Note: Descriptions are shown in the official language in which they were submitted.
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NATURAL GAS LIQUEFACTION
SPECIFICATION
BACKGROUND OF THE INVENTION
This invention relates to a process for processing natural gas or other
methane-rich gas streams to produce a liquefied natural gas (LNG) stream that
has a high
methane purity and a liquid stream containing predominantly hydrocarbons
heavier than
methane.
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
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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 while producing as a co-product a liquid stream consisting
primarily of
hydrocarbons heavier than methane, such as natural gas liquids (NGL) composed
of
ethane, propane, butanes, and heavier hydrocarbon components, liquefied
petroleum gas
(LPG) composed of propane, butanes, and heavier hydrocarbon components, or
condensate composed of butanes and heavier hydrocarbon components. Producing
the
co-product liquid stream has two important benefits: the LNG produced has a
high
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methane purity, and the co-product liquid is a valuable product that may be
used for
many other purposes. A typical analysis of a natural gas stream to be
processed in
accordance with this invention would be, in approximate mole percent, 84.2%
methane,
7.9% ethane and other C2 components, 4.9% propane and other C3 components,
1.0%
iso-butane, 1.1% normal butane, 0.8% 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 and Kikkawa, Yoshitsugi, Masaaki Ohishi, and Noriyoshi
Nozawa,
"Optimize the Power System of Baseload LNG Plant", Proceedings of the
Eightieth
Annual Convention of the Gas Processors Association, San Antonio, Texas,
March 12-14, 2001 for surveys of a number of such processes. U.S. Pat. Nos.
4,445,917;
4,525,185; 4,545,795; 4,755,200; 5,291,736; 5,363,655; 5,365,740; 5,600,969;
5,615,561;
5,651,269; 5,755,114; 5,893,274; 6,014,869; 6,062,041; 6,119,479; 6,125,653;
6,250,105 B1; 6,269,655 B1; 6,272,882 B1; 6,308,531 B1; 6,324,867 B1; and
6,347,532 B 1 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
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refrigerants having successively lower boiling points, such as propane,
ethane, and
methane. As an alternative, this heat exchange can be accomplished using a
single
refrigerant by evaporating the refrigerant at several different pressure
levels.
"Multi-component refrigeration" employs heat exchange of the natural gas with
one or
more refrigerant fluids composed of several refrigerant components in lieu of
multiple
single-component refrigerants. Expansion of the natural gas can be
accomplished both
isenthalpically (using Joule-Thomson expansion, for instance) and
isentropically (using a
work-expansion turbine, for instance).
Regardless of the method used to liquefy the natural gas stream, it is
common to require removal of a significant fraction of the hydrocarbons
heavier than
methane before the methane-rich stream is liquefied. The reasons for this
hydrocarbon
removal step are numerous, including the need to control the heating value of
the LNG
stream, and the value of these heavier hydrocarbon components as products in
their own
right. Unfortunately, little attention has been focused heretofore on the
efficiency of the
hydrocarbon removal step.
In accordance with the present invention, it has been found that careful
integration of the hydrocarbon removal step into the LNG liquefaction process
can
produce both LNG and a separate heavier hydrocarbon liquid product using
significantly
less energy than prior art processes. The present invention, although
applicable at lower
pressures, is particularly advantageous when processing feed gases in the
range of 400 to
1500 psia [2,758 to 10,342 kPa(a)] or higher.
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For a better understanding of the present invention, reference is made to
the following exam a and drawings. Referring to the drawings:
FIG. 1 is a flow diagram of a natural gas liquefaction plant adapted for
co-production of NGL in accordance with the present invention;
FIG. 2 is a pressure-enthalpy phase diagram for methane used to illustrate
the advantages of the present invention over prior art processes;
FIG. 3 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of NGL in accordance with the present invention;
FIG. 4 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of LPG in accordance with the present invention;
FIG. 5 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of condensate in accordance with the present
invention;
FIG. 6 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention;
FIG. 7 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention;
FIG. 8 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention;
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FIG. 9 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention;
FIG. 10 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention;
FIG. 11 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention;
FIG. 12 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention;
FIG. 13 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention;
FIG. 14 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention;
FIG. 15 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention;
FIG. 16 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention;
FIG. 17 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention;
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FIG. 18 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention;
FIG. 19 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention;
FIG. 20 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention;
and
FIG. 21 is a flow diagram of an alternative natural gas liquefaction plant
adapted for co-production of a liquid stream in accordance with the present
invention.
In the following explanation of the above figures, tables are provided
summarizing flow rates calculated for representative process conditions. In
the tables
appearing herein, the values for flow rates (in moles per hour) have been
rounded to the
nearest whole number for convenience. The total stream rates shown in the
tables
include all non-hydrocarbon components and hence are generally larger than the
sum of
the stream flow rates for the hydrocarbon components. Temperatures indicated
are
approximate values rounded to the nearest degree. It should also be noted that
the
process design calculations performed for the purpose of comparing the
processes
depicted in the figures are based on the assumption of no heat leak from (or
to) the
surroundings to (or from) the process. The quality of commercially available
insulating
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materials makes this a very reasonable assumption and one that is typically
made by
those skilled in the art.
For convenience, process parameters are reported in both the traditional
British units and in the units of the International System of Units (SI). The
molar flow
rates given in the tables may be interpreted as either pound moles per hour or
kilogram
moles per hour. The energy consumptions reported as horsepower (HP) and/or
thousand
British Thermal Units per hour (MBTU/Hr) correspond to the stated molar flow
rates in
pound moles per hour. The energy consumptions reported as kilowatts (kW)
correspond
to the stated molar flow rates in kilogram moles per hour. The production
rates reported
as pounds per hour (Lb/Hr) correspond to the stated molar flow rates in pound
moles per
hour. The production rates reported as kilograms per hour (kg/Hr) correspond
to the
stated molar flow rates in kilogram moles per hour.
DESCRIPTION OF THE INVENTION
Example 1
Referring now to FIG. 1, we begin with an illustration of a process in
accordance with the present invention where it is desired to produce an NGL co-
product
containing the majority of the ethane and heavier components in the natural
gas feed
stream. In this simulation of the present invention, inlet gas enters the
plant at 90 F
[32 C] and 1285 psia [8,860 kPa(a)] as stream 31. If the inlet gas contains a
concentration of carbon dioxide and/or sulfur compounds which would prevent
the
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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
refrigerant streams and demethanizer side reboiler liquids at -68 F [-55 C]
(stream 40).
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 -30 F [-34 C] and 1278 psia [8,812 kPa(a)] where
the vapor
(stream 32) is separated from the condensed liquid (stream 33).
The vapor (stream 32) from separator 11 is divided into two streams, 34
and 36. Stream 34, containing about 20% of the total vapor, is combined with
the
condensed liquid, stream 33, to form stream 35. Combined stream 35 passes
through
heat exchanger 13 in heat exchange relation with refrigerant stream 71 e,
resulting in
cooling and substantial condensation of stream 35a. The substantially
condensed stream
35a at -120 F [-85 C] is then flash expanded through an appropriate expansion
device,
such as expansion valve 14, to the operating pressure (approximately 465 psia
[3,206 kPa(a)]) of fractionation tower 19. During expansion a portion of the
stream is
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vaporized, resulting in cooling of the total stream. In the process
illustrated in FIG. 1, the
expanded stream 35b leaving expansion valve 14-reaches a temperature of -122 F
[-86 C], and is supplied at a mid-point feed position in demethanizing section
19b of
fractionation. tower 19.
