Note : Les descriptions sont présentées dans la langue officielle dans laquelle elles ont été soumises.
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HYDROCARBON GAS PROCESSING
SPECIFICATION
BACKGROUND OF THE INVENTION
[0001] Ethylene, ethane, propylene, propane, and/or heavier
hydrocarbons can
be recovered from a variety of gases, such as natural gas, refinery gas, and
synthetic
gas streams obtained from other hydrocarbon materials such as coal, crude oil,
naphtha, oil shale, tar sands, and lignite. Natural gas usually has a major
proportion
of methane and ethane, i.e., methane and ethane together comprise at least 50
mole
percent of the gas. The gas also contains relatively lesser amounts of heavier
hydrocarbons such as propane, butanes, pentanes, and the like, as well as
hydrogen,
nitrogen, carbon dioxide, and other gases.
[0002] The present invention is generally concerned with the recovery
of
ethylene, ethane, propylene, propane, and heavier hydrocarbons from such gas
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streams. A typical analysis of a gas stream to be processed in accordance with
this
invention would be, in approximate mole percent, 90.3% methane, 4.0% ethane
and
other C2 components, 1.7% propane and other C3 components, 0.3% iso-butane,
0.5%
normal butane, and 0.8% pentanes plus, with the balance made up of nitrogen
and
carbon dioxide. Sulfur containing gases are also sometimes present.
[0003] The historically cyclic fluctuations in the prices of both
natural gas and
its natural gas liquid (NGL) constituents have at times reduced the
incremental value
of ethane, ethylene, propane, propylene, and heavier components as liquid
products.
This has resulted in a demand for processes that can provide more efficient
recoveries
of these products and for processes that can provide efficient recoveries with
lower
capital investment. Available processes for separating these materials include
those
based upon cooling and refrigeration of gas, oil absorption, and refrigerated
oil
absorption. Additionally, cryogenic processes have become popular because of
the
availability of economical equipment that produces power while simultaneously
expanding and extracting heat from the gas being processed. Depending upon the
pressure of the gas source, the richness (ethane, ethylene, and heavier
hydrocarbons
content) of the gas, and the desired end products, each of these processes or
a
combination thereof may be employed.
[0004] The cryogenic expansion process is now generally preferred for
natural
gas liquids recovery because it provides maximum simplicity with ease of
startup,
operating flexibility, good efficiency, safety, and good reliability. U.S.
Patent Nos.
3,292,380; 4,061,481; 4,140,504; 4,157,904; 4,171,964; 4,185,978; 4,251,249;
4,278,457; 4,519,824; 4,617,039; 4,687,499; 4,689,063; 4,690,702; 4,854,955;
4,869,740; 4,889,545; 5,275,005; 5,555,748; 5,566,554; 5,568,737; 5,771,712;
5,799,507; 5,881,569; 5,890,378; 5,983,664; 6,182,469; 6,578,379; 6,712,880;
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6,915,662; 7,191,617; 7,219,513; reissue U.S. Patent No. 33,408; and co-
pending
application nos. 11/430,412; 11/839,693; 11/971,491; 12/206,230; 12/689,616;
12/717,394; 12/750,862; 12/772,472; 12/781,259; 12/868,993; 12/869,007;
12/869,139; 12/979,563; 13/048,315; and 13/051,682 describe relevant processes
(although the description of the present invention in some cases is based on
different
processing conditions than those described in the cited U.S. Patents).
[0005] In a
typical cryogenic expansion recovery process, a feed gas stream
under pressure is cooled by heat exchange with other streams of the process
and/or
external sources of refrigeration such as a propane compression-refrigeration
system.
As the gas is cooled, liquids may be condensed and collected in one or more
separators as high-pressure liquids containing some of the desired C2+
components.
Depending on the richness of the gas and the amount of liquids formed, the
high-pressure liquids may be expanded to a lower pressure and fractionated.
The
vaporization occurring during expansion of the liquids results in further
cooling of the
stream. Under some conditions, pre-cooling the high pressure liquids prior to
the
expansion may be desirable in order to further lower the temperature resulting
from
the expansion. The expanded stream, comprising a mixture of liquid and vapor,
is
fractionated in a distillation (demethanizer or deethanizer) column. In the
column, the
expansion cooled stream(s) is (are) distilled to separate residual methane,
nitrogen,
and other volatile gases as overhead vapor from the desired C2 components, C3
components, and heavier hydrocarbon components as bottom liquid product, or to
separate residual methane, C2 components, nitrogen, and other volatile gases
as
overhead vapor from the desired C3 components and heavier hydrocarbon
components
as bottom liquid product.
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[0006] If the feed gas is not totally condensed (typically it is not),
the vapor
remaining from the partial condensation can be split into two streams. One
portion of
the vapor is passed through a work expansion machine or engine, or an
expansion
valve, to a lower pressure at which additional liquids are condensed as a
result of
further cooling of the stream. The pressure after expansion is essentially the
same as
the pressure at which the distillation column is operated. The combined vapor-
liquid
phases resulting from the expansion are supplied as feed to the column.
