Note: Descriptions are shown in the official language in which they were submitted.
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HYDROCARBON GAS PROCESSING
BACKGROUND OF THE INVENTION
100011 This invention relates to a process for the separation of a gas
containing
hydrocarbons. The applicants claim the benefits under Title 35, United States
Code,
Section 119(e) of prior U.S. Provisional Application Numbers 60/848,299 which
was
filed on September 28, 2006 and 60/897,683 which was filed on January 25,
2007.
10002] 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,
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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.
[0003] 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 approximate mole percent, 90.5% methane, 4.1% ethane and other C2
components, 1.3% propane and other C3 components, 0.4% iso-butane, 0.3% normal
butane, 0.5% pentanes plus, and 2.6% carbon dioxide, with the balance made up
of
nitrogen. Sulfur containing gases are also sometimes present.
[0004] 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, for processes that can provide efficient recoveries with lower
capital investment
and lower operating costs, and for processes that can be easily adapted or
adjusted to vary
the recovery of a specific component over a broad range. 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.
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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.
[0005] 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; 5,568,737; 5,771,712; 5,799,507; 5,881,569; 5,890,378;
5,983,664;
6,182,469; 6,712,880; 6,915,662; 7,191,617; 7,219,513; reissue U.S. Patent No.
33,408;
and co-pending application no. 11/430,412 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 patents and applications).
[0006] 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+ or 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
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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.
[0007] If the feed gas is not totally condensed (typically it is not), a
portion of 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 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.
[0008] 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.
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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 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.
[0009] 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, however, this
ideal
situation is not obtained for two main reasons. The first reason is that 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 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 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 C3 components, C4 components, and heavier
hydrocarbon components from the vapors.
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[0010] The second reason that this ideal situation cannot be obtained is
that
carbon dioxide contained in the feed gas fractionates in the demethanizer and
can build
up to concentrations of as much as 5% to 10% or more in the tower even when
the feed
gas contains less than 1% carbon dioxide. At such high concentrations,
formation of
solid carbon dioxide can occur depending on temperatures, pressures, and the
liquid
solubility. It is well known that natural gas streams usually contain carbon
dioxide,
sometimes in substantial amounts. If the carbon dioxide concentration in the
feed gas is
high enough, it becomes impossible to process the feed gas as desired due to
blockage of
the process equipment with solid carbon dioxide (unless carbon dioxide removal
equipment is added, which would increase capital cost substantially). The
present
invention provides a means for generating a liquid reflux stream that will
improve the
recovery efficiency for the desired products while simultaneously
substantially mitigating
the problem of carbon dioxide icing.
[0011] 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
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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 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, 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. Unfortunately, these processes
require
the use of a large amount of compression power to provide the motive force for
recycling
the reflux stream to the demethanizer, adding to both the capital cost and the
operating
cost of facilities using these processes.
[0012] The
present invention also employs an upper rectification section (or a
separate rectification column in some embodiments). However, the reflux stream
for this
rectification section is provided by using a side draw of the vapors rising in
a lower
portion of the tower. By modestly elevating its pressure, a significant
quantity of liquid
can be condensed in this side draw stream, often using only the refrigeration
available in
the cold vapor leaving the upper rectification section. This condensed liquid,
which is
predominantly liquid methane, can then be used to absorb C2 components, C3
components, C4 components, and heavier hydrocarbon components from the vapors
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rising through the upper rectification section and thereby capture these
valuable
components in the bottom liquid product from the demethanizer.
[0013] Heretofore, such a side draw feature has been employed in C2+
recovery
systems, as illustrated in the assignee's U.S. Patent No. 7,191,617.
Surprisingly,
applicants have found that elevating the pressure of the side draw feature of
the assignee's
U.S. Patent No. 7,191,617 invention improves C3+ recoveries without
sacrificing C2
component recovery levels and improves the system efficiency, while
simultaneously
substantially mitigating the problem of carbon dioxide icing.
[0014] In accordance with the present invention, it has been found that
C3 and
C4+ recoveries in excess of 99 percent can be obtained with no loss in C2+
component
recovery. The present invention provides the further advantage of being able
to maintain
in excess of 99 percent recovery of the C3 and C4+ components as the recovery
of C2
components is adjusted from high to low values. In addition, the present
invention makes
possible essentially 100 percent separation of methane and lighter components
from the
C2 components and heavier components while maintaining the same recovery
levels as
the prior art and improving the safety factor with respect to the danger of
carbon dioxide
icing. 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.
[0015] For a better understanding of the present invention, reference is
made to
the following examples and drawings. Referring to the drawings:
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[0016] FIG. 1 is a flow diagram of a prior art natural gas processing
plant in
accordance with United States Patent No. 7,191,617;
[0017] FIG. 2 is a flow diagram of a natural gas processing plant in
accordance
with the present invention;
[0018] FIG. 3 is a concentration-temperature diagram for carbon dioxide
showing
the effect of the present invention;
[0019] FIG. 4 is a flow diagram illustrating an alternative means of
application of
the present invention to a natural gas stream;
[0020] FIG. 5 is a concentration-temperature diagram for carbon dioxide
showing
the effect of the present invention with respect to the process of FIG. 4;
[0021] FIGS. 6 through 9 are flow diagrams illustrating alternative
means of
application of the present invention to a natural gas stream; and
[0022] FIG. 10 is a partial flow diagram illustrating alternative means
of
accomplishing the splitting of the vapor feed in accordance with the present
invention.
[0023] 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.
[0024] 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
[0025] 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
assignee's U.S.
