Title of Invention

A PROCESS FOR SEPARATION OF A GAS STREAM

Abstract In a process for the separation of a gas stream containing methane, C2 components, C3 components and heavier hydrocarbon components into a volatile residue gas fraction containing a major portion of said methane and a relatively less volatile fraction containing a major portion of said C2 components, C3 components and heavier hydrocarbon components, in which process said gas stream is treated in one or more heat exchange steps to produce at least a first feed stream that has been cooled under pressure; said cooled first feed stream is expanded to a lower pressure, and thereafter supplied to a fractionation tower at a top feed point; and said cooled expanded first feed stream is fractionated at said lower pressure whereby the components of said relatively less volatile fraction are recovered; characterized in that a liquid distillation stream is withdrawn from said fractionation tower and heated; said heated distillation stream is returned to a lower point on said fractionation tower that is separated from said withdrawal point by at least one theoretical stage; and the quantities and temperatures of said feed streams to said fractionation tower are effective to maintain the overhead temperature of said fractionation tower at a temperature whereby the major portions of the components in said relatively less volatile fraction are recovered.
Full Text This invention relate to a process fot the sepa ration og ages stream.
BACKGROUND OF INVENTION
This invention relates to a process for the separation of a gas containing
hydrocarbons.
Ethlene, ethane, propylene, propane and/or heavier hydrocarbons can be
recovered from a variety of gases, such as natural gas, refinery gas, and
synthetic gas streams obtained from other hydrocarbon materials such as coal,
crude oil. naphtha, oil shale, tar sands, and lignite. Natural gas usually has a
major proportion of methane and ethane, i.e., methane and ethane together
comprise at least 50 mole percent of the gas. The gas also contains relatively
lesser amounts of heavier hydrocarbons such as propane, butanes, pentanes
and the like, as well as hydrogen, nitrogen, nitrogen, carbon dioxide and other gases.
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, 88.41% methane, 6.65%
ethane and other C2 components, 2.26% propane and other C3 components,
0.36% iso-butane, 0,45% normal butane, 0.31% pentanes plus, with the balance
made up of nitrogen and carbon dioxide. Sulfur containing gases are also
sometimes present.
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. Competition for processing rights has forced plant operators to
maximize the processing capacity and recovery efficiency of their existing gas
processing plants. 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.
The cryogenic expansion process is now generally preferred for natural
gas liquids recovery because it provides maximum simplicity with ease of start up,
operating flexibility, good efficiency, safety, and good reliability. U.S. Pat. Nos.
3,292,380; 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; reissue U.S. Pat. No. 33,408; and co-pending application no. 09/439,508
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 and patent applications).
In a typical cryogenic expansion recovery process, a feed gas stream
under pressure is cooled by heat exchange with other streams of the process and/or
external sources of refrigeration such as a propane compression-refrigeration system.
As the gas is cooled, liquids may be condensed and collected in one or more
separators as high-pressure liquids containing some of the desired C2+ components.
Depending on the richness of the gas and the amount of liquids formed, the
high-pressure liquids may be expanded to a lower pressure and fractionated. The
vaporization occurring during expansion of the liquids results in further cooling of the
stream. Under some conditions, pre-cooling the high pressure liquids prior to the
expansion may be desirable in order to ftirther lower the temperature resulting from
the expansion. The expanded stream, comprising a mixture of liquid and vapor, is
fractionated in a distillation (demethanizer) 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 liqmd product.
If the feed gas is not totally condensed (typically it is not), at least 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 a feed to the column. In recent years, the preferred
processes for hydrocarbon separation involve feeding this expanded vapor-liquid
stream at a mid-column feed point, with an upper absorber section providing
additional rectification of the vapor phase. There are, however, processes wherein
this expanded vapor-liquid stream is used as the top column feed. 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 colunm feed.
For those processes that include an upper rectification section, a reflux
stream must be provided for the section. One manner for accomplishing this is to
withdraw a vapor distillation stream from the upper section of the demethanizer
tower, cool it to partially condense it by heat exchange with other process streams,
e.g., part of the feed gas that has been cooled to substantial condensation and then
expanded to cool it further. The liquid condensed from the vapor distillation stream is
then supplied as the top feed to the demethanizer.
The purpose of this process is to perform a separation that produces a
residue gas leaving the process which contains substantially all of the methane in the
feed gas with essentially none of the C2 components and heavier hydrocarbon
components, and a bottoms fraction leaving the demethanizer which contains
substantially all of the C2 components and heavier hydrocarbon components with
essentially no methane or more volatile components while meeting plant
specifications for maximum permissible carbon dioxide content. The present
invention provides a means for providing a new plant or modifying an existing