[0007] The remaining portion of the vapor is cooled to substantial
condensation by heat exchange with other process streams, e.g., the cold
fractionation
tower overhead. Some or all of the high-pressure liquid may be combined with
this
vapor portion prior to cooling. The resulting cooled stream is then expanded
through
an appropriate expansion device, such as an expansion valve, to the pressure
at which
the demethanizer is operated. During expansion, a portion of the liquid will
vaporize,
resulting in cooling of the total stream. The flash expanded stream is then
supplied as
top feed to the demethanizer. Typically, the vapor portion of the flash
expanded
stream and the demethanizer overhead vapor combine in an upper separator
section in
the fractionation tower as residual methane product gas. Alternatively, the
cooled and
expanded stream may be supplied to a separator to provide vapor and liquid
streams.
The vapor is combined with the tower overhead and the liquid is supplied to
the
column as a top column feed.
[0008] In the ideal operation of such a separation process, the
residue gas
leaving the process will contain substantially all of the methane in the feed
gas with
essentially none of the heavier hydrocarbon components and the bottoms
fraction
leaving the demethanizer will contain substantially all of the heavier
hydrocarbon
components with essentially no methane or more volatile components. In
practice,
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however, this ideal situation is not obtained because the conventional
demethanizer is
operated largely as a stripping column. The methane product of the process,
therefore, typically comprises vapors leaving the top fractionation stage of
the
column, together with vapors not subjected to any rectification step.
Considerable
losses of C2, C3, and C4+ components occur because the top liquid feed
contains
substantial quantities of these components and heavier hydrocarbon components,
resulting in corresponding equilibrium quantities of C2 components, C3
components,
C4 components, and heavier hydrocarbon components in the vapors leaving the
top
fractionation stage of the demethanizer. The loss of these desirable
components could
be significantly reduced if the rising vapors could be brought into contact
with a
significant quantity of liquid (reflux) capable of absorbing the C2
components, C3
components, C4 components, and heavier hydrocarbon components from the vapors.
[0009] In recent years, the preferred processes for hydrocarbon
separation use
an upper absorber section to provide additional rectification of the rising
vapors. The
source of the reflux stream for the upper rectification section is typically a
recycled
stream of residue gas supplied under pressure. The recycled residue gas stream
is
usually cooled to substantial condensation by heat exchange with other process
streams, e.g., the cold fractionation tower overhead. The resulting
substantially
condensed stream is then expanded through an appropriate expansion device,
such as
an expansion valve, to the pressure at which the demethanizer is operated.
During
expansion, a portion of the liquid will usually vaporize, resulting in cooling
of the
total stream. The flash expanded stream is then supplied as top feed to the
demethanizer. Typically, the vapor portion of the expanded stream and the
demethanizer overhead vapor combine in an upper separator section in the
fractionation tower as residual methane product gas. Alternatively, the cooled
and
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expanded stream may be supplied to a separator to provide vapor and liquid
streams,
so that thereafter the vapor is combined with the tower overhead and the
liquid is
supplied to the column as a top column feed. Typical process schemes of this
type are
disclosed in U.S. Patent Nos. 4,889,545; 5,568,737; and 5,881,569, co-pending
application nos. 11/430,412; 11/971,491; and 12/717,394, and in Mowrey, E.
Ross,
"Efficient, High Recovery of Liquids from Natural Gas Utilizing a High
Pressure
Absorber", Proceedings of the Eighty-First Annual Convention of the Gas
Processors
Association, Dallas, Texas, March 11-13, 2002.
[0010] The present invention employs a novel means of perfoiming the
various steps described above more efficiently and using fewer pieces of
equipment.
This is accomplished by combining what heretofore have been individual
equipment
items into a common housing, thereby reducing the plot space required for the
processing plant and reducing the capital cost of the facility. Surprisingly,
applicants
have found that the more compact arrangement also significantly reduces the
power
consumption required to achieve a given recovery level, thereby increasing the
process efficiency and reducing the operating cost of the facility. In
addition, the
more compact arrangement also eliminates much of the piping used to
interconnect
the individual equipment items in traditional plant designs, further reducing
capital
cost and also eliminating the associated flanged piping connections. Since
piping
flanges are a potential leak source for hydrocarbons (which are volatile
organic
compounds, VOCs, that contribute to greenhouse gases and may also be
precursors to
atmospheric ozone foimation), eliminating these flanges reduces the potential
for
atmospheric emissions that can damage the environment.
[0011] In accordance with the present invention, it has been found
that C2
recoveries in excess of 95% can be obtained. Similarly, in those instances
where
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recovery of C2 components is not desired, C3 recoveries in excess of 95% can
be
maintained. In addition, the present invention makes possible essentially 100%
separation of methane (or C2 components) and lighter components from the C2
components (or C3 components) and heavier components at lower energy
requirements compared to the prior art while maintaining the same 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.
[0012] For a better understanding of the present invention, reference
is made
to the following examples and drawings. Referring to the drawings:
[0013] FIG. 1 is a flow diagram of a prior art natural gas processing
plant in
accordance with United States Patent No. 5,568,737;
[0014] FIG. 2 is a flow diagram of a natural gas processing plant in
accordance with the present invention; and
[0015] FIGS. 3 through 17 are flow diagrams illustrating alternative
means of
application of the present invention to a natural gas stream.
