Note: Descriptions are shown in the official language in which they were submitted.
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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 streams. A
typical
analysis of a gas stream to be processed in accordance with this invention
would be, in
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approximate mole percent, 90.0% 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; 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;
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12/869,139; and 12/979,563 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.
[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
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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] The present invention employs a novel means of performing 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,
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VOCs, that contribute to greenhouse gases and may also be precursors to
atmospheric ozone
formation), eliminating these flanges reduces the potential for atmospheric
emissions that can
damage the environment.
[0009] In accordance with the present invention, it has been found that
C2 recoveries
in excess of 88% can be obtained. Similarly, in those instances where recovery
of C2
components is not desired, C3 recoveries in excess of 93% 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.
[0010] For a better understanding of the present invention, reference is
made to the
following examples and drawings. Referring to the drawings:
[0011] FIG. 1 is a flow diagram of a prior art natural gas processing
plant in
accordance with United States Patent No. 4,157,904;
[0012] FIG. 2 is a flow diagram of a natural gas processing plant in
accordance with
the present invention; and
[0013] FIGS. 3 through 17 are flow diagrams illustrating alternative
means of
application of the present invention to a natural gas stream.
[0014] 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
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non-hydrocarbon components and hence are generally larger than the sum of the
stream flow
rates for the hydrocarbon components. Temperatures indicated are approximate
values
rounded to the nearest degree. It should also be noted that the process design
calculations
performed for the purpose of comparing the processes depicted in the figures
are based on the
assumption of no heat leak from (or to) the surroundings to (or from) the
process. The quality
of commercially available insulating materials makes this a very reasonable
assumption and
one that is typically made by those skilled in the art.
[0015] 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
[0016] 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.
4,157,904. In this simulation of the process, inlet gas enters the plant at
101 F [39 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.
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[0017] The feed stream 31 is divided into two portions, streams 32 and
33. Stream 32
is cooled to -31 F [-35 C] in heat exchanger 10 by heat exchange with cool
residue gas
(stream 41a), while stream 33 is cooled to -37 F [-38 C] in heat exchanger 11
by heat
exchange with demethanizer reboiler liquids at 43 F [6 C] (stream 43) and side
reboiler
liquids at -47 F [-44 C] (stream 42). Streams 32a and 33a recombine to form
stream 31a,
which enters separator 12 at -33 F [-36 C] and 893 psia [6,155 kPa(a)] where
the vapor
(stream 34) is separated from the condensed liquid (stream 35).
[0018] The vapor (stream 34) from separator 12 is divided into two
streams, 36 and
39. Stream 36, containing about 32% 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 the cold residue gas (stream 41) where it is cooled to
substantial condensation.
The resulting substantially condensed stream 38a at -131 F [-90 C] is then
flash expanded
through expansion valve 14 to the operating pressure (approximately 410 psia
[2,827 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 -137 F [-94 C] and is
supplied to
separator section 18a in the upper region of fractionation tower 18. The
liquids separated
therein become the top feed to demethanizing section 18b.
[0019] The remaining 68% 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 -97 F [-72 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
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compressor (such as item 16) that can be used to re-compress the residue gas
(stream 41b),
for example. The partially condensed expanded stream 39a is thereafter
supplied as feed to
fractionation tower 18 at a mid-column feed point.
[0020] 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 -136 F [-93 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.
[0021] The liquid product stream 44 exits the bottom of the tower at 65 F
[19 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 residue gas (demethanizer overhead vapor stream 41) passes
countercurrently to the incoming feed gas in heat exchanger 13 where it is
heated to -44 F
[-42 C] (stream 41a) and in heat exchanger 10 where it is heated to 96 F [36
C] (stream
41b). The residue gas 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
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pressure. After cooling to 120 F [49 C] in discharge cooler 21, the residue
gas product
(stream 41e) 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).
[0022] 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 12,359 546 233 229 13,726
32 8,404 371 159 155 9,334
33 3,955 175 74 74 4,392
34 12,117 493 172 70 13,196
35 242 53 61 159 530
36 3,829 156 54 22 4,170
38 4,071 209 115 181 4,700
39 8,288 337 118 48 9,026
41 12,350 62 5 1 12,620
44 9 484 228 228 1,106
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Recoveries*
Ethane 88.54%
Propane 97.70%
Butanes+ 99.65%
Power
Residue Gas Compression 5,174 HP [ 8,506 kW]
* (Based on un-rounded flow rates)
DESCRIPTION OF THE INVENTION
[0023] 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.
[0024] 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 region of feed cooling section 118a. Stream 32 is cooled while
further heating
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the distillation vapor stream, with stream 32a leaving the heat exchange means
at -26 F
[-32 C].
[0025] The second portion, stream 33, enters a heat and mass transfer
means in
demethanizing section 118d 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 118c inside processing assembly 118, so that stream 33 is
cooled while
heating the distillation liquid stream, cooling stream 33a to -38 F [-39 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.
[0026] Streams 32a and 33a recombine to form stream 31a, which enters
separator
section 118e inside processing assembly 118 at -30 F [-34 C] and 898 psia
[6,189 kPa(a)],
whereupon the vapor (stream 34) is separated from the condensed liquid (stream
35).
Separator section 118e has an internal head or other means to divide it from
demethanizing
section 118d, so that the two sections inside processing assembly 118 can
operate at different
pressures.
[0027] The vapor (stream 34) from separator section 118e is divided into
two streams,
36 and 39. Stream 36, containing about 32% of the total vapor, is combined
with the
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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.