processing plant to achieve this separation at significantly lower capital cost by
reducing the size of or eliminating the need for a product treating system for removal
of carbon dioxide. Alternatively, the present invention, whether applied in a new
facility or as a modification to an existing processing plant, can be used to recover
more C2 components and heavier hydrocarbon components in the bottom liquid
product for a given carbon dioxide concentration in the feed gas than other processing
schemes.
In accordance with the present invention, it has been found that C2
recoveries in excess of 66 percent can be maintained while maintaining the carbon
dioxide content of the bottom liquid product within specifications and providing
essentially complete rejection of methane to the residue gas stream. The present
invention, although applicable at lower pressures and warmer temperatures, is
particularly advantageous when processing feed gases at pressures in the range of 600
to 1000 psia or higher under conditions requiring column overhead temperatures of
-120°F or colder.
The present invention uses a modified reboiler scheme which can be
applied to any type of NGL recovery system. In a typical reboiler or side reboiler
application in a distillation column, the entire column down-flowing liquid stream is
withdrawn from the tower and passed through a heat exchanger, then returned to the
column at essentially the same point in the coliunn. In this modified reboiler system,
a portion of the column down-flowing liquid is withdrawn from a point higher in the
column, i.e., separated from the return point by at least one theoretical stage. Even
though the flow rate of the liquid may be lower, it is usually much colder and can
have advantages in improving recovery or reducing exchanger size.
It has been found that when the present invention is applied to prior art
processes for NGL recovery, the recovery of C2 components and heavier components
is improved by one to two percent. The improvement in recovery is much greater,
however, when it is desirable to reduce the carbon dioxide content in the recovered
NGL product. Recovery of ethane in a typical NGL recovery plant also results in
recovery of at least some of the carbon dioxide contained in the feed gas because
carbon dioxide falls in between methane and ethane in relative volatility. Therefore,


as ethane recovery increases, so does the recovery of carbon dioxide in the NGL
product By applying the modified reboiler scheme of the present invention, the
applicants have found that it is possible to significantly improve recovery of ethane in
the NGL product compared to use of the conventional reboiler or side reboiler
systems when the column is reboiled to meet the desired carbon dioxide content in the
NGL product.
For a better understanding of the present invention, reference is made
to the following examples and drawings. Referring to the drawings:
FIG. 1 is a flow diagram of a prior art cryogenic natural gas processing
plant;
FIG. 2 is a flow diagram illustrating how the processing plant of FIG. 1
can be adapted to be a natural gas processing plant in accordance with the present
invention; ,
FIG. 3 is a flow diagram illustrating an alternative adaptation of FIG. 1
to be a natural gas processing plant in accordance with the present invention;
FIG. 4 is a flow diagram illustrating an alternative adaptation of FIG. 1
to be a natural gas processing plant in accordance with the present invention;
FIG. 5 is a flow diagram illustrating how an alternative prior art
process can be adapted to be a natural gas processing plant in accordance with the
present invention;
FIG. 6 is a diagram illustrating the modified reboiler scheme of the
present invention for a processing plant wherein the scheme includes a thermosiphon
system;
FIG. 7 is a diagram illustrating the modified reboiler scheme of the
present invention for a processing plant wherein the scheme includes a pumped
system;
FIG. 8 is a diagram illustrating the modified reboiler scheme of the
present invention for a processing plant wherein the scheme includes a pumped
system; and
, FIG. 9 is a diagram illustrating the modified reboiler scheme of the
present invention for a processing plant wherein the scheme includes a split column
system.
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 pound moles par 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.
DESCRIPTION OF THE PRIOR ART
FIG. 1 is a process flow diagram showing the design of a processing
plant to recover C2+ components from natural gas using prior art according to U.S.
Pat. No. 3,292,380. In this simulation of the process, inlet gas enters the plant at 90°F
and 915 psia 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)
fonnation under cryogenic conditions. Solid desiccant has typically been used for this
purpose.
The feed stream 31 is cooled in exchanger 10 by heat exchange with
cold residue gas at -108°F (stream 37), demethanizer reboiler liquids at 59°F (stream
42), and demethanizer side reboiler liquids at 30°F (stream 40). Note that in all cases
exchange 10 is representative of either a multitude of individual heat exchangers or a
single multi-pass heat exchanger, or any combination thereof (The decision as to
whether to use more than one heat exchanger for the indicated cooling services will
depend on a number of factors including, but not limited to, inlet gas flow rate, heat
exchanger size, stream temperatures, etc.) Note also that heat exchanger 10 was
intended to use demethanizer liquid product (stream 43a) to provide a portion of the
feed gas cooling, but as will be explained later this stream is too warm to be used for
this purpose. The cooled stream 31a enters separator 11 at -30°F and 905 psia where
the vapor (stream 32) is separated from the condensed liquid (stream 35).
The vapor (stream 32) from separator 11 enters a work expansion
machine 14 in which mechanical energy is extracted from this portion of the high
pressure feed. The machine 14 expands the vapor substantially isentropically from a
pressure of about 905 psia to the operating pressure (approximately 315 psia) of
demethanizer column 17, with the work expansion cooling the expanded stream 32a
to a temperature of approximately -108°F. The typical commercially available
expanders are capable of recovering on the order of 80-85% of the work theoretically
available in an ideal isentropic expansion. The work recovered is often used to drive
a centrifugal compressor (such as item 15), that can be used to re-compress the
residue gas (stream 37a), for example. The expanded and partially condensed stream
32a is supplied to separator section 17a in the upper region of demethanizer tower 17.
The liquids separated therein become the top feed to theoretical stage 1 in
demethanizing section 17b.
The liquid (stream 35) from separator 11 is flash expanded through an
appropriate expansion device, such as expansion valve 16, to the operating pressvure of
demethanizer tower 17. During expansion a portion of the stream is vaporized,
resulting in cooling of the total stream. In the process illustrated in FIG. 1, the
expanded stream 35a leaving expansion valve 16 reaches a temperature of-73°F and
is supplied to a mid-column feed point on demethanizer tower 17.
The demethanizer 17 is a conventional distillation column containing a
plurality of vertically spaced trays, one or more packed beds, or some combination of
trays and packing. As is often the case in natural gas processing plants, the
demethanizer tower may consist of two sections. The upper section 17a is a separator
wherein the partially condensed top feed is divided into its respective vapor and liquid
portions, and wherein the vapor rising from the lower distillation or demethanizing
section 17b is combined with the vapor portion of the top feed to form the cold


