Title of Invention

PROCESS AND APPARATUS FOR LIQUEFYING A NATURAL GAS STREAM

Abstract 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a natural gas stream containing; methane and heavier hydrocarbon components wherein (a) said natural gas stream is withdrawn from a cryogenic natural gas processing plant recovering natural gas liquids; (b) said natural gas stream is cooled under pressure sufficiently to partially condense it; (c) a distillation stream is withdrawn from said plant to supply at least a portion of said cooling of said natural gas stream; (d) said partially condensed natural gas stream is separated into a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 it; (f) said condensed expanded stream is directed to a distillation column at a mid-column feed point; (g) a liquid distillation stream is withdrawn from a lower region of said distillation column and directed to said plant; (h) a vapor distillation stream is withdrawn from an upper region of said distillation column and cooled under pressure to condense at least a portion of it and form a condensed stream; (i) said condensed stream is divided into at least two portions, with a first portion directed to said distillation column at a top feed position; (j) a second portion of said condensed stream is expanded to lower pressure to form said liquefied natural gas stream; and (k) me temperature of said partially condensed natural gas stream and the quantities and temperatures of said feed streams to said distillation column are effective to maintain the overhead temperature of said distillation column at a temperature whereby the major portion of said heavier hydrocarbon components is recovered in said liquid stream and said liquid distillation stream.
Full Text LNG PRODUCTION IN CRYOGENIC NATURAL GAS PROCESSING
PLANTS
SPECIFICATION
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 were 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, 92.6% methane, 4.7%
ethane and other C, components, 1.0% propane and other C3 components, 0.2%
iso-butane, 0.2% normal butane, 0.1% 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, Georgia,
March 13-15,2000 for a survey of a number of such processes. U.S. Pat. Nos.
5,363,655; 5,600,969; and 5,615,561 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 a single refrigerant fluid 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 increasing its energy requirements and without reducing
the NGL recovery level. 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 United States Patent 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
United States Patent 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 the present invention;
FIG. 5 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;
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;
FIG. 7 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. 8 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.
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 man 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/H) 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 hour (m3/H) 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°? [-55°C] (stream 36a), bottom liquid
product at 56°F [13°C] (stream 41a) 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 31a 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 33a. The substantially
condensed stream 33a 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 33b leaving expansion valve 13
reaches a temperature of-153°F [-103°C], and is supplied to separator section 17a in
the upper region of fractionation tower 17. The liquids separated therein become the
top feed to demethanizing section 17b.
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 34a 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 34a is supplied as
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 35a 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 17a 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 17b 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 17b 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 41a) in pump 18. Stream 41a, now at about 56°F [13°C], is
warmed to 85°F [29°C] (stream 41b) 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 41b as
it is wanned 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 36a), and heat exchanger 10 where it is heated to 68°F [20°C] (stream 36b).
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 38b) to sales line pressure. After cooling to 120°F [49°C] in discharge
cooler 20, the residue gas product (stream 38c) 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:
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 (Price, Brian C. "LNG
Production for Peak Shaving Operations", Proceedings of the Seventy-Eighth Annual
Convention of the Gas Processors Association, pp. 273-280, Atlanta, Georgia, March
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
m3/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 38c) produced by the NGL recovery plant is divided into two portions.
One portion (stream 42) 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 tide 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 74b). 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 72a 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 72a would typically exit heat exchanger 51 as a subcooled liquid stream.)
Stream 72a 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 72b 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 trie
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.(5%
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°Cj by the flashed refrigerant stream 74b. The subcooled
liquid stream 74a 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 74b). The flash expanded stream 74b then reenters heat exchanger
51 where it provides cooling to the feed gas (stream 72) and the refrigerant liquid
(stream 74) as it is vaporized and superheated.
The superheated refrigerant vapor (stream 74c) leaves heat exchanger
51 at 110°F [43°C] and flows to refrigerant compressor 55, driven by a supplemental
power source. Compressor 55 compresses the refrigerant to 145 psia [1,000 kPa(a)],
whereupon the compressed stream 74d returns to partial condenser 56 to complete the
cycle.
A summary of stream flow rates and energy consumption for the
process illustrated in FIG. 2 is set forth in the following table:

* (Based on un-rounded flow rates)
As stated earlier, the NGL recovery plant operates exactly the same in
the FIG. 2 process as it does for the FIG. 1 process, so the recovery levels for ethane,
propane, and butanes+ displayed in Table II are exactly the same as those displayed in
Table I. The only significant difference is the amount of plant fuel gas (stream 37)
used in the two processes. As can be seen by comparing Tables I and II, the plant fuel
gas consumption is higher for the FIG. 2 process because of the additional power
consumption of refrigerant compressor 55 (which is assumed to be driven by a gas
engine or turbine). There is consequently a correspondingly lesser amount of gas
entering residue gas compressor 19 (stream 38a), so the power consumption of this
compressor is slightly less for the FIG. 2 process compared to the FIG. 1 process.
The net increase in compression power for the FIG. 2 process
compared to the FIG. 1 process is 2,249 HP [3,697 kW], which is used to produce a
nominal 50,000 gallons/D [417 m3/D] of LNG. Since the density of LNG varies
considerably depending on its storage conditions, it is more consistent to evaluate the
power consumption per unit mass of LNG. The LNG production rate is 7,397 Lb/H
[3,355 kg/H] in this case, so the specific power consumption for the FIG. 2 process is
0.304 HP-H/Lb [0.500 kW-H/kg].
For this adaptation of the prior art LNG production process where the
NGL recovery plant residue gas is used as the source of feed gas for LNG production,
no provisions have been included for removing heavier hydrocarbons from the LNG
feed gas. Consequently, all of the heavier hydrocarbons present in the feed gas
become part of the LNG product, reducing the purity (i.e. methane concentration) of
the LNG product. If higher LNG purity is desired, or if the source of feed gas
contains higher concentrations of heavier hydrocarbons (inlet gas stream 31, for
instance), the feed stream 72 would need to be withdrawn from heat exchanger 51
after cooling to an intermediate temperature so that condensed liquid could be
separated, with the uncondensed vapor thereafter returned to heat exchanger 51 for
cooling to the final outlet temperature. These condensed liquids would preferentially
contain the majority of the heavier hydrocarbons, along with a considerable fraction of
liquid methane, which could then be re-vaporized and used to supply a part of the
plant fuel gas requirements. Unfortunately, this means that the C2 components, C3
components, and heavier hydrocarbon components removed from the LNG feed
stream would not be recovered in the NGL product from the NGL recovery plant, and
their value as liquid products would be lost to the plant operator. Further, for feed
streams such as the one considered in this example, condensation of liquid from the
feed stream may not be possible due to the process operating conditions (i.e.,
operating at pressures above the cricondenbar of the stream), meaning that removal of
heavier hydrocarbons could not be accomplished in such instances.
The process of FIG. 2 is essentially a stand-alone LNG production
facility that takes no advantage of the process streams or equipment in the NGL
recovery plant FIG. 3 shows another manner in which the NGL recovery plant in
FIG. 1 can be adapted for co-production of LNG, in this case by application of the
prior art process for LNG production according to U.S. Pat. No. 5,615,561, which
integrates the LNG production process with the NGL recovery plant. The inlet gas
composition and conditions considered in the process presented in FIG. 3 are the same
as those in FIGS. 1 and 2.
In the simulation of the FIG. 3 process, the inlet gas cooling,
separation, and expansion scheme for the NGL recovery plant is essentially the same
as that used in FIG. 1. The main differences are in the disposition of the cold
demethanizer overhead vapor (stream 36) and the compressed and cooled
demethanizer overhead vapor (stream 45c) produced by the NGL recovery plant. Inlet
gas enters the plant at 90°F [32°C] and 740 psia [5,102 kPa(a)] as stream 31 and is
cooled in heat exchanger 10 by heat exchange with cool demethanizer overhead vapor
at -69°F [-56°C] (stream 36b), bottom liquid product at 48°F [9°C] (stream 41a) from
demethanizer bottoms pump 18, demethanizer reboiler liquids at 26°F [-3°C] (stream
40), and demethanizer side reboiler liquids at -50°F [-46°C] (stream 39). The cooled
stream 31a enters separator 11 at -46°F [-43°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 gaseous first
and second streams, 33 and 34. Stream 33, containing about 25 percent of the total
vapor passes through heat exchanger 12 in heat exchange relation with the cold
demethanizer overhead vapor stream 36a where it is cooled to -142°F [-97°C]. The
resulting substantially condensed stream 33a is then flash expanded through
expansion valve 13 to the operating pressure (approximately 291 psia [2,006 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. 3, the
expanded stream 33b leaving expansion valve 13 reaches a temperature of -158°F
[-105°C] and is supplied to fractionation tower 17 as the top column feed. The vapor
portion (if any) of stream 33b combines with the vapors rising from the top
fractionation stage of the column to form demethanizer overhead vapor stream 36,
which is withdrawn from an upper region of the tower.
Returning to the gaseous second stream 34, the remaining 75 percent of
the vapor from separator 11 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 34a to a temperature of approximately -116°F [-82°C]. The
expanded and partially condensed stream 34a is thereafter supplied as feed to
fractionation tower 17 at an intermediate point. The separator liquid (stream 35) is
likewise expanded to the tower operating pressure by expansion valve 16, cooling
stream 35a to -80°F [-62°C] before it is supplied to fractionation tower 17 at a lower
mid-column feed point.
The liquid product (stream 41) exits the bottom of tower 17 at 42°F
[6°C]. This stream is pumped to approximately 650 psia [4,482 kPa(a)] (stream 41a)
in pump 18 and wanned to 83°F [28°C] (stream 41b) in heat exchanger 10 as it
provides cooling to stream 31. The distillation vapor stream forming the tower
overhead (stream 36) leaves demethanizer 17 at -154°F [-103°C] and is divided into
two portions. One portion (stream 43) is directed to heat exchanger 51 in the LNG
production section to provide most of the cooling duty in this exchanger as it is
warmed to -42°F [-41 °C] (stream 43a). The remaining portion (stream 42) bypasses
heat exchanger 51, with control valve 21 adjusting the quantity of this bypass in order
to regulate the cooling accomplished in heat exchanger 51. The two portions
recombine at -146°F [-99°C] to form stream 36a, which passes countercurrently to the
incoming feed gas in heat exchanger 12 where it is heated to -69°F [-56°C] (stream
36b) and heat exchanger 10 where it is heated to 72°F [22°C] (stream 36c). Stream
36c combines with warm HP flash vapor (stream 73a) from the LNG production
section, forming stream 44 at 72°F [22°C]. A portion of this stream is withdrawn
(stream 37) to serve as part of the fuel gas for the plant. The remainder (stream 45) is
re-compressed in two stages, compressor 15 driven by expansion machine 14 and
compressor 19 driven by a supplemental power source, and cooled to 120°F [49°C] in
discharge cooler 20. The cooled compressed stream (stream 45c) is then divided into
two portions. One portion is the residue gas product (stream 46), which flows to the
sales gas pipeline at 740 psia [5,102 kPa(a)]. The other portion (stream 71) is the feed
stream for the LNG production section.
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 (0.005 mole percent). As
for the FIG. 2 process, the feed stream 71 must be processed in carbon dioxide
removal section 50 (which may also include dehydration of the treated gas stream)
before entering the LNG production section to avoid operating problems due to
carbon dioxide freezing.
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 LP flash vapor at -200°F [-129°C] (stream 75), HP flash vapor at
-164°F [-109°C] (stream 73), and a portion of the demethanizer overhead vapor
(stream 43) at -154°F [-103°C] from the NGL recovery plant. The purpose of heat
exchanger 51 is to cool the feed stream to substantial condensation, and preferably to
subcool the stream so as to reduce the quantity of flash vapor generated in subsequent
expansion steps in the LNG cool-down section. 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 72a leaves heat exchanger 51 at -148°F
[-100°C] as a dense-phase fluid. At pressures below the cricondenbar, stream 72a
would typically exit heat exchanger 51 as a condensed (and possibly subcooled) liquid
stream.
Stream 72a is flash expanded substantially isenthalpically in expansion
valve 52 from about 727 psia [5,012 kPa(a)] to the operating pressure of HP flash
drum 53, about 279 psia [1,924 kPa(a)]. During expansion a portion of the stream is
vaporized, resulting in cooling of the total stream to -164°F [-109°C] (stream 72b).
The flash expanded stream 72b then enters HP flash drum 53 where the HP flash
vapor (stream 73) is separated and directed to heat exchanger 51 as described
previously. The operating pressure of the HP flash drum is set so that the heated HP
flash vapor (stream 73a) leaving heat exchanger 51 is at sufficient pressure to allow it
to join the heated demethanizer overhead vapor (stream 36c) leaving the NGL
recovery plant and subsequently be compressed by compressors 15 and 19.
The HP flash liquid (stream 74) from HP flash drum 53 is flash
expanded substantially isenthalpically in expansion valve 54 from the operating
pressure of the HP flash drum to the operating pressure of LP flash drum 55, about
118 psia (814 kPa(a)J. During expansion a portion of the stream is vaporized,
resulting in cooling of the total stream to -200°F [-129°C] (stream 74a). The flash
expanded stream 74a then enters LP flash drum 55 where the LP flash vapor (stream
75) is separated and directed to heat exchanger 51 as described previously. The
operating pressure of the LP flash drum is set so that the heated LP flash vapor
(stream 75a) leaving heat exchanger 51 is at sufficient pressure to allow its use as
plant fuel gas.
The LP flash liquid (stream 76) from LP flash drum 55 is flash
expanded substantially isenthalpically in expansion valve 56 from the operating
pressure of the LP flash drum to the LNG storage pressure (18 psia [124 kPa(a)j),
slightly above atmospheric pressure. During expansion a portion of the stream is
vaporized, resulting in cooling of the total stream to -254°F [-159°C] (stream 76a),
whereupon it is then directed to LNG storage tank 57 where the flash vapor resulting
from expansion (stream 77) is separated from the LNG product (stream 78).
The flash vapor (stream 77) from LNG storage tank 57 is at too low a
pressure to be used for plant fuel gas, and is too cold to enter directly into a
compressor. Accordingly, it is first heated to -30CF [-34°C] (stream 77a) in heater 58,
then compressors 59 and 60 (driven by supplemental power sources) are used to
compress the stream (stream 77c). Following cooling in aftercooler 61, stream 77d at
115 psia [793 kPa(a)] is combined with streams 37 and 75a to become the fuel gas for
the plant (stream 79).
A summary of stream flow rates and energy consumption for the
process illustrated in FIG. 3 is set forth in the following table:
The process of FIG. 3 uses a portion (stream 43) of the cold
demethanizer overhead vapor (stream 36) to provide refrigeration to the LNG
production process, which robs the NGL recovery plant of some of its refrigeration.
Comparing the recovery levels displayed in Table III for the FIG. 3 process to those in
Table II for the FIG. 2 process shows that the NGL recoveries have been maintained
at essentially the same levels for both processes. However, this comes at the expense
of increasing the utility consumption for the FIG. 3 process. Comparing the utility
consumptions in Table III with those in Table II shows that the residue gas
compression for the FIG. 3 process is nearly 18% higher than for the FIG. 2 process.
Thus, the recovery levels could be maintained for the FIG, 3 process only by lowering
the operating pressure of demethanizer 17, increasing the work expansion in machine
14 and thereby reducing the temperature of the demethanizer overhead vapor (stream
36) to compensate for the refrigeration lost to the NGL recovery plant in stream 43.
As can be seen by comparing Tables I and III, the plant fuel gas
consumption is higher for the FIG. 3 process because of the additional power
consumption of flash vapor compressors 59 and 60 (which are assumed to be driven
by gas engines or turbines). There is consequently a correspondingly lesser amount of
gas entering residue gas compressor 19 (stream 45a), but the power consumption of
this compressor is still higher for the FIG. 3 process compared to the FIG. 1 process
because of the higher compression ratio. The net increase in compression power for
