US20110067441A1 - Hydrocarbon Gas Processing - Google Patents
Hydrocarbon Gas Processing Download PDFInfo
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- US20110067441A1 US20110067441A1 US12/868,993 US86899310A US2011067441A1 US 20110067441 A1 US20110067441 A1 US 20110067441A1 US 86899310 A US86899310 A US 86899310A US 2011067441 A1 US2011067441 A1 US 2011067441A1
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- F—MECHANICAL ENGINEERING; LIGHTING; HEATING; WEAPONS; BLASTING
- F25—REFRIGERATION OR COOLING; COMBINED HEATING AND REFRIGERATION SYSTEMS; HEAT PUMP SYSTEMS; MANUFACTURE OR STORAGE OF ICE; LIQUEFACTION SOLIDIFICATION OF GASES
- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
- F25J3/00—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification
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- F—MECHANICAL ENGINEERING; LIGHTING; HEATING; WEAPONS; BLASTING
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- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
- F25J3/00—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification
- F25J3/02—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream
- F25J3/0204—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream characterised by the feed stream
- F25J3/0209—Natural gas or substitute natural gas
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- F25J3/02—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream
- F25J3/0228—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream characterised by the separated product stream
- F25J3/0233—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream characterised by the separated product stream separation of CnHm with 1 carbon atom or more
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- F25J3/02—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream
- F25J3/0228—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream characterised by the separated product stream
- F25J3/0238—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream characterised by the separated product stream separation of CnHm with 2 carbon atoms or more
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- F25J2205/02—Processes or apparatus using other separation and/or other processing means using simple phase separation in a vessel or drum
- F25J2205/04—Processes or apparatus using other separation and/or other processing means using simple phase separation in a vessel or drum in the feed line, i.e. upstream of the fractionation step
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- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
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- F25J2215/00—Processes characterised by the type or other details of the product stream
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- F25J2235/60—Processes or apparatus involving steps for increasing the pressure or for conveying of liquid process streams the fluid being (a mixture of) hydrocarbons
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- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
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- F25J2240/40—Expansion without extracting work, i.e. isenthalpic throttling, e.g. JT valve, regulating valve or venturi, or isentropic nozzle, e.g. Laval
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- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
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- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
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- F25J2290/40—Vertical layout or arrangement of cold equipments within in the cold box, e.g. columns, condensers, heat exchangers etc.
Definitions
- This invention relates to a process and an apparatus for the separation of a gas containing hydrocarbons.
- Ethylene, ethane, propylene, propane, and/or heavier hydrocarbons can be recovered from a variety of gases, such as natural gas, refinery gas, and synthetic gas streams obtained from other hydrocarbon materials such as coal, crude oil, naphtha, oil shale, tar sands, and lignite.
- Natural gas usually has a major proportion of methane and ethane, i.e., methane and ethane together comprise at least 50 mole percent of the gas.
- the gas also contains relatively lesser amounts of heavier hydrocarbons such as propane, butanes, pentanes, and the like, as well as hydrogen, nitrogen, carbon dioxide, and other gases.
- 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, 80.8% methane, 9.4% ethane and other C 2 components, 4.7% propane and other C 3 components, 1.2% iso-butane, 2.1% normal butane, and 1.1% pentanes plus, with the balance made up of nitrogen and carbon dioxide. Sulfur containing gases are also sometimes present.
- 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.
- liquids may be condensed and collected in one or more separators as high-pressure liquids containing some of the desired C 2 + components.
- the high-pressure liquids may be expanded to a lower pressure and fractionated. The vaporization occurring during expansion of the liquids results in further cooling of the stream. Under some conditions, pre-cooling the high pressure liquids prior to the expansion may be desirable in order to further lower the temperature resulting from the expansion.
- the expanded stream comprising a mixture of liquid and vapor, is fractionated in a distillation (demethanizer or deethanizer) column.
- the expansion cooled stream(s) is (are) distilled to separate residual methane, nitrogen, and other volatile gases as overhead vapor from the desired C 2 components, C 3 components, and heavier hydrocarbon components as bottom liquid product, or to separate residual methane, C 2 components, nitrogen, and other volatile gases as overhead vapor from the desired C 3 components and heavier hydrocarbon components as bottom liquid product.
- the vapor remaining from the partial condensation can be split into two streams.
- One portion of the vapor is passed through a work expansion machine or engine, or an expansion valve, to a lower pressure at which additional liquids are condensed as a result of further cooling of the stream.
- the pressure after expansion is essentially the same as the pressure at which the distillation column is operated.
- the combined vapor-liquid phases resulting from the expansion are supplied as feed to the column.
- the remaining portion of the vapor is cooled to substantial condensation by heat exchange with other process streams, e.g., the cold fractionation tower overhead.
- Some or all of the high-pressure liquid may be combined with this vapor portion prior to cooling.
- the resulting cooled stream is then expanded through an appropriate expansion device, such as an expansion valve, to the pressure at which the demethanizer is operated. During expansion, a portion of the liquid will vaporize, resulting in cooling of the total stream.
- the flash expanded stream is then supplied as top feed to the demethanizer.
- the vapor portion of the flash expanded stream and the demethanizer overhead vapor combine in an upper separator section in the fractionation tower as residual methane product gas.
- the cooled and expanded stream may be supplied to a separator to provide vapor and liquid streams.
- the vapor is combined with the tower overhead and the liquid is supplied to the column as a top column feed.
- the residue gas leaving the process will contain substantially all of the methane in the feed gas with essentially none of the heavier hydrocarbon components, and the bottoms fraction leaving the demethanizer will contain substantially all of the heavier hydrocarbon components with essentially no methane or more volatile components.
- this ideal situation is not obtained because the conventional demethanizer is operated largely as a stripping column.
- the methane product of the process therefore, typically comprises vapors leaving the top fractionation stage of the column, together with vapors not subjected to any rectification step.
- the preferred processes for hydrocarbon separation use an upper absorber section to provide additional rectification of the rising vapors.
- the source of the reflux stream for the upper rectification section is typically a recycled stream of residue gas supplied under pressure.
- the recycled residue gas stream is usually cooled to substantial condensation by heat exchange with other process streams, e.g., the cold fractionation tower overhead.
- the resulting substantially condensed stream is then expanded through an appropriate expansion device, such as an expansion valve, to the pressure at which the demethanizer is operated. During expansion, a portion of the liquid will usually vaporize, resulting in cooling of the total stream.
- the flash expanded stream is then supplied as top feed to the demethanizer.
- 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.
- the cooled and expanded stream may be supplied to a separator to provide vapor and liquid streams, so that thereafter the vapor is combined with the tower overhead and the liquid is supplied to the column as a top column feed.
- Typical process schemes of this type are disclosed in U.S. Pat. Nos. 4,889,545; 5,568,737; and 5,881,569, assignee's co-pending application Ser. No. 12/717,394, and in Mowrey, E.
- the present invention also employs an upper rectification section (or a separate rectification column if plant size or other factors favor using separate rectification and stripping columns).
- the reflux stream for this rectification section is provided by using a side draw of the vapors rising in a lower portion of the tower. Because of the relatively high concentration of C 2 components in the vapors lower in the tower, a significant quantity of liquid can be condensed in this side draw stream without elevating its pressure, often using only the refrigeration available in the cold vapor leaving the upper rectification section and the flash expanded substantially condensed stream.
- This condensed liquid which is predominantly liquid methane, can then be used to absorb C 2 components, C 3 components, C 4 components, and heavier hydrocarbon components from the vapors rising through the upper rectification section and thereby capture these valuable components in the bottom liquid product from the demethanizer.
- the present invention provides the further advantage of being able to maintain in excess of 99% recovery of the C 3 and C 4 + components as the recovery of C 2 components is adjusted from high to low values.
- the present invention makes possible essentially 100% separation of methane and lighter components from the C 2 components and heavier components at the same energy requirements compared to the prior art while increasing the recovery levels.
- the present invention although applicable at lower pressures and warmer temperatures, is particularly advantageous when processing feed gases in the range of 400 to 1500 psia [2,758 to 10,342 kPa(a)] or higher under conditions requiring NGL recovery column overhead temperatures of ⁇ 50° F. [ ⁇ 46° C.] or colder.
- FIG. 1 is a flow diagram of a prior art natural gas processing plant in accordance with U.S. Pat. No. 5,890,378;
- FIG. 2 is a flow diagram of a prior art natural gas processing plant in accordance with U.S. Pat. No. 7,191,617;
- FIG. 3 is a flow diagram of a prior art natural gas processing plant in accordance with assignee's co-pending application Ser. No. 12/206,230;
- FIG. 4 is a flow diagram of a natural gas processing plant in accordance with the present invention.
- FIGS. 5 through 8 are flow diagrams illustrating alternative means of application of the present invention to a natural gas stream.
- FIG. 1 is a process flow diagram showing the design of a processing plant to recover C 2 + components from natural gas using prior art according to U.S. Pat. No. 5,890,378.
- inlet gas enters the plant at 85° F. [29° C.] and 970 psia [6,688 kPa(a)] as stream 31 .
- the sulfur compounds are removed by appropriate pretreatment of the feed gas (not illustrated).
- the feed stream is usually dehydrated to prevent hydrate (ice) formation under cryogenic conditions. Solid desiccant has typically been used for this purpose.
- the feed stream 31 is cooled in heat exchanger 10 by heat exchange with cool residue gas (stream 45 b ), demethanizer lower side reboiler liquids at 32° F. [0° C.] (stream 40 ), and propane refrigerant.
- exchanger 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.)
- the cooled stream 31 a enters separator 11 at 0° F.
- the vapor (stream 32 ) from separator 11 is further cooled in heat exchanger 13 by heat exchange with cool residue gas (stream 45 a ) and demethanizer upper side reboiler liquids at ⁇ 39° F. [ ⁇ 39° C.] (stream 39 ).
- the cooled stream 32 a enters separator 14 at ⁇ 31° F. [ ⁇ 35° C.] and 950 psia [6,550 kPa(a)] where the vapor (stream 34 ) is separated from the condensed liquid (stream 37 ).
- the separator liquid (stream 37 ) is expanded to the tower operating pressure by expansion valve 19 , cooling stream 37 a to ⁇ 66° F. [ ⁇ 54° C.] before it is supplied to fractionation tower 20 at a second lower mid-column feed point.
- the vapor (stream 34 ) from separator 14 is divided into two streams, 35 and 36 .
- Stream 35 containing about 39% of the total vapor, passes through heat exchanger 15 in heat exchange relation with the cold residue gas (stream 45 ) where it is cooled to substantial condensation.
- the resulting substantially condensed stream 35 a at ⁇ 123° F. [ ⁇ 86° C.] is then flash expanded through expansion valve 16 to slightly above the operating pressure of fractionation tower 20 . During expansion a portion of the stream is vaporized, resulting in cooling of the total stream.
- the expanded stream 35 b leaving expansion valve 16 reaches a temperature of ⁇ 130° F. [ ⁇ 90° C.].
- the expanded stream 35 b is warmed to ⁇ 126° F.
- the remaining 61% of the vapor from separator 14 enters a work expansion machine 17 in which mechanical energy is extracted from this portion of the high pressure feed.
- the machine 17 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 36 a to a temperature of approximately ⁇ 86° F. [ ⁇ 66° C.].
- the typical commercially available expanders are capable of recovering on the order of 80-85% of the work theoretically available in an ideal isentropic expansion.
- the work recovered is often used to drive a centrifugal compressor (such as item 18 ) that can be used to re-compress the residue gas (stream 45 c ), for example.
- the partially condensed expanded stream 36 a is thereafter supplied as feed to fractionation tower 20 at a mid-column feed point.
- the demethanizer in tower 20 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing.
- the demethanizer tower consists of two sections: an upper absorbing (rectification) section 20 a that contains the trays and/or packing to provide the necessary contact between the vapor portions of the expanded streams 35 c and 36 a rising upward and cold liquid falling downward to condense and absorb the C 2 components, C 3 components, and heavier components; and a lower, stripping section 20 b that contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward.