The remaining 80% of the vapor from separator 11 (stream 36) enters a
work expansion machine 15 in which mechanical energy is extracted from this
portion of
the high pressure feed. The machine 15 expands the vapor substantially
isentropically
from a pressure of about 1278 psia [8,812 kPa(a)] to the tower operating
pressure, with
the work expansion cooling the expanded stream 36a to a temperature of
approximately
-103 F [-75 C]. The typical commercially available expanders are capable of
recovering
on the order of 80-85% of the work theoretically available in an ideal
isentropic
expansion. The work recovered is often used to drive a centrifugal compressor
(such as
item 16) that can be used to re-compress the tower overhead gas (stream 38),
for
example. The expanded and partially condensed stream 36a is supplied as feed
to
distillation column 19 at a lower mid-column feed point.
The demethanizer in fractionation tower 19 is a conventional distillation
column containing a plurality of vertically spaced trays, one or more packed
beds, or
some combination of trays and packing. As is often the case in natural gas
processing
plants, the fractionation tower may consist of two sections. The upper section
19a is a
separator wherein the top feed is divided into its respective vapor and liquid
portions, and
wherein the vapor rising from the lower distillation or demethanizing section
19b is
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combined with the vapor portion (if any) of the top feed to form the cold
demethanizer
overhead vapor (stream 37) which exits the top of the tower at -135 F [-93 C].
The
lower, demethanizing section 19b 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 one or more reboilers (such as
reboiler 20)
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 115 F [46 C], based on a typical specification of a
methane to
ethane ratio of 0.020:1 on a molar basis in the bottom product.
The demethanizer overhead vapor (stream 37) is warmed to 90 F [32 C]
in heat exchanger 24, and a portion of the warmed demethanizer overhead vapor
is
withdrawn to serve as fuel gas (stream 48) for the plant. (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 refrigerant
compressors 64, 66,
and 68 in this example.) The remainder of the warmed demethanizer overhead
vapor
(stream 38) is compressed by compressor 16 driven by expansion machines 15,
61, and
63. After cooling to 100 F [38 C] in discharge cooler 25, stream 38b is
further cooled to
-123 F [-86 C] in heat exchanger 24 by cross exchange with the cold
demethanizer
overhead vapor, stream 37.
Stream 38c then enters heat exchanger 60 and is further cooled by
refrigerant stream 71d. After cooling to an intermediate temperature, stream
38c is
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divided into two portions. The first portion, stream 49, is further cooled in
heat
exchanger 60 to -257 F [-160 C] to condense and subcool it, whereupon it
enters a work
expansion machine 61 in which mechanical energy is extracted from the stream.
The
machine 61 expands liquid stream 49 substantially isentropically from a
pressure of about
562 psia [3,878 kPa(a)] to the LNG storage pressure (15.5 psia [107 kPa(a)]),
slightly
above atmospheric pressure. The work expansion cools the expanded stream 49a
to a
temperature of approximately -258 F [-161 C], whereupon it is then directed
to the LNG
storage tank 62 which holds the LNG product (stream 50).
Stream 39, the other portion of stream 38c, is withdrawn from heat
exchanger 60 at -160 F [-107 C] and flash expanded through an appropriate
expansion
device, such as expansion valve 17, to the operating pressure of fractionation
tower 19.
In the process illustrated in FIG. 1, there is no vaporization in expanded
stream 39a, so its
temperature drops only slightly to -161 F [-107 C] leaving expansion valve 17.
The
expanded stream 39a is then supplied to separator section 19a in the upper
region of
fractionation tower 19. The liquids separated therein become the top feed to
demethanizing section 19b.
All of the cooling for streams 35 and 38c is provided by a closed cycle
refrigeration loop. The working fluid for this cycle is a mixture of
hydrocarbons and
nitrogen, with the composition of the mixture adjusted as needed to provide
the required
refrigerant temperature while condensing at a reasonable pressure using the
available
cooling medium- In this case, condensing with cooling water has been assumed,
so a
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refrigerant mixture composed of nitrogen, methane, ethane, propane, and
heavier
hydrocarbons is used in the simulation of the FIG. 1 process. The composition
of the
stream, in approximate mole percent, is 7.5% nitrogen, 41.0% methane, 41.5%
ethane,
and 10.0% propane, with the balance made up of heavier hydrocarbons.
The refrigerant stream 71 leaves discharge cooler 69 at 100 F [38 C] and
607 psia [4,185 kPa(a)]. It enters heat exchanger 10 and is cooled to -31 F [-
35 C] and
partially condensed by the partially warmed expanded refrigerant stream 71f
and by other
refrigerant streams. For the FIG. 1 simulation, it has been assumed that these
other
refrigerant streams are commercial-quality propane refrigerant at three
different
temperature and pressure levels. The partially condensed refrigerant stream
71a then
enters heat exchanger 13 for further cooling to -114 F [-81 C] by partially
warmed
expanded refrigerant stream 71e, condensing and partially subcooling the
refrigerant
(stream 71b). The refrigerant is further subcooled to -257 F [-160 C] in heat
exchanger
60 by expanded refrigerant stream 71d. The subcooled liquid stream 71c enters
a work
expansion machine 63 in which mechanical energy is extracted from the stream
as it is
expanded substantially isentropically from a pressure of about 586 psia [4,040
kPa(a)] to
about 34 psia [234 kPa(a)]. During expansion a portion of the stream is
vaporized,
resulting in cooling of the total stream to -263 F [-164 C] (stream 71d). The
expanded
stream 71d then reenters heat exchangers 60, 13, and 10 where it provides
cooling to
stream 38c, stream 35, and the refrigerant (streams 71, 71a, and 71b) as it is
vaporized
and superheated.
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The superheated refrigerant vapor (stream 71 g) leaves heat exchanger 10
at 93 F [34 C] and is compressed in three stages to 617 psia [4,254 kPa(a)].
Each of the
three compression stages (refrigerant compressors 64, 66, and 68) is driven by
a
supplemental power source and is followed by a cooler (discharge coolers 65,
67, and 69)
to remove the heat of compression. The compressed stream 71 from discharge
cooler 69
returns to heat exchanger 10 to complete the cycle.
A summary of stream flow rates and energy consumption for the process
illustrated in FIG. 1 is set forth in the following table:
Table I
(FIG. 1)
Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]
Stream Methane Ethane Propane Butanes+ Total
31 40,977 3,861 2,408 1,404 48,656
32 32,360 2,675 1,469 701 37,209
33 8,617 1,186 939 703 11,447
34 6,472 535 294 140 7,442
36 25,888 2,140 1,175 561 29,767
37 47,771 223 0 0 48,000
39 6,867 32 0 0 6,900
41 73 3,670 2,408 1,404 7,556
48 3,168 15 0 0 3,184
50 37,736 176 0 0 37,916
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Recoveries in NGL*
Ethane 95.06%
Propane 100.00%
Butanes+ 100.00%
Production Rate 308,147 Lb/Hr [ 308,147 kg/Hr]
LNG Product
Production Rate 610,813 Lb/Hr [ 610,813 kg/Hr]
Purity* 99.52%
Lower Heating Value 912.3 BTU/SCF [ 33.99 MJ/m3]
Power
Refrigerant Compression 103,957 HP [ 170,904 kW]
Propane Compression 33,815 HP [ 55,591 kW]
Total Compression 137,772 HP [ 226,495 kW]
Utility Heat
Demethanizer Reboiler 29,364 MBTU/Hr [ 18,969 kW]
* (Based on un-rounded flow rates)
The efficiency of LNG production processes is typically compared using
the "specific power consumption" required, which is the ratio of the total
refrigeration
compression power to the total liquid production rate. Published information
on the
specific power consumption for prior art processes for producing LNG indicates
a range
of 0.168 HP-Hr/Lb [0.276 kW-Hr/kg] to 0.182 HP-Hr/Lb [0.300 kW-Hr/kg], which
is
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believed to be based on an on-stream factor of 340 days per year for the LNG
production
plant. On this same basis, the specific power consumption for the FIG. 1
embodiment of
the present invention is 0.161 HP-Hr/Lb [0.265 kW-Hr/kg], which gives an
efficiency
improvement of 4-13% over the prior art processes. Further, it should be noted
that the
specific power consumption for the prior art processes is based on co-
producing only an
LPG (C3 and heavier hydrocarbons) or condensate (C4 and heavier hydrocarbons)
liquid
stream at relatively low recovery levels, not an NGL (C2 and heavier
hydrocarbons)
liquid stream as shown for this example of the present invention. The prior
art processes
require considerably more refrigeration power to co-produce an NGL stream
instead of
an LPG stream or a condensate stream.