Pat. No. 7,191,617. In this simulation of the process, inlet gas enters the
plant at 120 F
[49 C] and 1040 psia [7,171 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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[0026] The feed stream 31 is cooled in heat exchanger 10 by heat
exchange with
cool residue gas at -28 F [-33 C] (stream 48a), demethanizer reboiler liquids
at 35 F
[2 C] (stream 41), demethanizer lower side reboiler liquids at -10 F [-23 C]
(stream 40),
and demethanizer upper side reboiler liquids at -79 F [-62 C] (stream 39). The
cooled
stream 31a enters separator 11 at -15 F [-26 C] and 1030 psia [7,102 kPa(a)]
where the
vapor (stream 32) is separated from the condensed liquid (stream 33). The
separator
liquid (stream 33) is expanded to the operating pressure (approximately 432
psia
[2,976 kPa(a)]) of fractionation tower 19 by expansion valve 12, cooling
stream 33a to
-39 F [-39 C] before it is supplied to fractionation tower 19 at a lower mid-
column feed
point.
[0027] The vapor (stream 32) from separator 11 is divided into two
streams, 35
and 36. Stream 35, containing about 36% of the total vapor, passes through
heat
exchanger 15 in heat exchange relation with the cold residue gas at -127 F [-
88 C]
(stream 48) where it is cooled to substantial condensation. The resulting
substantially
condensed stream 35a at -123 F [-86 C] is then flash expanded through
expansion valve
16 to the operating pressure of fractionation tower 19. During expansion a
portion of the
stream is vaporized, resulting in cooling of the total stream to -134 F [-92
C]. The
expanded stream 35h is supplied to fractionation tower 19 at an upper mid-
column feed
point.
[0028] The remaining 64% of the vapor from separator 11 (stream 36)
enters a
work expansion machine 17 in which mechanical energy is extracted from this
portion of
the high pressure feed. The machine 17 expands the vapor substantially
isentropically to
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the tower operating pressure, with the work expansion cooling the expanded
stream 36a
to a temperature of approximately -90 F [-68 C]. The typical commercially
available
expanders are capable of recovering on the order of 80-88% of the work
theoretically
available in an ideal isentropic expansion. The work recovered is often used
to drive a
centrifugal compressor (such as item 18) that can be used to re-compress the
residue gas
(stream 48b), for example. The partially condensed expanded stream 36a is
thereafter
supplied as feed to fractionation tower 19 a second lower mid-column feed
point.
[0029] The demethanizer in tower 19 is a conventional distillation
column
containing a plurality of vertically spaced trays, one or more packed beds, or
some
combination of trays and packing. The demethanizer tower consists of two
sections: an
upper absorbing (rectification) section 19a that contains the trays and/or
packing to
provide the necessary contact between the vapor portion of the expanded
streams 35b and
36a rising upward and cold liquid falling downward to condense and absorb the
C2
components, C3 components, and heavier components; and a lower stripping
(demethanizing) section 19b that contains the trays and/or packing to provide
the
necessary contact between the liquids falling downward and the vapors rising
upward.
The stripping section 19b also includes reboilers (such as trim reboiler 20
and the reboiler
and side reboilers 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 42, of methane and lighter components. Stream
36a
enters demethanizer 19 at an intermediate feed position located in the lower
region of
absorbing section 19a of demethanizer 19. The liquid portion of the expanded
stream
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commingles with liquids falling downward from the absorbing section 19a and
the
combined liquid continues downward into the stripping section 19b of
demethanizer 19.
The vapor portion of the expanded stream rises upward through absorbing
section 19a
and is contacted with cold liquid falling downward to condense and absorb the
C2
components, C3 components, and heavier components.
[0030] A portion of the distillation vapor (stream 43) is withdrawn from
the upper
region of stripping section 19b. This stream is then cooled from -112 F [-80
C] to
-130 F [-90 C] and partially condensed (stream 43a) in heat exchanger 22 by
heat
exchange with the cold demethanizer overhead stream 38 exiting the top of
demethanizer
19 at -134 F [-92 C]. The cold demethanizer overhead stream is warmed slightly
to
-126 F [-88 C] (stream 38a) as it cools and condenses at least a portion of
stream 43.
[0031] The operating pressure in reflux separator 23 (428 psia [2,951
kPa(a)]) is
maintained slightly below the operating pressure of demethanizer 19. This
provides the
driving force which causes distillation vapor stream 43 to flow through heat
exchanger 22
and thence into the reflux separator 23 wherein the condensed liquid (stream
45) is
separated from the uncondensed vapor (stream 44). Stream 44 then combines with
the
warmed demethanizer overhead stream 38a from heat exchanger 22 to form cold
residue
gas stream 48 at -127 F [-88 C].
[0032] The liquid stream 45 from reflux separator 23 is pumped by pump
24 to a
pressure slightly above the operating pressure of demethanizer 19, and stream
45a is then
supplied as cold top column feed (reflux) to demethanizer 19. This cold liquid
reflux
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absorbs and condenses the propane and heavier components rising in the upper
rectification region of absorbing section 19a of demethanizer 19.
[0033] In stripping section 19b of demethanizer 19, the feed streams are
stripped
of their methane and lighter components. The resulting liquid product (stream
42) exits
the bottom of tower 19 at 52 F [11 C], based on a typical specification of a
methane to
ethane ratio of 0.025:1 on a molar basis in the bottom product. The
distillation vapor
stream forming the tower overhead (stream 38) is warmed in heat exchanger 22
as it
provides cooling to distillation stream 43 as described previously, then
combines with
stream 44 to form the cold residue gas stream 48. The residue gas passes
countercurrently to the incoming feed gas in heat exchanger 15 where it is
heated to
-28 F [-33 C] (stream 48a), and in heat exchanger 10 where it is heated to 107
F [42 C]
(stream 48b) as it provides cooling as previously described. The residue gas
is then
re-compressed in two stages, compressor 18 driven by expansion machine 17 and
compressor 27 driven by a supplemental power source. After stream 48d is
cooled to
120 F [49 C] in discharge cooler 28, the residue gas product (stream 48e)
flows to the
sales gas pipeline at 1040 psia [7,171 kPa(a)].