the FIG. 3 process compared to the FIG. 1 process is 2,696 HP [4,432 kW] to produce
the nominal 50,000 gallons/D [417 m3/D] of LNG. The specific power consumption
for the FIG. 3 process is 0.366 HP-H/Lb [0.602 kW-H/kg], or about 20% higher than
for the FIG. 2 process.
The FIG. 3 process has no provisions for removing heavier
hydrocarbons from the feed gas to its LNG production section. Although some of the
heavier hydrocarbons present in the feed gas leave in the flash vapor (streams 73 and
75) from separators 53 and 55, most of the heavier hydrocarbons become part of the
LNG product and reduce its purity. The FIG. 3 process is incapable of increasing the
LNG purity, and if a feed gas containing higher concentrations of heavier
hydrocarbons (for instance, inlet gas stream 31, or even residue gas stream 45c when
the NGL recovery plant is operating at reduced recovery levels) is used to supply the
feed gas for the LNG production plant, the LNG purity would be even less than shown
in this example.
DESCRIPTION OF THE INVENTION
Example 1
FIG. 4 illustrates a flow diagram of a process in accordance with the
present invention. The inlet gas composition and conditions considered in the process
presented in FIG. 4 are the same as those in FIGS. I through 3. Accordingly, the FIG.
4 process can be compared with that of the FIG. 2 and FIG. 3 processes to illustrate
the advantages of the present invention.
In the simulation of the FIG. 4 process, the inlet gas cooling,
separation, and expansion scheme for the NGL recovery plant is essentially the same
as that used in FIG. 1. The main difference is that the inlet gas (stream 30) is divided
into two portions, and only the first portion (stream 31) is supplied to the NGL
recovery plant. The other portion (stream 71) is the feed gas for the LNG production
section which employs the present invention.
Inlet gas enters the plant at 90°F [32°C] and 740 psia [5,102 kPa(a)] as
stream 30. The feed gas for the LNG section is withdrawn (stream 71) and the
remaining portion (stream 31) is cooled in heat exchanger 10 by heat exchange with
cool distillation vapor at -66°F [-54°C] (stream 36a), bottom liquid product at 51°F
[10°C] (stream 41a) from demethanizer bottoms pump 18, demethanizer reboiler
liquids at 30°F [-1°C] (stream 40), and demethanizer side reboiler liquids at -39°F
[-39°C] (stream 39). The cooled stream 31a enters separator 11 at -44°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 gaseous first
and second streams, 33 and 34. Stream 33, containing about 26 percent of the total
vapor passes through heat exchanger 12 in heat exchange relation with cold
distillation vapor stream 36 where it is cooled to -148°F [-100°CJ. The resulting
substantially condensed stream 33a is then flash expanded through expansion valve
13 to the operating pressure (approximately 301 psia [2,075 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. 4, the expanded stream 33b
leaving expansion valve 13 reaches a temperature of -156°F [-105°C] and is supplied
to fractionation tower 17 as the top column feed. The vapor portion (if any) of stream
33b combines with the vapors rising from the top fractionation stage of the column to
form distillation vapor stream 42, which is withdrawn from an upper region of lie
tower.
Returning to the gaseous second stream 34, the remaining 74 percent of
the vapor from separator 11 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 34a to a temperature of approximately -111 °F [-80°C]. The
expanded and partially condensed stream 34a is thereafter supplied as feed to
fractionation tower 17 at an intermediate point. The separator liquid (stream 35) is
likewise expanded to the tower operating pressure by expansion valve 16, cooling
stream 35a to -75°F [-59°C] before it is supplied to fractionation tower 17 at a lower
mid-column feed point.
The liquid product (stream 41) exits the bottom of tower 17 at 45 °F
[7°C]. This stream is pumped to approximately 650 psia [4,482 kPa(a)] (stream 41a)
in pump 18 and warmed to 84°F [29°C] (stream 41b) in heat exchanger 10 as it
provides cooling to stream 31. The distillation vapor stream forming the tower
overhead at -152°F [-102°C] (stream 42) is divided into two portions. One portion
(stream 86) is directed to the LNG production section. The remaining portion (stream
36) passes countercurrently to the incoming feed gas in heat exchanger 12 where it is
heated to -66°F [-54°C] (stream 36a) and in heat exchanger 10 where it is heated to
72°F [22°C] (stream 36b). A portion of the warmed distillation vapor stream is
withdrawn (stream 37) to serve as part of the fuel gas for the plant, with the remainder
becoming the first residue gas (stream 43). The first residue gas is then re-compressed
in two stages, compressor 15 driven by expansion machine 14 and compressor 19
driven by a supplemental power source to form the compressed first residue gas
(stream 43b).
Turning now to the LNG production section that employs the present
invention, feed stream 71 enters heat exchanger 50 at 90°F [32°C] and 740 psia [5,102
kPa(a)]. Note that in all cases heat exchanger 50 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, feed stream flow rate, heat exchanger size, stream
temperatures, etc.) In heat exchanger 50, the feed stream 71 is cooled by heat
exchange with cool LNG flash vapor (stream 83a) and the distillation vapor stream
from the NGL recovery plant (stream 86). The cooled stream 71a enters separator 51
at -36°F [-38°C] and 737 psia [5,081 kPa(a)] where the vapor (stream 72) is separated
from the condensed liquid (stream 73).
The vapor (stream 72) from separator 51 enters a work expansion
machine 52 in which mechanical energy is extracted from this portion of the high
pressure feed. The machine 52 expands the vapor substantially isentropically from a
pressure of about 737 psia [5,081 kPa(a)] to slightly above the operating pressure (440
psia [3,034 kPa(a)]) of distillation column 56, with the work expansion cooling the
expanded stream 72a to a temperature of approximately -79°F [-62°C]. The expanded
and partially condensed stream 72a is directed to heat exchanger 50 and further cooled
and condensed by heat exchange with cool LNG flash vapor (stream 83a) and the
distillation vapor stream from the NGL recovery plant (stream 86) as described
earlier, and by flash liquids (stream 80) and distillation column reboiler liquids at
-135°F [-93°C] (stream 76). The condensed stream 72b, now at -135°F [-93°C], is
thereafter supplied as feed to distillation column 56 at an intermediate point.
Distillation column 56 serves as an LNG purification tower. It is a
conventional distillation column containing a plurality of vertically spaced trays, one
or more packed beds, or some combination of trays and packing. This tower recovers
nearly all of the hydrocarbons heavier than methane present in its feed stream (stream
72b) as its bottom product (stream 77) so that the only significant impurity in its
overhead (stream 74) is the nitrogen contained in the feed stream. Equally important,
this tower also captures in its bottom product nearly all of the carbon dioxide feeding
the tower, so that carbon dioxide does not enter the downstream LNG cool-down
section where the extremely low temperatures would cause the formation of solid
carbon dioxide, creating operating problems. The lower section of LNG purification
tower 56 includes a reboiler which heats and vaporizes a portion of the liquids flowing
down the column (by cooling stream 72a in heat exchanger 50 as described earlier) to
provide stripping vapors which flow up the column to strip some of the methane from
the liquids. This reduces the amount of methane in the bottom product from the tower
(stream 77) so that less methane must be rejected by fractionation tower 17 when this
stream is supplied to it (as described later).
Reflux for distillation column 56 is created by cooling and condensing
the tower overhead vapor (stream 74 at -142°F [-96°C]) in heat exchanger 50 by heat
exchange with cool LNG flash vapor at -147°F [-99°C] (stream 83a) and flash liquids
at -152°F [-102°C] (stream 80). The condensed stream 74a, now at -144°F [-98°C], is
divided into two portions. One portion (stream 78) becomes the feed to the LNG
cool-down section. The other portion (stream 75) enters reflux pump 55. After
pumping, stream 75a at -143°F [-97°C] is supplied to LNG purification tower 56 at a
top feed point to provide the reflux liquid for the tower. This reflux liquid rectifies
the vapors rising up the tower so that the tower overhead vapor (stream 74) and
consequently feed stream 78 to the LNG cool-down section contain minimal amounts
of carbon dioxide and hydrocarbons heavier than methane. The amount of reboiling
in the bottom of the column is adjusted as necessary to generate sufficient overhead
vapor from the column, so that there is enough reflux liquid from heat exchanger 50 to
provide the desired rectification in the tower.
The feed stream for the LNG cool-down section (condensed liquid
stream 78) enters heat exchanger 58 at -144°F [-98°C] and is subcooled by heat
exchange with cold LNG flash vapor at -255°F [-160°C] (stream 83) and cold flash
liquids (stream 79a). The cold flash liquids are produced by withdrawing a portion of
the partially subcooled feed stream (stream 79) from heat exchanger 58 and flash
expanding the stream through an appropriate expansion device, such as expansion
valve 59, to slightly above the operating pressure of fractionation tower 17. During
expansion a portion of the stream is vaporized, resulting in cooling of the total stream
from -157°F [-105°C] to -161°F [-107°C] (stream 79a). The flash expanded stream
79a is then supplied to heat exchanger 58 as previously described.
The remaining portion of the partially subcooled feed stream is further
subcooled in heat exchanger 58 to -170°F [-112°C] (stream 82). It then enters a work
expansion machine 60 in which mechanical energy is extracted from this intermediate
pressure stream. The machine 60 expands the subcooled liquid substantially
isentropically from a pressure of about 434 psia [2,992 kPa(a)] to the LNG storage
pressure (18 psia [124 kPa(a)]), slightly above atmospheric pressure. The work
expansion cools the expanded stream 82a to a temperature of approximately -255°F
[-160°C], whereupon it is then directed to LNG storage tank 61 where the flash vapor
resulting from expansion (stream 83) is separated from the LNG product (stream 84).
Tower bottoms stream 77 from LNG purification tower 56 is flash
expanded to slightly above the operating pressure of fractionation tower 17 by
expansion valve 57. During expansion a portion of the stream is vaporized, resulting
in cooling of the total stream from -133°F [-92°C] to -152°F [-102°C] (stream 77a).
The flash expanded stream 77* is then combined with warmed flash liquid stream 79b
leaving heat exchanger 58 at -I47°F [-99°C] to form a combined flash liquid stream
(stream 80) at -152°F [-102°C] which is supplied to heat exchanger 50. It is heaited to
-88°F [-67°C] (stream 80a) as it supplies cooling to expanded stream 72a and tower
overhead vapor stream 74 as described earlier.
The separator liquid (stream 73) is flash expanded to the operating
pressure of fractionation tower 17 by expansion valve 54, cooling stream 73a to -65°F
[-54°C]. The expanded stream 73a is combined with heated flash liquid stream 80a to
form stream 81, which is supplied to fractionation tower 17 at a lower mid-column
feed point. If desired, stream 81 can be combined with flash expanded stream 35a
described earlier and the combined stream supplied to a single lower mid-column feed
point on the tower.
The flash vapor (stream 83) from LNG storage tank 61 passes
countercurrently to the incoming liquid in heat exchanger 58 where it is heated to
-147°F [-99°C] (stream 83a). It then enters heat exchanger 50 where it is heated to
87°F [31 °C] (stream 83b) as it supplies cooling to feed stream 71, expanded stream
72a, and tower overhead stream 74. Since this stream is at low pressure (15.5 psia
[107 kPa(a)]), it must be compressed before it can be used as plant fuel gas.
Compressors 63 and 65 (driven by supplemental power sources) with intercooler 64
are used to compress the stream (stream 83e). Following cooling in aftercooler 66,
stream 83f at 115 psia [793 kPa(a)] is combined with stream 37 to become the fuel gas
for the plant (stream 85).
The cold distillation vapor stream from the NGL recovery plant
(stream 86) is heated to 86°F [30°C] as it supplies cooling to feed stream 71 and
expanded stream 72a in heat exchanger 50, becoming the second residue gas (stream
86a). The second residue gas is then re-compressed in two stages, compressor 53
driven by expansion machine 52 and compressor 62 driven by a supplemental power
source. The compressed second residue gas (stream 86c) combines with the
compressed first residue gas (stream 43b) to form residue gas stream 38. After
cooling to 120°F [49°C] in discharge cooler 20, the residue gas product (stream 38a)