There are two primary factors that account for the improved efficiency of
the present invention. The first factor can be understood by examining the
thermodynamics of the liquefaction process when applied to a high pressure gas
stream
such as that considered in this example. Since the primary constituent of this
stream is
methane, the thermodynamic properties of methane can be used for the purposes
of
comparing the liquefaction cycle employed in the prior art processes versus
the cycle
used in the present invention. FIG. 2 contains a pressure-enthalpy phase
diagram for
methane. In most of the prior art liquefaction cycles, all cooling of the gas
stream is
accomplished while the stream is at high pressure (path A-B), whereupon the
stream is
then expanded (path B-C) to the pressure of the LNG storage vessel (slightly
above
atmospheric pressure). This expansion step may employ a work expansion
machine,
which is typically capable of recovering on the order of 75-80% of the work
theoretically
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available in an ideal isentropic expansion. In the interest of simplicity,
fully isentropic
expansion is displayed in FIG. 2 for path B-C. :Even so, the enthalpy
reduction provided
by this work expansion is quite small, because the lines of constant entropy
are nearly
vertical in the liquid region of the phase diagram.
Contrast this now with the liquefaction cycle of the present invention.
After partial cooling at high pressure (path A-A'), the gas stream is work
expanded (path
A'-A") to an intermediate pressure. (Again, fully isentropic expansion is
displayed in the
interest of simplicity.) The remainder of the cooling is accomplished at the
intermediate
pressure (path A"-B'), and the stream is then expanded (path B'-C) to the
pressure of the
LNG storage vessel. Since the lines of constant entropy slope less steeply in
the vapor
region of the phase diagram, a significantly larger enthalpy reduction is
provided by the
first work expansion step (path A'-A") of the present invention. Thus, the
total amount of
cooling required for the present invention (the sum of paths A-A' an d A"-B')
is less than
the cooling required for the prior art processes (path A-B), reducing the
refrigeration (and
hence the refrigeration compression) required to liquefy the gas stream.
The second factor accounting for the improved efficiency of the present
invention is the superior performance of hydrocarbon distillation systems at
lower
operating pressures. The hydrocarbon removal step in most of the prior art
processes is
performed at high pressure, typically using a scrub column that employs a cold
hydrocarbon liquid as the absorbent stream to remove the heavier hydrocarbons
from the
incoming gas stream. Operating the scrub column at high pressure is not very
efficient,
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as it results in the co-absorption of a significant fraction of the methane
and ethane from
the gas stream, which must subsequently be stripped from the absorbent liquid
and cooled
to become part of the LNG product. In the present invention, the hydrocarbon
removal
step is conducted at the intermediate pressure where the vapor-liquid
equilibrium is much
more favorable, resulting in very efficient recovery of the desired heavier
hydrocarbons
in the co-product liquid stream.
Example 2
If the specifications for the LNG product will allow more of the ethane
contained in the feed gas to be recovered in the LNG product, a simpler
embodiment of
the present invention may be employed. FIG. 3 illustrates such an alternative
embodiment. The inlet gas composition and conditions considered in the process
presented in FIG. 3 are the same as those in FIG. 1. Accordingly, the FIG. 3
process can
be compared to the embodiment displayed in FIG. 1.
In the simulation of the FIG. 3 process, the inlet gas cooling, separation,
and expansion scheme for the NGL recovery section is essentially the same as
that used
in FIG. 1. Inlet gas enters the plant at 90 F [32 C] and 1285 psia [8,860
kPa(a)] as
stream 31 and is cooled in heat exchanger 10 by heat exchange with refrigerant
streams
and demethanizer side reboiler liquids at -35 F [-37 C] (stream 40). The
cooled stream
31a enters separator 11 at -30 F [-34 C] and 1278 psia [8,812 kPa(a)] where
the vapor
(stream 32) is separated from the condensed liquid (stream 33).
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S The vapor (stream 32) from separator 11 is divided into two streams, 34
and 36. Stream 34, containing about 20% of the total vapor, is combined with
the
condensed liquid, stream 33, to form stream 35. Combined stream 35 passes
through
heat exchanger 13 in heat exchange relation with refrigerant stream 71 e,
resulting in
cooling and substantial condensation of stream 35a. The substantially
condensed stream
35a at -120 F [-85 C] is then flash expanded through an appropriate expansion
device,
such as expansion valve 14, to the operating pressure (approximately 465 psia
[3,206 kPa(a)]) of fractionation tower 19. During expansion a portion of the
stream is
vaporized, resulting in cooling of the total stream. In the process
illustrated in FIG. 3, the
expanded stream 35b leaving expansion valve 14 reaches a temperature of -122 F
[-86 C], and is supplied to the separator section in the upper region of
fractionation tower
19. The liquids separated therein become the top feed to the demethanizing
section in the
lower region of fractionation tower 19.
The remaining 80% of the vapor from separator 11 (stream 36) enters a
work expansion machine 15 in which mechanical energy is extracted from this
portion of
the high pressure feed. The machine 15 expands the vapor substantially
isentropically
from a pressure of about 1278 psia [8,812 kPa(a)] to the tower operating
pressure, with
the work expansion cooling the expanded stream 36a to a temperature of
approximately
-103 F [-75 C]. The expanded and partially condensed stream 36a is supplied as
feed to
distillation column 19 at a mid-column feed point.
-19-
CA 02746624 2011-07-15
.1'
The cold demethanizer overhead vapor (stream 37) exits the top of
fractionation tower 19 at -123 F [-86 C]. The liquid product stream 41 exits
the bottom
of the tower at 118 F [48 C], based on a typical specification of a methane to
ethane ratio
of 0.020:1 on a molar basis in the bottom product.
The demethanizer overhead vapor (stream 37) is warmed to 90 F [32 C]
in heat exchanger 24, and a portion (stream 48) is then withdrawn to serve as
fuel gas for
the plant. The remainder of the warmed demethanizer overhead vapor (stream 49)
is
compressed by compressor 16. After cooling to 100 F [38 C] in discharge cooler
25,
stream 49b is further cooled to -112 F [-80 C] in heat exchanger 24 by cross
exchange
with the cold demethanizer overhead vapor, stream 37.
Stream 49c then enters heat exchanger 60 and is further cooled by
refrigerant stream 71d to -257 F [-160 C] to condense and subcool it,
whereupon it
enters a work expansion machine 61 in which mechanical energy is extracted
from the
stream. The machine 61 expands liquid stream 49d substantially isentropically
from a
pressure of about 583 psia [4,021 kPa(a)] to the LNG storage pressure (15.5
psia
[107 kPa(a)]), slightly above atmospheric pressure. The work expansion cools
the
expanded stream 49e to a temperature of approximately -258 F [-161 C],
whereupon it is
then directed to the LNG storage tank 62 which holds the LNG product (stream
50).
