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] 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
propylene, propane, and heavier hydrocarbons from such gas streams. A typical
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analysis of a gas stream to be processed in accordance with this invention
would be,
in approximate mole percent, 88.4% methane, 6.2% ethane and other C2
components,
2.6% propane and other C3 components, 0.3% iso-butane, 0.6% 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 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; and 13/048,315 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 C3+
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 (deethanizer) column. In the column, the
expansion
cooled stream(s) is (are) distilled 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 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
expanded
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stream then enters an absorbing section in the column and is contacted with
cold
liquids to absorb the C3 components and heavier components from the vapor
portion
of the expanded stream. The liquids from the absorbing section are then
directed to
the deethanizing section of the column.
[0007] A distillation vapor stream is withdrawn from the upper region of the
deethanizing section and is cooled by heat exchange relation with the overhead
vapor
stream from the absorbing section, condensing at least a portion of the
distillation
vapor stream. The condensed liquid is separated from the cooled distillation
vapor
stream to produce a cold liquid reflux stream that is directed to the upper
region of the
absorbing section, where the cold liquids can contact the vapor portion of the
expanded stream as described earlier. The vapor portion (if any) of the cooled
distillation vapor stream and the overhead vapor from the absorbing section
combine
to form the residual methane and C2 component product gas.
[0008] The separation that takes place in this process (producing a residue
gas
leaving the process which contains substantially all of the methane and C2
components in the feed gas with essentially none of the C3 components and
heavier
hydrocarbon components, and a bottoms fraction leaving the deethanizer which
contains substantially all of the C3 components and heavier hydrocarbon
components
with essentially no methane, C2 components or more volatile components)
consumes
energy for feed gas cooling, for reboiling the deethanizing section, for
refluxing the
absorbing section, and/or for re-compressing the residue gas.
[0009] 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
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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 formation), eliminating these flanges reduces the potential
for
atmospheric emissions that can damage the environment.
[0010] In accordance with the present invention, it has been found that C3
recoveries in excess of 99.6% can be obtained while providing essentially
complete
rejection of C2 components to the residue gas stream. In addition, the present
invention makes possible essentially 100% separation of C2 components and
lighter
components from the 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.
[0011] For a better understanding of the present invention, reference is made
to the following examples and drawings. Referring to the drawings:
[0012] FIG. 1 is a flow diagram of a prior art natural gas processing plant in
accordance with United States Patent No. 5,799,507;
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[0013] FIG. 2 is a flow diagram of a natural gas processing plant in
accordance with the present invention; and
[0014] FIGS. 3 through 21 are flow diagrams illustrating alternative means of
application of the present invention to a natural gas stream.
[0015] 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 materials makes this a very reasonable assumption and one that is
typically
made by those skilled in the art.
[0016] 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.
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DESCRIPTION OF THE PRIOR ART
[0017] FIG. 1 is a process flow diagram showing the design of a processing
plant to recover C3+ components from natural gas using prior art according to
U.S.
Pat. No. 5,799,507. In this simulation of the process, inlet gas enters the
plant at
110 F [43 C] and 885 psia [6,100 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.
[0018] The feed stream 31 is cooled in heat exchanger 10 by heat exchange
with cool residue gas (stream 44), flash expanded separator liquids (stream
35a), and
distillation liquids at -105 F [-76 C] (stream 43). The cooled stream 31a
enters
separator 11 at -34 F [-36 C] and 875 psia [6,031 kPa(a)] where the vapor
(stream 34)
is separated from the condensed liquid (stream 35). The separator liquid
(stream 35)
is expanded to slightly above the operating pressure (approximately 375 psia
[2,583 kPa(a)]) of fractionation tower 15 by expansion valve 12, cooling
stream 35a
to -65 F [-54 C]. Stream 35a enters heat exchanger 10 to supply cooling to the
feed
gas as described previously, heating stream 35b to 105 F [41 C] before it is
supplied
to fractionation tower 15 at a lower mid-column feed point.
