG. 1b), wherein the pressure of the gas is reduced from
about 2350 psia to about 610 psia with a corresponding reduction in
temperature from about 85.degree.F. to about -42.degree.F. The
turboexpander 28 is designed to operate at about 80 per cent
adiabatic efficiency such that a considerable amount of power is
developed as the high pressure gas expands. As discussed below,
this power is utilized as a partial drive for the third stage
refrigeration compressor 84.
In passing through the turboexpander 28, about 8 mole percent of
the gas in conduit 26 condenses. Accordingly, a mixed liquid and
gas stream is removed from turboexpander 28 (shown in FIG. 1b) by
means of conduit 30 and is passed via conduit 30 to LNG
fractionator 18 (shown in FIG. 1a) at a temperature of about
-42.degree.F. and a pressure of about 610 psia. A small quantity of
triethylene glycol solution enters the fractionator 18 with this
mixed stream.
The LNG fractionator 18 is designed to separate between methane and
carbon dioxide with an overhead temperature of about -120.degree.F.
and a bottoms temperature of about 120.degree.F. Typically, the
fractionator 18 requires about 19 actual trays at 80 percent
efficiency with a 0.55/1 L/D ratio. The hydrocarbon condensate in
conduit 16 is generally fed into the fractionator 18 at about the
eighth tray, while the mixed liquid and gas stream in conduit 30 is
fed into the fractionator at about the fourteenth tray.
The overhead of the fractionator 18 passes through conduit 32 where
it is separated into two streams 34 and 36 for separate
condensation. The stream 34 is condensed in a conventional flat
plate cryogenic heat exchanger 38 designed to condense a sufficient
portion of the overhead vapors for gravity run back to the
fractionator 18. The reflux is returned at its bubble point.
The reflux exchanger 38 is cooled by part of the cold refrigerant
gas as from a second stage expansion turbine 94 of the
refrigeration cycle described hereinbelow. Flow of the cooling
refrigerant to the exchanger 38 is controlled by the top
temperature of the fractionator 18.
The main portion of the fractionator overhead not needed for reflux
flows through conduit 36 to the LNG condenser 40, also a flat plate
cryogenic exchanger. Prior to entering the exchanger 40, the
overhead portion in conduit 36 is combined with an ethane, propane
and butane fraction in line 124 and recovered from an LPG
fractionator 114 (FIG. 1b) in a later section of the process. The
combined stream, which becomes the LNG product stream, then enters
the heat exchanger 40 through conduit 42 at a temperature of about
-111.degree.F. At this point, the combined stream is a two phase
system: about 25 percent being liquid. After passing through the
exchanger 40, the temperature of the combined LNG stream is reduced
to about -116.degree.F. The pressure of the LNG stream at this
point is about 600 psia. The combined LNG stream then passes
through the heat exchanger-subcooler 44 wherein it is reduced in
temperature to about -212.degree.F. by a combined refrigerant gas
formed in a later stage of the process and described below.
Typically, the subcooler 44 comprises a standard plate type
cryogenic heat exchanger. The effluent from the subcooler 44 is let
down through conduit 46 to an expansion valve 48 whereby the
pressure is isenthalpically reduced from about 595 psia to about 25
psia. In passing through valve 48 about 7 mole percent of the
subcooled LNG stream flashes and the temperature of the mixture
drops to about -246.degree.F. The pressure of about 25 psia is
selected since it is just sufficient to drive the flash gas through
the heat exchangers and to the regeneration gas compression system
as will be described in more detail hereinafter.
The cold and expanded mixture from the expansion valve 48 is
conducted through conduit 50 to the LNG gas-liquid separating drum
52 which is normally operated about half full of liquid.
The cold LNG liquid stream is passed from the separating drum 52 to
a storage vessel (not shown) through conduit 54 at a flow rate of
about 25014 Mols per hour and at a pressure of about 25 psia. An
analysis of the recovered liquid reveals the following composition
in mol percent: 0.70% nitrogen, 0.00% carbon dioxide, 90.61%
methane, 4.80% ethane, 3.20% propane, 0.69% butanes and 0.00%
C.sub.5.sub.+ hydrocarbons.
The residual cold flash gas leaves the LNG separating drum 52
through conduit 56 at a flow rate of about 2196 Mols per hour and
at a temperature of about -246.degree.F. The pressure of the LNG
flash gas is about 25 psia. The residual flash gas contains about
5.10% nitrogen and 94.90% methane, recorded as mol percent.