- an upper absorbing (rectification) section 20 a that contains the trays and/or packing to provide the necessary contact between the vapor portions of the expanded streams 35 c and 36 a rising upward and cold liquid falling downward to condense and absorb the C 2 components, C 3 components, and heavier components
- a lower, stripping section 20 b that contains the trays and/or
- the demethanizing section 20 b also includes one or more reboilers (such as reboiler 21 and the side reboilers described previously) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column to strip the liquid product, stream 41 , of methane and lighter components.
- Stream 36 a enters demethanizer 20 at an intermediate feed position located in the lower region of absorbing section 20 a of demethanizer 20 .
- the liquid portion of the expanded stream 36 a comingles with liquids falling downward from absorbing section 20 a and the combined liquid continues downward into stripping section 20 b of demethanizer 20 .
- the vapor portion of the expanded stream 36 a rises upward through absorbing section 20 a and is contacted with cold liquid falling downward to condense and absorb the C 2 components, C 3 components, and heavier components.
- a portion of the distillation vapor (stream 42 ) is withdrawn from the upper region of stripping section 20 b .
- This stream is then cooled and partially condensed (stream 42 a ) in exchanger 22 by heat exchange with expanded substantially condensed stream 35 b as described previously, cooling stream 42 from ⁇ 96° F. [ ⁇ 71° C.] to about ⁇ 128° F. [ ⁇ 89° C.] (stream 42 a ).
- the operating pressure (441 psia [3,038 kPa(a)]) in reflux separator 23 is maintained slightly below the operating pressure of demethanizer 20 . This provides the driving force which causes distillation vapor stream 42 to flow through heat exchanger 22 and thence into the reflux separator 23 where the condensed liquid (stream 44 ) is separated from any uncondensed vapor (stream 43 ).
- the liquid stream 44 from reflux separator 23 is pumped by pump 24 to a pressure slightly above the operating pressure of demethanizer 20 , and stream 44 a is then supplied as cold top column feed (reflux) to demethanizer 20 at ⁇ 128° F. [ ⁇ 89° C.].
- This cold liquid reflux absorbs and condenses the C 3 components and heavier components rising in the upper rectification region of absorbing section 20 a of demethanizer 20 .
- the liquid product stream 41 exits the bottom of the tower at 112° F. [44° C.], based on a typical specification of a methane to ethane ratio of 0.025:1 on a molar basis in the bottom product.
- Cold demethanizer overhead stream 38 exits the top of demethanizer 20 at ⁇ 128° F. [ ⁇ 89° C.] and combines with vapor stream 43 to form cold residue gas stream 45 at ⁇ 128° F. [ ⁇ 89° C.].
- the cold residue gas stream 45 passes countercurrently to the incoming feed gas in heat exchanger 15 where it is heated to ⁇ 37° F. [ ⁇ 38° C.] (stream 45 a ), in heat exchanger 13 where it is heated to ⁇ 5° F.
- stream 45 b [ ⁇ 21° C.] (stream 45 b ), and in heat exchanger 10 where it is heated to 80° F. [27° C.] (stream 45 c ).
- the residue gas is then re-compressed in two stages.
- the first stage is compressor 18 driven by expansion machine 17 .
- the second stage is compressor 25 driven by a supplemental power source which compresses the residue gas (stream 45 d ) to sales line pressure.
- the residue gas product (stream 451 ) flows to the sales gas pipeline at 1015 psia [6,998 kPa(a)], sufficient to meet line requirements (usually on the order of the inlet pressure).
- FIG. 2 represents an alternative prior art process according to U.S. Pat. No. 7,191,617.
- the process of FIG. 2 has been applied to the same feed gas composition and conditions as described above for FIG. 1 .
- operating conditions were selected to minimize energy consumption for a given recovery level.
- inlet gas enters the plant as stream 31 and is cooled in heat exchanger 10 by heat exchange with cool residue gas (stream 45 b ), demethanizer lower side reboiler liquids at 33° F. [0° C.] (stream 40 ), and propane refrigerant.
- the cooled stream 31 a enters separator 11 at 0° F. [ ⁇ 18° C.] and 955 psia [6,584 kPa(a)] where the vapor (stream 32 ) is separated from the condensed liquid (stream 33 ).
- the separator liquid (stream 33 ) is expanded to the operating pressure (approximately 450 psia [3,103 kPa(a)]) of fractionation tower 20 by expansion valve 12 , cooling stream 33 a to ⁇ 27° F. [ ⁇ 33° C.] before it is supplied to fractionation tower 20 at a first lower mid-column feed point.
- the vapor (stream 32 ) from separator 11 is further cooled in heat exchanger 13 by heat exchange with cool residue gas (stream 45 a ) and demethanizer upper side reboiler liquids at ⁇ 38° F. [ ⁇ 39° C.] (stream 39 ).
- the cooled stream 32 a enters separator 14 at ⁇ 29° F. [ ⁇ 34° C.] and 950 psia [6,550 kPa(a)] where the vapor (stream 34 ) is separated from the condensed liquid (stream 37 ).
- the separator liquid (stream 37 ) is expanded to the tower operating pressure by expansion valve 19 , cooling stream 37 a to ⁇ 64° F. [ ⁇ 53° C.] before it is supplied to fractionation tower 20 at a second lower mid-column feed point.
- the vapor (stream 34 ) from separator 14 is divided into two streams, 35 and 36 .
- Stream 35 containing about 37% of the total vapor, passes through heat exchanger 15 in heat exchange relation with the cold residue gas (stream 45 ) where it is cooled to substantial condensation.
- the resulting substantially condensed stream 35 a at ⁇ 115° F. [ ⁇ 82° C.] is then flash expanded through expansion valve 16 to the operating pressure of fractionation tower 20 .
- expansion valve 16 During expansion a portion of the stream is vaporized, resulting in cooling of stream 35 b to ⁇ 129° F. [ ⁇ 89° C.] before it is supplied to fractionation tower 20 at an upper mid-column feed point.
- the remaining 63% of the vapor from separator 14 enters a work expansion machine 17 in which mechanical energy is extracted from this portion of the high pressure feed.
- the machine 17 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 36 a to a temperature of approximately ⁇ 84° F. [ ⁇ 65° C.].
- the partially condensed expanded stream 36 a is thereafter supplied as feed to fractionation tower 20 at a mid-column feed point.
- a portion of the distillation vapor (stream 42 ) is withdrawn from the upper region of the stripping section in fractionation tower 20 .
- This stream is then cooled from ⁇ 91° F. [ ⁇ 68° C.] to ⁇ 122° F. [ ⁇ 86° C.] and partially condensed (stream 42 a ) in heat exchanger 22 by heat exchange with the cold demethanizer overhead stream 38 exiting the top of demethanizer 20 at ⁇ 127° F. [ ⁇ 88° C.].
- the cold demethanizer overhead stream is warmed slightly to ⁇ 120° F. [ ⁇ 84° C.] (stream 38 a ) as it cools and condenses at least a portion of stream 42 .
- the operating pressure (447 psia [3,079 kPa(a)]) in reflux separator 23 is maintained slightly below the operating pressure of demethanizer 20 .
- This provides the driving force which causes distillation vapor stream 42 to flow through heat exchanger 22 and thence into the reflux separator 23 where the condensed liquid (stream 44 ) is separated from any uncondensed vapor (stream 43 ).
- Stream 43 then combines with the warmed demethanizer overhead stream 38 a from heat exchanger 22 to form cold residue gas stream 45 at ⁇ 120° F. [ ⁇ 84° C.].
- the liquid stream 44 from reflux separator 23 is pumped by pump 24 to a pressure slightly above the operating pressure of demethanizer 20 , and stream 44 a is then supplied as cold top column feed (reflux) to demethanizer 20 at ⁇ 121° F. [ ⁇ 85° C.].
- This cold liquid reflux absorbs and condenses the C 3 components and heavier components rising in the upper rectification region of the absorbing section of demethanizer 20 .
- the liquid product stream 41 exits the bottom of tower 20 at 114° F. [45° C.].
- the cold residue gas stream 45 passes countercurrently to the incoming feed gas in heat exchanger 15 where it is heated to ⁇ 36° F. [ ⁇ 38° C.] (stream 45 a ), in heat exchanger 13 where it is heated to ⁇ 5° F. [ ⁇ 20° C.] (stream 45 b ), and in heat exchanger 10 where it is heated to 80° F. [27° C.] (stream 45 c ) as it provides cooling as previously described.
- the residue gas is then re-compressed in two stages, compressor 18 driven by expansion machine 17 and compressor 25 driven by a supplemental power source. After stream 45 e is cooled to 120° F. [49° C.] in discharge cooler 26 , the residue gas product (stream 45 f ) flows to the sales gas pipeline at 1015 psia [6,998 kPa(a)].
- FIG. 3 represents an alternative prior art process according to co-pending application Ser. No. 12/206,230.
- the process of FIG. 3 has been applied to the same feed gas composition and conditions as described above for FIGS. 1 and 2 .
- operating conditions were selected to minimize energy consumption for a given recovery level.
- inlet gas enters the plant as stream 31 and is cooled in heat exchanger 10 by heat exchange with cool residue gas (stream 45 b ), demethanizer lower side reboiler liquids at 36° F. [2° C.] (stream 40 ), and propane refrigerant.
- the cooled stream 31 a enters separator 11 at 1° F. [ ⁇ 17° C.] and 955 psia [6,584 kPa(a)] where the vapor (stream 32 ) is separated from the condensed liquid (stream 33 ).
- the separator liquid (stream 33 ) is expanded to the operating pressure (approximately 452 psia [3,116 kPa(a)]) of fractionation tower 20 by expansion valve 12 , cooling stream 33 a to ⁇ 25° F. [ ⁇ 32° C.] before it is supplied to fractionation tower 20 at a first lower mid-column feed point.
- the vapor (stream 32 ) from separator 11 is further cooled in heat exchanger 13 by heat exchange with cool residue gas (stream 45 a ) and demethanizer upper side reboiler liquids at ⁇ 37° F. [ ⁇ 38° C.] (stream 39 ).
- the cooled stream 32 a enters separator 14 at ⁇ 31° F. [ ⁇ 35° C.] and 950 psia [6,550 kPa(a)] where the vapor (stream 34 ) is separated from the condensed liquid (stream 37 ).
- the separator liquid (stream 37 ) is expanded to the tower operating pressure by expansion valve 19 , cooling stream 37 a to ⁇ 65° F. [ ⁇ 54° C.] before it is supplied to fractionation tower 20 at a second lower mid-column feed point.
- the vapor (stream 34 ) from separator 14 is divided into two streams, 35 and 36 .
- Stream 35 containing about 38% of the total vapor, passes through heat exchanger 15 in heat exchange relation with the cold residue gas (stream 45 ) where it is cooled to substantial condensation.
- the resulting substantially condensed stream 35 a at ⁇ 119° F. [ ⁇ 84° C.] is then flash expanded through expansion valve 16 to the operating pressure of fractionation tower 20 .
- expansion valve 16 During expansion a portion of the stream is vaporized, resulting in cooling of stream 35 b to ⁇ 129° F. [ ⁇ 90° C.] before it is supplied to fractionation tower 20 at an upper mid-column feed point.
- the remaining 62% of the vapor from separator 14 enters a work expansion machine 17 in which mechanical energy is extracted from this portion of the high pressure feed.
- the machine 17 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 36 a to a temperature of approximately ⁇ 85° F. [ ⁇ 65° C.].
- the partially condensed expanded stream 36 a is thereafter supplied as feed to fractionation tower 20 at a mid-column feed point.
- a portion of the distillation vapor (stream 42 ) is withdrawn from an intermediate region of the absorbing section in fractionation column 20 , above the feed position of expanded stream 36 a in the lower region of the absorbing section.
- This distillation vapor stream 42 is then cooled from ⁇ 101° F. [ ⁇ 74° C.] to ⁇ 124° F. [ ⁇ 86° C.] and partially condensed (stream 42 a ) in heat exchanger 22 by heat exchange with the cold demethanizer overhead stream 38 exiting the top of demethanizer 20 at ⁇ 128° F. [ ⁇ 89° C.].
- the cold demethanizer overhead stream is warmed slightly to ⁇ 124° F. [ ⁇ 86° C.] (stream 38 a ) as it cools and condenses at least a portion of stream 42 .
- the operating pressure (448 psia [3,090 kPa(a)]) in reflux separator 23 is maintained slightly below the operating pressure of demethanizer 20 .