[0019] The vapor (stream 34) from separator 11 enters a work expansion
machine 13 in which mechanical energy is extracted from this portion of the
high
pressure feed. The machine 13 expands the vapor substantially isentropically
to the
operating pressure of fractionation tower 15, with the work expansion cooling
the
expanded stream 34a to a temperature of approximately -100 F [-74 C]. The
typical
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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 14)
that can be
used to re-compress the heated residue gas (stream 44a), for example. The
partially
condensed expanded stream 34a is thereafter supplied as feed to fractionation
tower
15 at an upper mid-column feed point.
[0020] The deethanizer in tower 15 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 deethanizer tower consists of two
sections: an
upper absorbing (rectification) section 15a that contains the trays and/or
packing to
provide the necessary contact between the vapor portion of the expanded stream
34a
rising upward and cold liquid falling downward to condense and absorb the C3
components and heavier components; and a lower stripping section 15b 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 15b also
includes
at least one reboiler (such as reboiler 16) which heats and vaporizes 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 37, of methane, C2 components, and
lighter
components. Stream 34a enters deethanizer 15 at a mid-column feed position
located
in the lower region of absorbing section 15a of deethanizer 15. The liquid
portion of
expanded stream 34a commingles with liquids falling downward from absorbing
section 15a and the combined liquid continues downward into stripping section
15b
of deethanizer 15. The vapor portion of expanded stream 34a rises upward
through
absorbing section 15a and is contacted with cold liquid falling downward to
condense
and absorb the C3 components and heavier components.
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[0021] A portion of the distillation vapor (stream 38) is withdrawn from the
upper region of stripping section 15b. This stream is then cooled and
partially
condensed (stream 38a) in exchanger 17 by heat exchange with cold deethanizer
overhead stream 36 which exits the top of deethanizer 15 at -109 F [-79 C].
The cold
deethanizer overhead stream is warmed to approximately -33 F [-66 C] (stream
36a)
as it cools stream 38 from -30 F [-35 C] to about -103 F [-75 C] (stream 38a).
[0022] The operating pressure in reflux separator 18 is maintained slightly
below the operating pressure of deethanizer 15. This pressure difference
provides the
driving force that allows distillation vapor stream 38 to flow through heat
exchanger
17 and thence into the reflux separator 18 wherein the condensed liquid
(stream 40) is
separated from the uncondensed vapor (stream 39). The uncondensed vapor stream
39 combines with the warmed deethanizer overhead stream 36a from exchanger 17
to
form cool residue gas stream 44 at -37 F [-38 C].
[0023] The liquid stream 40 from reflux separator 18 is pumped by pump 19
to a pressure slightly above the operating pressure of deethanizer 15. The
resulting
stream 40a is then divided into two portions. The first portion (stream 41) is
supplied
as cold top column feed (reflux) to the upper region of absorbing section 15a
of
deethanizer 15. This cold liquid causes an absorption cooling effect to occur
inside
the absorbing (rectification) section 15a of deethanizer 15, wherein the
saturation of
the vapors rising upward through the tower by vaporization of liquid methane
and
ethane contained in stream 41 provides refrigeration to the section. Note
that, as a
result, both the vapor leaving the upper region (overhead stream 36) and the
liquids
leaving the lower region (distillation liquid stream 43) of absorbing section
15a are
colder than the either of the feed streams (streams 41 and stream 34a) to
absorbing
section 15a. This absorption cooling effect allows the tower overhead (stream
36) to
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provide the cooling needed in heat exchanger 17 to partially condense the
distillation
vapor stream (stream 38) without operating stripping section 15b at a pressure
significantly higher than that of absorbing section 15a. This absorption
cooling effect
also facilitates reflux stream 41 condensing and absorbing the C3 components
and
heavier components in the distillation vapor flowing upward through absorbing
section 15a. The second portion (stream 42) of pumped stream 40a is supplied
to the
upper region of stripping section 15b of deethanizer 15 where the cold liquid
acts as
reflux to absorb and condense the C3 components and heavier components flowing
upward from below so that distillation vapor stream 38 contains minimal
quantities of
these components.
[0024] A distillation liquid stream 43 from deethanizer 15 is withdrawn from
the lower region of absorbing section 15a and is routed to heat exchanger 10
where it
is heated as it provides cooling of the incoming feed gas as described
earlier.
