Classifications

• • 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
  • 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
• • C—CHEMISTRY; METALLURGY
  • C07—ORGANIC CHEMISTRY
  • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
  • C07C7/00—Purification; Separation; Use of additives
  • C07C7/005—Processes comprising at least two steps in series
• • C—CHEMISTRY; METALLURGY
  • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
  • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
  • C10G5/00—Recovery of liquid hydrocarbon mixtures from gases, e.g. natural gas
  • C10G5/06—Recovery of liquid hydrocarbon mixtures from gases, e.g. natural gas by cooling or compressing
• • C—CHEMISTRY; METALLURGY
  • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
  • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
  • C10G70/00—Working-up undefined normally gaseous mixtures obtained by processes covered by groups C10G9/00, C10G11/00, C10G15/00, C10G47/00, C10G51/00
  • C10G70/04—Working-up undefined normally gaseous mixtures obtained by processes covered by groups C10G9/00, C10G11/00, C10G15/00, C10G47/00, C10G51/00 by physical processes
• • 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
  • 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
• • 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
  • 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
  • F25J3/0219—Refinery gas, cracking gas, coke oven gas, gaseous mixtures containing aliphatic unsaturated CnHm or gaseous mixtures of undefined nature
• • 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
  • 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
  • F25J3/0242—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 3 carbon atoms or more
• • 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
  • F25J2200/00—Processes or apparatus using separation by rectification
  • F25J2200/04—Processes or apparatus using separation by rectification in a dual pressure main column system
• • 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
  • F25J2200/00—Processes or apparatus using separation by rectification
  • F25J2200/70—Refluxing the column with a condensed part of the feed stream, i.e. fractionator top is stripped or self-rectified
• • 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
  • F25J2200/00—Processes or apparatus using separation by rectification
  • F25J2200/74—Refluxing the column with at least a part of the partially condensed overhead gas
• • 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
  • F25J2200/00—Processes or apparatus using separation by rectification
  • F25J2200/76—Refluxing the column with condensed overhead gas being cycled in a quasi-closed loop refrigeration cycle
• • 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
  • F25J2200/00—Processes or apparatus using separation by rectification
  • F25J2200/78—Refluxing the column with a liquid stream originating from an upstream or downstream fractionator column
• • 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
  • F25J2205/00—Processes or apparatus using other separation and/or other processing means
  • F25J2205/02—Processes or apparatus using other separation and/or other processing means using simple phase separation in a vessel or drum
• • 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
  • F25J2205/00—Processes or apparatus using other separation and/or other processing means
  • 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
• • 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
  • F25J2210/00—Processes characterised by the type or other details of the feed stream
  • F25J2210/12—Refinery or petrochemical off-gas
• • 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
  • F25J2230/00—Processes or apparatus involving steps for increasing the pressure of gaseous process streams
  • F25J2230/60—Processes or apparatus involving steps for increasing the pressure of gaseous process streams the fluid being hydrocarbons or a mixture of hydrocarbons
• • 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
  • F25J2235/00—Processes or apparatus involving steps for increasing the pressure or for conveying of liquid process streams
  • 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
• • 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
  • F25J2240/00—Processes or apparatus involving steps for expanding of process streams
  • F25J2240/02—Expansion of a process fluid in a work-extracting turbine (i.e. isentropic expansion), e.g. of the feed stream
• • 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
  • F25J2270/00—Refrigeration techniques used
  • F25J2270/02—Internal refrigeration with liquid vaporising loop
• • 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
  • F25J2270/00—Refrigeration techniques used
  • F25J2270/90—External refrigeration, e.g. conventional closed-loop mechanical refrigeration unit using Freon or NH3, unspecified external refrigeration

Definitions

• 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 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, 92.6% methane, 4.7% ethane and other C 2 components, 1.0% propane and other C 3 components, 0.2% iso-butane, 0.2% normal butane, 0.1% pentanes plus, with the balance made up of nitrogen and carbon dioxide. Sulfur containing gases are also sometimes present.