residue gas distillation stream 37 which exits the top of the tower. The lower,
demethanizing section 17b contains the trays and/or packing and provide the
necessary contact between the liquids falling downward and the vapors rising upward.
The demethanizer colunm 17 also includes reboilers which heat and vaporize portions
of the liquids flowing down the column to provide the stripping vapors which flow up
the colum.
In many cases, the temperature of the liquid product (stream 43)
exiting the bottom of the tower is controlled on the basis of maintaining the desired
ratio of methane to ethane in the liquid product A typical specification for this is a
methane to ethane ratio of 0.025:1 on a molar basis in the bottom product. In this
case, however, the concentration of carbon dioxide in the liquid product would exceed
the plant owner's specification for a carbon dioxide to ethane ratio of 0.05:1 on a
molar basis if the demethanizer was controlled to maintain this methane:ethane ratio.
Thus, if operated in this manner this plant design would require the addition of a
treating system to remove carbon dioxide from the hydrocarbons in order to produce a
marketable liquid product. There are many options for removing the carbon dioxide
(treating the incoming feed gas, treating the total liquid product, treating the ethane
product after fractionation, etc.), but all of these options will add not only to the
capital cost of the plant (due to the cost of installing the treating system) but also to
the operating expense of the plant (due to energy and chemical consumption in the
treating system).
One way to keep the ethane product within the carbon dioxide
specification is to operate the demethanizer in a maimer to strip the carbon dioxide
from the bottom liquid product, by adding more reboil heat to the column using the
side reboiler and/or the bottom reboiler as illustrated here for the FIG. 1 process. This
results in the liquid product (stream 43) exiting the bottom of the tower at 77°F,
whereupon it is pumped to approximately 480 psia (stream 43a) in pump 20. (The
discharge pressure of the pump is usually set by the ultimate destination of the liquid
product. Generally the liquid product flows to storage after being used for heat
exchange and the pump discharge pressure is set so as to prevent any vaporization of
stream 43a as it warms to ambient temperature.) Because stream 43a is so warm.
however, it cannot be used for feed gas cooling in heat exchanger 10. Accordingly,
block valve 21a must be closed and block valve 21b opened to bypass the stream
around heat exchanger 10 and send it directly to storage (stream 43d).
The residue gas (stream 37) passes countercurrently to the incoming
feed gas in heat exchanger 10 where it is heated to 33°F (stream 37a). The residue
gas is then re-compressed in two stages. The first stage is compressor 15 driven by
expansion machine 14, and the second stage is compressor 22 driven by a
supplemental poweir source. After stream 37c is cooled to 120°F by cooler 23, the
residue gas product (stream 37d) flows to the sales pipeline at 1015 psia, sufficient to
meet line requirements (usually on theorder of the inlet pressure).
A summary of stream flow rates and energy consumption for the
process illustrated in FlG.1 is set forth in the following table:
The carbon dioxide:ethane ratio in the bottom liquid product for the
FIG. 1 process is 0.05:1, complying with the plant owner's specification. Note,
however, that the methane:ethane ratio in the bottom product is 0.000003:1 on a
molar basis, versus the allowable ratio of 0.025:1, indicating the degree of
over-stripping required to control the carbon dioxide content of the liquid product at
the required level. Examination of the recovery levels displayed in Table I shows that