[0016] 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
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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.
[0017] For convenience, process parameters are reported in both the
traditional British units and in the units of the Systeme International
d'Unites (SI).
The molar flow rates given in the tables may be interpreted as either pound
moles per
hour or kilogram moles per hour. The energy consumptions reported as
horsepower
(HP) and/or thousand British Thermal Units per hour (MBTU/Hr) correspond to
the
stated molar flow rates in pound moles per hour. The energy consumptions
reported
as kilowatts (kW) correspond to the stated molar flow rates in kilogram moles
per
hour.
DESCRIPTION OF THE PRIOR ART
[0018] FIG. 1 is a process flow diagram showing the design of a
processing
plant to recover C2+ components from natural gas using prior art according to
U.S.
Pat. No. 5,568,737. In this simulation of the process, inlet gas enters the
plant at
110 F [43 C] and 915 psia [6,307 kPa(a)] as stream 31. If the inlet gas
contains a
concentration of sulfur compounds which would prevent the product streams from
meeting specifications, the sulfur compounds are removed by appropriate
pretreatment of the feed gas (not illustrated). In addition, the feed stream
is usually
dehydrated to prevent hydrate (ice) formation under cryogenic conditions.
Solid
desiccant has typically been used for this purpose.
[0019] The feed stream 31 is divided into two portions, streams 32 and
33.
Stream 32 is cooled to -26 F [-32 C] in heat exchanger 10 by heat exchange
with cool
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distillation vapor stream 41a, while stream 33 is cooled to -32 F [-35 C] in
heat
exchanger 11 by heat exchange with demethanizer reboiler liquids at 41 F [5 C]
(stream 43) and side reboiler liquids at -49 F [-45 C] (stream 42). Streams
32a and
33a recombine to form stream 31a, which enters separator 12 at -28 F [-33 C]
and
893 psia [6,155 kPa(a)] where the vapor (stream 34) is separated from the
condensed
liquid (stream 35).
[0020] The vapor (stream 34) from separator 12 is divided into two
streams,
36 and 39. Stream 36, containing about 27% of the total vapor, is combined
with the
separator liquid (stream 35), and the combined stream 38 passes through heat
exchanger 13 in heat exchange relation with cold distillation vapor stream 41
where it
is cooled to substantial condensation. The resulting substantially condensed
stream
38a at -139 F [-95 C] is then flash expanded through expansion valve 14 to the
operating pressure (approximately 396 psia [2,730 kPa(a)]) of fractionation
tower 18.
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 38b leaving
expansion valve 14 reaches a temperature of -140 F [-95 C] and is supplied to
fractionation tower 18 at a first mid-column feed point.
[0021] The remaining 73% of the vapor from separator 12 (stream 39)
enters a
work expansion machine 15 in which mechanical energy is extracted from this
portion
of the high pressure feed. The machine 15 expands the vapor substantially
isentropically to the tower operating pressure, with the work expansion
cooling the
expanded stream 39a to a temperature of approximately -95 F [-71 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
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recovered is often used to drive a centrifugal compressor (such as item 16)
that can be
used to re-compress the heated distillation vapor stream (stream 41b), for
example.
The partially condensed expanded stream 39a is thereafter supplied as feed to
fractionation tower 18 at a second mid-column feed point.
[0022] The recompressed and cooled distillation vapor stream 41e is
divided
into two streams. One portion, stream 46, is the volatile residue gas product.
The
other portion, recycle stream 45, flows to heat exchanger 10 where it is
cooled to
-26 F [-32 C] by heat exchange with cool distillation vapor stream 41a. The
cooled
recycle stream 45a then flows to exchanger 13 where it is cooled to -139 F [-
95 C]
and substantially condensed by heat exchange with cold distillation vapor
stream 41.
The substantially condensed stream 45b is then expanded through an appropriate
expansion device, such as expansion valve 22, to the demethanizer operating
pressure,
resulting in cooling of the total stream to -147 F [-99 C]. The expanded
stream 45c is
then supplied to fractionation tower 18 as the top column feed. The vapor
portion (if
any) of stream 45c combines with the vapors rising from the top fractionation
stage of
the column to form distillation vapor stream 41, which is withdrawn from an
upper
region of the tower.
[0023] The demethanizer in tower 18 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
18a 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 18b is combined with the vapor portion of the top feed
to form
the cold demethanizer overhead vapor (stream 41) which exits the top of the
tower at
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-144 F [-98 C]. The lower, demethanizing section 18b contains the trays and/or
packing and provides the necessary contact between the liquids falling
downward and
the vapors rising upward. The demethanizing section 18b also includes
reboilers
(such as the reboiler and the side reboiler described previously) which heat
and
vaporize a portion of the liquids flowing down the column to provide the
stripping
vapors which flow up the column to strip the liquid product, stream 44, of
methane
and lighter components.