• cryogenic expansion process is now generally preferred for propylene and propane recovery because it provides maximum simplicity with ease of start up, operating flexibility, good efficiency, safety, and good reliability.
• 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 3 + 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 (deethanizer) column.
• the expansion cooled stream(s) is (are) distilled 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 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 slightly below the pressure at which the distillation column is operated.
• the expanded stream then enters the lower section of an absorption column and is contacted with cold liquids to absorb the C 3 components and heavier components from the vapor portion of the expanded stream.
• the liquids from the absorption column are then pumped into the deethanizer column at an upper column feed position, perhaps after heating to partially vaporize the stream.
• the overhead distillation stream from the deethanizer passes in heat exchange relation with the residue gas from the absorber column and is cooled, condensing at least a portion of the distillation stream from the deethanizer.
• the cooled distillation stream then enters the upper section of the absorption column where the cold liquids contained in the stream can contact the vapor portion of the expanded stream as described earlier.
• the vapor portion (if any) of the cooled distillation stream and the absorber overhead vapor combine in an upper separator section in the absorber column as residual methane and C 2 component product gas.
• the cooled distillation stream may be supplied to a separator to provide vapor and liquid streams.
• the vapor is combined with the absorber column overhead and the liquid is supplied to the absorber column as a top column feed. It may also be advantageous to supply a portion of the cold liquid condensate to the deethanizer tower to serve as reflux.
• the separation that takes place in this process consumes energy for feed gas cooling, for reboiling the deethanizer, for refluxing the deethanizer, and/or for re-compressing the residue gas.
• the present invention provides a means for achieving this separation while substantially reducing the utility requirements (cooling, reboiling, refluxing, and/or re-compressing) needed for the recovery of the desired products.
• FIG. 1 is a flow diagram of a prior art cryogenic natural gas processing plant in accordance with United States Patent No. 5,771,712;
• FIG. 2 is a flow diagram of a prior art cryogenic natural gas processing plant of an alternative system in accordance with United States Patent No. 5,771,712;
• FIG. 3 is a flow diagram of a natural gas processing plant in accordance with the present invention.
• FIG. 4 is a flow diagram illustrating an alternative means of application of the present invention to a natural gas stream
• FIG. 5 is a flow diagram illustrating an alternative means of application of the present invention to a natural gas stream
• FIG. 6 is a flow diagram illustrating an alternative means of application of the present invention to a natural gas stream
• FIG. 7 is a flow diagram illustrating an alternative means of application of the present invention to a natural gas stream
• FIG. 8 is a flow diagram illustrating an alternative means of application of the present invention to a natural gas stream
• FIG. 9 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream
• FIG. 10 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream
• FIG. 11 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream
• FIG. 12 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream
• FIG. 13 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream
• FIG. 14 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream
• FIG. 15 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream
• FIG. 16 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream
• FIG. 17 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream
• FIG. 18 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream.
• inlet gas enters the plant at 80°F [27°C] and 1215 psia [8,377 kPa(a)] as stream 31. If the inlet gas contains a concentration of sulfur compounds which would prevent the product streams from meeting specifications, the sulfur compounds are removed by appropriate pretreatment of the feed gas (not illustrated). In addition, the feed stream is usually dehydrated to prevent hydrate (ice) formation under cryogenic conditions. Solid desiccant has typically been used for this purpose.
• the feed stream 31 is cooled in exchanger 10 by heat exchange with cool residue gas at -76°F [-60°C] (stream 34a).
• stream 34a 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 feed stream pressure is above the cricondenbar, so no liquid will condense as the stream is cooled.
• the cooled stream 31a (a dense-phase fluid at these conditions) is supplied directly to work expansion machine 13 at -14°F [-26°C].