operating the FIG. 1 process in this manner to reduce the carbon dioxide content in
the ethane product causes a substantial reduction in liquids recovery. When operated
at a methane:ethane ratio of 0.025:1 in the bottom product, calculations indicate that
the FIG. 1 process can achieve an ethane recovery of 69.64%, a propane recovery of
96.18%, and a butanes+ recovery of 99.66%. Unfortunately, the resulting carbon
dioxide:ethane ratio (0.087:1) is too high to meet the plant owner's specification when
the plant is operated in this maimer. Thus, the requirement to operate the FIG. 1
process to reduce the concentration of carbon dioxide in the liquid product causes
reductions in the ethane, propane, and butanes+ recoveries of over 28 percentage
points, 10 percentage points, and 1 percentage point, respectively, for the prior art
process.
There are two factors at work in the FIG. 1 process that result in less
liquids recovery from the bottom of demethanizer tower 17 when the tower is
operated to control the carbon dioxide content of the liquid product. First, when the
temperature at the bottom of demethanizer column 17 is raised to 77'F by reboiling
the column more, the temperatures at each point in the colunm increase. This reduces
the amount of cooling that the tower liquid streams (streams 40,42, and 43) can
supply to the feed gas in heat exchanger 10. As a result, the cooled feed stream
(stream 31a) entering separator 11 is wanner, which in turn results in the lower ethane
retention in demethanizer column 17. Second, the higher temperatures in the lower
section of demethanizer coliumn 17 cause the temperatures in the upper section to be
higher also, resulting in less methane liquid entering the lower section of
demethanizer column 19. When this liquid methane is subsequently vaporized by the
side reboiler and main reboiler attached to demethanizer column 17, the methane
vapor helps to strip the carbon dioxide from the liquids flowing down the column.
With less methane available in the FIG. 1 process to strip the carbon dioxide, more of
the ethane in the liquids must be vaporized to serve as stripping gas. Since the
relative volatilities for carbon dioxide and ethane are very similar, the ethane vapor is
a much less effective stripping agent than the methane vapor, which reduces the
stripping efficiency in the column and causes lower recovery.


DESCRIPTION OF THE INVENTION
Example
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.
In the simulation of the FIG. 2 process, inlet gas enters at 90°F and a
pressure of 915 psia as stream 31. The feed stream 31 is cooled in exchanger 10 by
heat exchange with cold residue gas at -130°F (stream 37), demethanizer liquid
product at 57°F (stream 43a), demethanizer reboiler liquids at 33°F (stream 42), and a
portion of the liquids from the upper section of demethanizer column 17 at -130°F
(stream 40). The cooled stream 31a enters separator 11 at -59°F and 905 psia where
the vapor (stream 32) is separated from the condensed liquid (stream 35).
The condensed liquid (stream 35) from separator 11 is flash expanded
through an appropriate expansion device, such as expansion valve 16, to the operating
pressure (approximately 315 psia) of demethanizer tower 17. 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 35a leaving expansion valve 16
reaches a temperature of-114°F and is supplied to demethanizer column 17 at a
mid-column feed point.
The vapor (stream 32) from separator 11 enters a work expansion
machine 14 in which mechanical energy is extracted from this portion of the high
pressure feed. The machine 14 expands the vapor substantially isentropically from a
pressure of about 905 psia to the operating pressure of demethanizer tower 17, with
the work expansion cooling the expanded stream 32a to a temperature of
approximately -132°F. The expanded and partially condensed stream 32a is thereafter
supplied to demethanizer colunan 17 as the top column feed. The vapor portion of
stream 32a combines with the vapors rising from the top fractionation stage of the