10024] The liquid product stream 44 exits the bottom of the tower at
64 F
[18 C], based on a typical specification of a methane to ethane ratio of
0.010:1 on a
mass basis in the bottom product. The demethanizer overhead vapor stream 41
passes
countercurrently to the incoming feed gas and recycle stream in heat exchanger
13
where it is heated to -40 F [-40 C] (stream 41a) and in heat exchanger 10
where it is
heated to 104 F [40 C] (stream 41b). The distillation vapor stream is then
re-compressed in two stages. The first stage is compressor 16 driven by
expansion
machine 15. The second stage is compressor 20 driven by a supplemental power
source which compresses the residue gas (stream 41d) to sales line pressure.
After
cooling to 110 F [43 C] in discharge cooler 21, stream 41e is split into the
residue gas
product (stream 46) and the recycle stream 45 as described earlier. Residue
gas
stream 46 flows to the sales gas pipeline at 915 psia [6,307 kPa(a)],
sufficient to meet
line requirements (usually on the order of the inlet pressure).
[0025] A summary of stream flow rates and energy consumption for the
process illustrated in FIG. 1 is set forth in the following table:
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Table I
(FIG. 1)
Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]
Stream Methane Ethane Propane Butanes+
Total
31 12,398 546 233 229
13,726
32 8,431 371 159 156
9,334
33 3,967 175 74 73
4,392
34 12,195 501 179 77
13,261
35 203 45 54 152
465
36 3,317 136 49 21
3,607
38 3,520 181 103 173
4,072
39 8,878 365 130 56
9,654
41 13,765 30 0 0
13,992
45 1,377 3 0 0
1,400
46 12,388 27 0 0
12,592
44 10 519 233 229
1,134
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Recoveries*
Ethane 94.99%
Propane 99.99%
Butanes+ 100.00%
Power
Residue Gas Compression 6,149 HP [
10,109 kW]
* (Based on un-rounded flow rates)
DESCRIPTION OF THE INVENTION
[00261 FIG. 2 illustrates a flow diagram of a process in accordance
with the
present invention. The feed gas composition and conditions considered in the
process
presented in FIG. 2 are the same as those in FIG. 1. Accordingly, the FIG. 2
process
can be compared with that of the FIG. 1 process to illustrate the advantages
of the
present invention.
[00271 In the simulation of the FIG. 2 process, inlet gas enters the
plant as
stream 31 and is divided into two portions, streams 32 and 33. The first
portion,
stream 32, enters a heat exchange means in the upper region of feed cooling
section
118a inside processing assembly 118. This heat exchange means may be comprised
of a fin and tube type heat exchanger, a plate type heat exchanger, a brazed
aluminum
type heat exchanger, or other type of heat transfer device, including multi-
pass and/or
multi-service heat exchangers. The heat exchange means is configured to
provide
heat exchange between stream 32 flowing through one pass of the heat exchange
means and a distillation vapor stream arising from separator section 118b
inside
processing assembly 118 that has been heated in a heat exchange means in the
lower
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region of feed cooling section 118a. Stream 32 is cooled while further heating
the
distillation vapor stream, with stream 32a leaving the heat exchange means at -
25 F
[-32 C].
[0028] The second portion, stream 33, enters a heat and mass transfer
means
in demethanizing section 118e inside processing assembly 118. This heat and
mass
transfer means may also be comprised of a fin and tube type heat exchanger, a
plate
type heat exchanger, a brazed aluminum type heat exchanger, or other type of
heat
transfer device, including multi-pass and/or multi-service heat exchangers.
The heat
and mass transfer means is configured to provide heat exchange between stream
33
flowing through one pass of the heat and mass transfer means and a
distillation liquid
stream flowing downward from absorbing section 118d inside processing assembly
118, so that stream 33 is cooled while heating the distillation liquid stream,
cooling
stream 33a to -47 F [-44 C] before it leaves the heat and mass transfer means.
As the
distillation liquid stream is heated, a portion of it is vaporized to form
stripping vapors
that rise upward as the remaining liquid continues flowing downward through
the heat
and mass transfer means. The heat and mass transfer means provides continuous
contact between the stripping vapors and the distillation liquid stream so
that it also
functions to provide mass transfer between the vapor and liquid phases,
stripping the
liquid product stream 44 of methane and lighter components.
[0029] Streams 32a and 33a recombine to form stream 31a, which enters
separator section 118f inside processing assembly 118 at -32 F [-36 C] and 900
psia
[6,203 kPa(a)], whereupon the vapor (stream 34) is separated from the
condensed
liquid (stream 35). Separator section 118f has an internal head or other means
to
divide it from demethanizing section 118e, so that the two sections inside
processing
assembly 118 can operate at different pressures.
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[0030] The vapor (stream 34) from separator section 118f is divided
into two
streams, 36 and 39. Stream 36, containing about 27% of the total vapor, is
combined
with the separated liquid (stream 35, via stream 37), and the combined stream
38
enters a heat exchange means in the lower region of feed cooling section 118a
inside
processing assembly 118. This heat exchange means may likewise be comprised of
a
fin and tube type heat exchanger, a plate type heat exchanger, a brazed
aluminum type
heat exchanger, or other type of heat transfer device, including multi-pass
and/or
multi-service heat exchangers. The heat exchange means is configured to
provide
heat exchange between stream 38 flowing through one pass of the heat exchange
means and the distillation vapor stream arising from separator section 118b,
so that
stream 38 is cooled to substantial condensation while heating the distillation
vapor
stream.