[0028] The resulting substantially condensed stream 38a at -130 F [-90 C]
is then
flash expanded through expansion valve 14 to the operating pressure
(approximately 415 psia
[2,861 kPa(a)]) of absorbing section 118c (an absorbing means) 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 38b leaving
expansion valve
14 reaches a temperature of -136 F [-94 C] and is supplied to separator
section 118b inside
processing assembly 118. The liquids separated therein are directed to
absorbing section
118c, while the remaining vapors combine with the vapors rising from absorbing
section
118c to form the distillation vapor stream that is heated in cooling section
118a.
[0029] The remaining 68% of the vapor from separator section 118e (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 118c, with the
work expansion
cooling the expanded stream 39a to a temperature of approximately -94 F [-70
C]. The
partially condensed expanded stream 39a is thereafter supplied as feed to the
lower region of
absorbing section 118c inside processing assembly 118.
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[0030] Absorbing section 118c contains a plurality of vertically spaced
trays, one or
more packed beds, or some combination of trays and packing. The trays and/or
packing in
absorbing section 118c 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 118c and the combined
liquid
continues downward into demethanizing section 118d. The stripping vapors
arising from
demethanizing section 118d combine with the vapor portion of the expanded
stream 39a and
rise upward through absorbing section 118c, to be contacted with the cold
liquid falling
downward to condense and absorb the C2 components, C3 components, and heavier
components from these vapors.
[0031] The distillation liquid flowing downward from the heat and mass
transfer
means in demethanizing section 118d inside processing assembly 118 has been
stripped of
methane and lighter components. The resulting liquid product (stream 44) exits
the lower
region of demethanizing section 118d and leaves processing assembly 118 at 67
F [20 C].
The distillation vapor stream arising from separator section 118b is warmed in
feed cooling
section 118a as it provides cooling to streams 32 and 38 as described
previously, and the
resulting residue gas stream 41 leaves processing assembly 118 at 96 F [36 C].
The residue
gas 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 120 F
[49 C] in discharge cooler 21, the residue gas product (stream 41c) flows to
the sales gas
pipeline at 915 psia [6,307 kPa(a)].
[0032] 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 12,359 546 233 229 13,726
32 8,651 382 163 160 9,608
33 3,708 164 70 69 4,118
34 12,139 498 176 74 13,234
35 220 48 57 155 492
36 3,860 158 56 24 4,208
37 220 48 57 155 492
38 4,080 206 113 179 4,700
39 8,279 340 120 50 9,026
41 12,350 62 5 1 12,625
44 9 484 228 228 1,101
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Recoveries*
Ethane 88.58%
Propane 97.67%
Butanes+ 99.64%
Power
Residue Gas Compression 4,829 HP [ 7,939 kW]
* (Based on un-rounded flow rates)
[0033] 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 nearly a 7% improvement over the
prior art of the
FIG. 1 process.
[0034] 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 118d 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, absorbing section 118c in processing assembly 118 of the present
invention can operate
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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
residue gas due to eliminating the interconnecting piping, results in a
significantly higher
pressure for the residue gas entering compressor 20, thereby reducing the
power required by
the present invention to restore the residue gas to pipeline pressure.
[0035] Second, using the heat and mass transfer means in demethanizing
section 118d
to simultaneously heat the distillation liquid leaving absorbing section 118c
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.
[0036] 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. 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.
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Other Embodiments
[0037] 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.
[0038] Some circumstances may favor supplying liquid stream 35 directly
to the
lower region of absorbing section 118c 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 118c and
the resulting
expanded liquid stream 40a is supplied as feed to the lower region of
absorbing section 118c
(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 (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
118c 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 118c as a single feed.
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[0039] 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 118d
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 118d 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 118d as stream 40a.
[0040] 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 118e
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 expanded substantially isentropically
to the
operating pressure of absorbing section 118c, whereupon expanded stream 34a is
supplied as
feed to the lower region of absorbing section 118c 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 and cooled to
substantial
condensation. The substantially condensed stream 38a is flash expanded through
expansion
valve 14 to the operating pressure of absorbing section 118c, whereupon
expanded stream
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38b is supplied to separator section 118b inside processing assembly 118. 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 118c via expansion valve 17. Other circumstances
may favor
sending all of liquid stream 35 to the lower region of absorbing section 118c
via expansion
valve 17.
[0041] 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 118e 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.
[0042] Depending on the quantity of heavier hydrocarbons in the feed gas
and the
feed gas pressure, the cooled feed stream 31a entering separator section 118e
in 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 118e 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 118e 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 separator section 118b in
processing
assembly 118. In such circumstances, separator section 118e 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.
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[0043] 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).
[0044] 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 118e (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 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 118e (FIGS. 2, 3, 10, and 11) or separator 12 (FIGS. 6, 7,
14, and 15) or
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stream 32a enters separator section 118e (FIGS. 4, 5, 12, and 13) or separator
12 (FIGS. 8, 9,
16, and 17).
[0045] 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 118d to
meet the product specifications. In such cases, the heat and mass transfer
means in
demethanizing section 118d 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
118d 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 118d.
[0046] 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, and the
distillation vapor stream in order to accomplish the desired cooling and
heating.
[0047] Some circumstances may favor providing additional mass transfer in
the upper
region of demethanizing section 118d. 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 118c
and above where cooled second portion 33a leaves the heat and mass transfer
means in
demethanizing section 118d.
[0048] 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
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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 118c.
[0049] 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 118c may
increase
recovery while decreasing power recovered from the expander and thereby
increasing the
recompression horsepower requirements. Increasing feed below absorbing section
118c
reduces the horsepower consumption but may also reduce product recovery.
[0050] 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.
[0051] 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
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feed, or other requirements without departing from the spirit of the present
invention as
defined by the following claims.
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