[0034] 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+ C. Dioxide
Total
31 25,382 1,161 362 332 743
28,055
32 25,241 1,131 336 220 733
27,736
33 141 30 26 112 10 319
35 9,087 407 121 79 264
9,985
36 16,154 724 215 141 469
17,751
43 3,598 96 5 1 113
3,816
44 2,963 33 0 0 59
3,058
45 635 63 5 1 54 758
38 22,395 164 5 0 262
22,897
48 25,358 197 5 0 321
25,955
42 24 964 357 332 422
2,100
Recoveries*
Ethane 83.05%
Propane 98.50%
Butanes+ 99.94%
Power
Residue Gas Compression 12,464 HP [ 20,490
kW]
* (Based on un-rounded flow rates)
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DESCRIPTION OF THE INVENTION
Example 1
[0035] 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.
[0036] In the simulation of the FIG. 2 process, inlet gas enters the
plant as stream
31 and is cooled in heat exchanger 10 by heat exchange with cool residue gas
at -66 F
[-54 C] (stream 38b), demethanizer reboiler liquids at 48 F [9 C] (stream 41),
demethanizer lower side reboiler liquids at 5 F [-15 C] (stream 40), and
demethanizer
upper side reboiler liquids at -70 F [-57 C] (stream 39). The cooled stream
31a enters
separator 11 at -38 F [-39 C] and 1030 psia [7,102 kPa(a)] where the vapor
(stream 32)
is separated from the condensed liquid (stream 33). The separator liquid
(stream 33) may
in some cases be divided into two streams, stream 47 and stream 37. In this
example of
the present invention, all of the separator liquid in stream 33 is directed to
stream 37 and
is expanded to the operating pressure (approximately 470 psia [3,238 kPa(a)])
of
fractionation tower 19 by expansion valve 12, cooling stream 37a to -68 F [-56
C] before
it is supplied to fractionation tower 19 at a lower mid-column feed point. In
other
embodiments of the present invention, all of the separator liquid in stream 33
may be
directed to stream 47, or a portion of stream 33 may be directed to stream 37
with the
remaining portion directed to stream 47.
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[0037] The vapor (stream 32) from separator 11 is divided into two
streams, 34
and 36. Stream 34, containing about 22% of the total vapor, may in some
embodiments
be combined with a portion (stream 47) of separator liquid stream 33 to form
combined
stream 35. Stream 34 or 35, as the case may be, passes through heat exchanger
15 in heat
exchange relation with the cold residue gas at -105 F [-76 C] (stream 38a)
where it is
cooled to substantial condensation. The resulting substantially condensed
stream 35a at
-101 F [-74 C] is then flash expanded through expansion valve 16 to the
operating
pressure of fractionation tower 19. During expansion a portion of the stream
is
vaporized, resulting in cooling of the total stream. In the process
illustrated in FIG. 2, the
expanded stream 35h leaving expansion valve 16 reaches a temperature of -128 F
[-89 C] and is supplied to fractionation tower 19 at an upper mid-column feed
point.
[0038] The remaining 78% of the vapor from separator 11 (stream 36)
enters a
work expansion machine 17 in which mechanical energy is extracted from this
portion of
the high pressure feed. The machine 17 expands the vapor substantially
isentropically to
the tower operating pressure, with the work expansion cooling the expanded
stream 36a
to a temperature of approximately -102 F [-74 C]. The partially condensed
expanded
stream 36a is thereafter supplied as feed to fractionation tower 19 a second
lower
mid-column feed point.
[0039] The demethanizer in tower 19 is a conventional distillation
column
containing a plurality of vertically spaced trays, one or more packed beds, or
some
combination of trays and packing. The demethanizer tower consists of two
sections: an
upper absorbing (rectification) section 19a that contains the trays and/or
packing to
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provide the necessary contact between the vapor portion of the expanded
streams 35b and
36a rising upward and cold liquid falling downward to condense and absorb the
C2
components, C3 components, and heavier components; and a lower stripping
(demethanizing) section 19b that contains the trays and/or packing to provide
the
necessary contact between the liquids falling downward and the vapors rising
upward.
The stripping section 19b also includes reboilers (such as trim reboiler 20
and the reboiler
and side reboilers 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 42, of methane and lighter components. Stream
36a
enters demethanizer 19 at an intermediate feed position located in the lower
region of
absorbing section 19a of demethanizer 19. The liquid portion of the expanded
stream
commingles with liquids falling downward from the absorbing section 19a and
the
combined liquid continues downward into the stripping section 19b of
demethanizer 19.
The vapor portion of the expanded stream rises upward through absorbing
section 19a
and is contacted with cold liquid falling downward to condense and absorb the
C2
components, C3 components, and heavier components.
[0040] A portion of the distillation vapor (stream 43) is withdrawn from
the upper
region of stripping section 19b at -108 F [-78 C] below expanded stream 36a
and is
compressed to approximately 609 psia [4,199 kPa(a)] by vapor compressor 21.
The
compressed stream 43a is then cooled from -78 F [-61 C] to -125 F [-87 C] and
substantially condensed (stream 43b) in heat exchanger 22 by heat exchange
with the
cold demethanizer overhead stream 38 exiting the top of demethanizer 19 at -
129 F
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[-89 C]. The cold demethanizer overhead stream is warmed to -105 F [-76 C]
(stream
38a) as it cools and condenses stream 43a.