[0031] The resulting substantially condensed stream 38a at -138 F [-95
C] is
then flash expanded through expansion valve 14 to the operating pressure
(approximately 400 psia [2,758 kPa(a)]) of rectifying section 118c (an
absorbing
means) and absorbing section 118d (another absorbing means) inside processing
assembly 118. During expansion a portion of the stream may be vaporized,
resulting
in cooling of the total stream. In the process illustrated in FIG. 2, the
expanded
stream 38b leaving expansion valve 14 reaches a temperature of -139 F [-95 C]
and
is supplied to processing assembly 118 between rectifying section 118c and
absorbing
section 118d. The liquids in stream 38b combine with the liquids falling from
rectifying section 118c and are directed to absorbing section 118d, while any
vapors
combine with the vapors rising from absorbing section 118d and are directed to
rectifying section 118c.
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[0032] The remaining 73% of the vapor from separator section 118f
(stream
39) enters a work expansion machine 15 in which mechanical energy is extracted
from this portion of the high pressure feed. The machine 15 expands the vapor
substantially isentropically to the operating pressure of absorbing section
118d, with
the work expansion cooling the expanded stream 39a to a temperature of
approximately -99 F [-73 C]. The partially condensed expanded stream 39a is
thereafter supplied as feed to the lower region of absorbing section 118d
inside
processing assembly 118.
[0033] The recompressed and cooled distillation vapor stream 41c is
divided
into two streams. One portion, stream 46, is the volatile residue gas product.
The
other portion, recycle stream 45, enters a heat exchange means in the feed
cooling
section 118a inside processing assembly 118. This heat exchange means may also
be
comprised of a fin and tube type heat exchanger, a plate type heat exchanger,
a brazed
aluminum type heat exchanger, or other type of heat transfer device, including
multi-pass and/or multi-service heat exchangers. The heat exchange means is
configured to provide heat exchange between stream 45 flowing through one pass
of
the heat exchange means and the distillation vapor stream arising from
separator
section 118b, so that stream 45 is cooled to substantial condensation while
heating the
distillation vapor stream.
[0034] The substantially condensed recycle stream 45a leaves the heat
exchange means in feed cooling section 118a at -138 F [-95 C] and is flash
expanded
through expansion valve 22 to the operating pressure of rectifying section
118c inside
processing assembly 118. During expansion a portion of the stream is
vaporized,
resulting in cooling of the total stream. In the process illustrated in FIG.
2, the
expanded stream 45b leaving expansion valve 22 reaches a temperature of -146 F
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[-99 C] and is supplied to separator section 118b inside processing assembly
118.
The liquids separated therein are directed to rectifying section 118c, while
the
remaining vapors combine with the vapors rising from rectifying section 118c
to form
the distillation vapor stream that is heated in cooling section 118a.
[0035] Rectifying section 118c and absorbing section 118d each contain
an
absorbing means consisting of a plurality of vertically spaced trays, one or
more
packed beds, or some combination of trays and packing. The trays and/or
packing in
rectifying section 118c and absorbing section 118d provide the necessary
contact
between the vapors rising upward and cold liquid falling downward. The liquid
portion of the expanded stream 39a commingles with liquids falling downward
from
absorbing section 118d and the combined liquid continues downward into
demethanizing section 118e. The stripping vapors arising from demethanizing
section
118e combine with the vapor portion of the expanded stream 39a and rise upward
through absorbing section 118d, to be contacted with the cold liquid falling
downward to condense and absorb most of the C2 components, C3 components, and
heavier components from these vapors. The vapors arising from absorbing
section
118d combine with any vapor portion of the expanded stream 38b and rise upward
through rectifying section 118c, to be contacted with the cold liquid portion
of
expanded stream 45b falling downward to condense and absorb most of the C2
components, C3 components, and heavier components remaining in these vapors.
The
liquid portion of the expanded stream 38b commingles with liquids falling
downward
from rectifying section 118c and the combined liquid continues downward into
absorbing section 118d.
[0036] The distillation liquid flowing downward from the heat and mass
transfer means in demethanizing section 118e inside processing assembly 118
has
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been stripped of methane and lighter components. The resulting liquid product
(stream 44) exits the lower region of demethanizing section 118e and leaves
processing assembly 118 at 65 F [18 C]. The distillation vapor stream arising
from
separator section 118b is warmed in feed cooling section 118a as it provides
cooling
to streams 32, 38, and 45 as described previously, and the resulting
distillation vapor
stream 41 leaves processing assembly 118 at 105 F [40 C]. The distillation
vapor
stream is then re-compressed in two stages, compressor 16 driven by expansion
machine 15 and compressor 20 driven by a supplemental power source. After
stream
41b is cooled to 110 F [43 C] in discharge cooler 21 to form stream 41c,
recycle
stream 45 is withdrawn as described earlier, forming residue gas stream 46
which
thereafter flows to the sales gas pipeline at 915 psia [6,307 kPa(a)].