• the cricondenbar is the maximum pressure at which a vapor phase can exist in a multi-phase fluid. At pressures below the cricondenbar, a separator or scrubber would typically be used to separate any condensed liquid contained in stream 31a from the vapor so that only the vapor is supplied to work expansion machine 13.)
• the work expansion machine 13 extracts mechanical energy from the high pressure feed by expanding the stream substantially isentropically from a pressure of about 1210 psia [8,343 kPa(a)] to apressure of about 435 psia [2,999 kPa(a)] (the operating pressure of separator/absorber tower 15), with the work expansion cooling the expanded stream 31b to a temperature of approximately -104°F [-76°C].
• the expanded and partially condensed stream 31b is supplied to absorbing section 15b in a lower region of separator/absorber tower 15.
• the liquid portion of the expanded stream commingles with liquids falling downward from the absorbing section and the combined liquid stream 35 exits the bottom of separator/absorber tower 15 at -106°F [-77°C].
• the vapor portion of the expanded stream rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C 3 components and heavier components.
• the separator/absorber tower 15 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. As is often the case in natural gas processing plants, the separator/absorber tower may consist of two sections.
• the upper section 15a is a separator wherein any vapor contained in the top feed is separated from its corresponding liquid portion, and wherein the vapor rising from the lower distillation or absorbing section 15b is combined with the vapor portion (if any) of the top feed to form the cold distillation stream 34 which exits the top of the tower.
• the lower, absorbing section 15b contains the trays and/or packing and provides the necessary contact between the liquids falling downward and the vapors rising upward to condense and absorb the C 3 components and heavier components.
• the combined liquid stream 35 from the bottom of the separator/absorber tower 15 is supplied to deethanizer 17 by pump 16, entering at a mid-column feed point at -105°F [-76°C] as stream 35a to be stripped of its methane and C 2 components.
• the deethanizer in tower 17, operating at about 450 psia [3,103 kPa(a)], is also a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing.
• the deethanizer tower may also consist of two sections: an upper section 17a wherein any vapor contained in the top feed is separated from its corresponding liquid portion, and wherein the vapor rising from the lower distillation or deethanizing section 17b is combined with the vapor portion (if any) of the top feed to form distillation stream 36 which exits the top of the tower; and a lower, deethanizing section 17b that contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward.
• the deethanizing section 17b also includes a reboiler 18 which heats and vaporizes a portion of the liquid at the bottom of the column to provide the stripping vapors which flow up the column to strip the liquid product, stream 37, of methane and C 2 components.
• a typical specification for the bottom liquid product is to have an ethane to propane ratio of 0.02: 1 on a molar basis.
• the liquid product stream 37 exits the bottom of the deethanizer at 207°F [97°C] and is cooled to 110°F [43°C] (sfream 37a) in heat exchanger 19 before flowing to storage.
• deethanizer 17 The operating pressure in deethanizer 17 is maintained slightly above the operating pressure of separator/absorber tower 15. This allows the deethanizer overhead vapor (stream 36) to pressure flow through heat exchanger 20 and thence into the upper section of separator/absorber tower 15.
• the deethanizer overhead at -36°F [-38°C] is directed in heat exchange relation with the overhead (stream 34) from separator/absorber tower 15, cooling the stream to -107°F [-77°C] (stream 36a) and partially condensing it.
• the partially condensed stream is then supplied to the separator section in separator/absorber tower 15 where the condensed liquid is separated from the uncondensed vapor.
• the uncondensed vapor combines with the vapor rising from the lower absorbing section to form the cold distillation stream 34 leaving the upper region of separator/absorber tower 15.
• the condensed liquid is divided into two portions. One portion, stream 39, is routed to the lower absorbing section of separator/absorber tower 15 as the cold liquid that contacts the vapors rising upward through the absorbing section.
• the other portion, stream 38, is supplied to deethanizer 17 as reflux by pump 21, with reflux stream 38a flowing to a top feed point on deethanizer 17 at -107°F [-77°C].