Similar to the FIG. I process, all of the cooling for streams 35 and 49c is
provided by a closed cycle refrigeration loop. The composition of the stream
used as the
working fluid in the cycle for the FIG. 3 process, in approximate mole
percent, is 7.5%
-20-
CA 02746624 2011-07-15
nitrogen, 40.0% methane, 42.5% ethane, and 10.0% propane, with the balance
made up
of heavier hydrocarbons. The refrigerant stream 71 leaves discharge cooler 69
at 100 F
[38 C] and 607 psia [4,185 kPa(a)]. It enters heat exchanger 10 and is cooled
to -31 F
[-35 C] and partially condensed by the partially warmed expanded refrigerant
stream 71f
and by other refrigerant streams. For the FIG. 3 simulation, it has been
assumed that
these other refrigerant streams are commercial-quality propane refrigerant at
three
different temperature and pressure levels. The partially condensed refrigerant
stream 71a
then enters heat exchanger 13 for further cooling to -121 F [-85 C] by
partially warmed
expanded refrigerant stream 71 e, condensing and partially subcooling the
refrigerant
(stream 71b). The refrigerant is further subcooled to -257 F [-160 C] in heat
exchanger
60 by expanded refrigerant stream 71d. The subcooled liquid stream 71c enters
a work
expansion machine 63 in which mechanical energy is extracted from the stream
as it is
expanded substantially isentropically from a pressure of about 586 psia [4,040
kPa(a)] to
about 34 psia [234 kPa(a)]. During expansion a portion of the stream is
vaporized,
resulting in cooling of the total stream to -263 F [-164 C] (stream 71d). The
expanded
stream 71d then reenters heat exchangers 60, 13, and 10 where it provides
cooling to
stream 49c, stream 35, and the refrigerant (streams 71, 71a, and 71b) as it is
vaporized
and superheated.
The superheated refrigerant vapor (stream 71g) leaves heat exchanger 10
at 93 F [34 C] and is compressed in three stages to 617 psia [4,254 kPa(a)].
Each of the
three compression stages (refrigerant compressors 64, 66, and 68) is driven by
a
supplemental power source and is followed by a cooler (discharge coolers 65,
67, and 69)
-21-
CA 02746624 2011-07-15
to remove the heat of compression. The compressed stream 71 from discharge
cooler 69
returns to heat exchanger 10 to complete the cycle.
A summary of stream flow rates and energy consumption for the process
illustrated in FIG. 3 is set forth in the following table:
Table H
(FIG. 3)
Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]
Stream Methane Ethane Pro amine Butanes+ Total
31 40,977 3,861 2,408 - 1,404 48,656
32 32,360 2,675 1,469 701 37,209
33 8,617 1,186 939 703 11,447
34 6,472 535 294 140 7,442
36 25,888 2,140 1,175 561 29,767
37 40,910 480 62 7 41,465
41 67 3,381 2,346 1,397 7,191
48 2,969 35 4 0 3,009
50 37,941 445 58 7 38,456
-22-
i
CA 02746624 2011-07-15
Recoveries in NGL*
Ethane 87.57%
Propane 97.41%
Butanes+ 99.47%
Production Rate 296,175 Lb/Hr [ 296,175 kg/Hr]
LNG Product
Production Rate 625,152 Lb/Hr [ 625,152 kg/Hr]
Purity* 98.66%
Lower Heating Value 919.7 BTU/SCF [ 34.27 MJ/m3]
Power -
Refrigerant Compression 96,560 HP [ 158,743 kW]
Propane Compression 34,724 HP [ 57,086 kW]
Total Compression 131,284 HP [ 215,829 kW]
Utility Heat
Demethanizer Reboiler 22,177 MBTU/Hr [ 14,326 kW]
* (Based on un-rounded flow rates)
Assuming an on-stream factor of 340 days per year for the LNG
production plant, the specific power consumption for the FIG. 3 embodiment of
the
present invention is 0.153 HP-Hr/Lb [0.251 kW-Hr/kg]. Compared to the prior
art
processes, the efficiency improvement is 10-20% for the FIG. 3 embodiment. As
noted
earlier for the FIG. I embodiment, this efficiency improvement is possible
with the
-23-
CA 02746624 2011-07-15
present invention even though an NGL co-product is produced rather than the
LPG or
condensate co-product produced by the prior art processes.
Compared to the FIG. I embodiment, the FIG. 3 embodiment of the
present invention requires about 5% less power per unit of liquid produced.
Thus, for a
given amount of available compression power, the FIG. 3 embodiment could
liquefy
about 5% more natural gas than the FIG. I embodiment by virtue of recovering
less of
the C2 and heavier hydrocarbons in the NGL co-product. The choice between the
FIG. I
and the FIG. 3 embodiments of the present invention for a particular
application will
generally be dictated either by the monetary value of the heavier hydrocarbons
in the
NGL product versus their corresponding value in the LNG product, or by the
heating
value specification for the LNG product (since the heating value of the LNG
produced by
the FIG. 1 embodiment is lower than that produced by the FIG. 3 embodiment).
Example 3
If the specifications for the LNG product will allow all of the ethane
contained in the feed gas to be recovered in the LNG product, or if there is
no market for
a liquid co-product containing ethane, an alternative embodiment of the
present invention
such as that shown in FIG. 4 may be employed to produce an LPG co-product
stream.
The inlet gas composition and conditions considered in the process presented
in FIG. 4
are the same as those in FIGS. 1 and 3. Accordingly, the FIG. 4 process can be
compared
to the embodiments displayed in FIGS. 1 and 3.
-24-
CA 02746624 2011-07-15
In the simulation of the FIG. 4 process, inlet gas enters the plant at 90 F
[32 C] and 1285 psia [8,860 kPa(a)] as stream 31 and is cooled in heat
exchanger 10 by
heat exchange with refrigerant streams and flashed separator liquids at -46 F
[-43 C]
(stream 33a). The cooled stream 31a enters separator 11 at -1 F [-18 C] and
1278 psia
[8,812 kPa(a)] where the vapor (stream 32) is separated from the condensed
liquid
(stream 33).
The vapor (stream 32) from separator 11 enters work expansion machine
in which mechanical energy is extracted from this portion of the high pressure
feed.
The machine 15 expands the vapor substantially isentropically from a pressure
of about
1278 psia [8,812 kPa(a)] 'to a pressure of about 440 psia [3,034 kPa(a)] (the
operating
15 pressure of separator/absorber tower 18), with the work expansion cooling
the expanded
stream 32a to a temperature of approximately -81 F [-63 C]. The expanded and
partially
condensed stream 32a is supplied to absorbing section 18b in a lower region of
separator/absorber tower 18. The liquid portion of the expanded stream
commingles with
liquids falling downward from the absorbing section and the combined liquid
stream 40
exits the bottom of separator/absorber tower 18 at -86 F [-66 C]. The vapor
portion of
the expanded stream rises upward through the absorbing section and is
contacted with
cold liquid falling downward to condense and absorb the C3 components and
heavier
components.
The separator/absorber tower 18 is a conventional distillation column
containing a plurality of vertically spaced trays, one or more packed beds, or
some
-25-
CA 02746624 2011-07-15
combination of trays and packing. As is often the case in natural gas
processing plants,
the separator/absorber tower may consist of two sections. The upper section
18a is a
separator wherein any vapor contained in the top feed is separated from its
corresponding
liquid portion, and wherein the vapor rising from the lower distillation or
absorbing
section 18b is combined with the vapor portion (if any) of the top feed to
form the cold
distillation stream 37 which exits the top of the tower. The lower, absorbing
section 18b
contains the trays and/or packing and provides the necessary contact between
the liquids
falling downward and the vapors rising upward to condense and absorb the C3
components and heavier components.