flows to the sales gas pipeline at 740 psia [5,102 kPa(a)J.
A summary of stream flow rates and energy consumption for the
process illustrated in FIG. 4 is set forth in the following table:
Comparing the recovery levels displayed in Table IV for the FIG. 4
process to those in Table I for the FIG. 1 process shows that the recoveries in the NGL
recovery plant have been maintained at essentially the same levels for both processes.
Comparison of the utility consumptions displayed in Table IV for the FIG. 4 process
with those in Table I for the FIG. 1 process shows that the residue gas compression
required for the NGL recovery plant is essentially the same for both processes. This
indicates that there is no loss in recovery efficiency despite using a portion (stream
86) of the cold distillation vapor stream (stream 42) from the NGL recovery plant to
provide refrigeration to the LNG production section. Thus, unlike the FIG. 3 process,
integrating the LNG production process of the present invention with the NGL
recovery plant can be accomplished without adverse impact on NGL recovery
efficiency.
The net increase in compression power for the FIG. 4 process
compared to the FIG. 1 process is 1,498 HP [2,463 kW] to produce the nominal
50,000 gallons/D [417 m3/D] of LNG, giving a specific power consumption of 0.204
HP-H/Lb [0.336 kW-H/kg] for the FIG. 4 process. Thus, the present invention has a
specific power consumption that is only 67% of the FIG. 2 prior art process and only
56% of the FIG. 3 prior art process. Further, the present invention does not require
carbon dioxide removal from the feed gas prior to entering the LNG production
section like the prior art processes do, eliminating the capital cost and operating cost
associated with constructing and operating the gas treatment processes required for the
FIG. 2 and FIG. 3 processes.
Not only is the present invention more efficient than either prior art
process, the LNG it produces is of higher purity due to the inclusion of LNG
purification tower 56. This higher LNG purity is even more noteworthy considering
that the source of the feed gas used for this example (inlet gas, stream 30) contains
much higher concentrations of heavier hydrocarbons than the feed gas used in the
FIG. 2 and FIG. 3 processes (i.e., the NGL recovery plant residue gas). The purity of
the LNG is in fact limited only by the concentration of gases more volatile than
methane (nitrogen, for instance) present in feed stream 71, as the operating parameters
of purification tower 56 can be adjusted as needed to keep the concentration of
heavier hydrocarbons in the LNG product as low as desired.
Example 2
FIG. 4 represents the preferred embodiment of the present invention
for the temperature and pressure conditions shown because it typically provides the
most efficient LNG production. A slightly less complex design that maintains the
same LNG production with somewhat higher utility consumption can be achieved
using another embodiment of the present invention as illustrated in the FIG. 5 process.
The inlet gas composition and conditions considered in the process presented in FIG.
5 are the same as those in FIGS. 1 through 4. Accordingly, the FIG. 5 process can be
compared with that of the FIG. 2 and FIG. 3 processes to illustrate the advantages of
the present invention, and can likewise be compared to the embodiment displayed in
FIG. 4.
In the simulation of the FIG. 5 process, the inlet gas cooling,
separation, and expansion scheme for the NGL recovery plant is essentially the same
as that used in FIG. 4. Inlet gas enters the plant at 90°F [32°C] and 740 psia [5, 102
kPa(a)] as stream 30. The feed gas for the LNG section is withdrawn (stream 71) and
the remaining portion (stream 31) is cooled in heat exchanger 10 by heat exchange
with cool distillation vapor at -65°F [-54°C] (stream 36a), bottom liquid product at
50°F [10°C] (stream 41a) from demethanizer bottoms pump 18, demethanizer reboiler
liquids at 29°F [-2°C] (stream 40), and demethanizer side reboiler liquids at -41°F
[-40°C] (stream 39). The cooled stream 31a 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 gaseous first
and second streams, 33 and 34. Stream 33, containing about 26 percent of the total
vapor passes through heat exchanger 12 in heat exchange relation with the cold
distillation vapor stream 36 where it is cooled to -148°F [-100°C]. The resulting
substantially condensed stream 33a is then flash expanded through expansion valve
13 to the operating pressure (approximately 296 psia [2,041 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. 5, the expanded stream 33b
leaving expansion valve' 13 reaches a temperature of-157°F [-105°C] and is supplied
to fractionation tower 17 as the top column feed. The vapor portion (if any) of stream
33b combines with the vapors rising from the top fractionation stage of the column to
form distillation vapor stream 42, which is withdrawn from an upper region of the
tower.
Returning to the gaseous second stream 34, the remaining 74 percent of
the vapor from separator 11 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 34a to a temperature of approximately -112°F [-80°C]. The
expanded and partially condensed stream 34a is thereafter supplied as feed to
fractionation tower 17 at an intermediate point. The separator liquid (stream 35) is
likewise expanded to the tower operating pressure by expansion valve 16, cooling
stream 35a to -75°F [-59°C] before it is supplied to fractionation tower 17 at a lower
mid-column feed point.
The liquid product (stream 41) exits the bottom of tower 17 at 44°F
[7°C]. This stream is pumped to approximately 650 psia [4,482 kPa(a)] (stream 41a)
in pump 18 and wanned to 83°F [28°C] (stream 41b) in heat exchanger 10 as it
provides cooling to stream 31. The distillation vapor stream forming the tower
overhead at -153°F [-103°C] (stream 42) is divided into two portions. One portion
(stream 86) is directed to the LNG production section. The remaining portion (stream
36) passes countercurrently to the incoming feed gas in heat exchanger 12 where it is
heated to -65 °F [-54°C] (stream 36a) and heat exchanger 10 where it is heated to 73 °F
[23°C] (stream 36b). A portion of the warmed distillation vapor stream is withdrawn
(stream 37) to serve as part of the fuel gas for the plant, with the remainder becoming
the first residue gas (stream 43). The first residue gas is then re-compressed in two
stages, compressor 15 driven by expansion machine 14 and compressor 19 driven by a
supplemental power source to form the compressed first residue gas (stream 43b).
Turning now to the LNG production section that employs an
alternative embodiment of the present invention, feed stream 71 enters heat exchanger
50 at 90°F [32°C] and 740 psia [5,102 kPa(a)]. The feed stream 71 is cooled to
-120°F [-84°C] in heat exchanger 50 by heat exchange with cool LNG flash vapor
(stream 83a), the distillation vapor stream from the NGL recovery plant at -153°F
[-103°C] (stream 86), flash liquids (stream 80), and distillation column reboiler
liquids at -134°F [-92°C] (stream 76). The resulting substantially condensed stream
71a is then flash expanded through an appropriate expansion device, such as
expansion valve 52, to the operating pressure (440 psia [3,034 kPa(a)]) of distillation
column 56. During expansion a portion of the stream is vaporized, resulting in
cooling of the total stream. In the process illustrated in FIG. 5, the expanded stream
71b leaving expansion valve 52 reaches a temperature of-134°F [-92°C] and is
thereafter supplied as feed to distillation column 56 at an intermediate point.
As in the FIG. 4 embodiment of the present invention, distillation
column 56 serves as an LNG purification tower, recovering nearly all of the carbon
dioxide and the hydrocarbons heavier than methane present in its feed stream (stream
71b) as its bottom product (stream 77) so that the only significant impurity in its
overhead (stream 74) is the nitrogen contained in the feed stream. Reflux for
distillation column 56 is created by cooling and condensing the tower overhead vapor
(stream 74 at -141 °F [-96°CJ) in heat exchanger 50 by heat exchange with cool LNG
flash vapor at -146°F [-99°C] (stream 83a) and flash liquids at -152°F [-102°C]
(stream 80). The condensed stream 74a, now at -144°F [-98°C], is divided into two
portions. One portion (stream 78) becomes the feed to the LNG cool-down section.
The other portion (stream 75) enters reflux pump 55. After pumping, stream 75a at
-143°F [-97°C] is supplied to LNG purification tower 56 at a top feed point to provide
the reflux liquid for the tower. This reflux liquid rectifies the vapors rising up the
tower so that the tower overhead (stream 74) and consequently feed stream 78 to the
LNG cool-down section contain minimal amounts of carbon dioxide and
hydrocarbons heavier than methane.
The feed stream for the LNG cool-down section (condensed liquid
stream 78) enters heat exchanger 58 at -144°F [-98°C] and is subcooled by heat
exchange with cold LNG flash vapor at -255°F [-160°C] (stream 83) and cold flash
liquids (stream 79a). The cold flash liquids are produced by withdrawing a portion of
the partially subcooled feed stream (stream 79) from heat exchanger 58 and flash
expanding the stream through an appropriate expansion device, such as expansion
valve 59, to slightly above the operating pressure of fractionation tower 17. During
expansion a portion of the stream is vaporized, resulting in cooling of the total stream
from -157°F [-105°C] to -162°F [-108°C] (stream 79a). The flash expanded stream
79a is then supplied to heat exchanger 58 as previously described.
The remaining portion of the partially subcooled feed stream is further
subcooled in heat exchanger 58 to -170°F [-112°C] (stream 82). It then enters a work
expansion machine 60 in which mechanical energy is extracted from this intermediate
pressure stream. The machine 60 expands the subcooled liquid substantially
isentropicalty from a pressure of about 434 psia [2,992 kPa(a)] to the LNG storage
pressure (18 psia [124 kPa(a)]), slightly above atmospheric pressure. The work
expansion cools the expanded stream 82a to a temperature of approximately -255°F
[-160°C], whereupon it is then directed to LNG storage tank 61 where the flash vapor
resulting from expansion (stream 83) is separated from the LNG product (stream 84).
Tower bottoms stream 77 from LNG purification tower 56 is flash
expanded to slightly above the operating pressure of fractionation tower 17 by
expansion valve 57. During expansion a portion of the stream is vaporized, resulting
in cooling of the total stream from -133°F [-91°C] to -152°F [-102°C] (stream 77a).
The flash expanded stream 77a is then combined with warmed flash liquid stream 79b
leaving heat exchanger 58 at -146°F [-99°C] to form a combined flash liquid stream
(stream 80) at -152°F [-102°C] which is supplied to heat exchanger 50. It is heated to
-87°F [-66°C] (stream 80a) as it supplies cooling to feed stream 71 and tower
overhead vapor stream 74 as described earlier, and thereafter supplied to fractionation
tower 17 at a lower mid-column feed point If desired, stream 80a can be combined
with flash expanded stream 35a described earlier and the combined stream supplied to
a single lower mid-column feed point on the tower.
The flash vapor (stream 83) from LNG storage tank 61 passes
countercurrently to the incoming liquid in heat exchanger 58 where it is heated to
-146°F [-99°C] (stream 83a). It then enters heat exchanger 50 where it is heated to
87°F [31 °C] (stream 83b) as it supplies cooling to feed stream 71 and tower overhead
stream 74. Since this stream is at low pressure (15.5 psia [107 kPa(a)]), it must be
compressed before it can be used as plant fuel gas. Compressors 63 and 65 (driven by
supplemental power sources) with intercooler 64 are used to compress the stream
(stream 83e). Following cooling in aftercooler 66, stream 83f at 115 psia [793 kPa(a)]
is combined with stream 37 to become the fuel gas for the plant (stream 85).
The cold distillation vapor stream from the NGL recovery plant
(stream 86) is heated to 87°F [31°C] as it supplies cooling to feed stream 71 in heat
exchanger 50, becoming the second residue gas (stream 86a) which is then
re-compressed in compressor 62 driven by a supplemental power source. The
compressed second residue gas (stream 86b) combines with the compressed first
residue gas (stream 43b) to form residue gas stream 38. After cooling to 120°F
[49°C] in discharge cooler 20, the residue gas product (stream 38a) flows to the sales
gas pipeline at 740 psia [5,102 kPa(a)].
A summary of stream flow rates and energy consumption for the
process illustrated in FIG. 5 is set forth in the following table:
As can be seen by comparing the recovery levels and utility
consumptions displayed in Table V for the FIG. 5 process with those in Table I and
Table IV for the FIG. 1 and FIG. 4 processes, respectively, the recovery efficiency of
the NGL recovery plant is undiminished when integrated with this embodiment of the