column to form distillation stream 37, which is withdrawn from an upper region of the
tower.
The liquid portion of stream 32a is used to contact the vapors rising
from the lower fractionation stages of demethanizer column 17 and rectify the desired
C2 components and heavier components from the vapors, and is then divided into two
• portions. One portion (stream 41), containing about 40% of the total liquid, is
directed onto the lower fractionation stages in demethanizer column 17 to further
contact and rectify the vapors rising upward.
The other portion (stream 40), containing the remaining 60% of the
liquid, is withdrawn from the tower and directed to heat exchanger 10 where it
supplies part of the feed gas cooling as it is heated to 30°F and partially vaporized.
The heated stream 40a is thereafter supplied to demethanizer column 17 at a
mid-column feed point, separated from the point where stream 40 was withdrawn
from the column by at least one theoretical stage. In this case, the partially vaporized
stream 40a flows to the same point on the column that was used for the side reboiler
return (theoretical stage 11 in demethanizer tower 17) in the FIG. 1 process, which is
the equivalent of ten theoretical stages lower than the liquid stream withdrawal point
in the fractionation system (theoretical stage 1 in demethanizer tower 17).
The liquid product (stream 43) exits the bottom of dranethanizer tower
17 at 54°F. This stream is pumped to approximately 480 psia (stream 43a) in pump
20 and then directed to heat exchanger 10 where it is heated to 72°F as it supplies part
of the feed gas cooling as described previously. The residue gas (stream 37) passes
countercuirently to the incoming feed gas in heat exchanger 10 where it is heated to
58°F (stream 37a). The residue gas is then re-compressed in two stages, compressor
15 driven by expansion machine 14 and compressor 22 driven by a supplemental
power source. After stream 37c is cooled to 120°F by cooler 23, the residue gas
product (stream 37d) flows to the sales pipeline at 1015 psia.
A summary of stream flow rates and energy consumption for the
process illustrated in FIG. 2 is set forth in the following table:
Unlike the prior art process shown in Fig. 1, both the carbon
dioxide:ethane ratio (0.05:1) and the methane:ethane ratio (0.025:1) in the bottom
liquid product can be controlled at the specifications required by the client in the
FIG. 2 process. Comparison of the recovery levels displayed in Tables I and II shows
that the present invention allows achieving much higher liquids recovery efficiency


than the FIG. 1 process when it is operated in a fashion to limit the carbon dioxide
content of its liquid product A comparison of Tables I and II shows that, compared
to the prior art, the present invention improves ethane recovery from 40.74% to
66.58%, propane recovery from 85.47% to 95.91%, and butanes+ recovery from
98.09% to 99.63%. Comparison of Tables I and II further shows that the higher the
product yields were not simply the result of increasing the horsepower (utility)
requirements. To the contrary, when the present invention is employed as in this
Example, not only do the ethane, propane, and butanes+ recoveries increase over
those of the prior art process, liquid recovery efficiency also increases by 41 percent
(in terms of ethane recovered per unit of horsepower expended). The FIG. 2 process
recovers 0.83 gallons per hour of ethane per unit of horsepower consumed, versus
0.59 gallons per hour per unit of horsepower for the FIG. 1 process.
A significant benefit achieved by the present invention illustrated in
FIG. 2 is that the modified reboiler scheme provides colder column liquids for use in
refrigerating the incoming feed streams. This increases the cooling available to the
inlet gas, as not only can considerably more duty be obtained from the liquid in this
case, but at a colder temperature level. At the same time, more methane is introduced
lower in demethanizer column 17 than would othowise be there when reboiling the
column to meet the carbon dioxide content. (Note that stream 40 in the FIG. 2
process contains 1334 Lb. Moles/Hr of methane, whereas stream 40 in the FIG. 1
process contains only 14 Lb. Moles/Hr of methane.) This additional methane
provided by the present invention in the FIG. 2 process helps to strip the carbon
dioxide from the liquids flowing downward in the stripping column. The quantity of
carbon dioxide in the NGL product from the FIG. 2 process can be adjusted by
appropriate control of the quantity of liquid withdrawn to feed the modified reboiler
system instead of being directed to the fractionation stages in the upper section of
demethanizer column 17.
Other Embodiments
FIGS. 3 and 4 are flow diagrams illustrating alternative manners in
which the process and apparatus described and depicted in U.S. Pat. No. 3,292,380


