EP3052586B1

EP3052586B1 - Split feed addition to iso-pressure open refrigeration lpg recovery

Info

Publication number: EP3052586B1
Application number: EP14852006.7A
Authority: EP
Inventors: Ayyalasomayajula KUMAR; Robert Huebel; Michael Malsam
Original assignee: Lummus Technology Inc
Current assignee: Lummus Technology LLC
Priority date: 2013-10-09
Filing date: 2014-10-08
Publication date: 2022-06-15
Anticipated expiration: 2034-10-08
Also published as: SA516370890B1; EP3052586A1; WO2015054379A1; MX2016004599A; KR20160067957A; CN106029848A; BR112016007567A2; AU2014331943B2; CA2926654A1; JP2021047003A; EP3052586A4; PE20160625A1; AU2014331943A1; JP2016539300A; US20160258675A1

Description

RELATED APPLICATIONS

This patent claims priority from U.S. Provisional Application No. 61/888,901 filed October 9, 2013 .

FIELD

The embodiments described herein relate to improved processes for recovery of natural gas liquids from gas feed streams containing hydrocarbons, and in particular to recovery of propane and ethane from gas feed streams.

BACKGROUND

Natural gas contains various hydrocarbons, including methane, ethane and propane. Natural gas usually has a major proportion of methane and ethane, i.e. methane and ethane together typically 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. In addition to natural gas, other gas streams containing hydrocarbons may contain a mixture of lighter and heavier hydrocarbons. For example, gas streams formed in the refining process can contain mixtures of hydrocarbons to be separated. Separation and recovery of these hydrocarbons can provide valuable products that may be used directly or as feedstocks for other processes. These hydrocarbons are typically recovered as natural gas liquids (NGL).
The embodiments described herein are primarily directed to recovery of C₃+ components in gas streams containing hydrocarbons, and in particular to recovery of propane from these gas streams. A typical natural gas feed to be processed in accordance with the processes described below typically may contain, in approximate mole percent, 92.12% methane, 3.96% ethane and other C₂ components, 1.05% propane and other C₃ components, 0.15% iso-butane, 0.21% normal butane, 0.11% pentanes or heavier, and the balance made up primarily of nitrogen and carbon dioxide. Refinery gas streams may contain less methane and higher amounts of heavier hydrocarbons.
Recovery of natural gas liquids from a gas feed stream has been performed using various processes, such as cooling and refrigeration of gas, oil absorption, refrigerated oil absorption or through the use of multiple distillation towers. More recently, cryogenic expansion processes utilizing Joule-Thompson valves or turbo expanders have become preferred processes for recovery of NGL from natural gas.
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 the desired components.
The high-pressure liquids may be expanded to a lower pressure and fractionated. The expanded stream, comprising a mixture of liquid and vapor, is fractionated in a distillation column. In the distillation column volatile gases and lighter hydrocarbons are removed as overhead vapors and heavier hydrocarbon components exit as liquid product in the bottoms.
The feed gas is typically not totally condensed, and the vapor remaining from the partial condensation may be passed through a Joule-Thompson valve or a turbo expander to a lower pressure at which further liquids are condensed as a result of further cooling of the stream. The expanded stream is supplied as a feed stream to the distillation column.
A reflux stream is provided to the distillation column, typically a portion of partially condensed feed gas after cooling but prior to expansion. Various processes have used other sources for the reflux, such as a recycled stream of residue gas supplied under pressure.
While various improvements to the general cryogenic processes described above have been attempted, these improvements continue to use a turbo expander or Joule-Thompson valve to expand the feed stream to the distillation column. It would be desirable to have an improved process for enhanced recovery of NGLs from a natural gas feed stream. Processes for recovery of natural gas liquids are disclosed in US 2010/0223950 A1 and US 2013/0219957 A1 , US 2008/0115532 A1 , US 5,325,673 ..

SUMMARY

The invention is defined herein by the process described in appended claim 1, and further advantageous embodiments thereof are provided in the dependent claims. The embodiments described herein relate to improved processes for recovery of NGLs from a feed gas stream. The process utilizes an open loop mixed refrigerant process to achieve the low temperatures necessary for high levels of NGL recovery. A single distillation column is utilized to separate heavier hydrocarbons from lighter components such as sales gas. The overhead stream from the distillation column is cooled to partially liquefy the overhead stream. The partially liquefied overhead stream is separated into a vapor stream comprising lighter hydrocarbons, such as sales gas, and a liquid component that serves as a mixed refrigerant. The mixed refrigerant provides process cooling and a portion of the mixed refrigerant is used as a reflux stream to enrich the distillation column with key components. With the gas in the distillation column enriched, the overhead stream of the distillation column condenses at warmer temperatures, and the distillation column runs at warmer temperatures than typically used for high recoveries of NGLs. The process achieves high recovery of desired NGL components without expanding the gas as in a Joule-Thompson valve or turbo expander based plant, and with only a single distillation column.
C₃+ hydrocarbons, and in particular propane, are recovered. Temperatures and pressures are maintained as required to achieve the desired recovery of C₃+ hydrocarbons based upon the composition of the incoming feed stream. Feed gas enters a main heat exchanger and is cooled. The cooled feed gas is fed to a distillation column, which in this embodiment functions as a deethanizer. Cooling for the feed stream may be provided primarily by a warm refrigerant such as propane. The overhead stream from the distillation column enters the main heat exchanger and is cooled to the temperature required to produce the mixed refrigerant and to provide the desired NGL recovery from the system.
The cooled overhead stream from the distillation column is combined with an overhead stream from a reflux drum and separated in a distillation column overhead drum. The overhead vapor from the distillation column overhead drum is sales gas (i.e. methane, ethane and inert gases) and the liquid bottoms are the mixed refrigerant. The mixed refrigerant is enriched in C₂ and lighter components as compared to the feed gas. The sales gas is fed through the main heat exchanger where it is warmed. The temperature of the mixed refrigerant is reduced to a temperature cold enough to facilitate the necessary heat transfer in the main heat exchanger. The temperature of the refrigerant is lowered by reducing the refrigerant pressure across a control valve. The mixed refrigerant is fed to the main heat exchanger where it is evaporated and superheated as it passes through the main heat exchanger.
After passing through the main heat exchanger, the mixed refrigerant is compressed. Preferably, the compressor discharge pressure is greater than the distillation column pressure so no reflux pump is necessary. The compressed gas passes through the main heat exchanger, where it is partially condensed. The partially condensed mixed refrigerant is routed to a reflux drum. The bottom liquid from the reflux drum is used as a reflux stream for the distillation column. The vapors from the reflux drum are combined with the distillation column over head stream exiting the main heat exchanger and the combined stream is routed to the distillation column overhead drum. In this embodiment, the process can achieve over 99 percent recovery of propane from the feed gas.
In another embodiment of the process, the feed gas is treated as described above and a portion of the mixed refrigerant is removed from the plant following compression and cooling. The portion of the mixed refrigerant removed from the plant is fed to a C₂ recovery unit to recover the ethane in the mixed refrigerant. Removal of a portion of the mixed refrigerant stream after it has passed through the main heat exchanger and been compressed and cooled has minimal effect on the process provided that enough C₂ components remain in the system to provide the required refrigeration. In some embodiments, as much as 95 percent of the mixed refrigerant stream may be removed for C₂ recovery. The removed stream may be used as a feed stream in an ethylene cracking unit.
In another embodiment of the process, an absorber column is used to separate the distillation column overhead stream. The overhead stream from the absorber is sales gas, and the bottoms are the mixed refrigerant.
In yet another embodiment, only one separator drum is used. In this embodiment, the compressed, cooled mixed refrigerant is returned to the distillation column as a reflux stream.
A process not according to the invention may achieve separation of hydrocarbons in any manner desired. For example, the plant may be operated such that the distillation column separates C.sub.4+ hydrocarbons, primarily butane, from C₃ and lighter hydrocarbons. In another embodiment, the plant may be operated to recover both ethane and propane. In this embodiment, the distillation column is used as a demethanizer, and the plant pressures and temperatures are adjusted accordingly. In this embodiment, the bottoms from the distillation tower contain primarily the C₂+ components, while the overhead stream contains primarily methane and inert gases. In this embodiment, recovery of as much as 55 percent of the C₂+ components in the feed gas can be obtained.
Among the advantages of the process is that the reflux to the distillation column is enriched, for example in ethane, reducing loss of propane from the distillation column. The reflux also increases the mole fraction of lighter hydrocarbons, such as ethane, in the distillation column making it easier to condense the overhead stream. This process uses the liquid condensed in the distillation column overhead twice, once as a low temperature refrigerant and the second time as a reflux stream for the distillation column. Other advantages of the processes described herein will be apparent to those skilled in the art based upon the detailed description of preferred embodiments provided below.

DESCRIPTION OF THE FIGURES

FIG. 1 is a schematic drawing of a plant providing control examples of a method in which the mixed refrigerant stream is compressed and returned to the reflux separator.
FIG. 2 is a schematic drawing of a plant providing control examples of a method in which a portion of the compressed mixed refrigerant stream is removed from the plant for ethane recovery.
FIG. 3 is a schematic drawing of a plant providing control examples in which an absorber is used to separate the distillation overhead stream.
FIG. 4 is a schematic drawing of a plant providing control examples in which only one separator drum is used.
FIG. 5 is a schematic drawing of a plant for performing embodiments of a method in which the feed stream to the distillation column is split and fed to different locations of the column, and the mixed refrigerant is compressed and returned to the reflux separator.
FIG. 6 is a schematic drawing of a plant for performing embodiments of a method in which the feed stream to the distillation column is split and fed to different locations of the column, and a portion of the compressed mixed refrigerant stream is removed from the plant for ethane recovery.
FIG. 7 is a schematic drawing of a plant for performing embodiments of a method in which the feed stream to the distillation column is split and fed to different locations of the column, and an absorber is used to separate the distillation column overhead stream.
FIG. 8 is a schematic drawing of a plant for performing embodiments of a method in which the feed stream to the distillation column is split and fed to different locations of the column, and in which only one separator drum is used.
FIG. 9 is a schematic drawing of a plant for performing embodiments of another method in which the feed stream to the distillation column is split and fed to different locations of the column.