[0041] Since substantially condensed stream 43b is at a pressure greater
than the
operating pressure of demethanizer 19, it is flash expanded through expansion
valve 25 to
the operating pressure of fractionation tower 19. During expansion a small
portion of the
stream is vaporized, resulting in cooling of the total stream to -132 F [-91
C]. The
expanded stream 43c is then supplied as cold top column feed (reflux) to
demethanizer
19. The vapor portion (if any) of stream 43c combines with the distillation
vapor rising
from the upper fractionation stage to form residue gas stream 38, while the
cold liquid
reflux portion absorbs and condenses the C2 components, C3 components, and
heavier
components rising in the upper rectification region of absorbing section 19a
of
demethanizer 19.
[0042] In stripping section 19b of demethanizer 19, the feed streams are
stripped
of their methane and lighter components. The resulting liquid product (stream
42) exits
the bottom of tower 19 at 66 F [19 C]. The distillation vapor stream forming
cold
residue gas stream 38 is warmed in heat exchanger 22 as it provides cooling to
compressed distillation stream 43a as described previously. The residue gas
(stream 38a)
passes countercurrently to the incoming feed gas in heat exchanger 15 where it
is heated
to -66 F [-54 C] (stream 38b), and in heat exchanger 10 where it is heated to
110 F
[43 C] (stream 38c) as it provides cooling as previously described. The
residue gas is
then re-compressed in two stages, compressor 18 driven by expansion machine 17
and
compressor 27 driven by a supplemental power source. After stream 38e is
cooled to
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120 F [49 C] in discharge cooler 28, the residue gas product (stream 38f)
flows to the
sales gas pipeline at 1040 psia [7,171 kPa(a)].
[0043] A summary of stream flow rates and energy consumption for the
process
illustrated in FIG. 2 is set forth in the following table:
Table II
(FIG. 2)
Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]
Stream Methane Ethane Propane Butanes+ C. Dioxide
Total
31 25,382 1,161 362 332 743
28,055
32 25,050 1,096 310 180 720
27,431
33 332 65 52 152 23 624
34/35 5,473 239 68 39 157
5,994
36 19,577 857 242 141 563
21,437
43 3,936 114 7 1 109
4,171
38 25,358 197 2 0 403
26,034
42 24 964 360 332 340
2,021
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Recoveries*
Ethane 83.06%
Propane 99.33%
Butanes+ 99.97%
Power
Residue Gas Compression 11,111 HP [
18,266 kW]
Vapor Compression 278 HP [ 457 kW]
Total Compression 11,389 HP [
18,723 kW]
* (Based on un-rounded flow rates)
[0044] A comparison of Tables I and II shows that, compared to the prior
art, the
present invention maintains essentially the same ethane recovery (83.05%
versus
83.06%), but improves both the propane recovery (99.33% versus 98.50%) and
butanes+
recovery (99.97% versus 99.94%). Comparison of Tables I and II further shows
that
these increased yields were achieved using less horsepower than the prior art
(11,389 HP
versus 12,464 HP, or more than 8% less).
[0045] There are three primary factors that account for the improved
efficiency of
the present invention. First, the boost in pressure provided by vapor
compressor 21
allows the column overhead (stream 38) to condense all of distillation vapor
stream 43,
unlike the prior art process which can condense only a fraction of the stream.
As a result,
the top reflux stream (stream 43c) for the present invention is more than 5
times greater
than that of the prior art (stream 45a), providing much more efficient
rectification in the
upper region of absorbing section 19a. Second, with the increase in the
quantity of the
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top reflux stream possible with the present invention, the quantity of
secondary reflux
stream 35b can be correspondingly less without reducing the product yields.
This in turn
results in more flow (stream 36) to expansion machine 17 and the resultant
increase in the
energy recovered to power compressor 18, thereby reducing the power
requirements of
compressor 27. Third, the more efficient rectification provided by stream 43c
in the
upper region of absorbing section 19a allows operating demethanizer 19 at a
higher
pressure without reducing the product yields, further reducing the power
requirements of
compressor 27.
[0046] A
further advantage of the present invention is a reduced likelihood of
carbon dioxide icing. FIG. 3 is a graph of the relation between carbon dioxide
concentration and temperature. Line 71 represents the equilibrium conditions
for solid
and liquid carbon dioxide in methane. (The liquid-solid equilibrium line in
this graph is
based on the data given in FIG. 16-33 on page 16-24 of the Engineering Data
Book,
Twelfth Edition, published in 2004 by the Gas Processors Suppliers
Association, which is
often used as a reference when checking for potential icing conditions.) A
liquid
temperature on or to the right of line 71, or a carbon dioxide concentration
on or above
this line, signifies an icing condition. Because of the variations which
normally occur in
gas processing facilities (e.g., feed gas composition, conditions, and flow
rate), it is
usually desired to design a demethanizer with a considerable safety factor
between the
expected operating conditions and the icing conditions. (Experience has shown
that the
conditions of the liquids on the fractionation stages of a demethanizer,
rather than the
conditions of the vapors, typically govern the allowable operating conditions
in most
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demethanizers. For this reason, the corresponding vapor-solid equilibrium line
is not
shown in FIG. 3.)
[0047] Also plotted in FIG. 3 is a line representing the conditions for
the liquids
on the fractionation stages of demethanizer 19 in the prior art FIG. 1 process
(line 72).
As can be seen, a portion of this operating line lies above the liquid-solid
equilibrium
line, indicating that the prior art FIG. 1 process cannot be operated at these
conditions
without encountering carbon dioxide icing problems. As a result, it is not
possible to use
the FIG. 1 process under these conditions, so the prior art FIG. 1 process
cannot actually
achieve the recovery efficiencies stated in Table I in practice without
removal of at least
some of the carbon dioxide from the feed gas. This would, of course,
substantially
increase capital cost.