[0037] A summary of stream flow rates and energy consumption for the
process illustrated in FIG. 2 is set forth in the following table:
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Table II
(FIG. 2)
Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]
Stream Methane Ethane Propane Butanes+
Total
31 12,398 546 233 229
13,726
32 8,679 382 163 160
9,608
33 3,719 164 70 69
4,118
34 12,164 495 174 72
13,213
35 234 51 59 157 513
36 3,248 132 46 19
3,528
37 234 51 59 157 513
38 3,482 183 105 176
4,041
39 8,916 363 128 53
9,685
40 0 0 0 0 0
41 13,863 30 0 0
14,095
45 1,475 3 0 0
1,500
46 12,388 27 0 0
12,595
44 10 519 233 229
1,131
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Recoveries*
Ethane 95.03%
Propane 99.99%
Butanes+ 100.00%
Power
Residue Gas Compression 5,787 HP 9,514
kW]
* (Based on un-rounded flow rates)
[0038] A comparison of Tables I and II shows that the present
invention
maintains essentially the same recoveries as the prior art. However, further
comparison of Tables I and II shows that the product yields were achieved
using
significantly less power than the prior art. In terms of the recovery
efficiency
(defined by the quantity of ethane recovered per unit of power), the present
invention
represents more than a 6% improvement over the prior art of the FIG. 1
process.
[0039] The improvement in recovery efficiency provided by the present
invention over that of the prior art of the FIG. 1 process is primarily due to
two
factors. First, the compact arrangement of the heat exchange means in feed
cooling
section 118a and the heat and mass transfer means in demethanizing section
118e in
processing assembly 118 eliminates the pressure drop imposed by the
interconnecting
piping found in conventional processing plants. The result is that the portion
of the
feed gas flowing to expansion machine 15 is at higher pressure for the present
invention compared to the prior art, allowing expansion machine 15 in the
present
invention to produce as much power with a higher outlet pressure as expansion
machine 15 in the prior art can produce at a lower outlet pressure. Thus,
rectifying
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section 118c and absorbing section 118d in processing assembly 118 of the
present
invention can operate at higher pressure than fractionation column 18 of the
prior art
while maintaining the same recovery level. This higher operating pressure,
plus the
reduction in pressure drop for the distillation vapor stream due to
eliminating the
interconnecting piping, results in a significantly higher pressure for the
distillation
vapor stream entering compressor 20, thereby reducing the power required by
the
present invention to restore the residue gas to pipeline pressure.
[0040] Second, using the heat and mass transfer means in demethanizing
section 118e to simultaneously heat the distillation liquid leaving absorbing
section
118d while allowing the resulting vapors to contact the liquid and strip its
volatile
components is more efficient than using a conventional distillation column
with
external reboilers. The volatile components are stripped out of the liquid
continuously, reducing the concentration of the volatile components in the
stripping
vapors more quickly and thereby improving the stripping efficiency for the
present
invention.
[0041] The present invention offers two other advantages over the
prior art in
addition to the increase in processing efficiency. First, the compact
arrangement of
processing assembly 118 of the present invention replaces five separate
equipment
items in the prior art (heat exchangers 10, 11, and 13; separator 12; and
fractionation
tower 18 in FIG. 1) with a single equipment item (processing assembly 118 in
FIG. 2). This reduces the plot space requirements and eliminates the
interconnecting
piping, reducing the capital cost of a process plant utilizing the present
invention over
that of the prior art. Second, elimination of the interconnecting piping means
that a
processing plant utilizing the present invention has far fewer flanged
connections
compared to the prior art, reducing the number of potential leak sources in
the plant.
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Hydrocarbons are volatile organic compounds (VOCs), some of which are
classified
as greenhouse gases and some of which may be precursors to atmospheric ozone
formation, which means the present invention reduces the potential for
atmospheric
releases that can damage the environment.
Other Embodiments
[0042] Some circumstances may favor eliminating feed cooling section
118a
from processing assembly 118, and using a heat exchange means external to the
processing assembly for feed cooling, such as heat exchanger 10 shown in FIGS.
10
through 17. Such an arrangement allows processing assembly 118 to be smaller,
which may reduce the overall plant cost and/or shorten the fabrication
schedule in
some cases. Note that in all cases exchanger 10 is representative of either a
multitude
of individual heat exchangers or a single multi-pass heat exchanger, or any
combination thereof. Each such heat exchanger may be comprised of a fin and
tube
type heat exchanger, a plate type heat exchanger, a brazed aluminum type heat
exchanger, or other type of heat transfer device, including multi-pass and/or
multi-service heat exchangers.
[0043] Some circumstances may favor supplying liquid stream 35
directly to
the lower region of absorbing section 118d via stream 40 as shown in FIGS. 2,
4, 6, 8,
10, 12, 14, and 16. In such cases, an appropriate expansion device (such as
expansion
valve 17) is used to expand the liquid to the operating pressure of absorbing
section
118d and the resulting expanded liquid stream 40a is supplied as feed to the
lower
region of absorbing section 118d (as shown by the dashed lines). Some
circumstances may favor combining a portion of liquid stream 35 (stream 37)
with the
vapor in stream 36 (FIGS. 2, 6, 10, and 14) or with cooled second portion 33a
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(FIGS. 4, 8, 12, and 16) to form combined stream 38 and routing the remaining
portion of liquid stream 35 to the lower region of absorbing section 118d via
streams
40/40a. Some circumstances may favor combining the expanded liquid stream 40a
with expanded stream 39a (FIGS. 2, 6, 10, and 14) or expanded stream 34a
(FIGS. 4,
8, 12, and 16) and thereafter supplying the combined stream to the lower
region of
absorbing section 118d as a single feed.