DETAILED DESCRIPTION OF EMBODIMENTS

The embodiments described herein relate to improved processes for recovery of natural gas liquids (NGL) from gas feed streams containing hydrocarbons, such as natural gas or gas streams from petroleum processing. In embodiments, the process runs at approximately constant pressures with no intentional reduction in gas pressures through the plant. The process uses a single distillation column to separate lighter hydrocarbons and heavier hydrocarbons. An open loop mixed refrigerant provides process cooling to achieve the temperatures required for high recovery of NGL gases. The mixed refrigerant is comprised of a mixture of the lighter and heavier hydrocarbons in the feed gas, and is generally enriched in the lighter hydrocarbons as compared to the feed gas.
The open loop mixed refrigerant is also used to provide an enriched reflux stream to the distillation column, which allows the distillation column to operate at higher temperatures and enhances the recovery of NGLs. The overhead stream from the distillation column is cooled to partially liquefy the overhead stream. The partially liquefied overhead stream is separated into a vapor stream comprising lighter hydrocarbons, such as sales gas, and a liquid component that serves as a mixed refrigerant.
In embodiments, the process may be used to obtain the desired separation of hydrocarbons in a mixed feed gas stream. In one embodiment, the process of the present application may be used to obtain high levels of propane recovery. Recovery of as much as 99 percent or more of the propane in the feed case may be recovered in the process. The process can also be operated in a manner not according to the invention to recover significant amounts of ethane with the propane . Alternatively, the process can be operated in a manner not according to the invention to recover a high percentage of C.sub.4+ components of the feed stream and discharge C₃ and lighter components.
According to the invention, a substantial reduction in energy usage can be obtained using the split feed configuration described herein. In embodiments, compressor duty can be reduced by at least 5%, or at least 10, or by 5-20% as compared to a system in which a split feed is not used. In embodiments, the reboiler duty can be reduced by at least 10%, or at least 20%, or at least 30% as compared to a system that does not have a split feed. In embodiments, the size of the distillation column also can be reduced, resulting in lower capital cost.
A plant for providing control examples is shown schematically in FIG. 1. It should be understood that the operating parameters for the plant, such as the temperature, pressure, flow rates and compositions of the various streams, are established to achieve the desired separation and recovery of the NGLs. The required operating parameters also depend on the composition of the feed gas. The required operating parameters can be readily determined by those skilled in the art using known techniques, including for example computer simulations. Accordingly, the descriptions and ranges of the various operating parameters provided below are intended to provide a description of specific embodiments, and they are not intended to limit the scope of the disclosure in any way.
Feed gas is fed through line (12) to main heat exchanger (10). The feed gas may be natural gas, refinery gas or other gas stream requiring separation. The feed gas is typically filtered and dehydrated prior to being fed into the plant to prevent freezing in the NGL unit. The feed gas is typically fed to the main heat exchanger at a temperature between about 43.3°C (110.degree. F.) and 54.4°C (130.degree. F.) and at a pressure between about 0.689 MPa (100 psia) and 3.10 MPa (450 psia). The feed gas is cooled and partially liquefied in the main heat exchanger (10) by making heat exchange contact with cooler process streams and with a refrigerant which may be fed to the main heat exchanger through line (15) in an amount necessary to provide additional cooling necessary for the process. A warm refrigerant such as propane may be used to provide the necessary cooling for the feed gas. The feed gas is cooled in the main heat exchanger to a temperature between about -18°C (0.degree. F.) and -40°C (-40.degree. F.).
The cool feed gas (12) exits the main heat exchanger (10) and enters the distillation column (20) through feed line (13). The distillation column operates at a pressure slightly below the pressure of the feed gas, typically at a pressure of between about 34 kPa (5 psi) and 69 kPa (10 psi) less than the pressure of the feed gas. In the distillation column, heavier hydrocarbons, such as for example propane and other C₃+ components, are separated from the lighter hydrocarbons, such as ethane, methane and other gases. The heavier hydrocarbon components exit in the liquid bottoms from the distillation column through line (16), while the lighter components exit through vapor overhead line (14). Preferably, the bottoms stream (16) exits the distillation column at a temperature of between about 65.6°C (150.degree. F.) and 149°C (300.degree. F.), and the overhead stream (14) exits the distillation column at a temperature of between about - 23°C (-10.degree. F.) and -62 °C (-80.degree. F.).
The bottoms stream (16) from the distillation column is split, with a product stream (18) and a recycle stream (22) directed to a reboiler (30) which receives heat input (Q). Optionally, the product stream (18) may be cooled in a cooler to a temperature between about 16 °C (60.degree. F.) and 54 °C (130.degree. F.). The product stream (18) is highly enriched in the heavier hydrocarbons in the feed gas stream. In the embodiment shown in FIG. 1, the product stream may highly enriched in propane and heavier components, and ethane and lighter gases are removed as sales gas as described below. Alternatively, the plant may be operated such that the product stream is heavily enriched in C.sub.4+ hydrocarbons, and the propane is removed with the ethane in the sales gas. The recycle stream (22) is heated in reboiler (30) to provide heat to the distillation column. Any type of reboiler typically used for distillation columns may be used.
The distillation column overhead stream (14) passes through main heat exchanger (10), where it is cooled by heat exchange contact with process gases to partially liquefy the stream. The distillation column overhead stream exits the main heat exchanger through line (19) and is cooled sufficiently to produce the mixed refrigerant as described below. Preferably, the distillation column overhead stream is cooled to between about -34 °C (-30.degree. F.) and -90 °C (-130.degree. F.) in the main heat exchanger.
In the control example of the process shown in FIG. 1, the cooled and partially liquefied stream (19) is combined with the overhead stream (28) from reflux separator (40) in mixer (100) and is then fed through line (32) to the distillation column overhead separator (60). Alternatively, stream (19) may be fed to the distillation column overhead separator (60) without being combined with the overhead stream (28) from reflux separator (40). Overhead stream (28) may be fed to the distillation column overhead separator directly, or in other embodiments of the process, the overhead stream (28) from reflux separator (40) may be combined with the sales gas (42). Optionally, the overhead stream from reflux separator (40) may be fed through control valve (75) prior to being fed through line (28a) to be mixed with distillation column overhead stream (19). Depending upon the feed gas used and other process parameters, control valve (75) may be used to hold pressure on the ethane compressor (80), which can ease condensing this stream and to provide pressure to transfer liquid to the top of the distillation column. Alternatively, a reflux pump can be used to provide the necessary pressure to transfer the liquid to the top of the column.
In the control example shown in FIG. 1, the combined distillation column overhead stream and reflux drum overhead stream (32) is separated in the distillation column overhead separator (60) into an overhead stream (42) and a bottoms stream (34). The overhead stream (42) from the distillation column overhead separator (60) contains product sales gas (e.g. methane, ethane and lighter components). The bottoms stream (34) from the distillation column overhead separator is the liquid mixed refrigerant used for cooling in the main heat exchanger (10).
The sales gas flows through the main heat exchanger (10) through line (42) and is warmed. In a typical plant, the sales gas exits the deethanizer overhead separator at a temperature of between about -40°C (-40.degree. F.) and -84 °C (-120.degree. F.) and a pressure of between about 0.59 MPa (85 psia) and 3.00 MPa (435 psia), and exits the main heat exchanger at a temperature of between about 37.8 °C (100.degree. F.) and 48.9 °C (120.degree. F.). The sales gas is sent for further processing through line (43).
The mixed refrigerant flows through the distillation column overhead separator bottoms line (34). The temperature of the mixed refrigerant may be lowered by reducing the pressure of the refrigerant across control valve (65). The temperature of the mixed refrigerant is reduced to a temperature cold enough to provide the necessary cooling in the main heat exchanger (10). The mixed refrigerant is fed to the main heat exchanger through line (35). The temperature of the mixed refrigerant entering the main heat exchanger is typically between about -51 °C (-60.degree. F.) to -115 °C (-175.degree. F.). Where the control valve (65) is used to reduce the temperature of the mixed refrigerant, the temperature is typically reduced by between about -6.7 °C (20.degree. F.) to 10 °C (50.degree. F.) and the pressure is reduced by between about 0.62 MPa (90 psi) to 1.72 MPa (250 psi). The mixed refrigerant is evaporated and superheated as it passes through the main heat exchanger (10) and exits through line (35a). The temperature of the mixed refrigerant exiting the main heat exchanger is between about 27 °C (80.degree. F.) and 37.8 °C (100.degree. F.).
After exiting the main heat exchanger, the mixed refrigerant is fed to ethane compressor (80). The mixed refrigerant is compressed to a pressure about 0.10 MPa (15 psi) to 0.17 MPa (25 psi) greater than the operating pressure of the distillation column at a temperature of between about 110 °C (230.degree. F.) to 177 °C (350.degree. F.). By compressing the mixed refrigerant to a pressure greater than the distillation column pressure, there is no need for a reflux pump. The compressed mixed refrigerant flows through line (36) to cooler (90) where it is cooled to a temperature of between about 21 °C (70.degree. F.) and 54.4 °C (130.degree. F.). Optionally, cooler (90) may be omitted and the compressed mixed refrigerant may flow directly to main heat exchanger (10) as described below. The compressed mixed refrigerant then flows through line (38) through the main heat exchanger (10) where it is further cooled and partially liquefied. The mixed refrigerant is cooled in the main heat exchanger to a temperature of between about -9.4 °C (15.degree. F.) to -57 °C (-70.degree. F.). The partially liquefied mixed refrigerant is introduced through line (39) to the reflux separator (40). As described previously, in the embodiment of FIG. 1, the overhead (28) from reflux separator (40) is combined with the overheads (14) from the distillation column and the combined stream (32) is fed to the distillation column overhead separator. The liquid bottoms (26) from the reflux separator (40) are fed back to the distillation column as a reflux stream (26a). Control valves (75, 85) may be used to hold pressure on the compressor to promote condensation.