can be adapted to be natural gas processing plants in accordance with the present
invention. It should be noted that in the FIG. 3 embodiment of the present invention,
the distillation stream (stream 40) used for the modified reboiler scheme is produced
by dividing the liquids formed in stream 32a during expansion (stream 34 from
separator 19) external to demethanizer tower 17. This could also have been
accomplished by routing all of the expanded stream (stream 32a) from work
expansion machine 14 to a separator section in the upper part of demethanizer tower
17 to separate the liquids, then dividing the liquids to produce the reflux stream for
the tower (stream 41) and the distillation stream for the modified reboiler scheme
(stream 40). FIG. 5 is a flow diagram illustrating one manner in which the process
and apparatus described and depicted in U.S. Pat. No. 4,854,955 can be adapted to be
a natural gas processing plant in accordance with the present invention.
FIGS. 6,7,8, and 9 are diagrams showing some of the alternative
methods for implementing the modified reboiler scheme, FIG. 6 shows a typical
thermosiphon type application wherein the partial flow of liquid from fractionation
tower 50 to reboiler 57 could be controlled via valve 58 in liquid draw line 61. The
liquid portion not withdrawn from the column simply overflows chixnney tray 51 onto
distributor 52 for packing (or trays) 53 below. The heated stream in line 61a from
reboiler 57 is returned to fractionation tower 50 at a lower point which contains an
appropriate feed distribution mechanism, such as chimney tray 54 and distributor 55,
to mix the heated stream with the down-flowing tower liquids from packing (or trays)
53 and supply the mixture to packing (or trays) 56. FIGS. 7 and 8 show typical,
pumped adaptations wherein the total liquid down-flow is withdrawn in liquid draw
line 61 and pumped to higher pressure by pump 60. The flow of the pumped liquid in
line 61a is then divided via appropriate control valves 58 and 59 to arrive at the
desired quantity of liquid in line 62 flowing to reboiler 57. The heated stream in line
62a from reboiler 57 is returned to fractionation tower 50 at a lower point as described
previously for the FIG. 6 embodiment. In the FIG. 7 embodiment, the liquid that does
not flow to the reboiler (in line 63) is returned to chimney tray 51 from which the
liquid was initially withdrawn, whereupon it can overflow chimney tray 51 onto
distributor 52 for packing (or trays) 53 below. In the FIG. 8 embodiment, the liquid


that does not flow to the reboiler (in line 63) is returned below chimney tray 51 from
which the liquid was initially withdrawn, directly to distributor 52 that supplies the
liquid to packing (or trays) 53 below. FIG. 9 shows how the pumped system
described for FIG. 8 can be implemented in a split column approach, such as upper
column 65 and lower column 50.
One skilled in the art will recognize that the present invention gains
some of its benefit by providing a colder stream to the side reboiler(s) and/or
reboiler(s), allowing additional cooling of the column feed or feeds. This additional
cooling reduces utility requirements for a given product recovery level, or improves
product recovery levels for a given utility consumption, or some combination thereof
Further, one skilled in the art will recognize that the present invention also benefits by
introducing greater quantities of methane lower in the demethanizer to assist in
stripping carbon dioxide from the down-flowing liquids. With more methane
available for stripping the liquids, correspondingly less ethane is needed for stripping,
allowing more retention of ethane in the bottom liquid product. Therefore, the present
invention is generally applicable to any process dependent on cooling any number of
feed streams and supplying the resulting feed stream(s) to the column for distillation.
In accordance with this invention, the cooling of the demethanizer feed
streams may be accomplished in many ways. In the process of FIGS. 2,3, and 4, cold
residue gas (stream 37) and the demethanizer liquids (streams 40,42, and 43) are used
only for gas stream cooling. In the process of FIG. 5, feed stream 36 is cooled and
substantially condensed by cold residue gas (stream 37), distillation column overhead
vapor (stream 47) is cooled and partially condensed by expanded stream 36b, while
the expanded separator liquid (stream 35a) and the demethanizer liquid (stream 40)
are used only for gas cooling. However, demethanizer liquids could be used to supply
some or all of the cooling and substantial condensation of stream 36 in FIG. 5 or the
cooling and partial condensation of stream 47 in FIG. 5 in addition to or instead of gas
stream cooling. Further, any stream at a temperature colder than the feed stream
being cooled may be utilized. For instance, a side draw of vapor from the
demethanizer could be withdrawn and used for cooling. Other potential sources of
cooling include, but are not limited to, flashed high pressure separator liquids (such as