[0044] If the feed gas is richer, the quantity of liquid separated in
stream 35
may be great enough to favor placing an additional mass transfer zone in
demethanizing section 118e between expanded stream 39a and expanded liquid
stream 40a as shown in FIGS. 3, 7, 11, and 15, or between expanded stream 34a
and
expanded liquid stream 40a as shown in FIGS. 5, 9, 13, and 17. In such cases,
the
heat and mass transfer means in demethanizing section 118e may be configured
in
upper and lower parts so that expanded liquid stream 40a can be introduced
between
the two parts. As shown by the dashed lines, some circumstances may favor
combining a portion of liquid stream 35 (stream 37) with the vapor in stream
36
(FIGS. 3, 7, 11, and 15) or with cooled second portion 33a (FIGS. 5, 9, 13,
and 17) to
form combined stream 38, while the remaining portion of liquid stream 35
(stream 40)
is expanded to lower pressure and supplied between the upper and lower parts
of the
heat and mass transfer means in demethanizing section 118e as stream 40a.
[0045] Some circumstances may favor not combining the cooled first and
second portions (streams 32a and 33a) as shown in FIGS. 4, 5, 8, 9, 12, 13,
16, and
17. In such cases, only the cooled first portion 32a is directed to separator
section
118f inside processing assembly 118 (FIGS. 4, 5, 12, and 13) or separator 12
(FIGS. 8, 9, 16, and 17) where the vapor (stream 34) is separated from the
condensed
liquid (stream 35). Vapor stream 34 enters work expansion machine 15 and is
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expanded substantially isentropically to the operating pressure of absorbing
section
118d, whereupon expanded stream 34a is supplied as feed to the lower region of
absorbing section 118d inside processing assembly 118. The cooled second
portion
33a is combined with the separated liquid (stream 35, via stream 37), and the
combined stream 38 is directed to the heat exchange means in the lower region
of
feed cooling section 118a inside processing assembly 118 (or in heat exchanger
10
external to processing assembly 118) and cooled to substantial condensation.
The
substantially condensed stream 38a is flash expanded through expansion valve
14 to
the operating pressure of rectifying section 118c and absorbing section 118d,
whereupon expanded stream 38b is supplied to processing assembly 118 between
rectifying section 118c and absorbing section 118d. Some circumstances may
favor
combining only a portion (stream 37) of liquid stream 35 with the cooled
second
portion 33a, with the remaining portion (stream 40) supplied to the lower
region of
absorbing section 118c1 via expansion valve 17. Other circumstances may favor
sending all of liquid stream 35 to the lower region of absorbing section 118d
via
expansion valve 17.
[0046] In some circumstances, it may be advantageous to use an
external
separator vessel to separate cooled feed stream 31a or cooled first portion
32a, rather
than including separator section 118f in processing assembly 118. As shown in
FIGS. 6, 7, 14, and 15, separator 12 can be used to separate cooled feed
stream 31a
into vapor stream 34 and liquid stream 35. Likewise, as shown in FIGS. 8, 9,
16, and
17, separator 12 can be used to separate cooled first portion 32a into vapor
stream 34
and liquid stream 35.
[0047] Depending on the quantity of heavier hydrocarbons in the feed
gas and
the feed gas pressure, the cooled feed stream 31a entering separator section
118f in
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FIGS. 2, 3, 10, and 11 or separator 12 in FIGS. 6, 7, 14, and 15 (or the
cooled first
portion 32a entering separator section 118f in FIGS. 4, 5, 12, and 13 or
separator 12
in FIGS. 8, 9, 16, and 17) may not contain any liquid (because it is above its
dewpoint, or because it is above its cricondenbar). In such cases, there is no
liquid in
streams 35 and 37 (as shown by the dashed lines), so only the vapor from
separator
section 118f in stream 36 (FIGS. 2, 3, 10, and 11), the vapor from separator
12 in
stream 36 (FIGS. 6, 7, 14, and 15), or the cooled second portion 33a (FIGS. 4,
5, 8, 9,
12, 13, 16, and 17) flows to stream 38 to become the expanded substantially
condensed stream 38b supplied to processing assembly 118 between rectifying
section 118c and absorbing section 118d. In such circumstances, separator
section
118f in processing assembly 118 (FIGS. 2 through 5 and 10 through 13) or
separator
12 (FIGS. 6 through 9 and 14 through 17) may not be required.
[0048] Feed gas conditions, plant size, available equipment, or other
factors
may indicate that elimination of work expansion machine 15, or replacement
with an
alternate expansion device (such as an expansion valve), is feasible. Although
individual stream expansion is depicted in particular expansion devices,
alternative
expansion means may be employed where appropriate. For example, conditions may
warrant work expansion of the substantially condensed portion of the feed
stream
(stream 38a) or the substantially condensed recycle stream (stream 45a).