The open loop mixed refrigerant used as reflux enriches the distillation column with gas phase components. With the gas in the distillation column enriched, the overhead stream of the column condenses at warmer temperatures, and the distillation column runs at warmer temperatures than normally required for high recovery of NGLs.
The reflux to the distillation column also reduces losses of heavier hydrocarbons from the column. For example, in processes for recovery of propane, the reflux increases the mole fraction of ethane in the distillation column, which makes it easier to condense the overhead stream. The process uses the liquid condensed in the distillation column overhead drum twice, once as a low temperature refrigerant and the second time as a reflux stream for the distillation column.
In another control example shown in FIG. 2, in which like numbers indicate like components and flow streams described above, the process is used to separate propane and other C₃+ hydrocarbons from ethane and light components. A tee (110) is provided in line (38) after the mixed refrigerant compressor (80) and the mixed refrigerant cooler to split the mixed refrigerant into a return line (45) and an ethane recovery line (47). The return line (45) returns a portion of the mixed refrigerant to the process through main heat exchanger (10) as described above. Ethane recovery line (47) supplies a portion of the mixed refrigerant to a separate ethane recovery unit for ethane recovery. Removal of a portion of the mixed refrigerant stream has minimal effect on the process provided that enough C₂ components remain in the system to provide the required refrigeration. In some embodiments, as much as 95 percent of the mixed refrigerant stream may be removed for C₂ recovery. The removed stream may be used, for example, as a feed stream in an ethylene cracking unit.
In another control example, the NGL recovery unit can recover significant amounts of ethane with the propane. In this control example of the process, the distillation column is a demethanizer, and the overhead stream contains primarily methane and inert gases, while the column bottoms contain ethane, propane and heavier components.
In another control example of the process, the deethanizer overhead drum may be replaced by an absorber. As shown in FIG. 3, in which like numbers indicate like components and flow streams described above, in this embodiment, the overhead stream (14) from the distillation column (20) passes through main heat exchanger (10) and the cooled stream (19) is fed to absorber (120). The overhead stream (28) from reflux separator (40) is also fed to the absorber (120). The overhead stream (42) from the absorber is the sales gas and the bottoms stream (34) from the absorber is the mixed refrigerant. The other streams and components shown in FIG. 3 have the same flow paths as described above.
In yet another control example shown in FIG. 4, in which like numbers indicate like components and flow streams described above, the second separator and the cooler are not used in the process. In this embodiment, the compressed mixed refrigerant (36) is fed through the main heat exchanger (10) and fed to the distillation tower through line (39) to provide reflux flow.
In the embodiment shown in FIG. 5, the gas feed stream (112) is split to create a first feed stream (112a) and a second feed stream (112b). The first feed stream (112a) enters the heat exchanger (110) for cooling to form a cold feed stream (113) from the heat exchanger (110) that is partially liquefied to form a stream (113) containing a mixture of liquid and vapor. The second feed stream (112b) is a warm gas by-pass feed stream that is not pre-cooled and typically is entirely in a gas phase, with no liquid. A valve (195) is provided for the second feed stream (112b) for process control purposes, including controlling the relative flow rates of the first and second feed streams (112a) and (112b) into the distillation column (120). When liquids condense in the first feed stream (112a), a portion of the condensed liquid is methane and ethane. Normally methane and ethane are overhead vapor products from the process. The second feed stream (112b) has the same overall composition as first feed stream (112a) and the cooled feed stream (113) but typically does not contain any liquid. As a result, there is a higher concentration of propane vapor and butane vapor in the by-pass gas of the second feed stream (112b) than in the cold feed stream (113). Feeding the warm by-pass gas of the second feed stream (112b) to the distillation column one or more, or 1 to 10, or 1 to 7, or 1 to 4 vapor-liquid equilibrium stages below the cold feed stream (113) allows mass transfer to exchange liquid methane and ethane for propane and butane vapor within the distillation column (120). This vaporizes methane and ethane while condensing propane and butane. By doing this, it unexpectedly increases the overall efficiency of the process by substantially reducing the refrigeration duty and reboiler duty. In embodiments, the size of the distillation column can be reduced as compared to a system that does not include second feed stream (112b).
The feed gas (112) may be natural gas, refinery gas or other gas stream requiring separation. The feed gas is typically filtered and dehydrated prior to being fed into the plant to prevent freezing in the NGL unit. The feed gas in the first feed stream (112a) is typically fed to the main heat exchanger at a temperature between about 43.3 °C (110.degree. F.) and 54.4°C (130.degree. F.) and at a pressure between about 0.689 MPa (100 psia) and 3.10 MPa (450 psia). The feed gas is cooled and partially liquefied in the main heat exchanger (110) by making heat exchange contact with cooler process streams and with a refrigerant which may be fed to the main heat exchanger through line (115) in an amount necessary to provide additional cooling necessary for the process. A warm refrigerant such as propane may be used to provide the necessary cooling for the feed gas. The feed gas is cooled in the main heat exchanger to a temperature between about -17 °C (0.degree. F.) and -40 °C (-40.degree. F.).
The cool feed gas (112a) exits the main heat exchanger (110) and enters the distillation column (120) through feed line (113). The distillation column operates at a pressure slightly below the pressure of the feed gas, typically at a pressure of between about 5 psi and 10 psi less than the pressure of the feed gas. In the distillation column, heavier hydrocarbons, such as for example propane and other C₃+ components, are separated from the lighter hydrocarbons, such as ethane, methane and other gases. The heavier hydrocarbon components exit in the liquid bottoms from the distillation column through line (116), while the lighter components exit through vapor overhead line (114). Preferably, the bottoms stream (116) exits the distillation column at a temperature of between about 65.6 °C (150.degree. F.) and 149 °C (300.degree. F.), and the overhead stream (114) exits the distillation column at a temperature of between about -23 °C (-10.degree. F.) and -62 °C (-80.degree. F.).
The bottoms stream (116) from the distillation column is split, with a product stream (118) and a recycle stream (122) directed to a reboiler (130) which receives heat input (Q). Optionally, the product stream (118) may be cooled in a cooler to a temperature between about 16 °C (60.degree. F.) and 54.4 °C (130.degree. F.) The product stream (118) is highly enriched in the heavier hydrocarbons in the feed gas stream. In the embodiment shown in FIG. 5, the product stream may be highly enriched in propane and heavier components, and ethane and lighter gases are removed as sales gas in the sales gas line (143) as described below. Alternatively, the plant may be operated such that the product stream is heavily enriched in C.sub.4+ hydrocarbons, and the propane is removed with the ethane in the sales gas. The recycle stream (122) is heated in reboiler (130) to provide heat to the distillation column. Any type of reboiler typically used for distillation columns may be used.
The distillation column overhead stream (114) passes through main heat exchanger (110), where it is cooled by heat exchange contact with process gases to partially liquefy the stream. The distillation column overhead stream exits the main heat exchanger through line (119) and is cooled sufficiently to produce the mixed refrigerant as described below. Preferably, the distillation column overhead stream is cooled to between about -30.degree. F. and -130.degree. F. in the main heat exchanger.
In the embodiment of the process shown in FIG. 5, the cooled and partially liquefied stream (119) is combined with the overhead stream (128) from reflux separator (140) in mixer (200) and is then fed through line (132) to the distillation column overhead separator (160). Alternatively, stream (119) may be fed to the distillation column overhead separator (160) without being combined with the overhead stream (128) from reflux separator (140). Overhead stream (128) may be fed to the distillation column overhead separator directly, or in other embodiments of the process, the overhead stream (128) from reflux separator (140) may be combined with the sales gas (142). Optionally, the overhead stream from reflux separator (140) may be fed through control valve (175) prior to being fed through line (128a) to be mixed with distillation column overhead stream (119). Depending upon the feed gas used and other process parameters, control valve (175) may be used to hold pressure on the ethane compressor (180), which can ease condensing this stream and to provide pressure to transfer liquid to the top of the distillation column. Alternatively, a reflux pump can be used to provide the necessary pressure to transfer the liquid to the top of the column.
In the embodiment shown in FIG. 5, the combined distillation column overhead stream and reflux drum overhead stream (132) is separated in the distillation column overhead separator (160) into an overhead stream (142) and a bottoms stream (134). The overhead stream (142) from the distillation column overhead separator (160) contains product sales gas (e.g. methane, ethane and lighter components). The bottoms stream (134) from the distillation column overhead separator is the liquid mixed refrigerant used for cooling in the main heat exchanger (110).