Typically the flow of this liquid from the deethanizer is via thermosiphon
circulation,
but a pump could be used. The liquid stream is heated to -4 F [-20 C],
partially
vaporizing stream 43a before it is returned as a mid-column feed to
deethanizer 15, in
the middle region of stripping section 15b.
[0025] In stripping section 15b of deethanizer 15, the feed streams are
stripped of their methane and C2 components. The resulting liquid product
stream 37
exits the bottom of the tower at 201 F [94 C] based on a typical
specification of an
ethane to propane ratio of 0.048:1 on a molar basis in the bottom product. The
cool
residue gas (stream 44) passes countercurrently to the incoming feed gas in
heat
exchanger 10 where it is heated to 98 F [37 C] (stream 44a). The residue gas
is then
re-compressed in two stages. The first stage is compressor 14 driven by
expansion
machine 13. The second stage is compressor 20 driven by a supplemental power
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source which compresses the residue gas (stream 44c) to sales line pressure.
After
cooling to 120 F [49 C] in discharge cooler 21, residue gas stream 44d 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).
[0026] 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 19,419 1,355 565 387 21,961
34 18,742 1,149 360 98 20,573
35 677 206 205 289 1,388
36 18,400 1,242 3 0 19,869
38 2,759 1,758 15 0 4,602
39 1,019 86 0 0 1,116
40 1,740 1,672 15 0 3,486
41 1,044 1,003 9 0 2,092
42 696 669 6 0 1,394
43 1,388 911 365 98 2,796
44 19,419 1,328 3 0 20,985
37 0 27 562 387 976
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Recoveries*
Propane 99.56%
Butanes+ 100.00%
Power
Residue Gas Compression 9,868 HP [ 16,223 kW]
Reflux Pump 19 HP [ 31 kW]
Totals 9,887 HP [ 16,254 kW]
* (Based on un-rounded flow rates)
DESCRIPTION OF THE INVENTION
[0027] 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. I process to illustrate the advantages
of the
present invention.
[0028] In the simulation of the FIG. 2 process, inlet gas enters the plant as
stream 31 and enters a heat exchange means in feed cooling section 115a inside
processing assembly 115. 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 31 flowing through one pass of the heat exchange
means and flash expanded separator liquids (stream 35a) and a residue gas
stream
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from condensing section 115b inside processing assembly 115. Stream 31 is
cooled
while heating the flash expanded separator liquids and the residue gas stream.
A first
portion (stream 32) of stream 31 is withdrawn from the heat exchange means
after
stream 31 has been partially cooled to 25 F [-4 C], while the remaining second
portion (stream 33) is further cooled so that it leaves the heat exchange
means at
-20 F [-29 C].
[0029] Separator section 115e has an internal head or other means to divide it
from deethanizing section 115d, so that the two sections inside processing
assembly
115 can operate at different pressures. The first portion (stream 32) of
stream 31
enters the lower region of separator section 115e at 875 psia [6,031 kPa(a)]
where any
condensed liquid is separated from the vapor before the vapor is directed into
a heat
and mass transfer means inside separator section 115e. 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 the vapor
portion of
stream 32 flowing upward through one pass of the heat and mass transfer means
and
distillation liquid stream 43 from absorbing section 115c inside processing
assembly
115 flowing downward, so that the vapor is cooled while heating the
distillation liquid
stream. As the vapor stream is cooled, a portion of it may be condensed and
fall
downward while the remaining vapor continues flowing upward through the heat
and
mass transfer means. The heat and mass transfer means provides continuous
contact
between the condensed liquid and the vapor so that it also functions to
provide mass
transfer between the vapor and liquid phases to provide partial rectification
of the
vapor.
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[00301 The second portion (stream 33) of stream 31 enters separator section
115e inside processing assembly 115 above the heat and mass transfer means.
Any
condensed liquid is separated from the vapor and commingles with any liquid
that is
condensed from the vapor portion of stream 32 flowing up through the heat and
mass
transfer means. The vapor portion of stream 33 combines with the vapor leaving
the
heat and mass transfer means to form stream 34, which exits separator section
115e at
-31 F [-35 C]. The liquid portions (if any) of streams 32 and 33 and any
liquid
condensed from the vapor portion of stream 32 in the heat and mass transfer
means
combine to form stream 35, which exits separator section 115e at -15 F [-26
C]. It is
expanded to slightly above the operating pressure (approximately 383 psia
[2,639 kPa(a)]) of deethanizing section 115d inside processing assembly 115 by
expansion valve 12, cooling stream 35a to -42 F [-41'C]. Stream 35a enters the
heat
exchange means in feed cooling section 115a to supply cooling to the feed gas
as
described previously, heating stream 35b to 103 F [39 C] before it is supplied
to
deethanizing section 115d inside processing assembly 115 at a lower mid-column
feed point.