The residual flash gas is then combined at point 58 with 63215 Mols
per hour of a work-expanded (isentropically expanded) and cooled
refrigerant gas of line 61 from a third stage expansion turbine 60
so as to form the combined refrigerant gas employed in subcooler
44. The combined refrigerant gas thus passes through conduit 62 and
enters the subcooler 44 at a flow rate of about 65411 Mols per hour
and at a temperature and pressure of about -246.degree.F. and 25
psia, respectively. At a designated temperature and pressure, the
combined refrigerant gas remains in essentially a gaseous state, so
as to eliminate erosion of the turbo-expander blades and simplify
distribution. Although it is preferred that no liquid be present, a
small quantity, i.e., 1 to 3 percent, liquid in the form of a mist
can be tolerated.
The combined refrigerant gas of conduit 62 emerges from subcooler
44 through conduit 64 at -126.degree.F. and 21 psia, hereinafter
referred to as the refrigerant gas, is passed to the first stage
expander precooler 70 at a flow rate of about 63215 Mols per hour
and at the same temperature and pressure as indicated
hereinabove.
The low pressure refrigerant gas fraction in conduit 64 enters the
expander precooler 70 to cool the countercurrent flow of warm, high
pressure, refrigerant gas from the compressor system. The warmed
low pressure, refrigerant gas then leaves precooler 70 through
conduit 66 at a temperature of about 90.degree.F. and a pressure of
about 16 psia. At point 68 a minor bleed or regeneration-fuel
fraction emerges through conduit 72 at a flow rate approximately
equal to the flow rate of the residual flash gas from the LNG
gas-liquid separating drum 52, i.e., about 2196 Mols per hour. As
described more fully below, this latter fraction is employed as
fuel gas to supply part of the power required to operate the
compressor system utilized in the process. The remaining fraction
is passed by means of line 74 into a first stage compressor 76
wherein the pressure of the refrigerant gas is increased to about
61 psia. The partially compressed refrigerant gas is then forced
through conduit 78 to a second stage compressor 80, and then
through conduit 82 to a third stage compressor 84 (shown in FIG.
1b). As illustrated in the drawing, intercoolers 85 and 86 are
utilized to cool the refrigerant gas in conduits 78 and 82,
respectively. After leaving the third stage compressor 84 through
conduit 88, the compressed refrigerant gas is at a pressure of
about 572 psia. The compressed refrigerant gas is then passed
through after cooler 90 wherein the temperature is reduced to about
95.degree.F., the pressure dropping slightly to about 562 psia. The
compressed and cooled refrigerant gas is then passed to the
expanded precooler 70 (shown in FIG. 1a) to be precooled by the
incoming low pressure refrigerant gas from conduit 64. The
precooled, compressed refrigerant gas from the precooler 70 then
passes through conduit 92 to a second stage enpansion turbine 94.
The flow rate of the gas entering the second stage expansion
turbine is about 63215 Mols per hour at a temperature of
-81.degree.F. and a pressure of 560 psia. The second stage
refrigerant expander gas leaves the second stage expansion turbine
94 at about 190 psia and -170.degree.F. and is passed through
conduit 96 to the cryogenic heat exchangers 38 and 40 where it is
utilized as the refrigerant. After emerging from the heat
exchangers 38 and 40, the warmed refrigerant gas is at a
temperature of -126.degree.F. and a pressure of 185 psia. The
refrigerant gas is then passed through conduit 98 to a third stage
expansion turbine 60 wherein the pressure is further reduced to
about 25 psia with a corresponding decrease in temperature to about
-246.degree.F. As pointed out above, this low pressure, low
temperature refrigerant gas is passed via conduit 61 and combined
at point 58 with the LNG flash gas in conduit 56 to form the
combined refrigerant gas which is passed via conduit 62 through the
subcooler 44 to reduce the temperature and thereby completely
liquify the LNG product stream carried in conduit 42.
The liquid bottoms from the fractionator 18 contains substantially
all of the carbon dioxide and higher boiling materials. This
bottoms leaves the fractionator 18 at the reboiler temperature of
140.degree.F. and flows under the fractionator pressure through
conduit 100 to point 102 where the glycol solution is removed
therefrom by conventional means (not shown).