The sales gas flows through the main heat exchanger (110) through line (142) and is warmed. In a typical plant, the sales gas exits the deethanizer overhead separator at a temperature of between about -40 °C (-40.degree. F.) and -84 °C (-120.degree. F.) and a pressure of between about 0.59 MPa (85 psia) and 3.00 MPa (435 psia), and exits the main heat exchanger at a temperature of between about 37.8 °C (100.degree. F.) and 48.9 °C (120.degree. F.). The sales gas is sent for further processing through line (143).
The mixed refrigerant flows through the distillation column overhead separator bottoms line (134). The temperature of the mixed refrigerant may be lowered by reducing the pressure of the refrigerant across control valve (165). The temperature of the mixed refrigerant is reduced to a temperature cold enough to provide the necessary cooling in the main heat exchanger (110). The mixed refrigerant is fed to the main heat exchanger through line (135). The temperature of the mixed refrigerant entering the main heat exchanger is typically between about -51 °C (-60.degree. F.) to -115 °C (-175.degree. F.). Where the control valve (165) is used to reduce the temperature of the mixed refrigerant, the temperature is typically reduced by between about -6.7 °C (20.degree. F.) to 10 °C (50.degree. F.) and the pressure is reduced by between about 0.62 MPa (90 psi) to 1.72 MPa (250 psi). The mixed refrigerant is evaporated and superheated as it passes through the main heat exchanger (110) and exits through line (135a). The temperature of the mixed refrigerant exiting the main heat exchanger is between about 27 °C (80.degree. F.) and 38 °C (100.degree. F.).
After exiting the main heat exchanger, the mixed refrigerant is fed to ethane compressor (180). The mixed refrigerant is compressed to a pressure about 0.10 MPa 15 psi to 0.17 MPa (25 psi) greater than the operating pressure of the distillation column at a temperature of between about 110 °C (230.degree. F.) to 177 °C (350.degree. F.) By compressing the mixed refrigerant to a pressure greater than the distillation column pressure, there is no need for a reflux pump. The compressed mixed refrigerant flows through line (136) to cooler (190) where it is cooled to a temperature of between about 21 °C (70.degree. F.) and 54.4 °C (130.degree. F.). Optionally, cooler (190) may be omitted and the compressed mixed refrigerant may flow directly to main heat exchanger (110) as described below. The compressed mixed refrigerant then flows through line (138) through the main heat exchanger (110) where it is further cooled and partially liquefied. The mixed refrigerant is cooled in the main heat exchanger to a temperature of between about -9.4 °C (15.degree. F.) to -57 °C (-70.degree. F). The partially liquefied mixed refrigerant is introduced through line (139) to the reflux separator (140). As described previously, in the embodiment of FIG. 5, the overhead (128) from reflux separator (140) is combined with the overheads (114) from the distillation column and the combined stream (132) is fed to the distillation column overhead separator. The liquid bottoms (126) from the reflux separator (140) are fed back to the distillation column as a reflux stream (126). Control valves (175, 185) may be used to hold pressure on the compressor to promote condensation.
In the embodiment shown in FIG. 6 , the gas feed stream (212) is split to create a first feed stream (212a) and a second feed stream (212b). The first feed stream (212a) enters the heat exchanger (210) for cooling to form a cold feed stream (213) from the heat exchanger (210) that is partially liquefied. The second feed stream (212b) is a warm gas by-pass feed stream that is not pre-cooled and typically is in an entirely gas phase, with no liquid. A valve (295) is provided for the second feed stream (212b) for process control purposes. When liquids condense, a portion of the condensed liquid is methane and ethane. Normally methane and ethane are overhead vapor products from the process. The second feed stream (212b) has the same composition as the first feed stream (212a) but contains less liquid (and typically is entirely in the gas phase). As a result, there is a higher concentration of propane vapor and butane vapor in the by-pass gas of second feed stream (212b) than in the cold feed stream (213). Placing the warm by-pass gas of the second feed stream (212b) one or more, or 1 to 10, or 1 to 7, or 1 to 4 vapor-liquid equilibrium stages below the cold feed stream (213) allows mass transfer to exchange liquid methane and ethane for propane and butane vapor within the distillation column (220). This vaporizes methane and ethane while condensing propane and butane. By doing this, it unexpectedly increases the overall efficiency of the process by substantially reducing the refrigeration duty and reboiler duty. In embodiments, the size of the distillation column can be reduced as compared to a system that does not include stream (212b).
In the embodiment of Fig. 6 , the process is used to separate propane and other C₃+ hydrocarbons from ethane and light components. A tee (310) is provided in line (238) after the mixed refrigerant compressor (280) and the mixed refrigerant cooler to split the mixed refrigerant into a return line (245) and an ethane recovery line (247). The return line (245) returns a portion of the mixed refrigerant to the process through main heat exchanger (210) as described above. Ethane recovery line (247) supplies a portion of the mixed refrigerant to a separate ethane recovery unit for ethane recovery. Removal of a portion of the mixed refrigerant stream has minimal effect on the process provided that enough C₂ components remain in the system to provide the required refrigeration. In some embodiments, as much as 95 percent of the mixed refrigerant stream may be removed for C₂ recovery. The removed stream may be used, for example, as a feed stream in an ethylene cracking unit.
In the embodiment shown in FIG. 7 , the gas feed stream (312) is split to create a first feed stream (312a) that enters the heat exchanger (310) for cooling to form a cold feed stream (313) from the heat exchanger (310) that is partially liquefied, and a second feed stream (312b) that is a warm gas by-pass feed stream that is not pre-cooled. A valve (395) is provided for the second feed stream (312b) for process control purposes. When liquids condense in first feed stream (312a), a portion of the condensed liquid is methane and ethane. Normally methane and ethane are overhead vapor products from the process. The second feed stream (312b) has the same composition as the first feed stream (312a) but contains less liquid. As a result, there is a higher concentration of propane vapor and butane vapor in the by-pass gas of second feed stream (312b) than in the cold feed stream (313). Placing the warm by-pass gas of the second feed stream (312b) one or more, or 1 to 10, or 1 to 7, or 1 to 4, vapor-liquid equilibrium stages below the cold feed stream (313) allows mass transfer to exchange liquid methane and ethane for propane and butane vapor within the distillation column (320). This vaporizes methane and ethane while condensing propane and butane. By doing this, it unexpectedly increases the overall efficiency of the process by substantially reducing the refrigeration duty and reboiler duty. In embodiments, the size of the distillation column can be reduced as compared to a system that does not include stream (312b).
As is shown in Fig. 7, the deethanizer overhead drum may be replaced by an absorber. In this embodiment, the overhead stream (314) from the distillation column (320) passes through main heat exchanger (310) and the cooled stream (319) is fed to an absorber (321). The overhead stream (328) from reflux separator (340) is also fed to the absorber (321) through line (332). The overhead stream (342) from the absorber (321) is the sales gas and the bottoms stream (334) from the absorber (321) is the mixed refrigerant. The other streams and components shown in FIG. 7 have the same flow paths as described above.
In yet another embodiment shown in FIG. 8 the second separator and the cooler are not used in the process. In this embodiment, the compressed mixed refrigerant (436) is fed through the main heat exchanger (410) and fed to the distillation column (420) through line (439) to provide reflux flow.
In the embodiment shown in FIG. 8, the gas feed stream (412) is split to create a first feed stream (412a) that enters the heat exchanger (410) for cooling to form a cold feed stream (413) from the heat exchanger (410) that is partially liquefied, and a second feed stream (412b) that is a warm gas by-pass feed stream that is not pre-cooled. A valve (495) is provided for the second feed stream (412b) for process control purposes. When liquids condense, a portion of the condensed liquid is methane and ethane. Normally methane and ethane are overhead vapor products from the process. The second feed stream (412b) has the same composition as the first feed stream (412a) but contains less liquid. As a result, there is a higher concentration of propane vapor and butane vapor in the by-pass gas of second feed stream (412b) than in the cold feed stream (413). Placing the warm by-pass gas of the second feed stream (412b) one or more, or 1 to 10, or 1 to 7, or 1 to 4 vapor-liquid equilibrium stages below the cold feed stream (413) allows mass transfer to exchange liquid methane and ethane for propane and butane vapor within the distillation column (420). This vaporizes methane and ethane while condensing propane and butane. By doing this, it unexpectedly increases the overall efficiency of the process by substantially reducing the refrigeration duty and reboiler duty. In embodiments, the size of the distillation column can be reduced as compared to a system that does not include stream (412b).
In yet another embodiment shown in FIG. 9 , the split feed scheme is incorporated into a system that is somewhat similar to a process described in US Patent No. 8,627,681 , the contents of which are incorporated herein by reference in their entirety. The benefits of the embodiment of Fig.9 are surprisingly and unexpectedly discovered that there will be a decrease in refrigeration duty specification, decrease in deethanizer reboiler duty specification, decrease in deethanizer vapor and liquid traffic thus providing for a distillation column sizing decrease, and a decrease in the refrigeration and reboiler duty specification with high pressure feeds. The total propane and mixed refrigerant compressor duty is over 11 percent higher without the split feed. As is shown, considerable economic benefits from reduced total invested cost and operational costs can be obtained as a result of these unexpected improvements.
More specifically, the overall process of FIG. 9 is designated as 502. Feed stream (512) is split to create first feed stream (512a) and second feed stream (512b). First feed stream (512a) enters the heat exchanger (510) for cooling to form a cold or high pressure stream (513) from the heat exchanger (510) that is partially liquefied. Warm vapor by-pass stream (512b) is a second stream that is not pre-cooled. Stream (512b) passes through control valve (605) to reduce its pressure and is then fed to the middle of a distillation column (520) at a location that is one or more, or 1 to 10, or 1 to 7, or 1 to 4 vapor-liquid equilibrium stages below the entry point of stream 513.