The combined liquid stream 40 from the bottom of separator/absorber
tower 18 is routed to heat exchanger 13 by pump 26 where it (stream 40a) is
heated as it
provides cooling of deethanizer overhead (stream 42) and refrigerant (stream
71a). The
combined liquid stream is heated to -24 F [-31 C], partially vaporizing stream
40b before
it is supplied as a mid-column feed to deethanizer 19. The separator liquid
(stream 33) is
flash expanded to slightly above the operating pressure of deethanizer 19 by
expansion
valve 12, cooling stream 33 to -46 F [-43 C] (stream 33a) before it provides
cooling to
.the incoming feed gas as described earlier. Stream 33b, now at 85 F [29 C],
then enters
deethanizer 19 at a lower mid-column feed point. In the deethanizer, streams
40b and
33b are stripped of their methane and C2 components. The deethanizer in tower
19,
operating at about 453 psia [3,123 kPa(a)], is also a conventional
distillation column
containing a plurality of vertically spaced trays, one or more packed beds, or
some
combination of trays and packing. The deethanizer tower may also consist of
two
-26-
CA 02746624 2011-07-15
sections: an upper separator section 19a wherein any vapor contained in the
top feed is
separated from its corresponding liquid portion, and wherein the vapor rising
from the
lower distillation or deethanizing section 19b is combined with the vapor
portion (if any)
of the top feed to form distillation stream 42 which exits the top of the
tower; and a
lower, deethanizing section 19b that contains the trays and/or packing to
provide the
necessary contact between the liquids falling downward and the vapors rising
upward.
The deethanizing section 19b also includes one or more reboilers (such as
reboiler 20)
which heat and vaporize a portion of the liquid at the bottom of the column to
provide the
stripping vapors which flow up the column to strip the liquid product, stream
41, of
methane and C2 components. A typical specification for the bottom liquid
product is to
have an ethane to propane ratio of 0.020:1 on a molar basis. The liquid
product stream
41 exits the bottom of the deethanizer at 214 F [101 C].
The operating pressure in deethanizer 19 is maintained slightly above the
operating pressure of separator/absorber tower 18. This allows the deethanizer
overhead
vapor (stream 42) to pressure flow through heat exchanger 13 and thence into
the upper
section of separator/absorber tower 18. In heat exchanger 13, the deethanizer
overhead at
-19 F [-28 C] is directed in heat exchange relation with the combined liquid
stream
(stream 40a) from the bottom of separator/absorber tower 18 and flashed
refrigerant
stream 71e, cooling the stream to -89 F [-67 C] (stream 42a) and partially
condensing it.
The partially condensed stream enters reflux drum 22 where the condensed
liquid (stream
44) is separated from the uncondensed vapor (stream 43). Stream 43 combines
with the
distillation vapor stream (stream 37) leaving the upper region of
separator/absorber tower
-27-
CA 02746624 2011-07-15
18 to form cold residue gas stream 47. The condensed liquid (stream 44) is
pumped to
higher pressure by pump 23, whereupon stream 44a is divided into two portions.
One
portion, stream 45, is routed to the upper separator section of
separator/absorber tower 18
to serve as the cold liquid that contacts the vapors rising upward through the
absorbing
section. The other portion is supplied to deethanizer 19 as reflux stream 46,
flowing to a
top feed point on deethanizer 19 at -89 F [-67 C].
The cold residue gas (stream 47) is warmed from -94 F [-70 C] to 94 F
[34 C] in heat exchanger 24, and a portion (stream 48) is then withdrawn to
serve as fuel
gas for the plant. The remainder of the warmed residue gas (stream 49) is
compressed by
compressor 16. After cooling to 1009F [38 C] in discharge cooler 25, stream
49b is
further cooled to -78 F [-61 C] in heat exchanger 24 by cross exchange with
the cold
residue gas, stream 47.
Stream 49c then enters heat exchanger 60 and is further cooled by
refrigerant stream 71d to -255 F [-160 C] to condense and subcool it,
whereupon it
enters a work expansion machine 61 in which mechanical energy is extracted
from the
stream. The machine 61 expands liquid stream 49d substantially isentropically
from a
pressure of about 648 psia [4,465 kPa(a)] to the LNG storage pressure (15.5
psia
[107 kPa(a)]), slightly above atmospheric pressure. The work expansion cools
the
expanded stream 49e to a temperature of approximately -256 F [-160 C],
whereupon it is
then directed to the LNG storage tank 62 which holds the LNG product (stream
50).
-28-
CA 02746624 2011-07-15
Similar to the FIG. 1 and FIG. 3 processes, much of the cooling for stream
42 and all of the cooling for stream 49c is provided by a closed cycle
refrigeration loop.
The composition of the stream used as the working fluid in the cycle for the
FIG. 4
process, in approximate mole percent, is 8.7% nitrogen, 30.0% methane, 45.8%
ethane,
and 11.0% propane, with the balance made up of heavier hydrocarbons. The
refrigerant
stream 71 leaves discharge cooler 69 at 100 F [38 C] and 607 psia [4,185
kPa(a)]. It
enters heat exchanger 10 and is cooled to -17 F [-27 C] and partially
condensed by the
partially warmed expanded refrigerant stream 71f and by other refrigerant
streams. For
the FIG. 4 simulation, it has been assumed that these other refrigerant
streams are
commercial-quality propane refrigerant at three different temperature and
pressure levels.
The partially condensed refrigerant stream 71 a then enters heat exchanger 13
for further
cooling to -89 F [-67 C] by partially warmed expanded refrigerant stream 71e,
further
condensing the refrigerant (stream 71b). The refrigerant is totally condensed
and then
subcooled to -255 F [-160 C] in heat exchanger 60 by expanded refrigerant
stream 71d.
The subcooled liquid stream 71c enters a work expansion machine 63 in which
mechanical energy is extracted from the stream as it is expanded substantially
isentropically from a pressure of about 586 psia [4,040 kPa(a)] to about 34
psia
[234 kPa(a)]. During expansion a portion of the stream is vaporized, resulting
in cooling
of the total stream to -264 F [-164 C] (stream 71d). The expanded stream 71d
then
reenters heat exchangers 60, 13, and 10 where it provides cooling to stream
49c, stream
42, and the refrigerant (streams 71, 71 a, and 71 b) as it is vaporized and
superheated.
-29-
CA 02746624 2011-07-15
The superheated refrigerant vapor (stream 71g) leaves heat exchanger 10
at 90 F [32 C] and is compressed in three stages to 617 psia [4,254 kPa(a)].
Each of the
three compression stages (refrigerant compressors 64, 66, and 68) is driven by
a
supplemental power source and is followed by a cooler (discharge coolers 65,
67, and 69)
to remove the heat of compression. The compressed stream 71 from discharge
cooler 69
returns to heat exchanger 10 to complete the cycle.