The composition in mol percent of the glycol-free bottoms stream is
as follows: 11.89% carbon dioxide, 5.34% methane, 37.14% ethane,
16.57% propane, 9.01% butanes, and 20.05% C.sub.5.sub.+
hydrocarbons. This stream then passes at a rate of about 5003 Mols
per hour through conduit 104 to the CO.sub.2 fractionator 106
(shown in FIG. 1b). The fractionator 106 operates at 550 psia
distilling overhead at 20.degree.F. an azeotrope of approximately
equal molar quantities of ethane and carbon dioxide. In addition,
the overhead contains the methane left in the bottoms from the LNG
fractionator 18. The bottoms of the CO.sub.2 fractionator 106 is
substantially free from carbon dioxide. The fractionator 106 is
reboiled at about 195.degree.F., while overhead reflux is supplied
by a refrigerant such as propane at about 15.degree.F. in an
internal cooler 108. The use of the internal cooler 108 is
desirable since it saves equipment and space. The fractionator 106
typically comprises about 35 trays and operates at an L/D ratio of
about 2/1. The overhead from fractionator 106 is passed through
conduit 110 and may be employed as fuel gas. In a preferred
embodiment, the pressure of the overhead is reduced, e.g., to about
250 psia with a corresponding reduction in temperature to about
-14.degree.F. In this manner, the cooled overhead can be employed
to supply part of the cooling capacity required in throughout the
process being used as fuel gas.
The composition in mol percent of the liquid bottoms from the
CO.sub.2 fractionator 106 is approximately as follows: 35.48%
ethane, 23.45% propane, 12.73% butanes, and 28.34% C.sub.5.sub.+
hydrocarbons. This bottoms stream is transferred at a rate of about
3545 Mols per hour and a pressure of about 555 through conduit 112
to the LPG fractionator 114.
The LPG fractionator 114 operates at about 375 psia and
fractionates an overhead containing ethane, propane and as much
butane as can be tolerated to the solubility limit in LNG. To
increase the quantity of butane in the overhead, a rough cut may be
made between iso- and normal-butane as the iso- is more soluble in
LNG. As in the case of the CO.sub.2 fractionator 106, the overhead
of the LPG fractionator 114 is condensed by an internal cooler 116
to save equipment and space at the plant site. Typically, water is
employed as the refrigerant in cooler 116. The fractionator 114 is
designed to operate with about 25 trays and a L/D ratio of about
1/1.
The gas liquids bottoms from the fractionator 116 is withdrawn at a
temperature of about 470.degree.F. and is transferred to storage
through conduit 118 at a rate of about 1276 Mols per hour. The
composition in mol percent of this gas liquids product stream is
approximately as follows: 0.10% propane, 20.47% butanes, 16.42%
pentanes, 14.52% C.sub.6, and 48.49% C.sub.7.sub.+
hydrocarbons.
The overhead vapors from the LPG fractionator 116 pass through
conduit 120 to a compressor-condenser unit 122 where they are
compressed to about 635 psia and then condensed to a liquid by
cooling to about -10.degree.F. The condensed and cooled C.sub.2-4
fraction is then passed through conduit 124 for admixture with the
major portion of the overhead from the LNG fractionator 18 (shown
in FIG. 1a).
As discussed above, it is this mixed stream which passes through
conduit 42 and ultimately forms the LNG product stream.
As a result of the expansion in the turboexpanders 28, 60 and 94,
work is produced which furnishes a portion of the power necessary
to drive the compressors 84, 80 and 76. The power from the
turboexpanders is transferred to the compressors by connecting the
respective compressors and expanders with common shafts 126. Other
suitable mechanical means can also be used, such as gears, torque
converters, etc. The additional power required to drive the various
compressors is provided by separate conventional gas turbines 128
connected to each of the compressors 76, 80 and 84 by common shafts
130. Typically, the bleed gas or fuel gas removed from the process
in conduits 24, 72 or 110 is employed to drive the gas turbines
138. Steam turbines can also be used as an alternate drive. While
each of the gas turbines 128 is shown as being operatively
connected with a compressor by means of a common shaft 130, it will
be understood that any other suitable mechanical linkage, such as,
a gear chain, belt and pulley system, etc, can also be
employed.
Various modifications and alterations of this invention will become
apparent to those skilled in the art from the foregoing discussion
and accompanying drawing without departing from the scope and
spirit of this invention and this invention is not to be limited
unduly to that set forth herein for illustrative purposes.
* * * * *