Although a multi-pass heat exchanger (510) is illustrated, use of multiple heat exchangers may be used to achieve similar results, as is also the case with the embodiments shown in Figs. 5-8. The feed stream (512) may be natural gas, refinery gas or other gas stream requiring separation. The feed gas is typically filtered and dehydrated prior to being fed into the plant to prevent freezing in the NGL unit. In embodiments, the first feed stream (512a) is typically fed to the main heat exchanger at a temperature between about 43.degree. C. and 54.degree. C. (110.degree. F. and 130.degree. F.) and at a pressure between about 7 bar and 31 bar (100 psia and 450 psia). The first feed stream (512a) is cooled and partially liquefied in the main heat exchanger 510 via indirect heat exchange with cooler process streams and/or with a refrigerant which may be fed to the main heat exchanger via line (515) in an amount necessary to provide additional cooling necessary for the process. A warm refrigerant such as propane, for example, may be used to provide the necessary cooling for the feed gas. The feed gas may be cooled in the main heat exchanger to a temperature between about -18.degree. C. and - 40.degree. C. (0.degree. F. and -40.degree. F.).
The cool feed gas exits the main heat exchanger (510) and is fed to distillation column (520) via feed line (513). Distillation column (520) operates at a pressure slightly below the pressure of the feed gas, typically at a pressure about 0.3 to 0.7 bar (5 to 10 psi) less than the pressure of the feed gas. In the distillation column, heavier hydrocarbons, such as propane and other C.sub.3+ components, are separated from the lighter hydrocarbons, such as ethane, methane and other gases. The heavier hydrocarbon components exit in the liquid bottoms from the distillation column through line (516), while the lighter components exit through vapor overhead line (514). In embodiments, the bottoms stream 516 exits the distillation column at a temperature between about 65.degree. C. and 149.degree. C. (150.degree. F. and 300.degree. F.), and the overhead stream 14 exits the distillation column at a temperature of between about -23.degree. C. and -62.degree. C. (-10.degree. F. and -80.degree. F.).
The bottoms stream (516) from the distillation column is split, with a product stream (518) and a reboil stream (522) directed to a reboiler (530). Optionally, the product stream (518) may be cooled in a cooler (not shown) to a temperature between about 515.degree. C. and 554.degree. C. (60.degree. F. and 130.degree. F.). The product stream (518) is highly enriched in the heavier hydrocarbons in the feed gas stream. In the embodiment shown in FIG. 9, the product stream may be enriched in propane and heavier components, and ethane and lighter gases are further processed as described below. Alternatively, the plant may be operated such that the product stream is heavily enriched in C.sub.4+ hydrocarbons, and the propane is removed with the ethane in the sales gas produced. The reboil stream (522) is heated in reboiler (530) to provide heat to the distillation column. Any type of reboiler typically used for distillation columns may be used.
The distillation column overhead stream (514) passes through main heat exchanger (510), where it is cooled by indirect heat exchange with process gases to at least partially liquefy or completely (100%) liquefy the stream. The distillation column overhead stream exits the main heat exchanger (510) through line (519) and is cooled sufficiently to produce the mixed refrigerant as described below. In some embodiments, the distillation column overhead stream is cooled to between about -34.degree. C. and - 90.degree. C. (-30.degree. F. and -130.degree. F.) in main heat exchanger 510.
The cooled and partially liquefied stream (519) and the overhead stream (528) (stream 532 following control valve 575) from reflux separator (540) may be fed to distillation column overhead separator (585).
The components in distillation column overhead stream (519) and reflux drum overhead stream (532) are separated in overhead separator (585) into an overhead stream (542), a side draw fraction (551), and a bottoms stream (534). The overhead stream (542) from distillation column overhead separator (585) contains methane, ethane, nitrogen, and other lighter components, and is enriched in nitrogen content. Side draw fraction (551) may be of intermediate nitrogen content. The bottoms stream (534) from distillation column overhead separator (585) is the liquid mixed refrigerant used for cooling in the main heat exchanger (510), which may be depleted in nitrogen content. The side draw fraction may be reduced in pressure across flow valve (595), fed to heat exchanger (510) for use in the integrated heat exchange system, and recovered via flow line (552).
The components in overhead stream (542) are fed to main heat exchanger (510) and warmed. In a typical plant, the overhead fraction recovered via stream (542) from overhead separator (585) is at a temperature between about -40.degree. C. and - 84.degree. C. (-40.degree. F. and -120.degree. F.) and at a pressure between about 5 bar and 30 bar (85 psia and 435 psia). Following heat exchange in main heat exchanger (510), the overhead fraction recovered from heat exchanger 510 via stream (543) may be at a temperature between about 37.degree. C. and 49.degree. C. (100.degree. F. and 120.degree. F.). The overhead fraction is enriched in nitrogen content and may be recovered via stream (543) as a low-btu natural gas stream.
The mixed refrigerant, as mentioned above, is recovered from distillation column overhead separator (585) via bottoms line (534). The temperature of the mixed refrigerant may be lowered by reducing the pressure of the refrigerant across control valve (565). The temperature of the mixed refrigerant is reduced to a temperature cold enough to provide the necessary cooling in the main heat exchanger (510). The mixed refrigerant is fed to the main heat exchanger through line (535). The temperature of the mixed refrigerant entering the main heat exchanger is typically between about - 51.degree. C. and -115.degree. C. (-60.degree. F. to -175.degree. F.).
Where the control valve (565) is used to reduce the temperature of the mixed refrigerant, the temperature is typically reduced by about 6.degree. C. to 10.degree. C. (20.degree. F. to 50.degree. F.) and the pressure is reduced by about 6 bar to 17 bar (90 to 250 psi). The mixed refrigerant is evaporated and superheated as it passes through the main heat exchanger 510 and exits through line (535a).
The temperature of the mixed refrigerant exiting the main heat exchanger is between about 26.degree. C. and 38.degree. C. (80.degree. F. and 100.degree. F.).
After exiting main heat exchanger (510), the mixed refrigerant is fed to compressor (580). The mixed refrigerant is compressed to a pressure 1 bar to 2 bar (15 psi to 25 psi) greater than the operating pressure of the distillation column, and at a temperature between about 110.degree. C. to 177.degree. C. (230.degree. F. to 350.degree. F.). By compressing the mixed refrigerant to a pressure greater than the distillation column pressure, there is no need for a reflux pump. The compressed mixed refrigerant flows through line (536) to cooler (590) where it is cooled to a temperature between about 21.degree. C. and 54.degree. C. (70.degree. F. and 130.degree. F.). Optionally, cooler (590) may be omitted and the compressed mixed refrigerant may flow directly to main heat exchanger (510). The compressed mixed refrigerant then flows via line (538) through the main heat exchanger (510) where it is further cooled and partially liquefied.
The mixed refrigerant is cooled in the main heat exchanger to a temperature from about -9.degree. C. to -57.degree. C. (15.degree. F. to -70.degree. F.) The partially liquefied mixed refrigerant is introduced through line (539) to reflux separator (540). As described previously, the overheads (528) from reflux separator (540) and overheads (514) from the distillation column (520) are fed to the distillation column overhead separator (585). The liquid bottoms (526) from the reflux separator (540) are fed back to the distillation column (520) as a reflux stream (526). Control valves (575), (586) may be used to hold pressure on the compressor to promote condensation.
The mixed refrigerant used as reflux (fed via stream 526a) enriches distillation column (520) with gas phase components. With the gas in the distillation column enriched, the overhead stream of the column condenses at warmer temperatures, and the distillation column runs at warmer temperatures than normally required for a high recovery of NGLs.
The reflux to distillation column (520) also reduces heavier hydrocarbons in the overheads fraction. For example, in processes for recovery of propane, the reflux increases the mole fraction of ethane in the distillation column, which makes it easier to condense the overhead stream. The process uses the liquid condensed in the distillation column overhead separator twice, once as a low temperature refrigerant and the second time as a reflux stream for the distillation column.
At least a portion of the mixed refrigerant in flow line (528), having a very low nitrogen content, may be withdrawn via flow stream (532ex) prior to separator (585). In some embodiments, the portion withdrawn via flow stream (532ex) may be used for pipeline sales. In other embodiments, a mixed refrigerant stream (532ex), having less than 1 mole % nitrogen, may be mixed with a high or intermediate btu natural gas process stream having greater than 4% nitrogen to result in a pipeline sales stream having 4% or less nitrogen.
For example, mixed refrigerant stream (532ex) may be combined with intermediate btu natural gas in stream (551) (side draw) to result in a natural gas stream suitable for pipeline sales. The flow rates of streams (532ex) and (551) may be such that the resulting product stream (548) has a nitrogen (inert) content of less than 4 mole %. In some embodiments, flow stream (532ex) may be fed to main heat exchanger (510); and following heat transfer, the mixed refrigerant may be recovered from heat exchanger (510) via flow line (541) for admixture with intermediate btu stream (551). Other process streams may also be admixed with mixed refrigerant stream (532ex) in other embodiments.
Processes according to the embodiment of Fig. 9 allows for substantial process flexibility, providing for the ability to efficiently process feed gas streams having a wide range of nitrogen content. The embodiment described with regard to FIG. 9 allows for recovery of a majority of the feed gas btu value as a natural gas sales stream. Isopressure open refrigeration processes according to embodiments disclosed herein may additionally include separation of nitrogen from high or intermediate nitrogen content streams, allowing for additional recovery of btu value or additional flexibility with regard to process conditions and feed gas nitrogen content.
Examples of specific embodiments of the processes are described below. These examples are provided to further describe the processes described herein and they are not intended to limit the full scope of the disclosure in any way.