- This provides the driving force which causes distillation vapor stream 42 to flow through heat exchanger 22 and thence into the reflux separator 23 where the condensed liquid (stream 44 ) is separated from any uncondensed vapor (stream 43 ).
- Stream 43 then combines with the warmed demethanizer overhead stream 38 a from heat exchanger 22 to form cold residue gas stream 45 at ⁇ 124° F. [ ⁇ 86° C.].
- the liquid stream 44 from reflux separator 23 is pumped by pump 24 to a pressure slightly above the operating pressure of demethanizer 20 , and stream 44 a is then supplied as cold top column feed (reflux) to demethanizer 20 at ⁇ 123° F. [ ⁇ 86° C.].
- This cold liquid reflux absorbs and condenses the C 2 components, C 3 components, and heavier components rising in the upper rectification region of the absorbing section of demethanizer 20 .
- the liquid product stream 41 exits the bottom of tower 20 at 113° F. [45° C.].
- the cold residue gas stream 45 passes countercurrently to the incoming feed gas in heat exchanger 15 where it is heated to ⁇ 38° F. [ ⁇ 39° C.] (stream 45 a ), in heat exchanger 13 where it is heated to ⁇ 4° F. [ ⁇ 20° C.] (stream 45 b ), and in heat exchanger 10 where it is heated to 80° F. [27° C.] (stream 45 c ) as it provides cooling as previously described.
- the residue gas is then re-compressed in two stages, compressor 18 driven by expansion machine 17 and compressor 25 driven by a supplemental power source. After stream 45 e is cooled to 120° F. [49° C.] in discharge cooler 26 , the residue gas product (stream 45 f ) flows to the sales gas pipeline at 1015 psia [6,998 kPa(a)].
- FIG. 3 process improves the ethane recovery from 85.05% (for FIG. 1 ) and 85.08% (for FIG. 2 ) to 87.33%.
- the propane recovery for the FIG. 3 process (99.36%) is lower than that of the FIG. 1 process (99.57%) but higher than that of the FIG. 2 process (99.20%).
- the butanes+recovery is essentially the same for all three of these prior art processes.
- Comparison of Tables I, II, and III further shows that the FIG. 3 process using slightly less power than both prior art processes (more than 2% less than the FIG. 1 process and 0.4% less than the FIG. 2 process).
- FIG. 4 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. 4 are the same as those in FIGS. 1 , 2 , and 3 . Accordingly, the FIG. 4 process can be compared with that of the FIGS. 1 , 2 , and 3 processes to illustrate the advantages of the present invention.
- inlet gas enters the plant at 85° F. [29° C.] and 970 psia [6,688 kPa(a)] as stream 31 and is cooled in heat exchanger 10 by heat exchange with cool residue gas (stream 45 b ), demethanizer lower side reboiler liquids at 32° F. [0° C.] (stream 40 ), and propane refrigerant.
- the cooled stream 31 a enters separator 11 at 1° F. [ ⁇ 17° C.] and 955 psia [6,584 kPa(a)] where the vapor (stream 32 ) is separated from the condensed liquid (stream 33 ).
- the separator liquid (stream 33 ) is expanded to the operating pressure (approximately 452 psia [3,116 kPa(a)]) of fractionation tower 20 by expansion valve 12 , cooling stream 33 a to ⁇ 25° F. [ ⁇ 32° C.] before it is supplied to fractionation tower 20 at a first lower mid-column feed point (located below the feed point of stream 36 a described later in paragraph [0058]).
- the vapor (stream 32 ) from separator 11 is further cooled in heat exchanger 13 by heat exchange with cool residue gas (stream 45 a ) and demethanizer upper side reboiler liquids at ⁇ 38° F. [ ⁇ 39° C.] (stream 39 ).
- the cooled stream 32 a enters separator 14 at ⁇ 31° F. [ ⁇ 35° C.] and 950 psia [6,550 kPa(a)] where the vapor (stream 34 ) is separated from the condensed liquid (stream 37 ).
- the separator liquid (stream 37 ) is expanded to the tower operating pressure by expansion valve 19 , cooling stream 37 a to ⁇ 66° F. [ ⁇ 54° C.] before it is supplied to fractionation tower 20 at a second lower mid-column feed point (also located below the feed point of stream 36 a ).
- the vapor (stream 34 ) from separator 14 is divided into two streams, 35 and 36 .
- Stream 35 containing about 38% of the total vapor, passes through heat exchanger 15 in heat exchange relation with the cold residue gas (stream 45 ) where it is cooled to substantial condensation.
- the resulting substantially condensed stream 35 a at ⁇ 122° F. [ ⁇ 86° C.] is then flash expanded through expansion valve 16 to slightly above the operating pressure of fractionation tower 20 . During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in FIG. 4 , the expanded stream 35 b leaving expansion valve 16 reaches a temperature of ⁇ 130° F. [ ⁇ 90° C.].
- the expanded stream 35 b is warmed slightly to ⁇ 129° F. [ ⁇ 89° C.] and further vaporized in heat exchanger 22 as it provides a portion of the cooling of distillation vapor stream 42 .
- the warmed stream 35 c is then supplied at an upper mid-column feed point, in absorbing section 20 a of fractionation tower 20 .
- the remaining 62% of the vapor from separator 14 enters a work expansion machine 17 in which mechanical energy is extracted from this portion of the high pressure feed.
- the machine 17 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 36 a to a temperature of approximately ⁇ 86° F. [ ⁇ 65° C.].
- the partially condensed expanded stream 36 a is thereafter supplied as feed to fractionation tower 20 at a mid-column feed point (located below the feed point of stream 35 c ).
- the demethanizer in tower 20 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing.
- the demethanizer tower consists of two sections: an upper absorbing (rectification) section 20 a that contains the trays and/or packing to provide the necessary contact between the vapor portions of the expanded streams 35 c and 36 a rising upward and cold liquid falling downward to condense and absorb the C 2 components, C 3 components, and heavier components from the vapors rising upward; and a lower, stripping section 20 b that contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward.
- an upper absorbing (rectification) section 20 a that contains the trays and/or packing to provide the necessary contact between the vapor portions of the expanded streams 35 c and 36 a rising upward and cold liquid falling downward to condense and absorb the C 2 components, C 3 components, and heavier components from the vapors rising upward
- the demethanizing section 20 b also includes one or more reboilers (such as reboiler 21 and the side reboilers described previously) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column to strip the liquid product, stream 41 , of methane and lighter components.
- Stream 36 a enters demethanizer 20 at an intermediate feed position located in the lower region of absorbing section 20 a of demethanizer 20 .
- the liquid portion of the expanded stream 36 a comingles with liquids falling downward from absorbing section 20 a and the combined liquid continues downward into stripping section 20 b of demethanizer 20 .
- the vapor portion of the expanded stream 36 a rises upward through absorbing section 20 a and is contacted with cold liquid falling downward to condense and absorb the C 2 components, C 3 components, and heavier components.
- a portion of the distillation vapor (stream 42 ) is withdrawn from an intermediate region of absorbing section 20 a in fractionation column 20 , above the feed position of expanded stream 36 a in the lower region of absorbing section 20 a .
- This distillation vapor stream 42 is then cooled from ⁇ 103° F. [ ⁇ 75° C.] to ⁇ 128° F. [ ⁇ 89° C.] and partially condensed (stream 42 a ) in heat exchanger 22 by heat exchange with the cold demethanizer overhead stream 38 exiting the top of demethanizer 20 at ⁇ 129° F. [ ⁇ 89° C.] and with the expanded substantially condensed stream 35 b as described previously.
- the cold demethanizer overhead stream is warmed slightly to ⁇ 127° F. [ ⁇ 88° C.] (stream 38 a ) as it provides a portion of the cooling of distillation vapor stream 42 .
- the operating pressure (448 psia [3,090 kPa(a)]) in reflux separator 23 is maintained slightly below the operating pressure of demethanizer 20 .
- This provides the driving force which causes distillation vapor stream 42 to flow through heat exchanger 22 and thence into the reflux separator 23 where the condensed liquid (stream 44 ) is separated from any uncondensed vapor (stream 43 ).
- Stream 43 then combines with the warmed demethanizer overhead stream 38 a from heat exchanger 22 to form cold residue gas stream 45 at ⁇ 127° F. [ ⁇ 88° C.].
- the liquid stream 44 from reflux separator 23 is pumped by pump 24 to a pressure slightly above the operating pressure of demethanizer 20 , and stream 44 a is then supplied as cold top column feed (reflux) to demethanizer 20 at ⁇ 127° F. [ ⁇ 88° C.].
- This cold liquid reflux absorbs and condenses the C 2 components, C 3 components, and heavier components rising in the upper rectification region of absorbing section 20 a of demethanizer 20 .
- the feed streams are stripped of their methane and lighter components.
- the resulting liquid product (stream 41 ) exits the bottom of tower 20 at 113° F. [45° C.] (based on a typical specification of a methane to ethane ratio of 0.025:1 on a molar basis in the bottom product).
- the cold residue gas stream 45 passes countercurrently to the incoming feed gas in heat exchanger 15 where it is heated to ⁇ 40° F. [ ⁇ 40° C.] (stream 45 a ), in heat exchanger 13 where it is heated to ⁇ 4° F. [ ⁇ 20° C.] (stream 45 b ), and in heat exchanger 10 where it is heated to 80° F.
- stream 45 c [27° C.] (stream 45 c ) as it provides cooling as previously described.
- the residue gas is then re-compressed in two stages, compressor 18 driven by expansion machine 17 and compressor 25 driven by a supplemental power source.
- stream 45 e is cooled to 120° F. [49° C.] in discharge cooler 26
- the residue gas product flows to the sales gas pipeline at 1015 psia [6,998 kPa(a)].
- Tables I, II, III, and IV show that, compared to the prior art, the present invention matches or exceeds the propane and butanes+recoveries of all the prior art processes while significantly improving the ethane recovery.
- the ethane recovery for the present invention (87.56%) is higher than the FIG. 1 process (85.05%), the FIG. 2 process (85.08%), and the FIG. 3 process (87.33%).
- Comparison of Tables I, II, III, and IV further shows that the improvement in yields was achieved without using more power than the prior art, and in some cases using significantly less power.
- the present invention represents an improvement of 5%, 3%, and 0.3%, respectively, over the prior art of the FIG.
- FIG. 3 processes the power required for the present invention is essentially the same as that for the prior art FIG. 3 process, the present invention improves both the ethane recovery and the propane recovery by 0.2% compared to the FIG. 3 process without using more power.
- the present invention uses the expanded substantially condensed feed stream 35 c supplied to absorbing section 20 a of demethanizer 20 to provide bulk recovery of the C 2 components, C 3 components, and heavier hydrocarbon components contained in expanded feed 36 a and the vapors rising from stripping section 20 b , and the supplemental rectification provided by reflux stream 44 a to reduce the amount of C 2 components, C 3 components, and C 4 + components contained in the inlet feed gas that is lost to the residue gas.
- the present invention improves the rectification in absorbing section 20 a over that of the prior art processes by making more effective use of the refrigeration available in process streams 38 and 35 b to improve the recoveries and the recovery efficiency.
- the expanded substantially condensed stream 35 b (which is predominantly liquid methane) is a better refrigerant medium than demethanizer overhead vapor stream 38 (which is primarily methane vapor), so using stream 35 b to provide a portion of the cooling of distillation vapor stream 42 in heat exchanger 22 allows more methane to be condensed and used as reflux in the present invention.
- the absorbing (rectification) section of the demethanizer it is generally advantageous to design the absorbing (rectification) section of the demethanizer to contain multiple theoretical separation stages.
- the benefits of the present invention can be achieved with as few as two theoretical stages.
- all or a part of the pumped condensed liquid (stream 44 a ) from reflux separator 23 and all or a part of the warmed expanded substantially condensed stream 35 c from heat exchanger 22 can be combined (such as in the piping joining the pump and heat exchanger to the demethanizer) and if thoroughly intermingled, the vapors and liquids will mix together and separate in accordance with the relative volatilities of the various components of the total combined streams.
- Such comingling of the two streams, combined with contacting at least a portion of expanded stream 36 a shall be considered for the purposes of this invention as constituting an absorbing section.