A summary of stream flow rates and energy consumption for the process
illustrated in FIG. 4 is set forth in the following table:
-30-
CA 02746624 2011-07-15
Table III
(FIG. 4)--
Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]
Stream Methane Ethane Propane Butanes+ Total
31 40,977 3,861 2,408 1,404 48,656
32 38,431 3,317 1,832 820 44,405
33 2,546 544 576 584 4,251
37 36,692 3,350 19 0 40,066
40 5,324 3,386 1,910 820 11,440
41 0 48 2,386 1,404 3,837
42 10,361 6,258 168 0 16,789
43 4,285 463 3 0 4,753
44 6,076 5,795 165 0 12,036
45 3,585 3,419 97 0 7,101
46 2,491 2,376 68 0 4,935
47 40,977 3,813 22 0 44,819
48 2,453 228 1 0 2,684
50 38,524 3,585 21 0 42,135
-31-
CA 02746624 2011-07-15
Recoveries in LPG*
Propane 99.08%
Butanes+ 100.00%
Production Rate 197,051 Lb/Hr [ 197,051 kg/Hr]
LNG Product
Production Rate 726,918 Lb/Hr [ 726,918 kg/Hr]
Purity* 91.43%
Lower Heating Value 969.9 BTU/SCF [ 36.14 MJ/m3]
Power
Refrigerant Compression 95,424 HP - [ 156,876 kW]
Propane Compression 28,060 HP [ 46,130 kW]
Total Compression 123,484 HP [ 203,006 kW]
Utili Heat
Demethanizer Reboiler 55,070 MBTU/Hr [ 35,575 kW]
* (Based on un-rounded flow rates)
Assuming an on-stream factor of 340 days per year for the LNG
production plant, the specific power consumption for the FIG. 4 embodiment of
the
present invention is 0.143 HP-Hr/Lb [0.236 kW-Hr/kg]. Compared to the prior
art
processes, the efficiency improvement is 17-27% for the FIG. 4 embodiment.
Compared to the FIG. I and FIG. 3 embodiments, the FIG. 4 embodiment
of the present invention requires 6% to 11 % less power per unit of liquid
produced.
Thus, for a given amount of available compression power, the FIG. 4 embodiment
could
-32-
CA 02746624 2011-07-15
liquefy about 6% more natural gas than the FIG. 1 embodiment or about 11% more
natural gas than the FIG. 3 embodiment by virtue of recovering only the C3 and
heavier
hydrocarbons as an LPG co-product. The choice between the FIG. 4 embodiment
versus
either the FIG. I or FIG. 3 embodiments of the present invention for a
particular
application will generally be dictated either by the monetary value of ethane
as part of an
NGL product versus its corresponding value in the LNG product, or by the
heating value
specification for the LNG product (since the heating value of the LNG produced
by the
FIG. 1 and FIG. 3 embodiments is lower than that produced by the FIG. 4
embodiment).
Example 4
If the specifications for the LNG product will allow all of the ethane and
propane contained in the feed gas to be recovered in the LNG product, or if
there is no
market for a liquid co-product containing ethane and propane, an alternative
embodiment
of the present invention such as that shown in FIG. 5 may be employed to
produce a
condensate co-product stream. The inlet gas composition and conditions
considered in
the process presented in FIG. 5 are the same as those in FIGS. 1, 3, and 4.
Accordingly,
the FIG. 5 process can be compared to the embodiments displayed in FIGS. 1, 3,
and 4.
In the simulation of the FIG. 5 process, inlet gas enters the plant at 90 F
[32 C] and 1285 psia [8,860 kPa(a)] as stream 31 and is cooled in heat
exchanger 10 by
heat exchange with refrigerant streams, flashed high pressure separator
liquids at -37 F
[-38 C] (stream 33b), and flashed intermediate pressure separator liquids at -
37 F
[-38 C] (stream 39b). The cooled stream 31a enters high pressure separator 11
at -30 F
-33-
CA 02746624 2011-07-15
[-34 C] and 1278 psia [8,812 kPa(a)] where the vapor (stream 32) is separated
from the
condensed liquid (stream 33).
The vapor (stream 32) from high pressure separator 11 enters work
expansion machine 15 in which mechanical energy is extracted from this portion
of the
high pressure feed. The machine 15 expands the vapor substantially
isentropically from a
pressure of about 1278 psia [8,812 kPa(a)] to a pressure of about 635 psia
[4,378 kPa(a)],
with the work expansion cooling the expanded stream 32a to a temperature of
approximately -83 F [-64 C]. The expanded and partially condensed stream 32a
enters
intermediate pressure separator 18 where the vapor (stream 42) is separated
from the
condensed liquid (stream 39). The intermediate pressure separator liquid
(stream 39) is
flash expanded to slightly above the operating pressure of depropanizer 19 by
expansion
valve 17, cooling stream 39 to -108F [-78 C] (stream 39a) before it enters
heat exchanger
13 and is heated as it provides cooling to residue gas stream 49 and
refrigerant stream
71a, and thence to heat exchanger 10 to provide cooling to the incoming feed
gas as
described earlier. Stream 39c, now at -15 F [-26 C], then enters depropanizer
19 at an
upper mid-column feed point.
The condensed liquid, stream 33, from high pressure separator 11 is flash
expanded to slightly above the operating pressure of depropanizer 19 by
expansion valve
12, cooling stream 33 to -93F [-70 C] (stream 33a) before it enters heat
exchanger 13 and
is heated as it provides cooling to residue gas stream 49 and refrigerant
stream 71 a, and
thence to heat exchanger 10 to provide cooling to the incoming feed gas as
described
-34-
CA 02746624 2011-07-15
earlier. Stream 33c, now at 50 F [I O'C], then enters depropanizer 19 at a
lower
mid-column feed point. In the depropanizer, streams 39c and 33c are stripped
of their
methane, C2 components, and C3 components. The depropanizer in tower 19,
operating
at about 385 psia [2,654 kPa(a)], is a conventional distillation column
containing a
plurality of vertically spaced trays, one or more packed beds, or some
combination of
trays and packing. The depropanizer tower may consist of two sections: an
upper
separator section 19a wherein any vapor contained in the top feed is separated
from its
corresponding liquid portion, and wherein the vapor rising from the lower
distillation or
depropanizing section 19b is combined with the vapor portion (if any) of the
top feed to
form distillation stream 37 which exits the top of the tower; and a lower,
depropanizing
section 19b that contains the trays and/or packing to provide the necessary
contact
between the liquids falling downward and the vapors rising upward. The
depropanizing
section 19b also includes one or more reboilers (such as reboiler 20) which
heat and
vaporize a portion of the liquid at the bottom of the column to provide the
stripping
vapors which flow up the column to strip the liquid product, stream 41, of
methane, C2
components, and C3 components. A typical specification for the bottom liquid
product is
to have a propane to butanes ratio of 0.020:1 on a volume basis. The liquid
product
stream 41 exits the bottom of the deethanizer at 286 F [ 141 C].
The overhead distillation stream 37 leaves depropanizer 19 at 36 F [2 C]
and is cooled and partially condensed by commercial-quality propane
refrigerant in reflux
condenser 21. The partially condensed stream 37a enters reflux drum 22 at 2 F
[-17 C]
where the condensed liquid (stream 44) is separated from the uncondensed vapor
(stream
-35-
CA 02746624 2011-07-15
43). The condensed liquid (stream 44) is pumped by pump 23 to a top feed point
on
depropanizer 19 as reflux stream 44a.
The uncondensed vapor (stream 43) from reflux drum 22 is warmed to
94 F [34 C] in heat exchanger 24, and a portion (stream 48) is then withdrawn
to serve as
fuel gas for the plant. The remainder of the warmed vapor (stream 38) is
compressed by
compressor 16. After cooling to 100 F [38 C] in discharge cooler 25, stream
38b is
further cooled to 15 F [-9 C] in heat exchanger 24 by cross exchange with the
cool
vapor, stream 43.
Stream 38c then combines with the intermediate pressure separator vapor
(stream 42) to form cool residue gas stream 49. Stream 49 enters heat
exchanger 13 and
is cooled from -38 F [-39 C] to -102 F [-74 C] by separator liquids (streams
39a and
33a) as described earlier and by refrigerant stream 71e. Partially condensed
stream 49a
then enters heat exchanger 60 and is further cooled by refrigerant stream 71d
to -254 F
[-159 C] to condense and subcool it, whereupon it enters a work expansion
machine 61
in which mechanical energy is extracted from the stream. The machine 61
expands liquid
stream 49b substantially isentropically from a pressure of about 621 psia
[4,282 kPa(a)]
to the LNG storage pressure (15.5 psia [107 kPa(a)]), slightly above
atmospheric
pressure. The work expansion cools the expanded stream 49c to a temperature of
approximately -255 F [-159 C], whereupon it is then directed to the LNG
storage tank 62
which holds the LNG product (stream 50).