CONTROL EXAMPLE 1

In the following examples, operation of the processing plant shown in FIG. 1 with different types and compositions of feed gas were computer simulated using process the Apsen HYSYS simulator. In this example, the operating parameters for C₃+ recovery using a relatively lean feed gas are provided. Table 1 shows the operating parameters for propane recovery using a lean feed gas. The composition of the feed gas, the sales gas stream and the C₃+ product stream, and the mixed refrigerant stream in mole fractions are provided in Table 2. Energy inputs for this embodiment included about 3.92×10⁵ MJ/hr (3.717×10⁵5 Btu/hr) (Q) to the reboiler (30) and about 342 kWatt (459 horsepower) (P) to the ethane compressor (80).
As can be seen in Table 2, the product stream (18) from the bottom of the distillation column is highly enriched in C₃+ components, while the sales gas stream (43) contains almost entirely C₂ and lighter hydrocarbons and gases. Approximately 99.6% of the propane in the feed gas is recovered in the product stream. The mixed refrigerant is comprised primarily of methane and ethane, but contains more propane than the sales gas.

CONTROL EXAMPLE 2

In this example, operating parameters are provided for the processing plant shown in FIG. 1 using a refinery feed gas for recovery of C₃+ components in the product stream. Table 3 shows the operating parameters using the refinery feed gas. The composition of the feed gas, the sales gas stream and the C₃+ product stream, and the mixed refrigerant stream in mole fractions are provided in Table 4. Energy inputs for this embodiment included about 2.32×10⁶ MJ/hr (2.205×10⁶ Btu/hr) (Q) to the reboiler (30) and about 170 kWatt (228 horsepower) (P) to the ethane compressor (80).
As can be seen in Table 4, the product stream (18) from the bottom of the distillation column is highly enriched in C₃+ components, while the sales gas stream (43) contains almost entirely C₂ and lighter hydrocarbons and gases, in particular hydrogen. This stream could be used to feed a membrane unit or PSA to upgrade this stream to useful hydrogen. Approximately 97.2% of the propane in the feed gas is recovered in the product stream. The mixed refrigerant is comprised primarily of methane and ethane, but contains more propane than the sales gas.

CONTROL EXAMPLE 3

In this example, operating parameters are provided for the processing plant shown in FIG. 1 using a refinery feed gas for the recovery of C.sub.4+ components in the product stream, with the C₃ components removed in the sales gas stream. Table 5 shows the operating parameters for this embodiment of the process. The composition of the feed gas, the sales gas stream and the C.sub.4+ product stream, and the mixed refrigerant stream in mole fractions are provided in Table 6. Energy inputs for this embodiment included about 2.65 ×10⁶ MJ/hr (2.512×10⁶ Btu/hr) (Q) to the reboiler (30) and about 148 kWatt (198 horsepower) (P) to the ethane compressor (80).
As can be seen in Table 6, in this embodiment, the product stream (18) from the bottom of the distillation column is highly enriched in C.sub.4+ components, while the sales gas stream (43) contains almost entirely C₃ and lighter hydrocarbons and gases. Approximately 99.7% of the C.sub.4+ components in the feed gas is recovered in the product stream. The mixed refrigerant is comprised primarily of C₃ and lighter components, but contains more butane than the sales gas.

CONTROL EXAMPLE 4

In this example, operating parameters are provided for the processing plant shown in FIG. 2 using a refinery feed gas for recovery of C₃+ components in the product stream, with the C₂ and lighter components removed in the sales gas stream. In this embodiment, a portion of the mixed refrigerant is removed through line (47) and fed to an ethane recovery unit for further processing. Table 7 shows the operating parameters for this embodiment of the process. The composition of the feed gas, the sales gas stream and the C₃+ product stream, and the mixed refrigerant stream in mole fractions are provided in Table 8. Energy inputs for this embodiment included about 2.204 ×10⁶ MJ/hr (2.089×10⁶ Btu/hr) (Q) to the reboiler (30) and about 291 kWatt (391 horsepower) (P) to the ethane compressor (80).
As can be seen in Table 8, in this embodiment, the product stream (18) from the bottom of the distillation column is highly enriched in C₃+ components, while the sales gas stream (43) contains almost entirely C₂ and lighter hydrocarbons and gases. The mixed refrigerant is comprised primarily of C₂ and lighter components, but contains more propane than the sales gas.

CONTROL EXAMPLE 5

In this example, operating parameters are provided for the processing plant shown in FIG. 3 using a lean feed gas for recovery of C₃+ components in the product stream, with the C₂ and lighter components removed in the sales gas stream. In this embodiment, an absorber (120) is used to separate the distillation column overhead stream and the reflux separator overhead stream to obtain the mixed refrigerant. Table 9 shows the operating parameters for this embodiment of the process. The composition of the feed gas, the sales gas stream and the C₃+ product stream, and the mixed refrigerant stream in mole fractions are provided in Table 10. Energy inputs for this embodiment included about 3.940 ×10⁵ MJ/hr (3.734×10⁵ Btu/hr) (Q) to the reboiler (30) and about 235 kWatt (316 horsepower) (P) to the ethane compressor (80).
As can be seen in Table 10, in this embodiment, the product stream (18) from the bottom of the distillation column is highly enriched in C₃+ components, while the sales gas stream (43) contains almost entirely C₂ and lighter hydrocarbons and gases. The mixed refrigerant is comprised primarily of C₂ and lighter components, but contains more propane than the sales gas.

CONTROL EXAMPLE 6

In this example, operating parameters are provided for the processing plant shown in FIG. 1 using a rich feed gas for the recovery of C₃+ components in the product stream, with the C₂ components removed in the sales gas stream. Table 11 shows the operating parameters for this embodiment of the process. The composition of the feed gas, the sales gas stream and the C₃+ product stream, and the mixed refrigerant stream in mole fractions are provided in Table 12. Energy inputs for this embodiment included about 1.538×10⁶ MJ/hr (1.458×10⁶ Btu/hr) (Q) to the reboiler (30) and about 168 kWatt (226 horsepower) (P) to the ethane compressor (80).
As can be seen in Table 12, in this embodiment, the product stream (18) from the bottom of the distillation column is highly enriched in C₃+ components, while the sales gas stream (43) contains almost entirely C₂ and lighter hydrocarbons and gases. The mixed refrigerant is comprised primarily of C₂ and lighter components, but contains more propane than the sales gas.

EXAMPLE 7

In this example, operating parameters comparable to the prior control examples are provided for a simulated processing plant shown in FIG. 5 using the rich feed gas of Control Example 6 for the recovery of C₃+ components in the product stream, with the C₂ components removed in the sales gas stream. Energy inputs for this embodiment included about 1.178×10⁶ MJ/hr (1.117 ×10⁶ Btu/hr) (Q) to the reboiler (130) and a reduced horsepower to the ethane compressor (180). In this embodiment, about 15 weight % of the gas feed stream (112) formed the bypass stream (112b) and the remainder of stream (112) formed the first feed stream (112a).
As was the case in the prior control examples, in this embodiment, the product stream (118) from the bottom of the distillation column is highly enriched in C₃+ components, while the sales gas stream (143) contains almost entirely C₂ and lighter hydrocarbons and gases. The mixed refrigerant is comprised primarily of C₂ and lighter components, but contains more propane than the sales gas. Based on process simulation data, with the insertion of the feed stream (112b) into the column (120) as shown, it was discovered that there was an unexpected and surprising decrease in refrigeration duty specification, decrease in deethanizer reboiler duty specification, decrease in deethanizer vapor and liquid traffic thus providing for a distillation column sizing decrease, and a decrease in the refrigeration and reboiler duty specification with high pressure feeds. The total propane and mixed refrigerant compressor duty is over 12 percent higher without the split feed. This results in significant economic savings in both total invested cost (TIC) in the plant and operational costs. By way of illustration, in the USA, 200 MMSCFD gas plants are fairly typical. Such a plant may have about 15,000 HP of refrigeration compressors, depending of feed composition. Based on the following calculation, configuration results in about a 12 percent saving in compressor duty. 15,000 HP × 0.12 Savings × 0.746 kw/hp × 0.1 $/kwh= 134 $/hr, which is over a million dollars per year of electric power. For conversions of the units into SI-units see paragraph [00106].
For a 200 MMSCFD plant, a reboiler duty without the split feed is approximately 29.2 MMBTU/HR. In this Example, the reboiler duty is about 22.2 MMBTU/HR. Assuming an energy cost of US 5.00 MMBTU, annual savings would be about $307,000. For conversions of the units into SI-units see paragraph [00106]

EXAMPLE 8

In this set of examples, operating parameters comparable to the prior control examples were provided for a simulated processing plant shown in FIG. 5 using the same lean feed gas and product stream compositions as were used in Control Example 1 for the recovery of C₃+ components in the product stream, with the C₂ components removed in the sales gas stream. The by-pass feed stream contained about 10-15 weight % of the feed gas stream 112. Energy inputs for this embodiment were about 20-27 % lower than the energy input for Control Example 1. This set of examples resulted in significant economic savings in both total invested cost (TIC) in the plant and operational costs.

PROPHETIC EXAMPLE 9

In this example, operating parameters comparable to the prior control examples are provided for a simulated processing plant shown in FIG. 6 using the same lean feed gas and product stream compositions as were used in Control Example 4 for the recovery of C₃+ components in the product stream, with the C₂ components removed in the sales gas stream. The by-pass feed stream contains about 10-15 weight % of the feed gas stream 212. Energy inputs for this embodiment are lower than the energy input for Control Example 4. This embodiment results in significant economic savings in both total invested cost (TIC) in the plant and operational costs.

PROPHETIC EXAMPLE 10

In this example, operating parameters comparable to the prior control examples are provided for a simulated processing plant shown in FIG. 7 using the same lean feed gas and product stream compositions as were used in Control Example 5 for the recovery of C₃+ components in the product stream, with the C₂ components removed in the sales gas stream. The by-pass stream contains about 10-15 weight % of the feed gas stream 312. Energy inputs for this embodiment are lower than the energy input for Control Example 5. This embodiment results in significant economic savings in both total invested cost (TIC) in the plant and operational costs.
While specific embodiments have been described above, one skilled in the art will recognize that numerous variations or changes may be made to the process described above without departing from the scope as recited in the appended claims. Accordingly, the foregoing description of preferred embodiments is intended to describe the embodiments an exemplary, rather than a limiting, sense.

In the appended Tables 1, 3, 5, 7, 9 and 11, quantities are provided in angloamerican units and corresponding conversion to the SI system may be performed based on the following formula: 1 Btu/hr ≈ 0.29307 W; 1 psia ≈ 6894 Pa; 1 hp ≈ 745.70 W; 1°C = (1°F - 32) × 5/9; 1 barrel /day ≈ 159 I/24h, 1 MMSCFD ≈ 1178 Sm³/h; Ib/hr ≈ 0.4536 kg/hr.