- FIGS. 5 through 8 display other embodiments of the present invention.
- FIGS. 4 through 6 depict fractionation towers constructed in a single vessel.
- FIGS. 7 and 8 depict fractionation towers constructed in two vessels, absorber (rectifier) column 27 (a contacting and separating device) and stripper (distillation) column 20 .
- a portion of the distillation vapor (stream 54 ) is withdrawn from the lower section of absorber column 27 and routed to reflux condenser 22 to generate reflux for absorber column 27 .
- the overhead vapor stream 50 from stripper column 20 flows to the lower section of absorber column 27 (via stream 51 ) to be contacted by reflux stream 52 and warmed expanded substantially condensed stream 35 c .
- Pump 28 is used to route the liquids (stream 47 ) from the bottom of absorber column 27 to the top of stripper column 20 so that the two towers effectively function as one distillation system.
- the decision whether to construct the fractionation tower as a single vessel (such as demethanizer 20 in FIGS. 4 through 6 ) or multiple vessels will depend on a number of factors such as plant size, the distance to fabrication facilities, etc.
- distillation vapor stream 42 in FIGS. 5 and 6 may favor withdrawing the distillation vapor stream 42 in FIGS. 5 and 6 from the upper region of stripping section 20 b in demethanizer 20 (stream 55 ).
- a portion (stream 55 ) of overhead vapor stream 50 from stripper column 20 may be directed to heat exchanger 22 (optionally combined with distillation vapor stream 54 withdrawn from the lower section of absorber column 27 ), with the remaining portion (stream 51 ) flowing to the lower section of absorber column 27 .
- the distillation vapor stream 42 or the combined distillation vapor stream 42 is partially condensed and the resulting condensate used to absorb valuable C 2 components, C 3 components, and heavier components from the vapors rising through absorbing section 20 a of demethanizer 20 or through absorber column 27 .
- the present invention is not limited to this embodiment. It may be advantageous, for instance, to treat only a portion of these vapors in this manner, or to use only a portion of the condensate as an absorbent, in cases where other design considerations indicate portions of the vapors or the condensate should bypass absorbing section 20 a of demethanizer 20 or absorber column 27 .
- distillation vapor stream 42 may favor total condensation, rather than partial condensation, of distillation vapor stream 42 or combined distillation vapor stream 42 in heat exchanger 22 .
- Other circumstances may favor that distillation vapor stream 42 be a total vapor side draw from fractionation column 20 or absorber column 27 rather than a partial vapor side draw.
- it may be advantageous to use external refrigeration to provide partial cooling of distillation vapor stream 42 or combined distillation vapor stream 42 in heat exchanger 22 .
- Feed gas conditions, plant size, available equipment, or other factors may indicate that elimination of work expansion machine 17 , or replacement with an alternate expansion device (such as an expansion valve), is feasible.
- an alternate expansion device such as an expansion valve
- 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 35 a ).
- separator 11 in FIG. 4 may not be justified. In such cases, the feed gas cooling accomplished in heat exchangers 10 and 13 in FIG. 4 may be accomplished without an intervening separator as shown in FIGS. 5 through 8 .
- the decision of whether or not to cool and separate the feed gas in multiple steps will depend on the richness of the feed gas, plant size, available equipment, etc.
- the cooled feed stream 31 a leaving heat exchanger 10 in FIGS. 4 through 8 and/or the cooled stream 32 a leaving heat exchanger 13 in FIG. 4 may not contain any liquid (because it is above its dewpoint, or because it is above its cricondenbar), so that separator 11 shown in FIGS. 4 through 8 and/or separator 14 shown in FIG. 4 are not required.
- the high pressure liquid (stream 37 in FIG. 4 and stream 33 in FIGS. 5 through 8 ) need not be expanded and fed to a lower mid-column feed point on the distillation column. Instead, all or a portion of it may be combined with the portion of the separator vapor (stream 35 in FIG. 4 and stream 34 in FIGS. 5 through 8 ) flowing to heat exchanger 15 . (This is shown by the dashed stream 46 in FIGS. 5 through 8 .) Any remaining portion of the liquid may be expanded through an appropriate expansion device, such as an expansion valve or expansion machine, and fed to a lower mid-column feed point on the distillation column (stream 37 a in FIGS. 5 through 8 ). Stream 33 in FIG. 4 and stream 37 in FIGS. 4 through 8 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.
- 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 a rich inlet gas.
- the use and distribution of separator liquids and demethanizer side draw liquids for process heat exchange, and the particular arrangement of heat exchangers for inlet gas cooling must be evaluated for each particular application, as well as the choice of process streams for specific heat exchange services.
- Some circumstances may favor using a portion of the cold distillation liquid leaving absorbing section 20 a or absorber column 27 for heat exchange, such as dashed stream 49 in FIGS. 5 through 8 .
- a portion of the liquid from absorbing section 20 a or absorber column 27 can be used for process heat exchange without reducing the ethane recovery in demethanizer 20 or stripper column 20 , more duty can sometimes be obtained from these liquids than with liquids from stripping section 20 b or stripper column 20 . This is because the liquids in absorbing section 20 a of demethanizer 20 (or absorber column 27 ) are available at a colder temperature level than those in stripping section 20 b (or stripper column 20 ).
- stream 53 in FIGS. 5 through 8 it may be advantageous to split the liquid stream from reflux pump 24 (stream 44 a ) into at least two streams.
- a portion (stream 53 ) can then be supplied to the stripping section of fractionation tower 20 ( FIGS. 5 and 6 ) or the top of stripper column 20 ( FIGS. 7 and 8 ) to increase the liquid flow in that part of the distillation system and improve the rectification, thereby reducing the concentration of C 2 + components in stream 42 .
- the remaining portion (stream 52 ) is supplied to the top of absorbing section 20 a ( FIGS. 5 and 6 ) or absorber column 27 ( FIGS. 7 and 8 ).
- the splitting of the vapor feed may be accomplished in several ways.
- the splitting of vapor occurs following cooling and separation of any liquids which may have been formed.
- the high pressure gas may be split, however, prior to any cooling of the inlet gas or after the cooling of the gas and prior to any separation stages.
- vapor splitting may be effected in a separator.
- the relative amount of feed found in each branch of the split vapor feed will depend on several factors, including gas pressure, feed gas composition, the amount of heat which can economically be extracted from the feed, and the quantity of horsepower available. More feed to the top of the column may increase recovery while decreasing power recovered from the expander thereby increasing the recompression horsepower requirements. Increasing feed lower in the column reduces the horsepower consumption but may also reduce product recovery.
- the 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.
- 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.
- the present invention provides improved recovery of C 2 components, C 3 components, and heavier hydrocarbon components or of C 3 components and heavier hydrocarbon components per amount of utility consumption required to operate the process.
- An improvement in utility consumption required for operating the demethanizer or deethanizer process may appear in the form of reduced power requirements for compression or re-compression, reduced power requirements for external refrigeration, reduced energy requirements for tower reboilers, or a combination thereof.
Abstract
Description
- This invention relates to a process and an apparatus for the separation of a gas containing hydrocarbons. The applicants claim the benefits under
Title 35, United States Code, Section 119(e) of prior U.S. Provisional Applications No. 61/244,181 which was filed on Sep. 21, 2009, No. 61/346,150 which was filed on May 19, 2010, and No. 61/351,045 which was filed on Jun. 3, 2010. - Ethylene, ethane, propylene, propane, and/or heavier hydrocarbons can be recovered from a variety of gases, such as natural gas, refinery gas, and synthetic gas streams obtained from other hydrocarbon materials such as coal, crude oil, naphtha, oil shale, tar sands, and lignite. Natural gas usually has a major proportion of methane and ethane, i.e., methane and ethane together comprise at least 50 mole percent of the gas. The gas also contains relatively lesser amounts of heavier hydrocarbons such as propane, butanes, pentanes, and the like, as well as hydrogen, nitrogen, carbon dioxide, and other gases.
- 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, 80.8% methane, 9.4% ethane and other C2 components, 4.7% propane and other C3 components, 1.2% iso-butane, 2.1% normal butane, and 1.1% 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. This has resulted in a demand for processes that can provide more efficient recoveries of these products, for processes that can provide efficient recoveries with lower capital investment, and for processes that can be easily adapted or adjusted to vary the recovery of a specific component over a broad range. Available processes for separating these materials include those based upon cooling and refrigeration of gas, oil absorption, and refrigerated oil absorption. Additionally, cryogenic processes have become popular because of the availability of economical equipment that produces power while simultaneously expanding and extracting heat from the gas being processed. 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 startup, operating flexibility, good efficiency, safety, and good reliability. U.S. Pat. Nos. 3,292,380; 4,061,481; 4,140,504; 4,157,904; 4,171,964; 4,185,978; 4,251,249; 4,278,457; 4,519,824; 4,617,039; 4,687,499; 4,689,063; 4,690,702; 4,854,955; 4,869,740; 4,889,545; 5,275,005; 5,555,748; 5,566,554; 5,568,737; 5,771,712; 5,799,507; 5,881,569; 5,890,378; 5,983,664; 6,182,469; 6,578,379; 6,712,880; 6,915,662; 7,191,617; 7,219,513; reissue U.S. Pat. No. 33,408; and co-pending application Ser. Nos. 11/430,412; 11/839,693; 11/971,491; 12/206,230; 12/689,616; 12/717,394; 12/750,862; 12/772,472; and 12/781,259 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).
- In a typical cryogenic expansion recovery process, a feed gas stream under pressure is cooled by heat exchange with other streams of the process and/or external sources of refrigeration such as a propane compression-refrigeration system. As the gas is cooled, liquids may be condensed and collected in one or more separators as high-pressure liquids containing some of the desired C2+ components. Depending on the richness of the gas and the amount of liquids formed, the high-pressure liquids may be expanded to a lower pressure and fractionated. The vaporization occurring during expansion of the liquids results in further cooling of the stream. Under some conditions, pre-cooling the high pressure liquids prior to the expansion may be desirable in order to further lower the temperature resulting from the expansion. The expanded stream, comprising a mixture of liquid and vapor, is fractionated in a distillation (demethanizer or deethanizer) column. In the column, the expansion cooled stream(s) is (are) distilled to separate residual methane, nitrogen, and other volatile gases as overhead vapor from the desired C2 components, C3 components, and heavier hydrocarbon components as bottom liquid product, or to separate residual methane, C2 components, nitrogen, and other volatile gases as overhead vapor from the desired C3 components and heavier hydrocarbon components as bottom liquid product.
- If the feed gas is not totally condensed (typically it is not), the vapor remaining from the partial condensation can be split into two streams. One portion of the vapor is passed through a work expansion machine or engine, or an expansion valve, to a lower pressure at which additional liquids are condensed as a result of further cooling of the stream. The pressure after expansion is essentially the same as the pressure at which the distillation column is operated. The combined vapor-liquid phases resulting from the expansion are supplied as feed to the column.
- The remaining portion of the vapor is cooled to substantial condensation by heat exchange with other process streams, e.g., the cold fractionation tower overhead. Some or all of the high-pressure liquid may be combined with this vapor portion prior to cooling. The resulting cooled stream is then expanded through an appropriate expansion device, such as an expansion valve, to the pressure at which the demethanizer is operated. During expansion, a portion of the liquid will vaporize, resulting in cooling of the total stream. The flash expanded stream is then supplied as top feed to the demethanizer. Typically, the vapor portion of the flash expanded stream and the demethanizer overhead vapor combine in an upper separator section in the fractionation tower as residual methane product gas. Alternatively, the cooled and expanded stream may be supplied to a separator to provide vapor and liquid streams. The vapor is combined with the tower overhead and the liquid is supplied to the column as a top column feed.
- In the ideal operation of such a separation process, the residue gas leaving the process will contain substantially all of the methane in the feed gas with essentially none of the heavier hydrocarbon components, and the bottoms fraction leaving the demethanizer will contain substantially all of the heavier hydrocarbon components with essentially no methane or more volatile components. In practice, however, this ideal situation is not obtained because the conventional demethanizer is operated largely as a stripping column. The methane product of the process, therefore, typically comprises vapors leaving the top fractionation stage of the column, together with vapors not subjected to any rectification step. Considerable losses of C2, C3, and C4+ components occur because the top liquid feed contains substantial quantities of these components and heavier hydrocarbon components, resulting in corresponding equilibrium quantities of C2 components, C3 components, C4 components, and heavier hydrocarbon components in the vapors leaving the top fractionation stage of the demethanizer. The loss of these desirable components could be significantly reduced if the rising vapors could be brought into contact with a significant quantity of liquid (reflux) capable of absorbing the C2 components, C3 components, C4 components, and heavier hydrocarbon components from the vapors.