[0049] In accordance with the present invention, the use of external
refrigeration to supplement the cooling available to the inlet gas from the
distillation
vapor and liquid streams may be employed, particularly in the case of a rich
inlet gas.
In such cases, a heat and mass transfer means may be included in separator
section
118f (or a gas collecting means in such cases when the cooled feed stream 31a
or the
cooled first portion 32a contains no liquid) as shown by the dashed lines in
FIGS. 2
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through 5 and 10 through 13, or a heat and mass transfer means may be included
in
separator 12 as shown by the dashed lines in FIGS. 6 though 9 and 14 through
17.
This heat and mass transfer means may be comprised of a fin and tube type heat
exchanger, a plate type heat exchanger, a brazed aluminum type heat exchanger,
or
other type of heat transfer device, including multi-pass and/or multi-service
heat
exchangers. The heat and mass transfer means is configured to provide heat
exchange
between a refrigerant stream (e.g., propane) flowing through one pass of the
heat and
mass transfer means and the vapor portion of stream 31a (FIGS. 2, 3, 6, 7, 10,
11, 14,
and 15) or stream 32a (FIGS. 4, 5, 8, 9, 12, 13, 16, and 17) flowing upward,
so that
the refrigerant further cools the vapor and condenses additional liquid, which
falls
downward to become part of the liquid removed in stream 35. Alternatively,
conventional gas chiller(s) could be used to cool stream 32a, stream 33a,
and/or
stream 31a with refrigerant before stream 31a enters separator section 118f
(FIGS. 2,
3, 10, and 11) or separator 12 (FIGS. 6, 7, 14, and 15) or stream 32a enters
separator
section 118f (FIGS. 4, 5, 12, and 13) or separator 12 (FIGS. 8, 9, 16, and
17).
[0050] Depending on the temperature and richness of the feed gas and
the
amount of C2 components to be recovered in liquid product stream 44, there may
not
be sufficient heating available from stream 33 to cause the liquid leaving
demethanizing section 118e to meet the product specifications. In such cases,
the heat
and mass transfer means in demethanizing section 118e may include provisions
for
providing supplemental heating with heating medium as shown by the dashed
lines in
FIGS. 2 through 17. Alternatively, another heat and mass transfer means can be
included in the lower region of demethanizing section 118e for providing
supplemental heating, or stream 33 can be heated with heating medium before it
is
supplied to the heat and mass transfer means in demethanizing section 118e.
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100511 Depending on the type of heat transfer devices
selected for the heat
exchange means in the upper and lower regions of feed cooling section 118a, it
may
be possible to combine these heat exchange means in a single multi-pass and/or
multi-service heat transfer device. In such cases, the multi-pass and/or multi-
service
heat transfer device will include appropriate means for distributing,
segregating, and
collecting stream 32, stream 38, stream 45, and the distillation vapor stream
in order
to accomplish the desired cooling and heating.
100521 Some circumstances may favor providing additional
mass transfer in
the upper region of demethanizing section 118e. In such cases, a mass transfer
means
can be located below where expanded stream 39a (FIGS. 2, 3, 6, 7, 10, 11, 14,
and
15) or expanded stream 34a (FIGS. 4, 5, 8, 9, 12, 13, 16, and 17) enters the
lower
region of absorbing section 118d and above where cooled second portion 33a
leaves
the heat and mass transfer means in demethanizing section 118e.
100531 A less preferred option for the FIGS. 2, 3, 6, 7,
10, 11, 14, and 15
embodiments of the present invention is providing a separator vessel for
cooled first
portion 32a, a separator vessel for cooled second portion 33a, combining the
vapor
streams separated therein to form vapor stream 34, and combining the liquid
streams
separated therein to form liquid stream 35. Another less preferred option for
the
present invention is cooling stream 37 in a separate heat exchange means
inside feed
cooling section 118a in FIGS. 2, 3, 4, 5, 6, 7, 8, and 9 or a separate pass in
heat
exchanger 10 in FIGS. 10, 11, 12, 13, 14, 15, 16, and 17 (rather than
combining
stream 37 with stream 36 or stream 33a to form combined stream 38), expanding
the
cooled stream in a separate expansion device, and supplying the expanded
stream to
an intermediate region in absorbing section 118d.
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100541 It will be recognized that the relative amount of feed found in each
branch of the split vapor feed will depend on several factors, including gas
pressure,
feed gas composition, the amount of heat which can economically be extracted
from
the feed, and the quantity of horsepower available. More feed above absorbing
section 118d may increase recovery while decreasing power recovered from the
expander and thereby increasing the recompression horsepower requirements.
Increasing feed below absorbing section 118d reduces the horsepower
consumption
but may also reduce product recovery.
100551 The present invention provides improved recovery of C2 components,
C3 components, and heavier hydrocarbon components or of C3 components and
heavier hydrocarbon components per amount of utility consumption required to
operate the process. An improvement in utility consumption required for
operating
the process may appear in the form of reduced power requirements for
compression or
re-compression, reduced power requirements for external refrigeration, reduced
energy requirements for supplemental heating, or a combination thereof.
100561 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.
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