present invention for co-production of LNG. The LNG production efficiency of this
embodiment is not as high as for the preferred embodiment shown in FIG. 4 due to the
higher utility consumption of second residue gas compressor 62 that results from
eliminating the work expansion machine 52 that was used to drive compressor S3 in
the FIG. 4 embodiment. The net increase in compression power for the FIG. 5
process compared to the FIG. 1 process is 2,097 HP [3,447 kW] to produce the
nominal 50,000 gallons/D [417 m3/D] of LNG, giving a specific power consumption
of 0.286 HP-H/Lb [0.470 kW-H/kg] for the FIG. 5 process. Although this is about
40% higher than the preferred embodiment shown in FIG. 4, it is still lower than
either of the prior art processes displayed in FIGS. 2 and 3. Further, as for the FIG. 4
embodiment, the LNG purity is higher than for either prior art process, and carbon
dioxide removal from the feed gas to the LNG production section is not required.
The choice between the FIG. 4 embodiment and the FIG. 5
embodiment of the present invention depends on the relative value of the simpler
arrangement and lower capital cost of the FIG. 5 embodiment versus the lower utility
consumption of the FIG. 4 embodiment. The decision of which embodiment of the
present invention to use in a particular circumstance will often depend on factors such
as plant size, available equipment, and the economic balance of capital cost versus
operating cost.
Example 3
In FIGS. 4 and 5, a portion of the plant inlet gas is processed using the
present invention to co-produce LNG. Alternatively, the present invention can instead
be adapted to process a portion of the plant residue gas to co-produce LNG as
illustrated in FIG. 6. The inlet gas composition and conditions considered in the
process presented in FIG. 6 are the same as those in FIGS. I through 5. Accordingly,
the FIG. 6 process can be compared with that of the FIG. 2 and FIG. 3 processes to
illustrate the advantages of the present invention, and can likewise be compared to the
embodiments displayed in FIGS. 4 and 5.
In the simulation of the FIG. 6 process, the inlet gas cooling,
separation, and expansion scheme for the NGL recovery plant is essentially the same
as that used in FIG. 1. The main differences are in the disposition of the cold
distillation stream (stream 42) and the compressed and cooled third residue gas
(stream 44a) produced by the NGL recovery plant. Note that the third residue gas
(stream 44a) is divided into two portions, and only the first portion (stream 38)
becomes the residue gas product from the NGL recovery plant that flows to the sales
gas pipeline. The other portion (stream 71) is the feed gas for the LNG production
section which employs the present invention.
Inlet gas enters the plant at 90°F [32°C] and 740 psia [5,102 kPa(a)] as
stream 31 and is cooled in heat exchanger 10 by heat exchange with cool distillation
vapor stream 36a at -66°F [-55°C], bottom liquid product at 52°F [11°C] (stream 41a)
from demethanizer bottoms pump 18, demethanizer reboiler liquids at 31 °F [0°C]
(stream 40), and demethanizer side reboiler liquids at -42°F [-41°C] (stream 39). The
cooled stream 31a enters separator 11 at -44°F [-42°C] and 725 psia [4,999 kPa(a)]
where the vapor (stream 32) is separated mom the condensed liquid (stream 35).
The vapor (stream 32) from separator 11 is divided into gaseous first
and second streams, 33 and 34. Stream 33, containing about 26 percent of the total
vapor passes through heat exchanger 12 in heat exchange relation with the cold
distillation vapor stream 36 where it is cooled to -146°F [-99°C]. The resulting
substantially condensed stream 33a is then flash expanded through expansion valve
13 to the operating pressure (approximately 306 psia [2,110 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. 6, the expanded stream 33b
leaving expansion valve 13 reaches a temperature of-155°F [-104°C] and is supplied
to fractionation tower 17 as the top column feed. The vapor portion (if any) of stream
33b combines with the vapors rising from the top fractionation stage of the column to
form distillation vapor stream 42, which is withdrawn from an upper region of the
tower.
Returning to the gaseous second stream 34, the remaining 74 percent of
the vapor from separator 11 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 34a to a temperature of approximately -110°F [-79°C]. The
expanded and partially condensed stream 34a is thereafter supplied as feed to
fractionation tower 17 at an intermediate point. The separator liquid (stream 35) is
likewise expanded to the tower operating pressure by expansion valve 16, cooling
stream 35a to -75°F [-59°C] before it is supplied to fractionation tower 17 at a lower
mid-column feed point.
The liquid product (stream 41) exits the bottom of tower 17 at 47°F
[8°C]. This stream is pumped to approximately 650 psia [4,482 kPa(a)] (stream 41a)
in pump 18 and warmed to 83°F [28°C] (stream 41b) in heat exchanger 10 as it
provides cooling to stream 31. The distillation vapor stream forming the tower
overhead at -151°F [-102°C] (stream 42) is divided into two portions. One portion
(stream 86) is directed to the LNG production section. The remaining portion (stream
36) passes countercurrently to the incoming feed gas in heat exchanger 12 where it is
heated to -66°F [-55°C] (stream 36a) and heat exchanger 10 where it is heated 1o 72°F
[22°C] (stream 36b). A portion of the warmed distillation vapor stream is withdrawn
(stream 37) to serve as part of the fuel gas for the plant, with the remainder becoming
the first residue gas (stream 43). The first residue gas is then re-compressed in two
stages, compressor 15 driven by expansion machine 14 and compressor 19 driven by a
supplemental power source to form the compressed first residue gas (stream 43b).
Turning now to the LNG production section that employs an
alternative embodiment of the present invention, feed stream 71 enters heat exchanger
50 at 120°F [49°C] and 740 psia [5,102 kPa(a)]. The feed stream 71 is cooled to
-120°F [-84°C] in heat exchanger 50 by heat exchange with cool LNG flash vapor
(stream 83a), the distillation vapor stream from the NGL recovery plant at -151°F
[-102°C] (stream 86), flash liquids (stream 80), and distillation column reboiler
liquids at -142°F [-97°C] (stream 76). (For the conditions stated, the feed stream
pressure is above the cricondenbar, so no liquid will condense as the stream is cooled.
Instead, the cooled stream 71a leaves heat exchanger 50 as a dense-phase fluid. For
other processing conditions, it is possible that the feed gas pressure will be below its
cricondenbar pressure, in which case the feed stream will be cooled to substantial
condensation. In addition, it may be advantageous to withdraw the feed stream after
cooling to an intermediate temperature, separate any condensed liquid that may have
formed, and then expand the vapor stream in a work expansion machine prior to
cooling the expanded stream to substantial condensation, similar to the embodiment
displayed in FIG. 4. In this case, there was little advantage to work expanding the
dense-phase feed stream, so the simpler embodiment shown in FIG. 6 was employed
instead.) The resulting cooled stream 71a is then flash expanded through an
appropriate expansion device, such as expansion valve 52, to the operating pressure
(420 psia [2,896 kPa(a)]) of distillation column 56. During expansion a portion of the
stream is vaporized, resulting in cooling of the total stream. In the process illustrated
in FIG. 6, the expanded stream 71b leaving expansion valve 52 reaches a temperature
of-143°F [-97°C] and is thereafter supplied as feed to distillation column 56 at an
intermediate point
As for the FIG. 4 and FIG. 5 embodiments of the present invention,
distillation column 56 serves as an LNG purification tower, recovering nearly all of
the carbon dioxide and the hydrocarbons heavier than methane present in its feed
stream (stream 71b) as its bottom product (stream 77) so that the only significant
impurity in its overhead (stream 74) is the nitrogen contained in the feed stream.
Reflux for distillation column 56 is created by cooling and condensing the tower
overhead vapor (stream 74 at -144°F [-98°C]) in heat exchanger 50 by heat exchange
with cool LNG flash vapor at -155°F [-104°C] (stream 83a) and flash liquids at
-156°F [-105°C] (stream 80). The condensed stream 74a, now at -146°F [-99°C], is
divided into two portions. One portion (stream 78) becomes the feed to the LNG
cool-down section. The other portion (stream 75) enters reflux pump 55. After
pumping, stream 75a at -145°F [-98°C] is supplied to LNG purification tower 56 at a
top feed point to provide the reflux liquid for the tower. This reflux liquid rectifies
the vapors rising up the tower so that the tower overhead (stream 74) and
consequently feed stream 78 to the LNG cool-down section contain minimal amounts
of carbon dioxide and hydrocarbons heavier than methane.
The feed stream for the LNG cool-down section (condensed liquid
stream 78) enters heat exchanger 58 at -146°F [-99°C] and is subcooled by heat
exchange with cold LNG flash vapor at -255°F [-159°C] (stream 83) and cold flash
liquids (stream 79a). The cold flash liquids are produced by withdrawing a portion of
the partially subcooled feed stream (stream 79) from heat exchanger 58 and flash
expanding the stream through an appropriate expansion device, such as expansion
valve 59, to slightly above the operating pressure of fractionation tower 17. During
expansion a portion of the stream is vaporized, resulting in cooling of the total stream
from -156°F [-104°C] to -160°F [-106°C] (stream 79a). The flash expanded stream
79a is then supplied to heat exchanger 58 as previously described.
The remaining portion of the partially subcooled feed stream is further
subcooled in heat exchanger 58 to -169°F [-112°C] (stream 82). It then enters a work
expansion machine 60 in which mechanical energy is extracted from this intermediate
pressure stream. The machine 60 expands the subcooled liquid substantially
isentropically from a pressure of about 414 psia [2,858 kPa(a)] to the LNG storage
pressure (18 psia [124 kPa(a)J), slightly above atmospheric pressure. The work
expansion cools the expanded stream 82a to a temperature of approximately -255°F
[-159°C], whereupon it is then directed to LNG storage tank 61 where the flash vapor
resulting from expansion (stream 83) is separated from the LNG product (stream 84).
Tower bottoms stream 77 from LNG purification tower 56 is flash
expanded to slightly above the operating pressure of fractionation tower 17 by
expansion valve 57. During expansion a portion of the stream is vaporized, resulting
in cooling of the total stream from -141°F [-96°C] to -156°F [-105°C] (stream 77a).
The flash expanded stream 77a is then combined with warmed flash liquid stream 79b
leaving heat exchanger 58 at -155°F [-104°C] to form a combined flash liquid stream
(stream 80) at -156°F [-105°C] which is supplied to heat exchanger 50. It is heated to
-90°F [-68°C] (stream 80a) as it supplies cooling to feed stream 71 and tower
overhead vapor stream 74 as described earlier, and thereafter supplied to fractionation
tower 17 at a lower mid-column feed point If desired, stream 80a can be combined
with flash expanded stream 35a described earlier and the combined stream supplied to
a single lower mid-column feed point on the tower.
The flash vapor (stream 83) from LNG storage tank 61 passes
countercurrently to the incoming liquid in heat exchanger 58 where it is heated to
-155°F [-104°C] (stream 83a). It then enters heat exchanger 50 where it is heated to
115°F [46°C] (stream 83b) as it supplies cooling to feed stream 71 and tower
overhead stream 74. Since this stream is at low pressure (15.5 psia [107 kPa(a)]), it
must be compressed before it can be used as plant fuel gas. Compressors 63 and 65
(driven by supplemental power sources) with intercooler 64 are used to compress the
stream (stream 83e). Following cooling in aftercooler 66, stream 83f at 115 psia [793
kPa(a)] is combined with stream 37 to become the fuel gas for the plant (stream 85).
The cold distillation vapor stream from the NGL recovery plant
(stream 86) is heated to 115°F [46°C] as it supplies cooling to feed stream 71 in heat
exchanger 50, becoming the second residue gas (stream 86a) which is then
re-compressed in compressor 62 driven by a supplemental power source. The
compressed second residue gas (stream 86b) combines with the compressed firs;t
residue gas (stream 43b) to form third residue gas stream 44. After cooling to 120°F
[49°C] in discharge cooler 20, third residue gas stream 44a is divided into two
portions. One portion (stream 71) becomes the feed stream to the LNG production
section. The other portion (stream 38) becomes the residue gas product, which flows
to the sales gas pipeline at 740 psia [5,102 kPa(a)].
A summary of stream flow rates and energy consumption for the
process illustrated in FIG. 6 is set forth in the following table:
Comparing the recovery levels displayed in Table VI for the FIG. 6
process to those in Table I for the FIG. 1 process shows that the recoveries in the NGL
recovery plant have been maintained at essentially the same levels for both processes.
The net increase in compression power for the FIG. 6 process compared to the FIG. 1
process is 2,222 HP [3,653 kW] to produce the nominal 50,000 gallons/D [417 m3/D]
of LNG, giving a specific power consumption of 0.303 HP-H/Lb [0.498 kW-H/kg] for
the FIG. 6 process. Thus, the present invention has a specific power consumption that
is lower than both the FIG. 2 and the FIG. 3 prior art processes, with no need for
carbon dioxide removal from the feed gas prior to entering the LNG production
section like the prior art processes do.
This embodiment of the present invention, which uses the residue gas
from the NGL recovery plant as its feed gas, has a lower LNG production efficiency
that the FIG. 4 and FIG. 5 embodiments which process a portion of the NGL recovery
plant feed gas. This lower efficiency is mainly due to a reduction in the efficiency of
the NGL recovery plant as a result of using a portion (stream 86) of the cold
distillation vapor (stream 42) from the NGL recovery plant to supply some of the
process refrigeration in the LNG production section. Although stream 86 is used in a
similar fashion in the FIG. 4 and FIG. 5 embodiments, the NGL recovery plants in
these embodiments are processing a lesser quantity of the inlet gas since one portion
(stream 71 in FIGS. 4 and 5) is fed to the LNG production section rather than to the
NGL recovery plant The loss in NGL recovery plant efficiency is reflected in the
higher utility consumption of first residue gas compressor 19 shown in Table VI for
the FIG. 6 process versus the corresponding values in Table IV and Table V for the
FIG. 4 and FIG. 5 processes, respectively.
For most inlet gases, the plant inlet gas will be the preferred source of
the feed stream for processing according to the present invention, as illustrated in
Examples 1 and 2. In some cases, however, the NGL recovery plant residue gas may
be the better choice as the source of the feed stream as illustrated in Example 3. For
instance, if the inlet gas contains hydrocarbons that may solidify at cold temperatures,
such as heavy paraffins or benzene, the NGL recovery plant can serve as a feed
conditioning unit for the LNG production section by recovering these compounds in
the NGL product. The residue gas leaving the NGL recovery plant will not contain
significant quantities of heavier hydrocarbons, so processing a portion of the plant
residue gas for co-production of LNG using the present invention can be
accomplished in such instances without risk of solids formation in the heat exchangers
in the LNG production and LNG cool-down sections. The decision of which
embodiment of the present invention to use in a particular circumstance may also be
influenced by factors such as inlet gas and residue gas pressure levels, plant size,
available equipment, and the economic balance of capital cost versus operating cost.
Other Embodiments
One skilled in the art will recognize that the present invention can be
adapted for use with all types of NGL recovery plants to allow co-production of LNG.
The examples presented earlier have all depicted the use of the present invention with
an NGL recovery plant employing the process disclosed in United States Patent No.
4,278,457 in order to facilitate comparisons of the present invention with the prior art.
However, the present invention is generally applicable for use with any NGL
recovery process that produces a distillation vapor stream that is at temperatures of
-50°F [-46°C] or colder. Examples of such NGL recovery processes are described and
illustrated in United States Pat Nos. 3,292,380; 4,140,504; 4,157,904; 4,171,964;
4,185,978; 4,251,249; 4,278,457; 4,519,824; 4,617,039; 4,687,499; 4,689,063;
4,690,702; 4,854,955; 4,869,740; 4,889,545; 5,275,005; 5,555,748; 5,568,737;
5,771,712; 5,799,507; 5,881,569; 5,890,378; 5,983,664; 6,182,469; reissue U.S. Pat
No. 33,408; and co-pending application nos. 60/225,260 and 09/677,220, the full
disclosures of which are incorporated by reference herein in their entirety. Further,
the present invention is applicable for use with NGL recovery plants that are designed
to recover only C3 components and heavier hydrocarbon components in the NGL
product (i.e., no significant recovery of C2 components), or with NGL recovery plants
that are designed to recover C2 components and heavier hydrocarbon components in
the NGL product but are being operated to reject the C2 components to the residue gas
so as to recover only C3 components and heavier hydrocarbon components in the NGL
product (i.e., ethane rejection mode of operation). This feedstock flexibility is due to
LNG purification tower 56 shown in FIGS. 4 through 6, which ensures that only
methane (and other volatile gases when present) enters the LNG cool-down section.
In accordance with this invention, the cooling of the feed stream to the
LNG production section may be accomplished in many ways. In the processes of
FIGS. 4 through 6, feed stream 71, expanded stream 72a (for the FIG. 4 process only),
and distillation vapor stream 74 are cooled and condensed by a portion of the
demethanizer overhead vapor (stream 86) along with flash vapor, flash liquid, and
tower liquids produced in the LNG production and LNG cool-down sections.
However, demethanizer liquids (such as stream 39) could be used to supply some or
all of the cooling and condensation of streams 71 and 74 in FIGS. 4 through 6 and/or
stream 72a in FIG. 4, as could the flash expanded stream 73a as shown in FIG. 7.
Further, any stream at a temperature colder man the stream(s) being cooled may be
utilized. For instance, a side draw of vapor from the demethanizer could be
withdrawn and used for cooling. Other potential sources of cooling include, but are
not limited to, flashed high pressure separator liquids and mechanical refrigeration
systems. The selection of a source of cooling will depend on a number of factors
including, but not limited to, feed gas composition and conditions, plant size, heat
exchanger size, potential cooling source temperature, etc. One skilled in the art will
also recognize that any combination of the above cooling sources or methods of
cooling may be employed in combination to achieve the desired feed stream
temperature(s).
In accordance with this invention, external refrigeration may be
employed to supplement the cooling available to the feed gas from other process
streams, particularly in the case of a feed gas richer than that used in Examples 1 and
2. The use and distribution of LNG tower liquids for process heat exchange, and the
particular arrangement of heat exchangers for feed gas cooling, must be evaluated for
each particular application, as well as the choice of process streams for specific heat
exchange services.
It will also be recognized that the relative amount of the feed stream 71
that is directed to the LNG cool-down section (stream 78) and that is withdrawn to
become flash liquid (stream 79) will depend on several factors, including feed gas
pressure, feed gas composition, the amount of heat which can economically be
extracted from the feed, and the quantity of horsepower available. More feed to the
LNG cool-down section may increase LNG production while decreasing the purity of
the LNG (stream 84) because of the corresponding decrease in reflux (stream 75) to
the LNG purification tower. Increasing the amount that is withdrawn to become: flash
liquid reduces the power consumption for flash vapor compression but increases: the
power consumption for compression of the first residue gas by increasing the quantity
of recycle to demethanizer 17 in stream 79. Further, as shown by the dashed lines in
FIGS. 4 through 7, the flash liquid could be eliminated completely from heat
exchanger 58 (at the expense of increasing the quantity of flash vapor in stream 83
and increasing the power consumption for flash vapor compression).
Subcooling of condensed liquid stream 78 in heat exchanger 58
reduces the quantity of flash vapor (stream 83) generated during expansion of the
stream to the operating pressure of LNG storage tank 61. This generally reduces the
specific power consumption for producing the LNG by reducing the power
consumption of flash gas compressors 63 and 65. However, as illustrated in FIG. 8
and by the dashed lines in FIGS. 4 through 7, some circumstances may favor reducing
the capital cost of the facility by eliminating heat exchanger 58 in its entirety. As also
illustrated in FIG. 8 and by the dashed lines in FIGS. 4 through 7, the quantity of
tower bottoms stream 77 may be such that using the flash expanded stream 77a for
heat exchange may not be warranted. In such cases, the flash expanded stream 77a
could be supplied at an appropriate feed location directly to fractionation tower 17 as
shown.
Although individual stream expansion is depicted in particular
expansion devices, alternative expansion means may be employed where appropriate.
For example, conditions may warrant work expansion of the substantially condensed
feed stream (stream 71a in FIGS. S, 6, and 8) or the LNG purification tower bottoms
stream (stream 77 in FIGS. 4 through 8). Further, isenthalpic flash expansion may be
used in lieu of work expansion for subcooled liquid stream 82 in FIGS. 4 through 7 or
condensed liquid stream 78 in FIG. 8 (with the resultant increase in the relative
quantity of flash vapor produced by the expansion, increasing the power consumption
for flash vapor compression), or for vapor stream 72 in FIGS. 4 and 7 (with the
resultant increase in the power consumption for compression of the second residue
gas).
While there have been described what are believed to be preferred
embodiments of the invention, those skilled in the art will recognize that other and
further modifications may be made thereto, e.g. to adapt the invention to various
conditions, types of feed or other requirements without departing from the spirit of the
present invention.
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
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 it;
(f) said condensed expanded stream is directed to a
distillation column at a mid-column feed point;
(g) a liquid distillation stream is withdrawn from a lower
region of said distillation column and directed to said plant;
(h) a vapor distillation stream is withdrawn from an upper
region of said distillation column and cooled under pressure to condense at least a
portion of it and form a condensed stream;
(i) said condensed stream is divided into at least two
portions, with a first portion directed to said distillation column at a top feed position;
(j) a second portion of said condensed stream is expanded
to lower pressure to form said liquefied natural gas stream; and
(k) me temperature of said partially condensed natural gas
stream and the quantities and temperatures of said feed streams to said distillation
column are effective to maintain the overhead temperature of said distillation column
at a temperature whereby the major portion of said heavier hydrocarbon components
is recovered in said liquid stream and said liquid distillation 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;
(e) said liquid stream is expanded to an intermediate
pressure, heated, and thereafter directed to said plant;
(f) said vapor stream is expanded to an intermediate
pressure and further cooled at said intermediate pressure to condense it;
(g) said condensed expanded stream is directed to a
distillation column at a mid-column feed point;
(h) a liquid distillation stream is withdrawn from a lower
region of said distillation column and directed to said plant;
(i) a vapor distillation stream is withdrawn from an upper
region of said distillation column and cooled under pressure to condense at least a
portion of it and form a condensed stream;
(j) said condensed stream is divided into at least two
portions, with a first portion directed to said distillation column at a top feed position;
(k) a second portion of said condensed stream is expanded
to lower pressure to form said liquefied natural gas stream; and
0) the temperature of said partially condensed natural gas
stream and the quantities and temperatures of said feed streams to said distillation
column are effective to maintain the overhead temperature of said distillation column
at a temperature whereby the major portion of said heavier hydrocarbon components
is recovered in said liquid stream and said liquid distillation 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 to
substantially 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 condensed natural gas stream is expanded to an
intermediate pressure and directed to a distillation column at a mid-column feed, point;
(e) a liquid distillation stream is withdrawn from a lower
region of said distillation column and directed to said plant;
(f) a vapor distillation stream is withdrawn from an upper