-36-
CA 02746624 2011-07-15
Similar to the FIG. 1, FIG. 3, and FIG. 4 processes, much of the cooling
for stream 49 and all of the cooling for stream 49a is provided by a closed
cycle
refrigeration loop. The composition of the stream used as the working fluid in
the cycle
for the FIG. 5 process, in approximate mole percent, is 8.9% nitrogen, 34.3%
methane,
41.3% ethane, and 11.0% propane, with the balance made up of heavier
hydrocarbons.
The refrigerant stream 71 leaves discharge cooler 69 at 100 F [38 C] and 607
psia
[4,185 kPa(a)]. It enters heat exchanger 10 and is cooled to -30 F [-34 C] and
partially
condensed by the partially warmed expanded refrigerant stream 71f and by other
refrigerant streams. For the FIG. 5 simulation, it has been assumed that these
other
refrigerant streams are commercial-quality propane refrigerant at three
different
temperature and pressure levels. The partially condensed refrigerant stream 71
a then
enters heat exchanger 13 for further cooling to -102 F [-74 C] by partially
warmed
expanded refrigerant stream 71e, further condensing the refrigerant (stream
71b). The
refrigerant is totally condensed and then subcooled to -254 F [-159 C] in heat
exchanger
60 by expanded refrigerant stream 71d. The subcooled liquid stream 71c enters
a work
expansion machine 63 in which mechanical energy is extracted from the stream
as it is
expanded substantially isentropically from a pressure of about 586 psia [4,040
kPa(a)] to
about 34 psia [234 kPa(a)]. During expansion a portion of the stream is
vaporized,
resulting in cooling of the total stream to -264 F [-164 C] (stream 71d). The
expanded
stream 71 d then reenters heat exchangers 60, 13, and 10 where it provides
cooling to
stream 49a, stream 49, and the refrigerant (streams 71, 71a, and 71b) as it is
vaporized
and superheated.
-37-
CA 02746624 2011-07-15
The superheated refrigerant vapor (stream 71 g) leaves heat exchanger 10
at 93 F [34 C] and is compressed in three stages to 617 psia [4,254 kPa(a)].
Each of the
three compression stages (refrigerant compressors 64, 66, and 68) is driven by
a
supplemental power source and is followed by a cooler (discharge coolers 65,
67, and 69)
to remove the heat of compression. The compressed stream 71 from discharge
cooler 69
returns to heat exchanger 10 to complete the cycle.
A summary of stream flow rates and energy consumption for the process
illustrated in FIG. 5 is set forth in the following table:
Table IV
(FIG. 5)
Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]
Stream Methane Ethane Propane Butanes+ Total
31 40,977 3,861 2,408 1,404 48,656
32 32,360 2,675 1,469 701 37,209
33 8,617 1,186 939 703 11,447
38 13,133 2,513 1,941 22 17,610
39 6,194 1,648 1,272 674 9,788
41 0 0 22 1,352 1,375
42 26,166 1,027 197 27 27,421
43 14,811 2,834 2,189 25 19,860
48 1,678 321 248 3 2,250
50 39,299 3,540 2,138 49 45,031
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Recoveries in Condensate*
Butanes 95.04%
Pentanes+ 99.57%
Production Rate 88,390 Lb/Hr [ 88,390 kg/Hr]
LNG Product
Production Rate 834,183 Lb/Hr [ 834,183 kg/Hr]
Purity* 87.27%
Lower Heating Value 1033.8 BTU/SCF [ 38.52 MJ/m3]
Power -
Refrigerant Compression 84,974 HP [ 139,696 kW]
Propane Compression 39,439 HP [ 64,837 kW]
Total Compression 124,413 HP [ 204,533 kW]
Utility
Demethanizer Reboiler 52,913 MBTU/Hr [ 34,182 kW]
* (Based on un-rounded flow rates)
Assuming an on-stream factor of 340 days per year for the LNG
production plant, the specific power consumption for the FIG. 5 embodiment of
the
present invention is 0.145 HP-Hr/Lb [0.238 kW-Hr/kg]. Compared to the prior
art
processes, the efficiency improvement is 16-26% for the FIG. 5 embodiment.
Compared to the FIG. 1 and FIG. 3 embodiments, the FIG. 5 embodiment
of the present invention requires 5% to 10% less power per unit of liquid
produced.
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CA 02746624 2011-07-15
Compared to the FIG. 4 embodiment, the FIG. 5 embodiment of the present
invention
requires essentially the same power per unit of liquid produced. Thus, for a
given
amount of available compression power, the FIG. 5 embodiment could liquefy
about 5%
more natural gas than the FIG. 1 embodiment, about 10% more natural gas than
the
FIG. 3 embodiment, or about the same amount of natural gas as the FIG. 4
embodiment,
by virtue of recovering only the C4 and heavier hydrocarbons as a condensate
co-product.
The choice between the FIG. 5 embodiment versus either the FIG. 1, FIG. 3, or
FIG. 4
embodiments of the present invention for a particular application will
generally be
dictated either by the monetary values of ethane and propane as part of an NGL
or LPG
product versus their corresponding values in the LNG product, or by the
heating value
specification for the LNG product (since the heating value of the LNG produced
by the
FIG. 1, FIG. 3, and FIG. 4 embodiments is lower than that produced by the FIG.
5
embodiment).
Other Embodiments
One skilled in the art will recognize that the present invention can be
adapted for use with all types of LNG liquefaction plants to allow co-
production of an
NGL stream, an LPG stream, or a condensate stream, as best suits the needs at
a given
plant location. Further, it will be recognized that a variety of process
configurations may
be employed for recovering the liquid co-product stream. For instance, the
FIGS. 1 and 3
embodiments can be adapted to recover an LPG stream or a condensate stream as
the
liquid co-product stream rather than an NGL stream as described earlier in
Examples 1
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CA 02746624 2011-07-15
and 2. The FIG. 4 embodiment can be adapted to recover an NGL stream
containing a
significant fraction of the C2 components present in the feed gas, or to
recover a
condensate stream containing only the C4 and heavier components present in the
feed gas,
rather than producing an LPG co-product as described earlier for Example 3.
The FIG. 5
embodiment can be adapted to recover an NGL stream containing a significant
fraction of
the C2 components present in the feed gas, or to recover an LPG stream
containing a
significant fraction of the C3 components present in the feed gas, rather than
producing a
condensate co-product as described earlier for Example 4.
FIGS. 1, 3, 4, and 5 represent the preferred embodiments of the present
invention for the processing conditions indicated. FIGS. 6 through 21 depict
alternative
embodiments of the present invention that may be considered for a particular
application.