[0048] Line 73 in FIG. 3 represents the conditions for the liquids on
the
fractionation stages of demethanizer 19 in the present invention as depicted
in FIG. 2. In
contrast to the prior art FIG. 1 process, there is a minimum safety factor of
1.2 between
the carbon dioxide concentration in the column liquids for the anticipated
operating
conditions of the FIG. 2 process versus the concentrations at the liquid-solid
equilibrium
line. That is, it would require a 20 percent increase in the carbon dioxide
content of the
liquids to cause icing. Thus, the present invention could tolerate a 20%
higher
concentration of carbon dioxide in its feed gas than the prior art FIG. 1
process could
tolerate without risk of crossing the liquid-solid equilibrium line. Further,
whereas the
prior art FIG. 1 process cannot be operated to achieve the recovery levels
given in Table I
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because of icing, the present invention could in fact be operated at even
higher recovery
levels than those given in Table II without risk of icing.
[0049] The shift in the operating conditions of the FIG. 2 demethanizer
as
indicated by line 73 in FIG. 3 can be understood by comparing the
distinguishing features
of the present invention to the prior art process of FIG. 1. While the shape
of the
operating line for the prior art FIG. 1 process (line 72) is similar to the
shape of the
operating line for the present invention (line 73), there is a key difference.
The operating
temperatures of the critical upper fractionation stages in the demethanizer in
the FIG. 2
process are warmer than those of the corresponding fractionation stages in the
demethanizer in the prior art FIG. 1 process, effectively shifting the
operating line of the
FIG. 2 process away from the liquid-solid equilibrium line. The warmer
temperatures of
the fractionation stages in the FIG. 2 demethanizer are mainly the result of
operating the
tower at higher pressure than the prior art FIG. 1 process. However, the
higher tower
pressure does not cause a loss in C2+ component recovery levels because the
distillation
vapor stream 43 in the FIG. 2 process is in essence an open direct-contact
compression-refrigeration cycle for the demethanizer using a portion of the
inter-column
vapor as the working fluid, supplying refrigeration to the process needed to
overcome the
loss in recovery that normally accompanies an increase in demethanizer
operating
pressure.
[0050] Another advantage of the present invention is a reduction in the
amount of
carbon dioxide leaving demethanizer 19 in liquid product stream 42. Comparing
stream
42 in Table I for the prior art FIG. 1 process to stream 42 in Table II for
the FIG. 2
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embodiment of the present invention reveals that there is nearly a 20%
reduction in the
quantity of carbon dioxide captured in stream 42 with the present invention.
This
generally reduces the product treating requirements by a corresponding amount,
reducing
both the capital cost and the operating cost of the treating system.
[0051] One of the inherent features in the operation of a demethanizer
column to
recover C2 components is that the column must fractionate between the methane
that is to
leave the tower in its overhead product (vapor stream 38) and the C2
components that are
to leave the tower in its bottom product (liquid stream 42). However, the
relative
volatility of carbon dioxide lies between that of methane and C2 components,
causing the
carbon dioxide to appear in both terminal streams. Further, carbon dioxide and
ethane
form an azeotrope, resulting in a tendency for carbon dioxide to accumulate in
the
intermediate fractionation stages of the column and thereby cause large
concentrations of
carbon dioxide to develop in the tower liquids.
[0052] The reflux streams for absorbing section 19a in demethanizer 19
of the
prior art FIG. 1 process are streams 45a and 35b, while those for the present
invention
shown in the FIG. 2 process are streams 43c and 35b. Comparing these streams
in
Table I and Table II, note that the total amounts of C2 components and carbon
dioxide in
the reflux streams in the prior art FIG. 1 process are 470 and 318 Lb.
Moles/Hr [470 and
318 kg moles/Hr], respectively, versus 353 and 266 Lb. Moles/Hr [353 and
266 kg moles/Hr], respectively, for the reflux streams in the FIG. 2 process
of the present
invention. Thus, significantly less of the azeotrope forming components enter
absorbing
section 19a in the cold liquid reflux streams, entering instead into the
warmer, lower
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region of absorbing section 19a with stream 36a so that there is less
accumulation of
carbon dioxide in the fractionation stages of absorbing section 19a. This
allows more of
the carbon dioxide to escape in overhead stream 38 instead of being captured
in liquid
product stream 42.
Example 2
[0053] An alternative embodiment of the present invention is shown in
FIG. 4.
The feed gas composition and conditions considered in the process presented in
FIG. 4
are the same as those in FIGS. 1 and 2. Accordingly, FIG. 4 can be compared
with the
prior art FIG. 1 process to illustrate the advantages of the present
invention, and can
likewise be compared to the embodiment displayed in FIG. 2.
[0054] In the simulation of the FIG. 4 process, inlet gas enters the
plant as stream
31 and is cooled in heat exchanger 10 by heat exchange with cool residue gas
at -66 F
[-55 C] (stream 38b), demethanizer reboiler liquids at 51 F [11 C] (stream
41),
demethanizer lower side reboiler liquids at 10 F [-12 C] (stream 40), and
demethanizer
upper side reboiler liquids at -65 F [-54 C] (stream 39). The cooled stream
31a enters
separator 11 at -38 F [-39 C] and 1030 psia [7,102 kPa(a)] where the vapor
(stream 32)
is separated from the condensed liquid (stream 33). The separator liquid
(stream 33) may
in some cases be divided into two streams, stream 47 and stream 37. In this
example of
the present invention, all of the separator liquid in stream 33 is directed to
stream 37 and
is expanded to the operating pressure (approximately 480 psia [3,309 kPa(a)])
of
fractionation tower 19 by expansion valve 12, cooling stream 37a to -67 F [-55
C] before
it is supplied to fractionation tower 19 at a lower mid-column feed point. In
other
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embodiments of the present invention, all of the separator liquid in stream 33
may be
directed to stream 47, or a portion of stream 33 may be directed to stream 37
with the
remaining portion directed to stream 47.