- In recent years, the preferred processes for hydrocarbon separation use an upper absorber section to provide additional rectification of the rising vapors. The source of the reflux stream for the upper rectification section is typically a recycled stream of residue gas supplied under pressure. The recycled residue gas stream is usually cooled to substantial condensation by heat exchange with other process streams, e.g., the cold fractionation tower overhead. The resulting substantially condensed stream is then expanded through an appropriate expansion device, such as an expansion valve, to the pressure at which the demethanizer is operated. During expansion, a portion of the liquid will usually vaporize, resulting in cooling of the total stream. The flash expanded stream is then supplied as top feed to the demethanizer. Typically, the vapor portion of the expanded stream and the demethanizer overhead vapor combine in an upper separator section in the fractionation tower as residual methane product gas. Alternatively, the cooled and expanded stream may be supplied to a separator to provide vapor and liquid streams, so that thereafter the vapor is combined with the tower overhead and the liquid is supplied to the column as a top column feed. Typical process schemes of this type are disclosed in U.S. Pat. Nos. 4,889,545; 5,568,737; and 5,881,569, assignee's co-pending application Ser. No. 12/717,394, and in Mowrey, E. Ross, “Efficient, High Recovery of Liquids from Natural Gas Utilizing a High Pressure Absorber”, Proceedings of the Eighty-First Annual Convention of the Gas Processors Association, Dallas, Tex., Mar. 11-13, 2002. Unfortunately, these processes require the use of a compressor to provide the motive force for recycling the reflux stream to the demethanizer, adding to both the capital cost and the operating cost of facilities using these processes.
- The present invention also employs an upper rectification section (or a separate rectification column if plant size or other factors favor using separate rectification and stripping columns). However, the reflux stream for this rectification section is provided by using a side draw of the vapors rising in a lower portion of the tower. Because of the relatively high concentration of C2 components in the vapors lower in the tower, a significant quantity of liquid can be condensed in this side draw stream without elevating its pressure, often using only the refrigeration available in the cold vapor leaving the upper rectification section and the flash expanded substantially condensed stream. This condensed liquid, which is predominantly liquid methane, can then be used to absorb C2 components, C3 components, C4 components, and heavier hydrocarbon components from the vapors rising through the upper rectification section and thereby capture these valuable components in the bottom liquid product from the demethanizer.
- Heretofore, such a side draw feature has been employed in C3+ recovery systems, as illustrated in the assignee's U.S. Pat. No. 5,799,507, as well as in C2+ recovery systems, as illustrated in the assignee's U.S. Pat. No. 7,191,617 and co-pending application Ser. Nos. 12/206,230 and 12/781,259. Surprisingly, applicants have found that using the flash expanded substantially condensed stream to provide a portion of the cooling of the side draw feature disclosed in assignee's co-pending application Ser. Nos. 12/206,230 and 12/781,259 processes improves the C2+ recoveries and the system efficiency with no increase in operating cost.
- In accordance with the present invention, it has been found that C2 recovery in excess of 87% and C3 and C4+ recoveries in excess of 99% can be obtained without the need for compression of the reflux stream for the demethanizer. The present invention provides the further advantage of being able to maintain in excess of 99% recovery of the C3 and C4+ components as the recovery of C2 components is adjusted from high to low values. In addition, the present invention makes possible essentially 100% separation of methane and lighter components from the C2 components and heavier components at the same energy requirements compared to the prior art while increasing the recovery levels. The present invention, although applicable at lower pressures and warmer temperatures, is particularly advantageous when processing feed gases in the range of 400 to 1500 psia [2,758 to 10,342 kPa(a)] or higher under conditions requiring NGL recovery column overhead temperatures of −50° F. [−46° C.] or colder.
- 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 natural gas processing plant in accordance with U.S. Pat. No. 5,890,378; -
FIG. 2 is a flow diagram of a prior art natural gas processing plant in accordance with U.S. Pat. No. 7,191,617; -
FIG. 3 is a flow diagram of a prior art natural gas processing plant in accordance with assignee's co-pending application Ser. No. 12/206,230; -
FIG. 4 is a flow diagram of a natural gas processing plant in accordance with the present invention; and -
FIGS. 5 through 8 are flow diagrams illustrating alternative means of application of the present invention to a natural gas stream. - In the following explanation of the above figures, tables are provided summarizing flow rates calculated for representative process conditions. In the tables appearing herein, the values for flow rates (in moles per hour) have been rounded to the nearest whole number for convenience. The total stream rates shown in the tables include all non-hydrocarbon components and hence are generally larger than the sum of the stream flow rates for the hydrocarbon components. Temperatures indicated are approximate values rounded to the nearest degree. It should also be noted that the process design calculations performed for the purpose of comparing the processes depicted in the figures are based on the assumption of no heat leak from (or to) the surroundings to (or from) the process. The quality of commercially available insulating materials makes this a very reasonable assumption and one that is typically made by those skilled in the art.
- For convenience, process parameters are reported in both the traditional British units and in the units of the Systeme International d'Unités (SI). The molar flow rates given in the tables may be interpreted as either pound moles per hour or kilogram moles per hour. The energy consumptions reported as horsepower (HP) and/or thousand British Thermal Units per hour (MBTU/Hr) correspond to the stated molar flow rates in pound moles per hour. The energy consumptions reported as kilowatts (kW) correspond to the stated molar flow rates in kilogram moles per hour.
-
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. 5,890,378. In this simulation of the process, inlet gas enters the plant at 85° F. [29° C.] and 970 psia [6,688 kPa(a)] asstream 31. If the inlet gas contains a concentration of sulfur compounds which would prevent the product streams from meeting specifications, the sulfur compounds are removed by appropriate pretreatment of the feed gas (not illustrated). In addition, the feed stream is usually dehydrated to prevent hydrate (ice) formation under cryogenic conditions. Solid desiccant has typically been used for this purpose. - The
feed stream 31 is cooled inheat exchanger 10 by heat exchange with cool residue gas (stream 45 b), demethanizer lower side reboiler liquids at 32° F. [0° C.] (stream 40), and propane refrigerant. Note that in all cases exchanger 10 is representative of either a multitude of individual heat exchangers or a single multi-pass heat exchanger, or any combination thereof. (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.) The cooledstream 31 a entersseparator 11 at 0° F. [−18° C.] and 955 psia [6,584 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33). The separator liquid (stream 33) is expanded to the operating pressure (approximately 444 psia [3,061 kPa(a)]) offractionation tower 20 byexpansion valve 12, coolingstream 33 a to −27° F. [−33° C.] before it is supplied tofractionation tower 20 at a first lower mid-column feed point. - The vapor (stream 32) from
separator 11 is further cooled inheat exchanger 13 by heat exchange with cool residue gas (stream 45 a) and demethanizer upper side reboiler liquids at −39° F. [−39° C.] (stream 39). The cooledstream 32 a entersseparator 14 at −31° F. [−35° C.] and 950 psia [6,550 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 37). The separator liquid (stream 37) is expanded to the tower operating pressure byexpansion valve 19, coolingstream 37 a to −66° F. [−54° C.] before it is supplied tofractionation tower 20 at a second lower mid-column feed point. - The vapor (stream 34) from
separator 14 is divided into two streams, 35 and 36.Stream 35, containing about 39% of the total vapor, passes throughheat exchanger 15 in heat exchange relation with the cold residue gas (stream 45) where it is cooled to substantial condensation. The resulting substantially condensedstream 35 a at −123° F. [−86° C.] is then flash expanded throughexpansion valve 16 to slightly above the operating pressure offractionation tower 20. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated inFIG. 1 , the expandedstream 35 b leavingexpansion valve 16 reaches a temperature of −130° F. [−90° C.]. The expandedstream 35 b is warmed to −126° F. [−88° C.] and further vaporized inheat exchanger 22 as it provides cooling and partial condensation ofdistillation vapor stream 42 withdrawn from strippingsection 20 b offractionation tower 20. The warmedstream 35 c is then supplied at an upper mid-column feed point, in absorbingsection 20 a offractionation tower 20. - The remaining 61% of the vapor from separator 14 (stream 36) enters a
work expansion machine 17 in which mechanical energy is extracted from this portion of the high pressure feed. Themachine 17 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expandedstream 36 a to a temperature of approximately −86° F. [−66° C.]. The typical commercially available expanders are capable of recovering on the order of 80-85% of the work theoretically available in an ideal isentropic expansion. The work recovered is often used to drive a centrifugal compressor (such as item 18) that can be used to re-compress the residue gas (stream 45 c), for example. The partially condensed expandedstream 36 a is thereafter supplied as feed tofractionation tower 20 at a mid-column feed point. - The demethanizer in
tower 20 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. The demethanizer tower consists of two sections: an upper absorbing (rectification)section 20 a that contains the trays and/or packing to provide the necessary contact between the vapor portions of the expandedstreams section 20 b that contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. Thedemethanizing section 20 b also includes one or more reboilers (such asreboiler 21 and the side reboilers described previously) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column to strip the liquid product,stream 41, of methane and lighter components.Stream 36 a entersdemethanizer 20 at an intermediate feed position located in the lower region of absorbingsection 20 a ofdemethanizer 20. The liquid portion of the expandedstream 36 a comingles with liquids falling downward from absorbingsection 20 a and the combined liquid continues downward into strippingsection 20 b ofdemethanizer 20. The vapor portion of the expandedstream 36 a rises upward through absorbingsection 20 a and is contacted with cold liquid falling downward to condense and absorb the C2 components, C3 components, and heavier components. - A portion of the distillation vapor (stream 42) is withdrawn from the upper region of stripping
section 20 b. This stream is then cooled and partially condensed (stream 42 a) inexchanger 22 by heat exchange with expanded substantially condensedstream 35 b as described previously, coolingstream 42 from −96° F. [−71° C.] to about −128° F. [−89° C.] (stream 42 a). The operating pressure (441 psia [3,038 kPa(a)]) inreflux separator 23 is maintained slightly below the operating pressure ofdemethanizer 20. This provides the driving force which causesdistillation vapor stream 42 to flow throughheat exchanger 22 and thence into thereflux separator 23 where the condensed liquid (stream 44) is separated from any uncondensed vapor (stream 43). - The
liquid stream 44 fromreflux separator 23 is pumped bypump 24 to a pressure slightly above the operating pressure ofdemethanizer 20, and stream 44 a is then supplied as cold top column feed (reflux) todemethanizer 20 at −128° F. [−89° C.]. This cold liquid reflux absorbs and condenses the C3 components and heavier components rising in the upper rectification region of absorbingsection 20 a ofdemethanizer 20. - The