As shown in FIGS. 6 and 7, all or a portion of the condensed liquid (stream
33) from
separator 11 can be supplied to fractionation tower 19 at a separate lower mid-
column
feed position rather than combining with the portion of the separator vapor
(stream 34)
flowing to heat exchanger 13. FIG. 8 depicts an alternative embodiment of the
present
invention that requires less equipment than the FIG. 1 and FIG. 6 embodiments,
although
its specific power consumption is somewhat higher. Similarly, FIG. 9 depicts
an
alternative embodiment of the present invention that requires less equipment
than the
FIG. 3 and FIG. 7 embodiments, again at the expense of a higher specific power
consumption. FIGS. 10 through 14 depict alternative embodiments of the present
invention that may require less equipment than the FIG. 4 embodiment, although
their
specific power consumptions may be higher. (Note that as shown in FIGS. 10
through
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CA 02746624 2011-07-15
14, distillation columns or systems such as deethanizer 19 include both
reboiled absorber
tower designs and refluxed, reboiled tower designs.) FIGS. 15 and 16 depict
alternative
embodiments of the present invention that combine the functions of
separator/absorber
tower 18 and deethanizer 19 in the FIGS. 4 and 10 through 14 embodiments into
a single
fractionation column 19. Depending on the quantity of heavier hydrocarbons in
the feed
gas and the feed gas pressure, the cooled feed stream 31 a leaving heat
exchanger 10 may
not contain any liquid (because it is above its dewpoint, or because it is
above its
cricondenbar), so that separator 11 shown in FIGS. I and 3 through 16 is not
required,
and the cooled feed stream can flow directly to an appropriate expansion
device, such as
work expansion machine 15.
The disposition of the gas stream remaining after recovery of the liquid
co-product stream (stream 37 in FIGS. 1, 3, 6 through 11, 13, and 14, stream
47 in
FIGS. 4, 12, 15, and 16, and stream 43 in FIG. 5) before it is supplied to
heat exchanger
60 for condensing and subcooling may be accomplished in many ways. In the
processes
of FIGS. 1 and 3 through 16, the stream is heated, compressed to higher
pressure using
energy derived from one or more work expansion machines, partially cooled in a
discharge cooler, then further cooled by cross exchange with the original
stream. As
shown in FIG. 17, some applications may favor compressing the stream to higher
pressure, using supplemental compressor 59 driven by an external power source
for
example. As shown by the dashed equipment (heat exchanger 24 and discharge
cooler
25) in FIGS. 1 and 3 through 16, some circumstances may favor reducing the
capital cost
of the facility by reducing or eliminating the pre-cooling of the compressed
stream before
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CA 02746624 2011-07-15
it enters heat exchanger 60 (at the expense of increasing the cooling load on
heat
exchanger 60 and increasing the power consumption of refrigerant compressors
64, 66,
and 68). In such cases, stream 49a leaving the compressor may flow directly to
heat
exchanger 24 as shown in FIG. 18, or flow directly to heat exchanger 60 as
shown in
FIG. 19. If work expansion machines are not used for expansion of any portions
of the
high pressure feed gas, a compressor driven by an external power source, such
as
compressor 59 shown in FIG. 20, may be used in lieu of compressor 16. Other
circumstances may not justify any compression of the stream at all, so that
the stream
flows directly to heat exchanger 60 as shown in FIG. 21 and by the dashed
equipment
(heat exchanger 24, compressor 16, and discharge cooler 25) in FIGS. 1 and 3
through
16. If heat exchanger 24 is not included to heat the stream before the plant
fuel gas
(stream 48) is withdrawn, a supplemental heater 58 may be needed to warm the
fuel gas
before it is consumed, using a utility stream or another process stream to
supply the
necessary heat, as shown in FIGS. 19 through 21. Choices such as these must
generally
be evaluated for each application, as factors such as gas composition, plant
size, desired
co-product stream recovery level, and available equipment must all be
considered.
In accordance with the present invention, the cooling of the inlet gas
stream and the feed stream to the LNG production section may be accomplished
in many
ways. In the processes of FIGS. 1, 3, and 6 through 9, inlet gas stream 31 is
cooled and
condensed by external refrigerant streams and tower liquids from fractionation
tower 19.
In FIGS. 4, 5, and 10 through 14 flashed separator liquids are used for this
purpose along
with the external refrigerant streams. In FIGS. 15 and 16 tower liquids and
flashed
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CA 02746624 2011-07-15
separator liquids are used for this purpose along with the external
refrigerant streams.
And in FIGS. 17 through 21, only external refrigerant streams are used to cool
inlet gas
stream 31. However, the cold-process streams could also be used to supply some
of the
cooling to the high pressure refrigerant (stream 71 a), such as shown in FIGS.
4, 5, 10,
and 11. Further, any stream at a temperature colder than the stream(s) being
cooled may
be utilized. For instance, a side draw of vapor from separator/absorber tower
18 or
fractionation tower 19 could be withdrawn and used for cooling. The use and
distribution
of tower liquids and/or vapors for process heat exchange, and the particular
arrangement
of heat exchangers for inlet gas and feed gas cooling, must be evaluated for
each
particular application, as well as the choice of process streams for specific
heat exchange
services. 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).
Further, the supplemental external refrigeration that is supplied to the inlet
gas stream and the feed stream to the LNG production section may also be
accomplished
in many different ways. In FIGS. 1 and 3 through 21, boiling single-component
refrigerant has been assumed for the high level external refrigeration and
vaporizing
multi-component refrigerant has been assumed for the low level external
refrigeration,
with the single-component refrigerant used to pre-cool the multi-component
refrigerant
stream. Alternatively, both the high level cooling and the low level cooling
could be
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CA 02746624 2011-07-15
accomplished using single-component refrigerants with successively lower
boiling points
(i.e., "cascade refrigeration"), or one single-component refrigerant at
successively lower
evaporation pressures. As another alternative, both the high level cooling and
the low
level cooling could be accomplished using multi-component refrigerant streams
with
their respective compositions adjusted to provide the necessary cooling
temperatures.
The selection of the method for providing external refrigeration will depend
on a number
of factors including, but not limited to, feed gas composition and conditions,
plant size,
compressor driver size, heat exchanger size, ambient heat sink temperature,
etc. One
skilled in the art will also recognize that any combination of the methods for
providing
external refrigeration described above may be employed in combination to
achieve the
desired feed stream temperature(s).
Subcooling of the condensed liquid stream leaving heat exchanger 60
(stream 49 in FIGS. 1, 6, and 8, stream 49d in FIGS. 3, 4, 7, and 9 through
16, stream
49b in FIGS. 5, 19, and 20, stream 49e in FIG. 17, stream 49c in FIG. 18, and
stream 49a
in FIG. 21) reduces or eliminates the quantity of flash vapor that may be
generated during
expansion of the stream to the operating pressure of LNG storage'tank 62. This
generally
reduces the specific power consumption for producing the LNG by eliminating
the need
for flash gas compression. However, some circumstances may favor reducing the
capital
cost of the facility by reducing the size of heat exchanger 60 and using flash
gas
compression or other means to dispose of any flash gas that may be generated.
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CA 02746624 2011-07-15
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 35a in FIGS. 1, 3, 6, and 7) or the intermediate pressure reflux
stream (stream 39
in FIGS. 1, 6, and 8). Further, isenthalpic flash expansion may be used in
lieu of work
expansion for the subcooled liquid stream leaving heat exchanger 60 (stream 49
in
FIGS. 1, 6, and 8, stream 49d in FIGS. 3, 4, 7, and 9 through 16, stream 49b
in FIGS. 5,
19, and 20, stream 49e in FIG. 17, stream 49c in FIG. 18, and stream 49a in
FIG. 21), but
will necessitate either more subcooling in heat exchanger 60 to avoid forming
flash vapor
in the expansion, or else adding flash vapor compression or other means for
disposing of
the flash vapor that results. Similarly, isenthalpic flash expansion may be
used in lieu of
work expansion for the subcooled high pressure refrigerant stream leaving heat
exchanger 60 (stream 71c in FIGS. 1 and 3 through 21), with the resultant
increase in the
power consumption for compression of the refrigerant.
While there have been described what are believed to be preferred
embodiments of the invention, those skilled in the art will recognize that
other and further
modifications may be made thereto, e.g. to adapt the invention to various
conditions,
types of feed, or other requirements without departing from the spirit of the
present
invention as defined by the following claims.
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