[0055] The vapor (stream 32) from separator 11 is divided into two
streams, 34
and 36. Stream 34, containing about 23% of the total vapor, may in some
embodiments
be combined with a portion (stream 47) of separator liquid stream 33 to form
combined
stream 35. Stream 34 or 35, as the case may be, passes through heat exchanger
15 in heat
exchange relation with the cold residue gas at -106 F [-77 C] (stream 38a)
where it is
cooled to substantial condensation. The resulting substantially condensed
stream 35a at
-102 F [-74 C] is then flash expanded through expansion valve 16 to the
operating
pressure of fractionation tower 19. During expansion a portion of the stream
is
vaporized, resulting in cooling of the total stream. In the process
illustrated in FIG. 4, the
expanded stream 35h leaving expansion valve 16 reaches a temperature of -127 F
[-88 C] and is supplied to fractionation tower 19 at an upper mid-column feed
point.
[0056] The remaining 77% of the vapor from separator 11 (stream 36)
enters a
work expansion machine 17 in which mechanical energy is extracted from this
portion of
the high pressure feed. The machine 17 expands the vapor substantially
isentropically to
the tower operating pressure, with the work expansion cooling the expanded
stream 36a
to a temperature of approximately -101 F [-74 C]. The partially condensed
expanded
stream 36a is thereafter supplied as feed to fractionation tower 19 a second
lower
mid-column feed point.
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[0057] A portion of the distillation vapor (stream 43) is withdrawn from
the lower
region of absorbing section 19a of demethanizer 19 at -113 F [-81 C] above
expanded
stream 36a and is compressed to approximately 619 psia [4,266 kPa(a)] by vapor
compressor 21. The compressed stream 43a is then cooled from -84 F [-65 C] to -
124 F
[-87 C] and substantially condensed (stream 43b) in heat exchanger 22 by heat
exchange
with the cold demethanizer overhead stream 38 exiting the top of demethanizer
19 at
-128 F [-89 C]. The cold demethanizer overhead stream is warmed to -106 F [-77
C]
(stream 38a) as it cools and condenses stream 43a.
[0058] Since substantially condensed stream 43b is at a pressure greater
than the
operating pressure of demethanizer 19, it is flash expanded through expansion
valve 25 to
the operating pressure of fractionation tower 19. During expansion a small
portion of the
stream is vaporized, resulting in cooling of the total stream to -131 F [-91
C]. The
expanded stream 43c is then supplied as cold top column feed (reflux) to
demethanizer
19. The vapor portion (if any) of stream 43c combines with the distillation
vapor rising
from the upper fractionation stage to form residue gas stream 38, while the
cold liquid
reflux portion absorbs and condenses the C2 components, C3 components, and
heavier
components rising in the upper rectification region of absorbing section 19a
of
demethanizer 19.
[0059] In stripping section 19b of demethanizer 19, the feed streams are
stripped
of their methane and lighter components. The resulting liquid product (stream
42) exits
the bottom of tower 19 at 70 F [21 C]. The distillation vapor stream forming
cold
residue gas stream 38 is warmed in heat exchanger 22 as it provides cooling to
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compressed distillation stream 43a as described previously. The residue gas
(stream 38a)
passes countercurrently to the incoming feed gas in heat exchanger 15 where it
is heated
to -66 F [-55 C] (stream 38b), and in heat exchanger 10 where it is heated to
110 F
[43 C] (stream 38c) as it provides cooling as previously described. The
residue gas is
then re-compressed in two stages, compressor 18 driven by expansion machine 17
and
compressor 27 driven by a supplemental power source. After stream 38e is
cooled to
120 F [49 C] in discharge cooler 28, the residue gas product (stream 38f)
flows to the
sales gas pipeline at 1040 psia [7,171 kPa(a)].
[0060] A summary of stream flow rates and energy consumption for the
process
illustrated in FIG. 4 is set forth in the following table:
Table III
(FIG. 4)
Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]
Stream Methane Ethane Propane Butanes+ C. Dioxide
Total
31 25,382 1,161 362 332 743
28,055
32 25,050 1,096 310 180 720
27,431
33 332 65 52 152 23 624
34/35 5,636 247 70 40 162
6,172
36 19,414 849 240 140 558
21,259
43 3,962 100 3 0 125
4,200
38 25,358 197 2 0 425
26,055
42 24 964 360 332 318
2,000
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Recoveries*
Ethane 83.06%
Propane 99.50%
Butanes+ 99.98%
Power
Residue Gas Compression 10,784 HP [ 17,728 kW]
Vapor Compression 260 HP [ 428
kW]
Total Compression 11,044 HP [ 18,156 kW]
* (Based on un-rounded flow rates)
[0061] A comparison of Tables II and III shows that, compared to the
FIG. 2
embodiment of the present invention, the FIG. 4 embodiment maintains the same
ethane
recovery while improving the propane recovery (99.50% versus 99.33%) and
butanes+
recovery (99.98% versus 99.97%) slightly. However, comparison of Tables II and
III
further shows that these yields were achieved using about 3% less horsepower
than that
required by the FIG. 2 embodiment of the present invention. The drop in the
power
requirements for the FIG. 4 embodiment is mainly due to the lower content of
C2+
components in top reflux stream 43c, which provides more efficient
rectification in the
upper region of absorbing section 19a so that demethanizer 19 can be operated
at a
slightly higher operating pressure (thereby reducing compression requirements)
without
reducing product yields. Comparing distillation vapor stream 43 in Table III
for the
FIG. 4 embodiment of the present invention to stream 43 in Table II for the
FIG. 2
embodiment of the present invention, the concentrations of C2 components and
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particularly the C3+ components in stream 43 of the FIG. 4 embodiment are
significantly
lower, so that higher product yields are achieved using less power than the
FIG. 2
embodiment. The lower concentrations of C2 components and C3+ components in
stream
43 of the FIG. 4 embodiment are the result of withdrawing the distillation
vapor from the
lower region of absorbing section 19a rather than from the upper region of
stripping
section 19b as in the FIG. 2 embodiment. The distillation vapor at the higher
column
location has been subjected to more rectification than the distillation vapor
lower in the
column, and so is closer to being the pure methane stream that would be the
ideal reflux
stream for the top of the column. In the prior art process of FIG. 1, the
column overhead
(stream 38) could not condense a pure methane stream, but with the elevation
in pressure
provided by vapor compressor 21 of the present invention, column overhead
stream 38 is
cold enough to totally condense the distillation vapor stream 43 even though
it is almost
pure methane.