[00311 The vapor (stream 34) from separator section 115e enters a work
expansion machine 13 in which mechanical energy is extracted from this portion
of
the high pressure feed. The machine 13 expands the vapor substantially
isentropically
to the operating pressure (approximately 380 psia [2,618 kPa(a)]) of absorbing
section
115c, with the work expansion cooling the expanded stream 34a to a temperature
of
approximately -98 F [-72 C]. The partially condensed expanded stream 34a is
thereafter supplied as feed to the lower region of absorbing section 115c
inside
processing assembly 115.
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[0032] Absorbing section 115c contains 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 absorbing section 115c provide
the
necessary contact between the vapors rising upward and cold liquid falling
downward.
The vapor portion of expanded stream 34a rises upward through the absorbing
means
in absorbing section 115c to be contacted with the cold liquid falling
downward to
condense and absorb most of the C3 components and heavier components from
these
vapors. The liquid portion of expanded stream 34a commingles with liquids
falling
downward from the absorbing means in absorbing section 115c to form
distillation
liquid stream 43, which is withdraw from the lower region of absorbing section
115c
at -102 F [-74 C]. The distillation liquid is heated to -9 F [-23 C] as it
cools the
vapor portion of stream 32 in separator section 115e as described previously,
with the
heated distillation liquid stream 43a thereafter supplied to deethanizing
section 115d
inside processing assembly 115 at an upper mid-column feed point. Typically
the
flow of this liquid from absorbing section 115c through the heat and mass
transfer
means in separator section 115e to deethanizing section 115d is via
thermosiphon
circulation, but a pump could be used.
[0033] Absorbing section 115c has an internal head or other means to divide it
from deethanizing section 115d, so that the two sections inside processing
assembly
115 can operate with the pressure of deethanizing section 115d slightly higher
than
that of absorbing section 115c. This pressure difference provides the driving
force
that allows a first distillation vapor stream (stream 38) to be withdrawn from
the
upper region of deethanizing section 115d and directed to the heat exchange
means in
condensing section 115b inside processing assembly 115. This heat exchange
means
may likewise be comprised of a fin and tube type heat exchanger, a plate type
heat
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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 first distillation vapor
stream
38 flowing through one pass of the heat exchange means and a second
distillation
vapor stream arising from absorbing section 115c inside processing assembly
115.
The second distillation vapor stream is heated while it cools and at least
partially
condenses stream 38, which thereafter exits the heat exchange means and is
separated
into its respective vapor and liquid phases. The vapor phase (if any) combines
with
the heated second distillation vapor stream exiting the heat exchange means to
form
the residue gas stream that provides cooling in feed cooling section 115a as
described
previously. The liquid phase is divided into two portions, streams 41 and 42.
[0034] The first portion (stream 41) is supplied as cold top column feed
(reflux) to the upper region of absorbing section 115c inside processing
assembly 115
by gravity flow. This cold liquid causes an absorption cooling effect to occur
inside
absorbing (rectification) section 115a, wherein the saturation of the vapors
rising
upward through the tower by vaporization of liquid methane and ethane
contained in
stream 41 provides refrigeration to the section. This absorption cooling
effect allows
the second distillation vapor stream to provide the cooling needed in the heat
exchange means in condensing section 115b to partially condense the first
distillation
vapor stream (stream 38) without operating deethanizing section 115d at a
pressure
significantly higher than that of absorbing section 115c. This absorption
cooling
effect also facilitates reflux stream 41 condensing and absorbing the C3
components
and heavier components in the distillation vapor flowing upward through
absorbing
section 115c. The second portion (stream 42) of the liquid phase separated in
condensing section 115b is supplied as cold top column feed (reflux) to the
upper
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region of deethanizing section 115d inside processing assembly 115 by gravity
flow,
so that the cold liquid acts as reflux to absorb and condense the C3
components and
heavier components flowing upward from below so that distillation vapor stream
38
contains minimal quantities of these components.