region of said distillation column and cooled under pressure to condense at least a
portion of it and form a condensed stream;
(g) said condensed stream is divided into at least two
portions, with a first portion directed to said distillation column at a top feed position;
(h) . a second portion of said condensed stream is expanded
to lower pressure to form said liquefied natural gas stream; and
(i) the quantities and temperatures of said feed streams to
said distillation column are effective to maintain the overhead temperature of said
distillation column at a temperature whereby the major portion of said heavier
hydrocarbon components is recovered in said liquid distillation stream.
4. The improvement according to claims 1, 2, or 3 wherein said
second portion of said condensed stream is cooled before being expanded to said.
lower pressure.
5. The improvement according to claim 4 wherein a third portion
of said condensed stream is withdrawn, expanded to an intermediate pressure, and
directed in heat exchange relation with said second portion of said condensed stream
to supply at least a portion of said cooling.
6. The improvement according to claims 1, 2, or 3 wherein said
liquid distillation stream is expanded and heated before being directed to said plant
7. The improvement according to claim 4 wherein said liquid
distillation stream is expanded and heated before being directed to said plant
8. The improvement according to claim 5 wherein said liquid
distillation stream is expanded and heated before being directed to said plant.
9. 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) first 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 first 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 first 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 said first heat exchange means to supply
said expanded vapor stream to said first heat exchange means, with said first heat
exchange means being adapted to further cool said expanded vapor stream at said
intermediate pressure to substantially condense it;
(f) a distillation column connected to said first heat
exchange means to receive said substantially condensed expanded stream at a
mid-column feed point, with said distillation column adapted to withdraw a liquid
distillation stream from a lower region of said distillation column and direct it to said
plant, and to withdraw a vapor distillation stream from an upper region of said
distillation column, said distillation column being further connected to said first heat
exchange means to supply said vapor distillation stream to said first heat exchange
means, with said first heat exchange means being adapted to cool said vapor
distillation stream under pressure, thereby to condense at least a portion of it and form
a condensed stream;
(g) dividing means connected to said first heat exchange
means to receive said condensed stream and divide it into at least two portions, said
dividing means being further connected to said distillation column to direct a first
portion of said condensed stream to said distillation column at a top feed position;
(h) second expansion means connected to said dividing
means to receive a second portion of said condensed stream and expand it to lower
pressure to form said liquefied natural gas stream; and
(i) control means adapted to regulate the temperature of
said partially condensed natural gas stream and the quantities and temperatures of said
feed streams to said distillation column to maintain the overhead temperature of said
distillation column at a temperature whereby the major portion of said heavier
hydrocarbon components is recovered in said liquid stream and said liquid distillation
stream.
10. 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) first 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 first 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 first heat exchange
means to receive said partially condensed natural gas stream and to separate it into a
vapor stream and a liquid stream;
(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 said first heat exchange means to supply
said expanded vapor stream to said first heat exchange means, with said first heat
exchange means being adapted to further cool said expanded vapor stream at said
intermediate pressure to substantially condense it;
(f) a distillation column connected to said first heat
exchange means to receive said substantially condensed expanded stream at a
mid-column feed point, with said distillation column adapted to withdraw a liquid
distillation stream from a lower region of said distillation column and direct it to said
plant, and to withdraw a vapor distillation stream from an upper region of said
distillation column, said distillation column being further connected to said first heat
exchange means to supply said vapor distillation stream to said first heat exchange
means, with said first heat exchange means being adapted to cool said vapor
distillation stream under pressure, thereby to condense at least a portion of it and form
a condensed stream;
(g) dividing means connected to said first heat exchange
means to receive said condensed stream and divide it into at least two portions;, said
dividing means being further connected to said distillation column to direct a first
portion of said condensed stream to said distillation column at a top feed position;
(h) second expansion means connected to said dividing
means to receive a second portion of said condensed stream and expand it to lower
pressure to form said liquefied natural gas stream;
(i) third expansion means connected to said separation
means to receive said liquid stream and expand it to an intermediate pressure, said
third expansion means being further connected to said first heat exchange means to
heat said expanded liquid stream and thereby supply at least a portion of said cooling,
with said expanded heated liquid stream thereafter directed to said plant; and
(j) control means adapted to regulate the temperature of
said partially condensed natural gas stream and the quantities and temperatures of said
feed streams to said distillation column to maintain the overhead temperature of said
distillation column at a temperature whereby the major portion of said heavier
hydrocarbon components is recovered in said liquid stream and said liquid distillation
stream.
11. 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) first heat exchange means connected to said first
withdrawing means to receive said natural gas stream and cool it under pressure to
substantially condense it;
(c) second withdrawing means connected to said plant to
withdraw a distillation stream, said second withdrawing means being further
connected to said first 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 first heat
exchange means to receive said substantially condensed stream and expand it to an
intermediate pressure;
(e) a distillation column connected to said first expansion
means to receive said expanded stream at a mid-column feed point, with said
distillation column adapted to withdraw a liquid distillation stream from a lower
region of said distillation column and direct it to said plant, and to withdraw a vapor
distillation stream from an upper region of said distillation column, said distillation
column being further connected to said first heat exchange means to supply said vapor
distillation stream to said first heat exchange means, with said first heat exchange
means being adapted to cool said vapor distillation stream under pressure, thereby to
condense at least a portion of it and form a condensed stream;
(f) dividing means connected to said first heat exchange
means to receive said condensed stream and divide it into at least two portions, said
dividing means being further connected to said distillation column to direct a first
portion of said condensed stream to said distillation column at a top feed position;
(g) second expansion means connected to said dividing
means to receive a second portion of said condensed stream and expand it to lower
pressure to form said liquefied natural gas stream; and
(h) control means adapted to regulate the quantities and
temperatures of said feed streams to said distillation column to maintain the overhead
temperature of said distillation column at a temperature whereby the major portion of
said heavier hydrocarbon components is recovered in said liquid distillation stream.
12. The improvement according to claims 9 or 11 wherein a second
heat exchange means is connected to said dividing means to receive said second
portion of said condensed stream and cool it, said second heat exchange means being
further connected to supply said cooled second portion to said second expansion
means.
13. The improvement according to claim 10 wherein a second heat
exchange means is connected to said dividing means to receive said second portion of
said condensed stream and cool it, said second heat exchange means being further
connected to supply said cooled second portion to said second expansion means.
14. The improvement according to claim 12 wherein a third
withdrawing means is connected to said second heat exchange means to withdraw a
third portion of said condensed stream from said cooled second portion, said third
withdrawing means being further connected to supply said third portion to a third
expansion means and expand it to an intermediate pressure, said third expansion
means being further connected to supply said expanded third portion to said second
heat exchange means to supply at least a portion of said cooling.
15. The improvement according to claim 13 wherein a third
withdrawing means is connected to said second heat exchange means to withdraw a
third portion of said condensed stream from said cooled second portion, said third
withdrawing means being further connected to supply said third portion to a fourth
expansion means and expand it to an intermediate pressure, said fourth expansion
means being further connected to supply said expanded third portion to said second
heat exchange means to supply at least a portion of said cooling.
16. The improvement according to claims 9 or 11 wherein a third
expansion means is connected to said distillation column to receive said liquid
distillation stream and expand it, said third expansion means being further connected
to said first heat exchange means to heat said expanded liquid distillation stream and
thereby supply at least a portion of said cooling, with said expanded heated liquid
distillation stream thereafter directed to said plant
17. The improvement according to claim 10 wherein a fourth
expansion means is connected to said distillation column to receive said liquid
distillation stream and expand it, said fourth expansion means being further connected
to supply said expanded liquid distillation stream to said first heat exchange means to
heat said expanded liquid distillation stream and thereby supply at least a portion of
said cooling, with said expanded heated liquid distillation stream thereafter directed to
said plant.
18. The improvement according to claim 12 wherein a third
expansion means is connected to said distillation column to receive said liquid
distillation stream and expand it, said third expansion means being further connected
to supply said expanded liquid distillation stream to said first heat exchange means to
heat said expanded liquid distillation stream and thereby supply at least a portion of
said cooling, with said expanded heated liquid distillation stream thereafter directed to
said plant.
19. The improvement according to claim 13 wherein a fourth
expansion means is connected to said distillation column to receive said liquid
distillation stream and expand it, said fourth expansion means being further connected
to supply said expanded liquid distillation stream to said first heat exchange means to
heat said expanded liquid distillation stream and thereby supply at least a portion of
said cooling, with said expanded heated liquid distillation stream thereafter directed to
said plant
20. The improvement according to claim 14 wherein a fourth
expansion means is connected to said distillation column to receive said liquid
distillation stream and expand it, said fourth expansion means being further connected
to supply said expanded liquid distillation stream to said first heat exchange means to
heat said expanded liquid distillation stream and thereby supply at least a portion of
said cooling, with said expanded heated liquid distillation stream thereafter directed to
said plant
21. The improvement according to claim 15 wherein a fifth
expansion means is connected to said distillation column to receive said liquid
distillation stream and expand it, said fifth expansion means being further connected
to supply said expanded liquid distillation stream to said first heat exchange means to
heat said expanded liquid distillation stream and thereby supply at least a portion of
said cooling, with said expanded heated liquid distillation stream thereafter directed to
said plant.