[0062] When the present invention is employed as in Example 2, the
advantage
with respect to avoiding carbon dioxide icing conditions is maintained
compared to the
FIG. 2 embodiment. FIG. 5 is another graph of the relation between carbon
dioxide
concentration and temperature, with line 71 as before representing the
equilibrium
conditions for solid and liquid carbon dioxide in methane and line 72
representing the
conditions for the liquids on the fractionation stages of demethanizer 19 in
the prior art
process of FIG. 1. Line 74 in FIG. 5 represents the conditions for the liquids
on the
fractionation stages of demethanizer 19 in the present invention as depicted
in FIG. 4, and
shows a safety factor of 1.2 between the anticipated operating conditions and
the icing
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conditions for the FIG. 4 process. Thus, this embodiment of the present
invention could
also tolerate an increase of 20 percent in the concentration of carbon dioxide
without risk
of icing. In practice, this improvement in the icing safety factor could be
used to
advantage by operating the demethanizer at lower pressure (i.e., with colder
temperatures
on the fractionation stages) to raise the C2+ component recovery levels
without
encountering icing problems. The shape of line 74 in FIG. 5 for the FIG. 4
embodiment
is very similar to that of line 73 in FIG. 3 for the FIG. 2 embodiment. The
primary
difference is the significantly lower carbon dioxide concentrations of the
liquids on the
fractionation stages in the lower section of the FIG. 4 demethanizer due to
withdrawing
the distillation vapor stream at a higher location on the column in this
embodiment. As
can be seen by comparing stream 42 in Tables II and III, even less of the
carbon dioxide
in the feed gas is captured with the bottom liquid product in the FIG. 4
embodiment of
the present invention, which generally means still less product treating will
be required
compared to the FIG. 2 embodiment of the present invention.
Other Embodiments
[0063] In accordance with this invention, it is generally advantageous
to design
the absorbing (rectification) section of the demethanizer to contain multiple
theoretical
separation stages. However, the benefits of the present invention can be
achieved with as
few as one theoretical stage, and it is believed that even the equivalent of a
fractional
theoretical stage may allow achieving these benefits. For instance, all or a
part of the
expanded substantially condensed distillation stream 43c from expansion valve
25, all or
a part of the expanded substantially condensed stream 35b from expansion valve
16, and
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all or a part of the expanded stream 36a from work expansion machine 17 can be
combined (such as in the piping joining the expansion valve to the
demethanizer) and if
thoroughly intermingled, the vapors and liquids will mix together and separate
in
accordance with the relative volatilities of the various components of the
total combined
streams. Such commingling of the three streams shall be considered for the
purposes of
this invention as constituting an absorbing section.
[0064] In some cases it may be advantageous to split the substantially
condensed
distillation stream 43b into at least two streams as shown in FIGS. 6 through
9. This
allows a portion (stream 51) to be supplied above the location where vapor
distillation
stream 43 is withdrawn (and perhaps also above the feed location of expanded
stream
36a), either lower in the absorbing section of fractionation tower 19 (FIGS. 6
and 7) or
lower on absorber column 19 (FIGS. 8 and 9), to increase the liquid flow in
that part of
the distillation system and improve the rectification of stream 43. In such
cases,
expansion valve 26 is used to expand stream 51 to the column operating
pressure
(forming stream 51a), while expansion valve 25 is used to expand the remaining
portion
(stream 50) to the column operating pressure so that the resulting stream 50a
can then be
supplied to the top of the absorbing section in demethanizer 19 (FIGS. 6 and
7) or to the
top of absorber column 19 (FIGS. 8 and 9).
[0065] FIGS. 8 and 9 depict a fractionation tower constructed in two
vessels,
absorber (rectifier) column 19 (a contacting and separating device) and
stripper column
29 (a distillation column). In FIG. 8, the overhead vapor (stream 46) from
stripper
column 29 is split into two portions. One portion (stream 43) is routed to
compressor 21
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and thence to heat exchanger 22 to generate reflux for absorber column 19 as
described
earlier. The remaining portion (stream 49) flows to the lower section of
absorber column
19 to be contacted by expanded substantially condensed stream 35b and the
expanded
substantially condensed distillation stream (either stream 50a, or streams 50a
and 51a).
Pump 30 is used to route the liquids (stream 52) from the bottom of absorber
column 19
to the top of stripper column 29 so that the two towers effectively function
as one
distillation system. In FIG. 9, all of the overhead vapor (stream 46) flows to
the lower
section of absorber column 19, and distillation vapor stream 43 is withdrawn
from a
location higher in absorber column 19, above the feed location of expanded
stream 36a.