liquid product stream 41 exits the bottom of the tower at 112° F. [44° C.], based on a typical specification of a methane to ethane ratio of 0.025:1 on a molar basis in the bottom product. Cold demethanizeroverhead stream 38 exits the top ofdemethanizer 20 at −128° F. [−89° C.] and combines withvapor stream 43 to form coldresidue gas stream 45 at −128° F. [−89° C.]. The coldresidue gas stream 45 passes countercurrently to the incoming feed gas inheat exchanger 15 where it is heated to −37° F. [−38° C.] (stream 45 a), inheat exchanger 13 where it is heated to −5° F. [−21° C.] (stream 45 b), and inheat exchanger 10 where it is heated to 80° F. [27° C.] (stream 45 c). The residue gas is then re-compressed in two stages. The first stage iscompressor 18 driven byexpansion machine 17. The second stage iscompressor 25 driven by a supplemental power source which compresses the residue gas (stream 45 d) to sales line pressure. After cooling to 120° F. [49° C.] in discharge cooler 26, the residue gas product (stream 451) flows to the sales gas pipeline at 1015 psia [6,998 kPa(a)], sufficient to meet line requirements (usually on the order of the inlet pressure). - A summary of stream flow rates and energy consumption for the process illustrated in
FIG. 1 is set forth in the following table: -
TABLE I (FIG. 1) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 53,228 6,192 3,070 2,912 65,876 32 49,244 4,670 1,650 815 56,795 33 3,984 1,522 1,420 2,097 9,081 34 47,282 4,037 1,178 405 53,293 37 1,962 633 472 410 3,502 35 18,582 1,587 463 159 20,944 36 28,700 2,450 715 246 32,349 38 44,854 790 11 0 45,920 42 12,398 720 42 3 13,270 43 8,242 135 2 0 8,421 44 4,156 585 40 3 4,849 45 53,096 925 13 0 54,341 41 132 5,267 3,057 2,912 11,535 Recoveries* Ethane 85.05% Propane 99.57% Butanes+ 99.99% Power *(Based on un-rounded flow rates) -
FIG. 2 represents an alternative prior art process according to U.S. Pat. No. 7,191,617. The process ofFIG. 2 has been applied to the same feed gas composition and conditions as described above forFIG. 1 . In the simulation of this process, as in the simulation for the process ofFIG. 1 , operating conditions were selected to minimize energy consumption for a given recovery level. - In the simulation of the
FIG. 2 process, inlet gas enters the plant asstream 31 and is cooled inheat exchanger 10 by heat exchange with cool residue gas (stream 45 b), demethanizer lower side reboiler liquids at 33° F. [0° C.] (stream 40), and propane refrigerant. The cooledstream 31 a entersseparator 11 at 0° F. [−18° C.] and 955 psia [6,584 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33). The separator liquid (stream 33) is expanded to the operating pressure (approximately 450 psia [3,103 kPa(a)]) offractionation tower 20 byexpansion valve 12, coolingstream 33 a to −27° F. [−33° C.] before it is supplied tofractionation tower 20 at a first lower mid-column feed point. - The vapor (stream 32) from
separator 11 is further cooled inheat exchanger 13 by heat exchange with cool residue gas (stream 45 a) and demethanizer upper side reboiler liquids at −38° F. [−39° C.] (stream 39). The cooledstream 32 a entersseparator 14 at −29° F. [−34° C.] and 950 psia [6,550 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 37). The separator liquid (stream 37) is expanded to the tower operating pressure byexpansion valve 19, coolingstream 37 a to −64° F. [−53° C.] before it is supplied tofractionation tower 20 at a second lower mid-column feed point. - The vapor (stream 34) from
separator 14 is divided into two streams, 35 and 36.Stream 35, containing about 37% of the total vapor, passes throughheat exchanger 15 in heat exchange relation with the cold residue gas (stream 45) where it is cooled to substantial condensation. The resulting substantially condensedstream 35 a at −115° F. [−82° C.] is then flash expanded throughexpansion valve 16 to the operating pressure offractionation tower 20. During expansion a portion of the stream is vaporized, resulting in cooling ofstream 35 b to −129° F. [−89° C.] before it is supplied tofractionation tower 20 at an upper mid-column feed point. - The remaining 63% of the vapor from separator 14 (stream 36) enters a
work expansion machine 17 in which mechanical energy is extracted from this portion of the high pressure feed. Themachine 17 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expandedstream 36 a to a temperature of approximately −84° F. [−65° C.]. The partially condensed expandedstream 36 a is thereafter supplied as feed tofractionation tower 20 at a mid-column feed point. - A portion of the distillation vapor (stream 42) is withdrawn from the upper region of the stripping section in
fractionation tower 20. This stream is then cooled from −91° F. [−68° C.] to −122° F. [−86° C.] and partially condensed (stream 42 a) inheat exchanger 22 by heat exchange with the cold demethanizeroverhead stream 38 exiting the top ofdemethanizer 20 at −127° F. [−88° C.]. The cold demethanizer overhead stream is warmed slightly to −120° F. [−84° C.] (stream 38 a) as it cools and condenses at least a portion ofstream 42. - The operating pressure (447 psia [3,079 kPa(a)]) in
reflux separator 23 is maintained slightly below the operating pressure ofdemethanizer 20. This provides the driving force which causesdistillation vapor stream 42 to flow throughheat exchanger 22 and thence into thereflux separator 23 where the condensed liquid (stream 44) is separated from any uncondensed vapor (stream 43).Stream 43 then combines with the warmed demethanizeroverhead stream 38 a fromheat exchanger 22 to form coldresidue gas stream 45 at −120° F. [−84° C.]. - The
liquid stream 44 fromreflux separator 23 is pumped bypump 24 to a pressure slightly above the operating pressure ofdemethanizer 20, and stream 44 a is then supplied as cold top column feed (reflux) todemethanizer 20 at −121° F. [−85° C.]. This cold liquid reflux absorbs and condenses the C3 components and heavier components rising in the upper rectification region of the absorbing section ofdemethanizer 20. - The
liquid product stream 41 exits the bottom oftower 20 at 114° F. [45° C.]. The coldresidue gas stream 45 passes countercurrently to the incoming feed gas inheat exchanger 15 where it is heated to −36° F. [−38° C.] (stream 45 a), inheat exchanger 13 where it is heated to −5° F. [−20° C.] (stream 45 b), and inheat exchanger 10 where it is heated to 80° F. [27° C.] (stream 45 c) as it provides cooling as previously described. The residue gas is then re-compressed in two stages,compressor 18 driven byexpansion machine 17 andcompressor 25 driven by a supplemental power source. Afterstream 45 e is cooled to 120° F. [49° C.] in discharge cooler 26, the residue gas product (stream 45 f) flows to the sales gas pipeline at 1015 psia [6,998 kPa(a)]. - A summary of stream flow rates and energy consumption for the process illustrated in
FIG. 2 is set forth in the following table: -
TABLE II (FIG. 2) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 53,228 6,192 3,070 2,912 65,876 32 49,244 4,670 1,650 815 56,795 33 3,984 1,522 1,420 2,097 9,081 34 47,440 4,081 1,204 420 53,536 37 1,804 589 446 395 3,259 35 17,553 1,510 445 155 19,808 36 29,887 2,571 759 265 33,728 38 48,675 811 23 1 49,805 42 5,555 373 22 2 6,000 43 4,421 113 2 0 4,562 44 1,134 260 20 2 1,438 45 53,096 924 25 1 54,367 41 132 5,268 3,045 2,911 11,509 Recoveries* Ethane 85.08% Propane 99.20% Butanes+ 99.98% Power *(Based on un-rounded flow rates) - A comparison of Tables I and II shows that, compared to the
FIG. 1 process, theFIG. 2 process maintains essentially the same ethane recovery (85.08% versus 85.05%) and butanes+recovery (99.98% versus 99.99%), but the propane recovery drops from 99.57% to 99.20%. However, comparison of Tables I and II further shows that the power requirement for theFIG. 2 process is about 2% lower than that of theFIG. 1 process. -
FIG. 3 represents an alternative prior art process according to co-pending application Ser. No. 12/206,230. The process ofFIG. 3 has been applied to the same feed gas composition and conditions as described above forFIGS. 1 and 2 . In the simulation of this process, as in the simulation for the process ofFIGS. 1 and 2 , operating conditions were selected to minimize energy consumption for a given recovery level. - In the simulation of the
FIG. 3 process, inlet gas enters the plant asstream 31 and is cooled inheat exchanger 10 by heat exchange with cool residue gas (stream 45 b), demethanizer lower side reboiler liquids at 36° F. [2° C.] (stream 40), and propane refrigerant. The cooledstream 31 a entersseparator 11 at 1° F. [−17° C.] and 955 psia [6,584 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33). The separator liquid (stream 33) is expanded to the operating pressure (approximately 452 psia [3,116 kPa(a)]) offractionation tower 20 byexpansion valve 12, coolingstream 33 a to −25° F. [−32° C.] before it is supplied tofractionation tower 20 at a first lower mid-column feed point. - The vapor (stream 32) from
separator 11 is further cooled inheat exchanger 13 by heat exchange with cool residue gas (stream 45 a) and demethanizer upper side reboiler liquids at −37° F. [−38° C.] (stream 39). The cooledstream 32 a entersseparator 14 at −31° F. [−35° C.] and 950 psia [6,550 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 37). The separator liquid (stream 37) is expanded to the tower operating pressure byexpansion valve 19, coolingstream 37 a to −65° F. [−54° C.] before it is supplied tofractionation tower 20 at a second lower mid-column feed point. - The vapor (stream 34) from
separator 14 is divided into two streams, 35 and 36.Stream 35, containing about 38% of the total vapor, passes throughheat exchanger 15 in heat exchange relation with the cold residue gas (stream 45) where it is cooled to substantial condensation. The resulting substantially condensedstream 35 a at −119° F. [−84° C.] is then flash expanded throughexpansion valve 16 to the operating pressure offractionation tower 20. During expansion a portion of the stream is vaporized, resulting in cooling ofstream 35 b to −129° F. [−90° C.] before it is supplied tofractionation tower 20 at an upper mid-column feed point. - The remaining 62% of the vapor from separator 14 (stream 36) enters a
work expansion machine 17 in which mechanical energy is extracted from this portion of the high pressure feed. Themachine 17 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expandedstream 36 a to a temperature of approximately −85° F. [−65° C.]. The partially condensed expandedstream 36 a is thereafter supplied as feed tofractionation tower 20 at a mid-column feed point. - A portion of the distillation vapor (stream 42) is withdrawn from an intermediate region of the absorbing section in
fractionation column 20, above the feed position of expandedstream 36 a in the lower region of the absorbing section. Thisdistillation vapor stream 42 is then cooled from −101° F. [−74° C.] to −124° F. [−86° C.] and partially condensed (stream 42 a) inheat exchanger 22 by heat exchange with the cold demethanizeroverhead stream 38 exiting the top ofdemethanizer 20 at −128° F. [−89° C.]. The cold demethanizer overhead stream is warmed slightly to −124° F. [−86° C.] (stream 38 a) as it cools and condenses at least a portion ofstream 42. - The operating pressure (448 psia [3,090 kPa(a)]) in
reflux separator 23 is maintained slightly below the operating pressure ofdemethanizer 20. This provides the driving force which causesdistillation vapor stream 42 to flow throughheat exchanger 22 and thence into thereflux separator 23 where the condensed liquid (stream 44) is separated from any uncondensed vapor (stream 43).Stream 43 then combines with the warmed demethanizeroverhead stream 38 a fromheat exchanger 22 to form coldresidue gas stream 45 at −124° F. [−86° C.]. - The
liquid stream 44 fromreflux separator 23 is pumped bypump 24 to a pressure slightly above the operating pressure ofdemethanizer 20, and stream 44 a is then supplied as cold top column feed (reflux) todemethanizer 20 at −123° F. [−86° C.]. This cold liquid reflux absorbs and condenses the C2 components, C3 components, and heavier components rising in the upper rectification region of the absorbing section ofdemethanizer 20. - The
liquid product stream 41 exits the bottom oftower 20 at 113° F. [45° C.]. The coldresidue gas stream 45 passes countercurrently to the incoming feed gas inheat exchanger 15 where it is heated to −38° F. [−39° C.] (stream 45 a), inheat exchanger 13 where it is heated to −4° F. [−20° C.] (stream 45 b), and inheat exchanger 10 where it is heated to 80° F. [27° C.] (stream 45 c) as it provides cooling as previously described. The residue gas is then re-compressed in two stages,compressor 18 driven byexpansion machine 17 andcompressor 25 driven by a supplemental power source. Afterstream 45 e is cooled to 120° F. [49° C.] in discharge cooler 26, the residue gas product (stream 45 f) flows to the sales gas pipeline at 1015 psia [6,998 kPa(a)]. - A summary of stream flow rates and energy consumption for the process illustrated in
FIG. 3 is set forth in the following table: -
TABLE III (FIG. 3) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 53,228 6,192 3,070 2,912 65,876 32 49,340 4,702 1,672 831 56,962 33 3,888 1,490 1,398 2,081 8,914 34 47,289 4,040 1,179 404 53,301 37 2,051 662 493 427 3,661 35 17,828 1,523 444 152 20,094 36 29,461 2,517 735 252 33,207 38 49,103 691 19 0 50,103 42 4,946 285 8 0 5,300 43 3,990 93 1 0 4,119 44 956 192 7 0 1,181 45 53,093 784 20 0 54,222 41 135 5,408 3,050 2,912 11,654 Recoveries* Ethane 87.33% Propane 99.36% Butanes+ 99.99% Power *(Based on un-rounded flow rates) - A comparison of Tables I, II, and III shows that the