indicated by the dashed line in FIG. 3) and mechanical refrigeration systems. The
selection of a source of cooling will depend on a number of factors including, but not
limited to, inlet gas composition and conditions, plant size, heat exchanger size,
potential cooling source temperature, etc. One skilled in the art will also recognize
that any combination of the above cooling sources or methods of cooling may be
employed in combination to achieve the desired feed stream temperature(s).
In accordance with this invention, the use of external refrigeration to
supplement the cooling available to the inlet gas from other process streams may be
employed, particularly in the case of an inlet gas richer than that used in the Example.
The use and distribution of demethanizer 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.
The high pressure liquid in FIG. 5 (stream 35) can be combined with
the portion of the separator vapor (stream 33) flowing to heat exchanger 12.
Alternatively, this liquid stream (or a portion thereof) may be expanded through an
appropriate expansion device, such as expansion valve 16, and fed to a lower
mid-column feed point on the distillation column (demethanizer tower 17 in FIG. 5).
The liquid stream 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, as
illustrated in FIG. 5.
It will also be recognized that the relative amount of feed found in each
branch of the column feed streams 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
expansion machine thereby increasing the recompression horsepower requirements.
Increasing feed lower in the column reduces the horsepower consumption but may
also reduce product recovery. However, the 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. FIG. 2 is the preferred embodiment for the
compositions and pressure conditions shown. 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 36a in FIG. 5).
The fractionation towers depicted as single columns in FIGS. 2
through 5 can instead be constructed in two sections (an absorbing section and a
stripping section, for instance) because of the size of the plant. The decision whether
to construct the fractionation tower as a single vessel (such as tower 17 in FIGS. 2
through 5) or multiple vessels will depend on a number of factors such as plant size,
the distance to fabrication facilities, etc.
While there have been described what are beheved 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.
We claim. 1. In a process for the separation of a gas stream containing
-methane, C2 components, C3 components and heavier hydrocarbon components into a
volatile residue gas fraction containing a major portion of said methane and a
relatively less volatile fraction containing a major portion of said C2 components, C3
components and heavier hydrocarbon components, in which process
(a) said gas stream is treated in one or more heat exchange
I steps to produce at least a first feed stream that has been cooled under pressure;
(b) said cooled first feed stream is expanded to a lower
pressure, and thereafter supplied to a fractionation tower at a top feed point; and
(c) said cooled expanded first feed stream is fractionated at
said lower pressure whereby the components of said relatively less volatile fraction
are recovered;
the characterized inthat
(1) a liquid distillation stream is withdrawn from said
fractionation tower and heated;
(2) said heated distillation stream is returned to a lower
point on said fractionation tower that is separated from said withdrawal point by at
least one theoretical stage; and
(3) the quantities and temperatures of said feed streams to
said fractionation tower are effective to maintain the overhead temperature of said
fractionation tower at a temperature whereby the major portions of the components in
said relatively less volatile fraction are recovered.
2. In a process for the separation of a gas stream containing
methane, C2 components, C3 components and heavier hydrocarbon components into a
volatile residue gas fraction containing a major portion of said methane and a
relatively less volatile fraction containing a major portion of said C2 components, C3
components and heavier hydrocarbon components, in which process
(a) said gas sfream is freated in one or more heat exchange


steps and at least one division step to produce at least a first feed stream that has been
cooled under pressure to condense substantially all of it, and at least a second feed
stream that has been cooled under pressure;
(b) said substantially condensed first feed stream is
expanded to a lower pressure whereby it is further cooled, and thereafter directed in
heat exchanger relation with a warmer distillation stream which rises from
fractionation stages of a fractionation tower;
(c) said distillation stream is cooled by said first stream
sufGciently to partially condense it, whereupon said partially condensed distillation
stream is separated to provide said volatile residue gas fraction and a reflux stream,
with said reflux stream thereafter supplied to said fractionation tower at a top feed
point;
(d) said wanned first stream is supplied to said
fractionation tower at a first mid-column feed point;
(e) said cooled second feed stream is expanded to said
lower pressure, and thereafter supplied to said fractionation tower at a second
mid-column feed point; and
(f) said reflux stream, said heated first feed stream, and
said expanded second feed stream are fractionated at said lower pressure whereby the
components of said relatively less volatile fraction are recovered;
the improvement wherein
(1) a liquid distillation stream is withdrawn from said
fractionation tower and heated;
(2) said heated distillation stream is returned to a lower
point on said fractionation tower that is separated from said withdrawal point by at
least one theoretical stage; and
(3) the quantities and temperatures of said feed streams to
said fractionation tower are effective to maintain the overhead temperature of said
fractionation tower at a temperature whereby the major portions of the components in
said relatively less volatile fraction are recovered.
3. The Process as claimed in 1 or 2 wherein said
liquid distillation stream is pumped after being withdrawn from said fractionation
tower.
4. The Process as claimed in claim 3 wherein
(a) said pumped liquid distillation stream is divided into at
least a first portion and a second portion;
(b) said first portion is heated; and
(c) said heated first portion is returned to a lower point on
said fractionation tower that is separated from said withdrawal point by at least one
theoretical stage.
5. The Process as claimed in 1 or 2 wherein said
liquid distillation stream is directed in heat exchange relation with at least a portion of
said gas stream or said feed streams, to supply said cooling thereto and thereby heat
said liquid distillation stream.
6. Process as claimed in 3 wherein said pumped
liquid distillation stream is directed in heat exchange relation with at least a portion of
said gas stream or said feed streams, to supply said cooling thereto and thereby heat
said pumped liquid distillation stream.