The decision whether to construct the fractionation tower as a single vessel
(such as
demethanizer 19 in FIGS. 2, 4, 6, and 7) or multiple vessels will depend on a
number of
factors such as plant size, the distance to fabrication facilities, etc.
[0066] Feed gas conditions, plant size, available equipment, or other
factors may
indicate that elimination of work expansion machine 17, 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 35a) and/or
the
substantially condensed distillation stream (stream 43b).
[0067] As described in the earlier examples, distillation stream 43 is
substantially
condensed and the resulting condensate used to absorb valuable C2 components,
C3
components, and heavier components from the vapors rising through the upper
region of
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absorbing section 19a of demethanizer 19 (FIGS. 2, 4, 6, and 7) or absorber
column 19
(FIGS. 8 and 9). However, the present invention is not limited to this
embodiment. It
may be advantageous, for instance, to treat only a portion of these vapors 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 vapors or the condensate should bypass
absorbing
section 19a of demethanizer 19 (FIGS. 2, 4, 6, and 7) or absorber column 19
(FIGS. 8
and 9). Some circumstances may favor partial condensation, rather than total
condensation, of distillation stream 43a in heat exchanger 22. Other
circumstances may
favor that distillation stream 43 be a total vapor side draw from
fractionation column 19
rather than a partial vapor side draw. 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 some portion of the cooling of distillation stream 43a in heat
exchanger 22.
[0068] Under
some circumstances, it may be advantageous to heat distillation
stream 43 before it is compressed, as this may reduce the capital cost of
compressor 21.
One means to accomplish this is to use compressed distillation stream 43a
(which is
warmer due to the heat of compression) to supply this heating using a cross
exchanger.
In such cases, it may be possible to supplement the cooling of compressed
distillation
stream 43a by the use of aerial cooling or other means, thereby reducing the
cooling that
must be supplied in heat exchanger 22 by overhead stream 38. The potential
reduction in
the capital cost of compressor 21 must be weighed against the capital cost of
the
additional heating and cooling means for each application to determine whether
this
embodiment is advantageous.
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[0069] In accordance with this invention, the splitting of the vapor
feed may be
accomplished in several ways. In some embodiments, vapor splitting may be
effected in
a separator. In the processes of FIGS. 2, 4, and 6 through 9, the splitting of
the vapor
occurs following cooling, and perhaps after separation of any liquids which
may have
been formed. The high pressure gas may be split, however, prior to any cooling
of the
inlet gas as shown in FIG. 10. Streams 35b, 36a, and 37a in FIG. 10 may all be
fed to a
distillation column (such as demethanizer 19 in FIGS. 2, 4, 6, and 7), or
streams 35b and
36a may be fed to a contacting and separating device and stream 37a may be fed
to a
distillation column (such as absorber column 19 and stripper column 29,
respectively, in
FIGS. 8 and 9). The cooling of stream 53 in heat exchanger 10 in FIG. 10 may
be
accomplished or supplemented by additional process streams (such as streams
39, 40, and
41 in FIGS. 2, 4, and 6 through 9) and/or external refrigeration.
[0070] When the inlet gas is leaner, separator 11 in FIGS. 2, 4, and 6
through 10
may not be needed. Depending on the quantity of heavier hydrocarbons in the
feed gas
and the feed gas pressure, the cooled feed stream 31a leaving heat exchanger
10 in
FIGS. 2, 4, and 6 through 9 or the cooled stream 53a leaving heat exchanger 10
in
FIG. 10 may not contain any liquid (because it is above its dewpoint, or
because it is
above its cricondenbar), so that separator 11 shown in FIGS. 2, 4, and 6
through 10 is not
required.
[0071] The high pressure liquid (stream 33) in FIGS. 2, 4, and 6 through
9 need
not be expanded and fed to a mid-column feed point on the distillation column.
Instead,
all or a portion of it (dashed stream 47) may be combined with the portion of
the
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separator vapor (stream 34) to form combined stream 35 that flows to heat
exchanger 15.
Any remaining portion of the liquid (dashed stream 37) may be expanded through
an
appropriate expansion device, such as expansion valve 12, to form stream 37a
which is
then fed to a mid-column feed point on distillation column 19 (FIGS. 2, 4, 6,
and 7) or
stripper column 29 (FIGS. 8 and 9). Stream 33 in FIGS. 2, 4, and 6 through 9
and/or
stream 37 in FIGS. 2, 4, and 6 through 10 may also be used for inlet gas
cooling or other
heat exchange service before or after the expansion step prior to flowing to
the
demethanizer.
[0072] In accordance with this invention, the use of external
refrigeration to
supplement the cooling available to the inlet gas and/or the distillation
stream from other
process streams may be employed, particularly in the case of a rich inlet gas.
The use
and distribution of separator liquids and demethanizer side draw liquids for
process heat
exchange, and the particular arrangement of heat exchangers for inlet gas
cooling must be
evaluated for each particular application, as well as the choice of process
streams for
specific heat exchange services.
[0073] It will also 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 to the top of the column
may
increase recovery while decreasing power recovered from the expander thereby
increasing the recompression horsepower requirements. Increasing feed lower in
the
column reduces the horsepower consumption but may also reduce product
recovery. The
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relative locations of the mid-column feeds may vary depending on inlet
composition or
other factors such as desired recovery levels and amount of liquid formed
during inlet gas
cooling. Moreover, two or more of the feed streams, or portions thereof, may
be
combined depending on the relative temperatures and quantities of individual
streams,
and the combined stream then fed to a mid-column feed position.
[0074] 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
demethanizer 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 reboilers, or a combination thereof.
[0075] 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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