FIG. 3 process improves the ethane recovery from 85.05% (forFIG. 1 ) and 85.08% (forFIG. 2 ) to 87.33%. The propane recovery for theFIG. 3 process (99.36%) is lower than that of theFIG. 1 process (99.57%) but higher than that of theFIG. 2 process (99.20%). The butanes+recovery is essentially the same for all three of these prior art processes. Comparison of Tables I, II, and III further shows that theFIG. 3 process using slightly less power than both prior art processes (more than 2% less than theFIG. 1 process and 0.4% less than theFIG. 2 process). -
FIG. 4 illustrates a flow diagram of a process in accordance with the present invention. The feed gas composition and conditions considered in the process presented inFIG. 4 are the same as those inFIGS. 1 , 2, and 3. Accordingly, theFIG. 4 process can be compared with that of theFIGS. 1 , 2, and 3 processes to illustrate the advantages of the present invention. - In the simulation of the
FIG. 4 process, inlet gas enters the plant at 85° F. [29° C.] and 970 psia [6,688 kPa(a)] asstream 31 and is cooled inheat exchanger 10 by heat exchange with cool residue gas (stream 45 b), demethanizer lower side reboiler liquids at 32° F. [0° C.] (stream 40), and propane refrigerant. The cooledstream 31 a entersseparator 11 at 1° F. [−17° C.] and 955 psia [6,584 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33). The separator liquid (stream 33) is expanded to the operating pressure (approximately 452 psia [3,116 kPa(a)]) offractionation tower 20 byexpansion valve 12, coolingstream 33 a to −25° F. [−32° C.] before it is supplied tofractionation tower 20 at a first lower mid-column feed point (located below the feed point ofstream 36 a described later in paragraph [0058]). - The vapor (stream 32) from
separator 11 is further cooled inheat exchanger 13 by heat exchange with cool residue gas (stream 45 a) and demethanizer upper side reboiler liquids at −38° F. [−39° C.] (stream 39). The cooledstream 32 a entersseparator 14 at −31° F. [−35° C.] and 950 psia [6,550 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 37). The separator liquid (stream 37) is expanded to the tower operating pressure byexpansion valve 19, coolingstream 37 a to −66° F. [−54° C.] before it is supplied tofractionation tower 20 at a second lower mid-column feed point (also located below the feed point ofstream 36 a). - The vapor (stream 34) from
separator 14 is divided into two streams, 35 and 36.Stream 35, containing about 38% of the total vapor, passes throughheat exchanger 15 in heat exchange relation with the cold residue gas (stream 45) where it is cooled to substantial condensation. The resulting substantially condensedstream 35 a at −122° F. [−86° C.] is then flash expanded throughexpansion valve 16 to slightly above the operating pressure offractionation tower 20. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated inFIG. 4 , the expandedstream 35 b leavingexpansion valve 16 reaches a temperature of −130° F. [−90° C.]. The expandedstream 35 b is warmed slightly to −129° F. [−89° C.] and further vaporized inheat exchanger 22 as it provides a portion of the cooling ofdistillation vapor stream 42. The warmedstream 35 c is then supplied at an upper mid-column feed point, in absorbingsection 20 a offractionation tower 20. - The remaining 62% of the vapor from separator 14 (stream 36) enters a
work expansion machine 17 in which mechanical energy is extracted from this portion of the high pressure feed. Themachine 17 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expandedstream 36 a to a temperature of approximately −86° F. [−65° C.]. The partially condensed expandedstream 36 a is thereafter supplied as feed tofractionation tower 20 at a mid-column feed point (located below the feed point ofstream 35 c). - The demethanizer in
tower 20 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. The demethanizer tower consists of two sections: an upper absorbing (rectification)section 20 a that contains the trays and/or packing to provide the necessary contact between the vapor portions of the expandedstreams section 20 b that contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. Thedemethanizing section 20 b also includes one or more reboilers (such asreboiler 21 and the side reboilers described previously) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column to strip the liquid product,stream 41, of methane and lighter components.Stream 36 a entersdemethanizer 20 at an intermediate feed position located in the lower region of absorbingsection 20 a ofdemethanizer 20. The liquid portion of the expandedstream 36 a comingles with liquids falling downward from absorbingsection 20 a and the combined liquid continues downward into strippingsection 20 b ofdemethanizer 20. The vapor portion of the expandedstream 36 a rises upward through absorbingsection 20 a and is contacted with cold liquid falling downward to condense and absorb the C2 components, C3 components, and heavier components. - A portion of the distillation vapor (stream 42) is withdrawn from an intermediate region of absorbing
section 20 a infractionation column 20, above the feed position of expandedstream 36 a in the lower region of absorbingsection 20 a. Thisdistillation vapor stream 42 is then cooled from −103° F. [−75° C.] to −128° F. [−89° C.] and partially condensed (stream 42 a) inheat exchanger 22 by heat exchange with the cold demethanizeroverhead stream 38 exiting the top ofdemethanizer 20 at −129° F. [−89° C.] and with the expanded substantially condensedstream 35 b as described previously. The cold demethanizer overhead stream is warmed slightly to −127° F. [−88° C.] (stream 38 a) as it provides a portion of the cooling ofdistillation vapor stream 42. - The operating pressure (448 psia [3,090 kPa(a)]) in
reflux separator 23 is maintained slightly below the operating pressure ofdemethanizer 20. This provides the driving force which causesdistillation vapor stream 42 to flow throughheat exchanger 22 and thence into thereflux separator 23 where the condensed liquid (stream 44) is separated from any uncondensed vapor (stream 43).Stream 43 then combines with the warmed demethanizeroverhead stream 38 a fromheat exchanger 22 to form coldresidue gas stream 45 at −127° F. [−88° C.]. - The
liquid stream 44 fromreflux separator 23 is pumped bypump 24 to a pressure slightly above the operating pressure ofdemethanizer 20, and stream 44 a is then supplied as cold top column feed (reflux) todemethanizer 20 at −127° F. [−88° C.]. This cold liquid reflux absorbs and condenses the C2 components, C3 components, and heavier components rising in the upper rectification region of absorbingsection 20 a ofdemethanizer 20. - In stripping
section 20 b ofdemethanizer 20, the feed streams are stripped of their methane and lighter components. The resulting liquid product (stream 41) exits the bottom oftower 20 at 113° F. [45° C.] (based on a typical specification of a methane to ethane ratio of 0.025:1 on a molar basis in the bottom product). The coldresidue gas stream 45 passes countercurrently to the incoming feed gas inheat exchanger 15 where it is heated to −40° F. [−40° C.] (stream 45 a), inheat exchanger 13 where it is heated to −4° F. [−20° C.] (stream 45 b), and inheat exchanger 10 where it is heated to 80° F. [27° C.] (stream 45 c) as it provides cooling as previously described. The residue gas is then re-compressed in two stages,compressor 18 driven byexpansion machine 17 andcompressor 25 driven by a supplemental power source. Afterstream 45 e is cooled to 120° F. [49° C.] in discharge cooler 26, the residue gas product (stream 45 f) flows to the sales gas pipeline at 1015 psia [6,998 kPa(a)]. - A summary of stream flow rates and energy consumption for the process illustrated in
FIG. 4 is set forth in the following table: -
TABLE IV (FIG. 4) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 53,228 6,192 3,070 2,912 65,876 32 49,407 4,712 1,676 832 57,046 33 3,821 1,480 1,394 2,080 8,830 34 47,346 4,041 1,176 401 53,354 37 2,061 671 500 431 3,692 35 17,991 1,536 447 152 20,274 36 29,355 2,505 729 249 33,080 38 49,756 713 14 0 50,779 42 4,688 249 7 0 5,000 43 3,336 57 0 0 3,420 44 1,352 192 7 0 1,580 45 53,092 770 14 0 54,199 41 136 5,422 3,056 2,912 11,677 Recoveries* Ethane 87.56% Propane 99.55% Butanes+ 99.99% Power * (Based on un-rounded flow rates) - A comparison of Tables I, II, III, and IV shows that, compared to the prior art, the present invention matches or exceeds the propane and butanes+recoveries of all the prior art processes while significantly improving the ethane recovery. The ethane recovery for the present invention (87.56%) is higher than the
FIG. 1 process (85.05%), theFIG. 2 process (85.08%), and theFIG. 3 process (87.33%). Comparison of Tables I, II, III, and IV further shows that the improvement in yields was achieved without using more power than the prior art, and in some cases using significantly less power. In terms of the recovery efficiency (defined by the quantity of ethane recovered per unit of power), the present invention represents an improvement of 5%, 3%, and 0.3%, respectively, over the prior art of theFIG. 1 ,FIG. 2 , andFIG. 3 processes. Although the power required for the present invention is essentially the same as that for the prior artFIG. 3 process, the present invention improves both the ethane recovery and the propane recovery by 0.2% compared to theFIG. 3 process without using more power. - Like the
FIGS. 1 , 2, and 3 prior art processes, the present invention uses the expanded substantiallycondensed feed stream 35 c supplied to absorbingsection 20 a ofdemethanizer 20 to provide bulk recovery of the C2 components, C3 components, and heavier hydrocarbon components contained in expandedfeed 36 a and the vapors rising from strippingsection 20 b, and the supplemental rectification provided byreflux stream 44 a to reduce the amount of C2 components, C3 components, and C4+ components contained in the inlet feed gas that is lost to the residue gas. However, the present invention improves the rectification in absorbingsection 20 a over that of the prior art processes by making more effective use of the refrigeration available in process streams 38 and 35 b to improve the recoveries and the recovery efficiency. - Comparing
reflux stream 44 in Table I for theFIG. 1 prior art process with that in Table IV for the present invention, it can be seen that although the compositions of the streams are similar, theFIG. 1 process has over 3 times as much supplemental reflux as the present invention. Surprisingly, however, theFIG. 1 process achieves much lower ethane recovery than the present invention despite the much greater quantity of reflux. The better recovery achieved by the present invention can be understood by comparing the condition of the warmed expanded substantially condensedstream 35 c in theFIG. 1 prior art process with that of the corresponding stream in theFIG. 4 embodiment of the present invention. Although the temperature of this stream is only slightly warmer in theFIG. 1 process, the proportion of this stream that has been vaporized before enteringdemethanizer 20 is vastly higher than that of the present invention (42% versus 12%). This means that not only is there less cold liquid instream 35 c of theFIG. 1 process available for rectification of the vapors rising in absorbingsection 20 a, there is much more vapor in the upper region of absorbingsection 20 a that must be rectified byreflux stream 44 a. The net result is thatreflux stream 44 a of theFIG. 1 process allows more of the C2 components to escape to demethanizeroverhead stream 38 than the present invention does, reducing both the recovery and the recovery efficiency of theFIG. 1 process compared to the present invention. The key improvement of the present invention over theFIG. 1 prior art process is that the cold demethanizeroverhead vapor stream 38 is used to provide a portion of the cooling ofdistillation vapor stream 42 inheat exchanger 22 so that sufficient methane can be condensed for use as reflux, without adding significant rectification load in absorbingsection 20 a due to the excessive vaporization ofstream 35 c that is inherent in theFIG. 1 prior art process. - Comparing
reflux stream 44 in Tables II and III for theFIGS. 2 and 3 prior art processes with that in Table IV for the present invention, it can be seen that the present invention produces both more reflux and a better reflux stream than these prior art processes. Not only is the quantity of reflux higher (10% higher than theFIG. 2 process and 34% higher than theFIG. 3 process), the concentration of C2+ components is significantly lower (12.6% for the present invention, versus 19.6% for theFIG. 2 process and 16.9% for theFIG. 3 process). This makesreflux stream 44 a of the present invention more effective for rectification in absorbingsection 20 a ofdemethanizer 20, improving both the recovery and the recovery efficiency of the present invention compared to theFIGS. 2 and 3 prior art processes. The key improvement of the present invention over theFIGS. 2 and 3 prior art processes is that the expanded substantially condensedstream 35 b (which is predominantly liquid methane) is a better refrigerant medium than demethanizer overhead vapor stream 38 (which is primarily methane vapor), so usingstream 35 b to provide a portion of the cooling ofdistillation vapor stream 42 inheat exchanger 22 allows more methane to be condensed and used as reflux in the present invention. - In accordance with this invention, it is generally advantageous to design the absorbing (rectification) section of the demethanizer to contain multiple theoretical separation stages. However, the benefits of the present invention can be achieved with as few as two theoretical stages. For instance, all or a part of the pumped condensed liquid (