[0035] Deethanizing section 115d inside processing assembly 115 contains a
mass transfer 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
deethanizing section 115d provide the necessary contact between the vapors
rising
upward and cold liquid falling downward. Deethanizing section 115d also
includes a
heat and mass transfer means beneath the mass transfer means. 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 a
heating
medium flowing through one pass of the heat and mass transfer means and a
distillation liquid stream flowing downward from the mass transfer means in
deethanizing section 115d, so that the distillation liquid stream is heated.
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 37 of methane, C2 components, and lighter components.
The
resulting liquid product (stream 37) exits the lower region of deethanizing
section
115d and leaves processing assembly 115 at 203 F [95 C].
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[0036] The second distillation vapor stream arising from absorbing section
115c is warmed in condensing section 115b as it provides cooling to stream 38
as
described previously. The warmed second distillation vapor stream combines
with
any vapor separated from the cooled first distillation vapor stream 38 as
described
previously. The resulting residue gas stream is heated in feed cooling section
115a as
it provides cooling to stream 31 as described previously, whereupon residue
gas
stream 44 leaves processing assembly 115 at 104 F [40 C]. The residue gas
stream is
then re-compressed in two stages, compressor 14 driven by expansion machine 13
and
compressor 20 driven by a supplemental power source. After cooling to 120 F
[49 C] in discharge cooler 21, residue gas stream 44c 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).
[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 19,419 1,355 565 387 21,961
32 4,855 339 141 97 5,490
33 14,564 1,016 424 290 16,471
34 18,870 1,135 348 104 20,683
35 549 220 217 283 1,278
38 2,398 1,544 13 0 4,015
41 1,018 868 8 0 1,924
42 737 628 5 0 1,394
43 1,112 723 353 104 2,320
44 19,419 1,328 3 0 20,984
37 0 27 562 387 977
Recoveries*
Propane 99.63%
Butanes+ 100.00%
Power
Residue Gas Compression 9,363 HP [ 15,393 kW]
* (Based on un-rounded flow rates)
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[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 propane recovered per unit of power), the present
invention represents more than a 5% 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
three
factors. First, the compact arrangement of the heat exchange means in feed
cooling
section 115a and condensing section 115b in processing assembly 115 eliminates
the
pressure drop imposed by the interconnecting piping found in conventional
processing
plants. The result is that the residue gas flowing to compressor 14 is at
higher
pressure for the present invention compared to the prior art, so that the
residue gas
entering compressor 20 is at significantly higher pressure, 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 deethanizing section
115d to simultaneously heat the distillation liquid leaving the mass transfer
means in
deethanizing section 115d 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.
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[0041] Third, using the heat and mass transfer means in separator section 115e
to simultaneously cool the vapor portion of stream 32 while condensing the
heavier
hydrocarbon components from the vapor provides partial rectification of stream
34
before it is subsequently expanded and supplied as feed to absorbing section
115c. As
a result, less reflux flow (stream 41) is required to rectify the expanded
stream 34a to
remove the C3 components and heavier hydrocarbon components from it, as seen
by
comparing the flow rate of stream 41 in Tables I and II.
[0042] 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 115 of the present invention replaces six separate
equipment
items in the prior art (heat exchangers 10 and 17, separator 11, reflux
separator 18,
reflux pump 19; and fractionation tower 15 in FIG. 1) with a single equipment
item
(processing assembly 115 in FIG. 2). This reduces the plot space requirements,
eliminates the interconnecting piping, and eliminates the power consumed by
the
reflux pump, reducing the capital cost and operating 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
[0043] Some circumstances may favor eliminating feed cooling section 115a
and condensing section 115b from processing assembly 115, and using one or
more
heat exchange means external to the processing assembly for feed cooling and
reflux
condensing, such as heat exchangers 23 and 17 shown in FIGS. 14 through 21.