TABLE 1

Material Streams

		12	13	19	15	17	14	18	32	34	42
Vapour Fraction		1.0000	0.9838	0.3989	0.0000	0.5000	1.0000	0.0000	0.6145	0.0000	1.0000
Temperature	F	120.0	-25.00	-129.0	-30.00	-29.68	-76.88	251.9	-118.6	-118.7	-118.7
Pressure	psia	415.0	410.0	400.0	21.88	20.88	405.0	410.0	400.0	400.0	400.0
Molar Flow	MMSCFD	10.00	10.00	11.76	1.317	1.317	11.76	0.2517	15.89	6.139	9.723
Mass Flow	Ib/hr	1.973e+004	1.973e+004	2.362e+004	6356	6356	2.362e+004	1671	3.220e+004	1.414e+004	1.800e+004
Liquid Volume Flow	barrel/day	4203	4203	5100	862.2	862.2	5100	196.3	6931	2925	3995

		43	35	35a	36	38	39	28	26	26a	28a
Vapour Fraction		1.0000	0.2758	1.0000	1.0000	1.0000	0.6723	1.0000	0.0000	0.0452	.09925
Temperature	F	110.0	-165.0	90.00	262.2	120.0	-63.00	-63.00	-63.00	-68.04	-69.27
Pressure	psia	395.0	149.9	144.9	470.0	465.0	460.0	460.0	460.0	415.0	400.0
Molar Flow	MMSCFD	9.723	6.139	6.139	6.139	6.139	6.139	4.127	2.011	2.011	4.127
Mass Flow	lb/hr	1.800e+004	1.414e+004	1.414e+004	1.414e+004	1.414e+004	1.414e+004	8573	5566	5566	8573
Liquid Volume Flow	barrel/day	3995	2925	2925	2925	2925	2925	1831	1094	1094	1831

TABLE 2

Mole Fractions of Components in Streams
	Feed Gas (12)	Product (18)	Sales Gas (43)	Mixed Refrigerant (35)
Methane	0.9212	0.0000	0.9453	0.6671
Ethane	0.0396	0.0082	0.0402	0.3121
Propane	0.0105	0.4116	0.0001	0.0046
Butane	0.0036	0.1430	0.0000	0.0000
Pentane	0.0090	0.3576	0.0000	0.0000
Heptane	0.0020	0.0795	0.0000	0.0000
CO₂	0.0050	0.0000	0.0051	0.0145
Nitrogen	0.0091	0.0000	0.0094	0.0017

TABLE 3

Material Streams

		12	13	19	15	17	14	18	32	34	42
Vapour Fraction		0.9617	0.7601	0.7649	0.0000	0.5000	1.0000	0.0000	0.7669	0.0000	1.0000
Temperature	F	120.0	-5.00	-85.00	-15.00	-14.37	-50.25	162.6	-84.09	-84.07	-84.07
Pressure	psia	200.0	195.0	185.0	30.12	29.12	190.0	195.0	185.0	185.0	185.0
Molar Flow	MMSCFD	10.00	10.00	9.821	8.498	8.498	9.821	2.377	9.937	2.314	7.617
Mass Flow	lb/hr	2.673e+004	2.673e+004	1.852e+004	4.102e+004	4.102e+004	1.852e+004	1.559e+004	1.883e+004	7696	1.112e+004
Liquid Volume Flow	barrel/day	4723	4723	4252	5564	5564	4252	1844	4314	1436	2876

		43	35	35a	36	38	39	28	26	26a	28a
Vapour Fraction		1.0000	0.0833	1.0000	1.0000	1.0000	0.0500	1.0000	0.0000	0.0032	1.0000
Temperature	F	110.0	-103.0	90.00	260.4	120.0	-29.77	-29.77	-29.77	-30.32	-33.30
Pressure	psia	180.0	50.8	45.8	215.0	210.0	205.0	205.0	205.0	200.0	185.0
Molar Flow	MMSCFD	7.617	2.314	2.314	2.314	2.314	2.314	0.1157	2.198	2.198	0.1157
Mass Flow	lb/hr	1.112e+004	7696	7696	7696	7696	7696	308.1	7388	7388	308.1
Liquid Volume Flow	barrel/day	2876	1436	1436	1436	1436	1436	62.34	1373	1373	62.34

TABLE 4

Mole Fractions of Components in Streams
	Feed Gas (12)	Product (18)	Sales Gas (43)	Mixed Refrigerant (35)
Hydrogen	0.3401	0.0000	0.4465	0.0038
Methane	0.2334	0.0000	0.3062	0.0658
Ethane	0.1887	0.0100	0.2439	0.8415
Propane	0.0924	0.3783	0.0034	0.0889
Butane	0.0769	0.3234	0.0000	0.0000
Pentane	0.0419	0.1760	0.0000	0.0000
Heptane	0.0267	0.1124	0.0000	0.0000
CO₂	0.0000	0.0000	0.0000	0.0000
Nitrogen	0.0000	0.0000	0.0000	0.0000

TABLE 5

Material Streams

		12	13	19	15	17	14	18	32	34	42
Vapour Fraction		0.9805	0.8125	0.8225	0.0000	0.5000	1.0000	0.0000	0.8234	0.0000	1.0000
Temperature	F	120.0	0.00	-43.00	-20.00	-19.46	-13.13	195.3	-42.52	-42.49	-42.49
Pressure	psia	135.0	130.0	120.0	27.15	26.15	125.0	130.0	120.0	120.0	120.0
Molar Flow	MMSCFD	10.00	10.00	10.31	8.058	8.058	10.31	1.462	10.38	1.840	8.557
Mass Flow	lb/hr	2.673e+004	2.673e+004	2.339e+004	3.890e+004	3.890e+004	2.339e+004	1.119e+004	2.360e+004	8068	1.561e+004
Liquid Volume Flow	barrel/day	4723	4723	4624	5276	5276	4624	1245	4661	1183	3490

		43	35	35a	36	38	39	28	26	26a	28a
Vapour Fraction		1.0000	0.0805	1.0000	1.0000	1.0000	0.0349	1.0000	0.0000	0.0038	1.0000
Temperature	F	110.0	-62.0	90.00	238.2	120.0	15.00	15.00	15.00	14.31	11.44
Pressure	psia	115.0	31.75	26.75	150.0	145.0	140.0	140.0	140.0	135.0	120.0
Molar Flow	MMSCFD	8.557	1.840	1.840	1.840	1.840	1.840	6.425e-002	1.776	1.776	6.425e-002
Mass Flow	lb/hr	1.561e+004	8068	8068	8068	8068	8068	211.4	7856	7856	211.4
Liquid Volume Flow	barrel/day	3490	1183	1183	1183	1183	1183	36.58	1147	1147	36.58

TABLE 6

Mole Fractions of Components in Streams
	Feed Gas (12)	Product (18)	Sales Gas (43)	Mixed Refrigerant (35)
Hydrogen	0.3401	0.0000	0.3975	0.0022
Methane	0.2334	0.0000	0.2728	0.0257
Ethane	0.1887	0.0000	0.2220	0.2461
Propane	0.0924	0.0100	0.1074	0.7188
Butane	0.0769	0.5212	0.0003	0.0071
Pentane	0.0419	0.2861	0.0000	0.0000
Heptane	0.0267	0.1828	0.0000	0.0000
CO₂	0.0000	0.0000	0.0000	0.0000
Nitrogen	0.0000	0.0000	0.0000	0.0000

TABLE 7

Material Streams

		12	13	19	15	17	14	18	32	34	42	43
Vapour Fraction		0.9617	0.7202	0.6831	0.0000	0.5000	1.0000	0.0000	0.6833	0.0000	1.0000	1.000
Temperature	F	120.0	-25.00	-145.0	-30.00	-29.68	-22.80	176.0	-144.9	-144.9	-144.9	110.0
Pressure	psia	200.0	195.0	185.0	21.88,	20.88	190.0	195.0	185.0	185.0	185.0	180.0
Molar Flow	MMSCFD	10.00	10.00	8.153	7.268-	7.628	8.153	1.970	8.160	2.589	5.576,	5.576
Mass Flow	lb/hr	2.673e+004	2.673e+004	1.367e+004	3.508e+004	3.508e+004	1.367e+004	1.348e+004	1.369e+004	8758	4943	4943
Liquid Volume Flow	barrel/day	4723	4723	3231	4758	475	3231	1567	3234	1570	1667	1667

		35	35a	36	38	39	28	26	26a	28a	45	47
Vapour Fraction		0.0957	1.0000	1.0000	1.0000	0.0500	1.0000	0.0000	0.0028	1.0000	1.000	1.0000
Temperature	F	-163.1	90.00	330.0	120.0	-61.75	-61.75	-61.75	-62.15	-64.65	120.0	120.0
Pressure	psia	28.00	23.00	215.0	210.0	205.0	205.0	205.0	200.0	185.0	210.0	210.0
Molar Flow	MMSCFD	2.589	2.589	2.589	2,589	0.1294	6.472e-003	0.12301	0.1230	6.472e-003	0.1294	2.459
Mass Flow	lb/hr	8758	8758	8758	8758	437.9	14.05		423.8	14.05	437.9	8320
Liquid Volume Flow	barrel/day	1570	1570	1570	1570	78.48	3.009	75.47	75.47	3.009	78.48	1491

TABLE 8

Mole Fractions of Components in Streams
	Feed Gas (12)	Product (18)	Sales Gas (43)	Mixed Refrigerant (35)
Hydrogen	0.3401	0.0000	0.6085	0.0034
Methane	0.2334	0.0000	0.3517	0.1520
Ethane	0.1887	0.0100	0.0392	0.6719
Propane	0.0924	0.2974	0.0006	0.1363
Butane	0.0769	0.3482	0.0000	0.0335
Pentane	0.0419	0.2087	0.0000	0.0028
Heptane	0.0267	0.1828	0.0000	0.0000
CO₂	0.0000	0.1357	0.0000	0.0000
Nitrogen	0.0000	0.0000	0.0000	0.0000

TABLE 9

Material Streams

		12	13	19	15	17	14	18	32	34	42
Vapour Fraction		1.0000	0.9838	0.6646	0.0000	0.5000	1.0000	0.0000	0.9925	0.0000	1.0000
Temperature	F	120.0	-25.00	-119.0	-30.00	-29.68.	-79.00	251.1	-77.01	-109.5	-118.9
Pressure	psia	415.0	410.0	400.0	21.88	20.88	405.0	410.0	405.0	405.0	400.0
Molar Flow	MMSCFD	10.00	10.00	11.83	1.263	1.263	11.83	0.2534	1.577	3.668	9.730
Mass Flow	lb/hr	1.973e+004	1.973e+004	2.369e+004	6096	6096	2.369e+004	1679	3206	8867	1.801e+004
Liquid Volume Flow	barrel/day	4203	4203	5115	826.9	826.9	5115	197.4	688.7	1804	3997