stream 44 a) fromreflux separator 23 and all or a part of the warmed expanded substantially condensedstream 35 c fromheat exchanger 22 can be combined (such as in the piping joining the pump and heat exchanger to the demethanizer) and if thoroughly intermingled, the vapors and liquids will mix together and separate in accordance with the relative volatilities of the various components of the total combined streams. Such comingling of the two streams, combined with contacting at least a portion of expandedstream 36 a, shall be considered for the purposes of this invention as constituting an absorbing section. -
FIGS. 5 through 8 display other embodiments of the present invention.FIGS. 4 through 6 depict fractionation towers constructed in a single vessel.FIGS. 7 and 8 depict fractionation towers constructed in two vessels, absorber (rectifier) column 27 (a contacting and separating device) and stripper (distillation)column 20. In such cases, a portion of the distillation vapor (stream 54) is withdrawn from the lower section ofabsorber column 27 and routed to refluxcondenser 22 to generate reflux forabsorber column 27. Theoverhead vapor stream 50 fromstripper column 20 flows to the lower section of absorber column 27 (via stream 51) to be contacted byreflux stream 52 and warmed expanded substantially condensedstream 35 c.Pump 28 is used to route the liquids (stream 47) from the bottom ofabsorber column 27 to the top ofstripper column 20 so that the two towers effectively function as one distillation system. The decision whether to construct the fractionation tower as a single vessel (such asdemethanizer 20 inFIGS. 4 through 6 ) or multiple vessels will depend on a number of factors such as plant size, the distance to fabrication facilities, etc. - Some circumstances may favor withdrawing the
distillation vapor stream 42 inFIGS. 5 and 6 from the upper region of strippingsection 20 b in demethanizer 20 (stream 55). In other cases, it may be advantageous to withdraw adistillation vapor stream 54 from the lower region of absorbingsection 20 a (above the feed point of expandedstream 36 a), withdraw adistillation vapor stream 55 from the upper region of strippingsection 20 b (below the feed point of expandedstream 36 a), combine streams 54 and 55 to form combineddistillation vapor stream 42, and direct combineddistillation vapor stream 42 toheat exchanger 22 to be cooled and partially condensed. Similarly, inFIGS. 7 and 8 a portion (stream 55) ofoverhead vapor stream 50 fromstripper column 20 may be directed to heat exchanger 22 (optionally combined withdistillation vapor stream 54 withdrawn from the lower section of absorber column 27), with the remaining portion (stream 51) flowing to the lower section ofabsorber column 27. - Some circumstances may favor mixing the remaining vapor portion (stream 43) of cooled
distillation vapor stream 42 a with the fractionation column overhead (stream 38), then supplying the mixed stream toheat exchanger 22 to provide a portion of the cooling ofdistillation vapor stream 42 or combineddistillation vapor stream 42. This is shown inFIGS. 6 and 8 , where themixed stream 45 resulting from combining the reflux separator vapor (stream 43) with the column overhead (stream 38) is routed toheat exchanger 22. - As described earlier, the
distillation vapor stream 42 or the combineddistillation vapor stream 42 is partially condensed and the resulting condensate used to absorb valuable C2 components, C3 components, and heavier components from the vapors rising through absorbingsection 20 a ofdemethanizer 20 or throughabsorber column 27. However, the present invention is not limited to this embodiment. It may be advantageous, for instance, to treat only a portion of these vapors in this manner, or to use only a portion of the condensate as an absorbent, in cases where other design considerations indicate portions of the vapors or the condensate should bypass absorbingsection 20 a ofdemethanizer 20 orabsorber column 27. Some circumstances may favor total condensation, rather than partial condensation, ofdistillation vapor stream 42 or combineddistillation vapor stream 42 inheat exchanger 22. Other circumstances may favor thatdistillation vapor stream 42 be a total vapor side draw fromfractionation column 20 orabsorber column 27 rather than a partial vapor side draw. It should also be noted that, depending on the composition of the feed gas stream, it may be advantageous to use external refrigeration to provide partial cooling ofdistillation vapor stream 42 or combineddistillation vapor stream 42 inheat exchanger 22. - Feed gas conditions, plant size, available equipment, or other factors may indicate that elimination of
work expansion machine 17, or replacement with an alternate expansion device (such as an expansion valve), is feasible. Although individual stream expansion is depicted in particular expansion devices, alternative expansion means may be employed where appropriate. For example, conditions may warrant work expansion of the substantially condensed portion of the feed stream (stream 35 a). - When the inlet gas is leaner,
separator 11 inFIG. 4 may not be justified. In such cases, the feed gas cooling accomplished inheat exchangers FIG. 4 may be accomplished without an intervening separator as shown inFIGS. 5 through 8 . The decision of whether or not to cool and separate the feed gas in multiple steps will depend on the richness of the feed gas, plant size, available equipment, etc. Depending on the quantity of heavier hydrocarbons in the feed gas and the feed gas pressure, the cooledfeed stream 31 a leavingheat exchanger 10 inFIGS. 4 through 8 and/or the cooledstream 32 a leavingheat exchanger 13 inFIG. 4 may not contain any liquid (because it is above its dewpoint, or because it is above its cricondenbar), so thatseparator 11 shown inFIGS. 4 through 8 and/orseparator 14 shown inFIG. 4 are not required. - The high pressure liquid (
stream 37 inFIG. 4 andstream 33 inFIGS. 5 through 8 ) need not be expanded and fed to a lower mid-column feed point on the distillation column. Instead, all or a portion of it may be combined with the portion of the separator vapor (stream 35 inFIG. 4 andstream 34 inFIGS. 5 through 8 ) flowing toheat exchanger 15. (This is shown by the dashedstream 46 inFIGS. 5 through 8 .) Any remaining portion of the liquid may be expanded through an appropriate expansion device, such as an expansion valve or expansion machine, and fed to a lower mid-column feed point on the distillation column (stream 37 a inFIGS. 5 through 8 ).Stream 33 inFIG. 4 andstream 37 inFIGS. 4 through 8 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. - In accordance with the present 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 a rich inlet gas. The use and distribution of separator liquids and demethanizer side draw liquids for process heat exchange, and the particular arrangement of heat exchangers for inlet gas cooling must be evaluated for each particular application, as well as the choice of process streams for specific heat exchange services.
- Some circumstances may favor using a portion of the cold distillation liquid leaving absorbing
section 20 a orabsorber column 27 for heat exchange, such as dashedstream 49 inFIGS. 5 through 8 . Although only a portion of the liquid from absorbingsection 20 a orabsorber column 27 can be used for process heat exchange without reducing the ethane recovery indemethanizer 20 orstripper column 20, more duty can sometimes be obtained from these liquids than with liquids from strippingsection 20 b orstripper column 20. This is because the liquids in absorbingsection 20 a of demethanizer 20 (or absorber column 27) are available at a colder temperature level than those in strippingsection 20 b (or stripper column 20). - As shown by dashed
stream 53 inFIGS. 5 through 8 , in some cases it may be advantageous to split the liquid stream from reflux pump 24 (stream 44 a) into at least two streams. A portion (stream 53) can then be supplied to the stripping section of fractionation tower 20 (FIGS. 5 and 6 ) or the top of stripper column 20 (FIGS. 7 and 8 ) to increase the liquid flow in that part of the distillation system and improve the rectification, thereby reducing the concentration of C2+ components instream 42. In such cases, the remaining portion (stream 52) is supplied to the top of absorbingsection 20 a (FIGS. 5 and 6 ) or absorber column 27 (FIGS. 7 and 8 ). - In accordance with the present invention, the splitting of the vapor feed may be accomplished in several ways. In the processes of
FIGS. 4 through 8 , the splitting of vapor occurs following cooling and separation of any liquids which may have been formed. The high pressure gas may be split, however, prior to any cooling of the inlet gas or after the cooling of the gas and prior to any separation stages. In some embodiments, vapor splitting may be effected in a separator. - It will also be recognized that the relative amount of feed found in each branch of the split vapor feed will depend on several factors, including gas pressure, feed gas composition, the amount of heat which can economically be extracted from the feed, and the quantity of horsepower available. More feed to the top of the column may increase recovery while decreasing power recovered from the expander thereby increasing the recompression horsepower requirements. Increasing feed lower in the column reduces the horsepower consumption but may also reduce product recovery. The 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.
- The present invention provides improved recovery of C2 components, C3 components, and heavier hydrocarbon components or of C3 components and heavier hydrocarbon components per amount of utility consumption required to operate the process. An improvement in utility consumption required for operating the demethanizer or deethanizer process may appear in the form of reduced power requirements for compression or re-compression, reduced power requirements for external refrigeration, reduced energy requirements for tower reboilers, or a combination thereof.
- While there have been described what are believed to be preferred embodiments of the invention, those skilled in the art will recognize that other and further modifications may be made thereto, e.g. to adapt the invention to various conditions, types of feed, or other requirements without departing from the spirit of the present invention as defined by the following claims.
Claims (60)
Priority Applications (60)
Application Number | Priority Date | Filing Date | Title |
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US12/868,993 US20110067441A1 (en) | 2009-09-21 | 2010-08-26 | Hydrocarbon Gas Processing |
PE2012000351A PE20121421A1 (en) | 2009-09-21 | 2010-08-27 | PROCESSING OF HYDROCARBON GASES |
CA2773211A CA2773211C (en) | 2009-09-21 | 2010-08-27 | Hydrocarbon gas processing |
BR112012006279A BR112012006279A2 (en) | 2009-09-21 | 2010-08-27 | hydrocarbon gas processing |
JP2012529780A JP5850838B2 (en) | 2009-09-21 | 2010-08-27 | Hydrocarbon gas treatment |
NZ599331A NZ599331A (en) | 2009-09-21 | 2010-08-27 | Hydrocarbon gas processing |
JP2012529781A JP5793145B2 (en) | 2009-09-21 | 2010-08-27 | Hydrocarbon gas treatment |
PCT/US2010/046966 WO2011034710A1 (en) | 2009-09-21 | 2010-08-27 | Hydrocarbon gas processing |
JP2012529779A JP5793144B2 (en) | 2009-09-21 | 2010-08-27 | Hydrocarbon gas treatment |
CN201080041904.9A CN102498360B (en) | 2009-09-21 | 2010-08-27 | Hydrocarbon gas processing |
CA2773157A CA2773157C (en) | 2009-09-21 | 2010-08-27 | Hydrocarbon gas processing |
PCT/US2010/046967 WO2011049672A1 (en) | 2009-09-21 | 2010-08-27 | Hydrocarbon gas processing |
EA201200521A EA028835B1 (en) | 2009-09-21 | 2010-08-27 | Hydrocarbon gas processing |
PE2012000352A PE20121420A1 (en) | 2009-09-21 | 2010-08-27 | PROCESSING OF HYDROCARBON GASES |
AU2010308519A AU2010308519B2 (en) | 2009-09-21 | 2010-08-27 | Hydrocarbon gas processing |
BR112012006277A BR112012006277A2 (en) | 2009-09-21 | 2010-08-27 | gaseous hydrocarbon processing |
MX2012002971A MX348674B (en) | 2009-09-21 | 2010-08-27 | Hydrocarbon gas processing. |
KR1020127009963A KR101619568B1 (en) | 2009-09-21 | 2010-08-27 | Hydrocarbon gas processing |
CN201080041508.6A CN102498359B (en) | 2009-09-21 | 2010-08-27 | Hydrocarbon gas processing |
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NZ599335A NZ599335A (en) | 2009-09-21 | 2010-08-27 | Hydrocarbon gas processing |
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CL2012000687A CL2012000687A1 (en) | 2009-09-21 | 2012-03-19 | Process and apparatus for separating a gas stream containing methane, c2, c3, and heavier hydrocarbons into a volatile off-gas fraction and a relatively less volatile fraction. |
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