7. The Process as claimed in 4 wherein said first
portion is directed in heat exchange relation with at least a portion of said gas stream
or said feed streams, to supply said cooling thereto and thereby heat said first portion.
8. The Process as claimed in 1 or 2 wherein the
quantity and temperature of said heated distillation stream and the heating supplied to
said fractionation tower are effective to maintain the bottom temperature of said
fractionation tower at a temperature to reduce the quantity of carbon dioxide
contained in said relatively less volatile fraction.
9. The Process as claimed in 3 wherein the quantity
and temperature of said heated distillation stream and the heating supplied to said
fractionation tower are effective to maintain the bottom temperature of said
fractionation tower at a temperature to reduce the quantity of carbon dioxide
contained in said relatively less volatile fraction.
10. The Process as claimed in 4 wherein the quantity
and temperature of said heated first portion and the heating supplied to said
fractionation tower are effective to maintain the bottom temperature of said
fractionation tower at a temperature to reduce the quantity of carbon dioxide
contained in said relatively less volatile fraction.
11. The Process as claimed in 5 wherein the quantity
and temperature of said heated distillation stream and the heating supplied said
fractionation tower are effective to maintain the bottom temperature of said
fractionation tower at a temperature to reduce the quantity of carbon dioxide
contained in said relatively less volatile fraction.
12. The Process as claimed in claim 6 wherein the quantity
and temperature of said heated distillation stream and the heating supplied to said
fractionation tower are effective to maintain the bottom temperature of said
fractionation tower at a temperature to reduce the quantity of carbon dioxide
contained in said relatively less volatile fraction.
13. The Process as claimed in 7 wherein the quantity
and temperature of said heated first portion and the heating supplied to said
fractionation tower are effective to maintain the bottom temperature of said
fractionation tower at a temperature to reduce the quantity of carbon dioxide
contained in said relatively less volatile fraction.

In a process for the separation of a gas stream containing methane, C2
components, C3 components and heavier hydrocarbon components into a
volatile residue gas fraction containing a major portion of said methane and a
relatively less volatile fraction containing a major portion of said C2 components,
C3 components and heavier hydrocarbon components, in which process said gas
stream is treated in one or more heat exchange steps to produce at least a first
feed stream that has been cooled under pressure; said cooled first feed stream is
expanded to a lower pressure, and thereafter supplied to a fractionation tower at
a top feed point; and said cooled expanded first feed stream is fractionated at
said lower pressure whereby the components of said relatively less volatile
fraction are recovered; characterized in that a liquid distillation stream is
withdrawn from said fractionation tower and heated; said heated distillation
stream is returned to a lower point on said fractionation tower that is separated
from said withdrawal point by at least one theoretical stage; and the quantities
and temperatures of said feed streams to said fractionation tower are effective to
maintain the overhead temperature of said fractionation tower at a temperature
whereby the major portions of the components in said relatively less volatile
fraction are recovered.

Documents:

369-kolnp-2003-abstract.pdf

369-KOLNP-2003-ASSIGNMENT 1.1.pdf

369-kolnp-2003-assignment.pdf

369-kolnp-2003-claims.pdf

369-KOLNP-2003-CORRESPONDENCE 1.1.pdf

369-kolnp-2003-correspondence.pdf

369-kolnp-2003-description (complete).pdf

369-kolnp-2003-drawings.pdf

369-kolnp-2003-examination report.pdf

369-kolnp-2003-form 18.pdf

369-kolnp-2003-form 2.pdf

369-kolnp-2003-form 26.pdf

369-kolnp-2003-form 3.pdf

369-kolnp-2003-form 5.pdf

369-kolnp-2003-form 6.pdf

369-KOLNP-2003-FORM-27.pdf

369-KOLNP-2003-OTHERS 1.1.pdf

369-KOLNP-2003-OTHERS.pdf

369-KOLNP-2003-PA.pdf

369-kolnp-2003-specification.pdf


Patent Number 242871
Indian Patent Application Number 369/KOLNP/2003
PG Journal Number 38/2010
Publication Date 17-Sep-2010
Grant Date 16-Sep-2010
Date of Filing 31-Mar-2003
Name of Patentee ELKCORP
Applicant Address WELLINGTON CENTRE, SUITE 1000, 14643 DALLAS PARKWAY, DALLAS, TX 75240-8871
Inventors:
# Inventor's Name Inventor's Address
1 WILKINSON JOHN D 2800 WEST DENGAR, MIDLAND, TX 79705
2 PIERCE MICHAEL C 1640 DOE LANE, ODESSA, TX 79762
3 HUDSON HANK M 2508 WEST SINCLAIR, MIDLAND, TX 79705
PCT International Classification Number F25J 3/02
PCT International Application Number PCT/US2001/30600
PCT International Filing date 2001-09-28
PCT Conventions:
# PCT Application Number Date of Convention Priority Country
1 NA