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
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 it;
(f) said condensed expanded stream is directed to a
distillation column at a mid-column feed point;
(g) a liquid distillation stream is withdrawn from a lower
region of said distillation column and directed to said plant;
(h) a vapor distillation stream is withdrawn from an upper
region of said distillation column and cooled under pressure to condense at least a
portion of it and form a condensed stream;
(i) said condensed stream is divided into at least two
portions, with a first portion directed to said distillation column at a top feed position;
(j) a second portion of said condensed stream is expanded
to lower pressure to form said liquefied natural gas stream; and
(k) me temperature of said partially condensed natural gas
stream and the quantities and temperatures of said feed streams to said distillation
column are effective to maintain the overhead temperature of said distillation column
at a temperature whereby the major portion of said heavier hydrocarbon components

is recovered in said liquid stream and said liquid distillation stream.

Documents:

1321-KOLNP-2003-ABSTRACT 1.2.pdf

1321-KOLNP-2003-ABSTRACT-1.1.pdf

1321-kolnp-2003-assignment.pdf

1321-kolnp-2003-assignment1.1.pdf

1321-KOLNP-2003-CANCELLED PAGES.pdf

1321-KOLNP-2003-CLAIMS-1.1.pdf

1321-kolnp-2003-claims.pdf

1321-kolnp-2003-correspondence.pdf

1321-kolnp-2003-correspondence1.1.pdf

1321-kolnp-2003-description (complete).pdf

1321-kolnp-2003-drawings.pdf

1321-KOLNP-2003-EXAMINATION REPORT REPLY RECIEVED 1.2.pdf

1321-kolnp-2003-examination report.pdf

1321-KOLNP-2003-FORM 1.1.pdf

1321-kolnp-2003-form 1.2.pdf

1321-kolnp-2003-form 1.pdf

1321-kolnp-2003-form 13.1.pdf

1321-kolnp-2003-form 13.pdf

1321-kolnp-2003-form 18.1.pdf

1321-kolnp-2003-form 18.pdf

1321-KOLNP-2003-FORM 2-1.1.pdf

1321-KOLNP-2003-FORM 2.2.pdf

1321-kolnp-2003-form 2.pdf

1321-kolnp-2003-form 26.1.pdf

1321-kolnp-2003-form 26.pdf

1321-KOLNP-2003-FORM 27.pdf

1321-kolnp-2003-form 3.1.pdf

1321-kolnp-2003-form 3.pdf

1321-KOLNP-2003-FORM 5-1.1.pdf

1321-kolnp-2003-form 5.1.pdf

1321-kolnp-2003-form 5.pdf

1321-kolnp-2003-form 6.pdf

1321-KOLNP-2003-FORM-27-1.pdf

1321-KOLNP-2003-FORM-27.pdf

1321-kolnp-2003-granted-abstract.pdf

1321-kolnp-2003-granted-claims.pdf

1321-kolnp-2003-granted-description (complete).pdf

1321-kolnp-2003-granted-drawings.pdf

1321-kolnp-2003-granted-form 1.pdf

1321-kolnp-2003-granted-form 2.pdf

1321-kolnp-2003-granted-letter patent.pdf

1321-kolnp-2003-granted-specification.pdf

1321-KOLNP-2003-OTHERS DOCUMENTS1.2.pdf

1321-KOLNP-2003-OTHERS-1.1.pdf

1321-kolnp-2003-others.pdf

1321-KOLNP-2003-PA.pdf

1321-KOLNP-2003-PETITION UNDER RULE 137.pdf

1321-KOLNP-2003-REPLY TO EXAMINATION REPORT.pdf

1321-kolnp-2003-reply to examination report1.1.pdf

1321-kolnp-2003-specification.pdf


Patent Number 244172
Indian Patent Application Number 1321/KOLNP/2003
PG Journal Number 48/2010
Publication Date 26-Nov-2010
Grant Date 22-Nov-2010
Date of Filing 14-Oct-2003
Name of Patentee ORTLOFF ENGINEERS LTD.
Applicant Address WELLINGTON CENTRE, SUITE 1000,14643 DALLAS PARKWAY, DALLAS TX 75240-8871
Inventors:
# Inventor's Name Inventor's Address
1 WILKINSON JOHN D 2800 W.DENGAR MIDLAND, TX 79705
2 CUELLAR KYLET 1611 COTTAGE POINT KATY, TX 77494
3 CAMPBELL ROY E 1600 WEST CUTHBERT, MIDLAND, TX 70701
4 HUDSON HANK M 1600 WEST CUTHBERT, MIDLAND, TX 70701
PCT International Classification Number F25J3/00
PCT International Application Number PCT/US2002/11793
PCT International Filing date 2002-04-15
PCT Conventions:
# PCT Application Number Date of Convention Priority Country
1 09/839,907 2001-04-20 U.S.A.