Such
an arrangement allows processing assembly 115 to be smaller, which may reduce
the
overall plant cost and/or shorten the fabrication schedule in some cases. Note
that in
all cases exchangers 23 and 17 are 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. In
some cases, it may be advantageous to combine the feed cooling and reflux
condensing in a single multi-service heat exchanger. With heat exchanger 17
external
to the processing assembly, reflux separator 18 and pump 19 will typically be
needed
to separate condensed liquid stream 40 and deliver at least a portion of it to
absorbing
section 115c as reflux.
[0044] As described earlier for the embodiment of the present invention
shown in FIG. 2, the first distillation vapor stream 38 is partially condensed
and the
resulting condensate used to absorb valuable C3 components and heavier
components
from the vapors leaving the work expansion machine. However, the present
invention
is not limited to this embodiment. It may be advantageous, for instance, to
treat only
a portion of the outlet vapor from the work expansion machine in this manner,
or to
use only a portion of the condensate as an absorbent, in cases where other
design
considerations indicate portions of the expansion machine outlet or the
condensate
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should bypass absorbing section 115c of processing assembly 115. Feed gas
conditions, plant size, available equipment, or other factors may indicate
that
elimination of work expansion machine 13, or replacement with an alternate
expansion device (such as an expansion valve), is feasible, or that total
(rather than
partial) condensation of first distillation vapor stream 38 in condensing
section 115b
inside processing assembly 115 (FIGS. 2 through 13) or heat exchanger 17
(FIGS. 14
through 21) is possible or is preferred. It should also be noted that,
depending on the
composition of the feed gas stream, it may be advantageous to use external
refrigeration to provide partial cooling of first distillation vapor stream 38
in
condensing section 115b (FIGS. 2 through 13) or heat exchanger 17 (FIGS. 14
through 21).
[0045] In some circumstances, it may be advantageous to use an external
separator vessel to separate cooled first and second portions 32 and 33 or
cooled feed
stream 31a, rather than including separator section 115e in processing
assembly 115.
As shown in FIGS. 8 and 18, a heat and mass transfer means in separator 11 can
be
used to separate cooled first and second portions 32 and 33 into vapor stream
34 and
liquid stream 35. Likewise, as shown in FIGS. 9 through 13 and 19 through 21,
separator 11 can be used to separate cooled feed stream 31 a into vapor stream
34 and
liquid stream 35.
[0046] The use and distribution of the liquid stream 35 from separator section
115e or separator 11 and distillation liquid stream 43 from absorbing section
115c for
process heat exchange, the particular arrangement of heat exchangers for
cooling feed
gas (streams 31 and/or 32) and first distillation vapor stream 38, and the
choice of
process streams for specific heat exchange services must be evaluated for each
particular application. For instance, FIGS. 4 through 6, 10 through 12, 16,
and 20
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depict using distillation liquid stream 43 to supply a portion of the cooling
of first
distillation vapor stream 38 in condensing section 115b (FIGS. 4, 5, 10, and
11), heat
exchanger 10 (FIGS. 6 and 12), or heat exchanger 17 (FIGS. 16 and 20). In such
cases, a heat and mass transfer means may not be needed in separator section
115e
(FIGS. 4 through 6 and 16) or separator 11 (FIGS. 10 through 12 and 20). In
the
embodiments shown in FIGS. 4 and 10, a pump 22 is used to deliver distillation
liquid
stream 43 to the heat exchange means in condensing section 115b. In the
embodiments shown in FIGS. 5 and 11, condensing section 115b is located below
absorbing section 115c in processing assembly 115 so that flow of distillation
liquid
stream 43 is via thermosiphon circulation. In the embodiments shown in FIGS. 6
and
12, a heat exchanger 10 external to processing assembly 115 is employed and
feed
cooling section 115a is located below absorbing section 115c in processing
assembly
115 so that flow of distillation liquid stream 43 is via thermosiphon
circulation. (The
embodiments shown in FIGS. 5, 6, 11, and 12 use reflux pump 19 to supply
reflux to
locations above the point in processing assembly 115 where the liquid phase
condensed from stream 38 is collected.) In the embodiments shown in FIGS. 16
and
20, thermosiphon circulation may be sufficient to allow distillation liquid
stream 43 to
flow through heat exchanger 17, or pump 22 may be needed to circulate stream
43.