		35	35a	36	38	39	28	26	26a	43
Vapour Fraction		0.3049	1.0000	1.0000	1.0000	0.4300	1.0000	0.0000	0.0464	1.000
Temperature	F	-162.0	90.00	280.9	120.0	-71.34	-71.34	-71.34	-76.54	110.0
Pressure	psia	128.30	123.30	470.0	465.0	460.0	460.0	460.0	415.0	395.0
Molar Flow	MMSCFD	3.668	3.668	3.668	3.668	3.688	1.577	2.091	2.091	9.730
Mass Flow	lb/hr	8867	8867	8867	8867	8867	3206	5661	5661	1.801e+004
Liquid Volume Flow	barrel/day	1804	1804	1804	1804	1804	688.7	1115	1115	3997

TABLE 10

Mole Fractions of Components in Streams
	Feed Gas (12)	Product (18)	Sales Gas (43)	Mixed Refrigerant (35)
Methane	0.9212	0.0000	0.9457	0.5987
Ethane	0.0396	0.0083	0.0397	0.3763
Propane	0.0105	0.4154	0.0001	0.0054
Butane	00036.	0.1421	0.0000	0.0000
Pentane	0.0090	0.3552	0.0000	0.0000
Heptane	0.0020	0.0789	0.0000	0.0000
CO₂	0.0050	0.0000	0.0051	0.0195
Nitrogen	0.0091	0.0000	0.0094	0.0001

TABLE 11

Material Streams

		12	13	19	15	17	14	18	32	34	42
Vapour Fraction		1.0000	0.8833	0.7394	0.0000	0.5000	1.0000	0.0000	0.7491	0.0000	1.0000
Temperature	F	120.0	-20.00	-85.5	-30.00	-29.68	-55.13	181.7	-84.23	-84.24	-84.24
Pressure	psia	315.0	310.0	305.0	21.88	20.88	310.0	315.0	305.0	305.0	305.0
Molar Flow	MMSCFD	10.00	10.00	11.37	5.018	5.018	11.37	1.139	11.81	2.952	8.844
Mass Flow	lb/hr	2.484e+004	2.484e+004	2.549e+004	2.422e+004	2.422e+004	2.549e+004	6778	2.648e+004	8419	1.802e+004
Liquid Volume Flow	barrel/day	4721	4721	5338	3285	3285	5338	834.5	5546	1660	3877

		43	35	35a	36	38	39	28	26	26a	28a
Vapour Fraction		1.0000	0.2044	1.0000	1.0000	1.0000	0.1500	1.0000	0.0000	0.0434	.09975
Temperature	F	110.0	-120.0	90.00	246.2	120.0	-49.05	-49.05	-49.05	-54.73	-57.22
Pressure	psia	300.0	113.9	108.9	375.0	370.0	365.0	365.0	365.0	320.0	305.0
Molar Flow	MMSCFD	8.844	2.952	2952	2952	2952	2952	0.4429	2.510	2.510	0.4429
Mass Flow	lb/hr	1.802e+004	8419	8419	8419	8419	8419	990.7	7429	7429	990.7
Liquid Volume Flow	barrel/day	3877	1660	1660	1660	1660	1660	207.9	1452	1452	207.9

TABLE 12

Mole Fractions of Components in Streams
	Feed Gas (12))	Product (18)	Sales Gas (43)	Mixed Refrigerant (35)
Methane	0.7304	0.0000	0.8252	0.3071
Ethane	0.1429	0.0119	0.1566	0.6770
Propane	0.0681	0.5974	0.0003	0.0071
Butane	0.0257	0.2256	0.0000	0.0000
Pentane	0.0088	0.0772	0.0000	0.0000
Heptane	0.0100	0.0878	0.0000	0.0000
CO₂	0.0050	0.0000	0.0056	0.0079
Nitrogen	0.0091	0.0000	0.0103	0.0009

Claims

A process for recovery of natural gas liquids from a feed gas stream (112, 212, 312, 412, 512), comprising:
(a) forming a first portion (112a, 212a, 312a, 412a, 512a) of the feed gas stream and a second portion (112b, 212b, 312b, 412b, 512b) of the feed gas stream, wherein the mass ratio of the first portion to the second portion is in the range of 90:10 to 85:15;

(b) cooling the first portion (112a, 212a, 312a, 412a, 512a) in a heat exchanger (110, 210, 310, 410, 510) and at least partially condensing the first portion;

(c) feeding the second portion (112b, 212b, 312b, 412b, 512b) that is not pre-cooled and the cooled and at least partially condensed first portion to a distillation column (120, 220, 320, 420, 520) wherein lighter components are removed from the distillation column (120, 220, 320, 420, 520) as an overhead vapor stream (114, 214, 314, 414, 514) and heavier components are removed from the distillation column (120, 220, 320, 420, 520) in the bottoms as a product stream (118, 218, 318, 418, 518), and wherein the second portion is fed into the distillation column (120, 220, 320, 420, 520) at a point one or more vapor-liquid equilibrium stages below the first portion, thereby allowing mass transfer exchange between liquids of the cooled first portion and vapors of the second portion within the column, wherein a valve (195, 295, 395, 495) is provided for the second portion for process control purposes;

(d) feeding the distillation column overhead stream (114, 214, 314, 414, 514) to the heat exchanger (110, 210, 310, 410, 510) and cooling the distillation column overhead stream (114, 214, 314, 414, 514) to at least partially liquefy the distillation column overhead stream;

(e) feeding the at least partially liquefied distillation column overhead stream to a first separator (160, 260, 321, 460, 585);

(f) separating the vapor and liquid in the first separator (160, 260, 321, 460, 585) to produce an overhead vapor stream (142, 242, 342, 442, 542) comprising sales gas and a bottoms stream (134, 234, 334, 434, 534) comprising a mixed refrigerant;

(g) feeding the mixed refrigerant stream to the heat exchanger (110, 210, 310, 410, 510) to provide cooling, wherein the mixed refrigerant stream vaporizes as it passes through the heat exchanger;

(h) compressing the vaporized mixed refrigerant stream and passing the compressed mixed refrigerant stream through the heat exchanger (110, 210, 310, 410, 510); and

(i) feeding at least a portion of the compressed mixed refrigerant stream to the distillation column (120, 220, 320, 420, 520) as a reflux stream (126, 226, 326, 526),

the process further comprising reboiling a portion of the distillation column bottoms in a distillation column reboiler (130, 230, 330, 430, 530), and

the process using the feed gas stream for the recovery of C₃+ components in the product stream, and with the C₂ components removed in the sales gas stream.
The process of claim 1, further comprising, before (i), feeding the compressed mixed refrigerant stream to a second separator (140, 240, 340, 540), and feeding the bottoms from the second separator (140, 240, 340, 540) to the distillation column (120, 220, 320, 420, 520) as the reflux stream (126, 226, 326, 526).
The process of claim 1, further comprising reducing the temperature of the mixed refrigerant stream before the mixed refrigerant stream enters the heat exchanger (110, 210, 310, 410, 510) in step g by reducing the pressure of the mixed refrigerant using a control valve (165, 265, 365, 465, 565).
The process of claim 2, further comprising combining the overhead stream (128, 228, 328, 528) from the second separator (140, 240, 340, 540) with the overhead stream from the distillation column (120, 220, 320, 420, 520) and feeding the combined stream (132, 232, 332) to the first separator (160, 260, 321, 460, 585).
The process of claim 1, further comprising cooling the compressed mixed refrigerant in a cooler (190, 290, 390) before passing the compressed mixed refrigerant stream through the heat exchanger (110, 210, 310, 410, 510).
The process of claim 1, wherein the first separator is an absorber (321).
The process of claim 1, wherein the feed gas stream (112, 212, 312, 412, 512) is one of natural gas or refinery gas.
The process of claim 1, wherein the distillation column (120, 220, 320, 420, 520) operates at a pressure of between about 0.689 MPa (100 psia) and 3.10 MPa (450 psia).
The process of claim 1, wherein the first and second portions of the feed gas stream have the same composition.
The process of claim 1, wherein a portion of compressed mixed refrigerant stream is removed as a supplemental product stream.
The process of claim 1, wherein separating the vapors and liquids in the first separator (585) further includes producing a side draw fraction (551).
The process of claim 11, wherein the overhead vapor stream is enriched in nitrogen and depleted in propane, the bottoms fraction is depleted in nitrogen and enriched in propane, and the side draw fraction has intermediate propane and nitrogen content.
The process of claim 1, wherein the energy input to the distillation column reboiler (130, 230, 330, 430, 530) is at least 5% lower than the energy input for thesame process with the same volumes and compositions of the feed gas stream, product stream and sales gas stream, and in which no second portion is formed from the feed gas stream.
The process of claim 1, wherein the energy input to the distillation column reboiler (130, 230, 330, 430, 530) is at least 10% lower than the energy input for the same process with the same volumes and compositions of the feed gas stream, product stream and sales gas stream, and in which no second portion is formed from the feed gas stream.
The process of claim 1, wherein a compressor (180, 280, 380, 480, 580) is used in step (h), wherein the total compressor duty in step (h) of the process is at least 5% lower than the compressor duty for the same process with the same volumes and compositions of the feed gas stream, product stream and sales gas stream, but in which no second portion is formed from the feed gas stream.
The process of claim 15, wherein the total compressor duty in step (h) of the process is at least 10% lower than the compressor duty for the same process with the same volumes and compositions of the feed gas stream, product stream and sales gas stream, but in which no second portion is formed from the feed gas stream.