Some circumstances may favor using distillation liquid stream 43 to cool
stream 32 in
a heat exchanger external to processing assembly 115, such as heat exchanger
10
depicted in FIGS. 3, 9, 15, and 19. Still other circumstances may favor no
heating of
distillation liquid stream 43 at all, and instead using distillation liquid
stream 43 as the
reflux to the upper region of deethanizing section 115d as shown in FIGS. 7,
13, 17,
and 21. (For the embodiment shown in FIGS. 13 and 21, pump 22 may be needed
because gravity flow of stream 43 may not be possible.)
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[0047] Depending on the quantity of heavier hydrocarbons in the feed gas and
the feed gas pressure, the cooled first and second portions 32 and 33 entering
separator section 115e in FIGS. 2 and 14 or separator 11 in FIGS. 8 and 18 (or
the
cooled feed stream 31a entering separator section 115e in FIGS. 3 through 7
and 15
through 17 or separator 11 in FIGS. 9 through 13 and 19 through 21) 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 stream 35 (as shown by the dashed lines).
In such
circumstances, separator section 115e in processing assembly 115 (FIGS. 2
through 7
and 14 through 17) or separator 11 (FIGS. 8 through 13 and 18 through 21) may
not
be required.
[0048] In accordance with the present invention, the use of external
refrigeration to supplement the cooling available to the inlet gas and/or the
first
distillation vapor stream from the second distillation vapor stream and the
distillation
liquid stream may be employed, particularly in the case of a rich inlet gas.
In such
cases where additional inlet gas cooling is desired, a heat and mass transfer
means
may be included in separator section 115e (or a gas collecting means in such
cases
when the cooled first and second portions 32 and 33 or the cooled feed stream
31a
contains no liquid) as shown by the dashed lines in FIGS. 3 through 7 and 15
through
17, or a heat and mass transfer means may be included in separator 11 as shown
by
the dashed lines in FIGS. 9 though 13 and 19 through 21. 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
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vapor portion of stream 31a 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. As shown by the dashed lines in FIGS. 2, 8, 14,
and 18,
the heat and mass transfer means in separator section 115e (FIGS. 2 and 14) or
separator 11 (FIGS. 8 and 18) may include provisions for providing
supplemental
cooling with refrigerant. Alternatively, conventional gas chiller(s) could be
used to
cool stream 32, stream 33, and/or stream 31a with refrigerant before streams
32 and
33 enter separator section 115e (FIGS. 2 and 14) or separator 11 (FIGS. 8 and
18) or
stream 31a enters separator section 115e (FIGS. 3 through 7 and 15 through 17)
or
separator 11 (FIGS. 9 through 13 and 19 through 21). In cases where additional
cooling of the first distillation vapor stream is desired, the heat exchange
means in
condensing section 115b of processing assembly 115 (FIGS. 2 through 5, 7
through
11, and 13), heat exchanger 10 (FIGS. 6 and 12), or heat exchanger 17 (FIGS.
14
through 21) may include provisions for providing supplemental cooling with
refrigerant as shown by the dashed lines.
[0049] Depending on the type of heat transfer devices selected for the heat
exchange means in feed cooling section 115a and condensing section 115b (FIGS.
2
through 5, 7 through 11, and 13), 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 31, stream 32,
stream 33,
first distillation vapor stream 38, any vapor separated from the cooled stream
38, and
the second distillation vapor stream in order to accomplish the desired
cooling and
heating.
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[0050] It will also be recognized that the relative amount of condensed liquid
that is split between streams 41 and 42 in FIGS. 2 through 6, 8 through 12, 14
through
16, and 18 through 20 will depend on several factors, including gas pressure,
feed gas
composition, and the quantity of horsepower available. The optimum split
generally
cannot be predicted without evaluating the particular circumstances for a
specific
application of the present invention. Some circumstances may favor feeding all
of the
condensed liquid to the upper region of absorbing section 115c in stream 41
and none
to the upper region of deethanizing section 115d in stream 42, as shown by the
dashed
lines for stream 42. In such cases, the heated distillation liquid stream 43a
may be
supplied to the upper region of deethanizing section 115d to serve as reflux.
[0051] The present invention provides improved recovery 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 tower reboiling, or a combination thereof.
[0052] 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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