• the distillation stream leaving the top of separator/absorber tower 15 at -112°F [-80°C] is the cold residue gas stream 34.
• the residue gas stream passes countercurrently to deethanizer overhead stream 36 in heat exchanger 20 and is warmed to -76 °F [-60°C] (stream 34a) as it provides cooling and partial condensation of the deethanizer overhead stream.
• the residue gas is further warmed to 54 °F [12°C] (stream 34b) as it passes countercurrently to the incoming feed stream in heat exchanger 10.
• the residue gas is then re-compressed in two stages.
• the first stage is compressor 14 driven by expansion machine 13.
• the second stage is compressor 22 driven by a supplemental power source which compresses the residue gas (sfream 34d) to sales line pressure.
• the residue gas product (stream 34e) flows to the sales gas pipeline at 110°F [43°C] and 1215 psia [8,377 kPa(a)].
• deethanizer 17 since the operating pressure of deethanizer 17 must of necessity be maintained somewhat higher than that of separator/absorber 15 so that its overhead stream 36 can pressure flow through heat exchanger 20 and into the separator section of separator/absorber 15, reducing the expansion ratio across work expansion machine 13 also means raising the operating pressure of deethanizer 17. Unfortunately, this is not advisable in this instance because of the detrimental effect on distillation performance in deethanizer 17 that would result from the higher operating pressure. This effect is manifested by poor mass transfer in deethanizer 17 due to the phase behavior of its vapor and liquid streams. Of particular concern are the physical properties that affect the vapor-liquid separation efficiency, namely the liquid surface tension and the differential in the densities of the two phases.
• FIG. 2 represents an alternative application of the prior art process in accordance with U.S. Pat. No.5,771,712.
• the process of FIG. 2 has been applied to the same feed gas composition as described above for FIG. 1, but in this simulation of the process the inlet gas (stream 31) enters the plant at 80°F [27°C] and 580 psia [3,999 kPa(a)].
• the feed stream 31 is cooled in exchanger 10 by heat exchange with cool residue gas at -95°F [-71 °C] (stream 34a), with separator liquids at -92°F [-69°C] (stream 33a), and with separator/absorber liquids at -107°F [-77°C] (stream 35a).
• the feed stream is below the cricondenbar, so the cooled stream 31a enters separator 11 at -77°F [-60°C] and 570 psia [3,930 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (sfream 33).
• the vapor (sfream 32) from separator 11 enters work expansion machine 13 in which mechanical energy is extracted from this portion of the high pressure feed.
• the machine 13 expands the vapor substantially isentropically from a pressure of about 570 psia [3,930 kPa(a)] to a pressure of about 380 psia [2,620 kPa(a)] (the operating pressure of separator/absorber 15), with the work expansion cooling the expanded sfream 32a to a temperature of approximately -107°F [-77°C].
• the expanded and partially condensed stream 32a enters the lower section r of separator/absorber 15.
• the liquid portion of the expanded sfream commingles with liquids falling downward from the absorbing section and the combined liquid stream 35 exits the bottom of separator/absorber 15 at -108°F [-78°C].
• the vapor portion of the expanded stream rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C 3 components and heavier components.
• the combined liquid stream 35 from the bottom of the separator/absorber 15 is routed to heat exchanger 10 by pump 16 where it (sfream 35a) is heated as it provides cooling of the incoming feed gas as described earlier.
• the combined liquid stream is heated to -85°F [-65°C], partially vaporizing stream
• stream 33 is flash expanded to slightly above the 395 psia [2,723 kPa(a)] operating pressure of deethanizer 17 by expansion valve 12, cooling stream 33 to -92°F [-69°C] (stream 33 a) before it provides cooling to the incoming feed gas as described earlier.
• Stream 33b, now at 65°F [18°C] then enters deethanizer 17 at a lower mid-column feed point.
• streams 35b and 33b are stripped of their methane and C 2 components.
• the resulting liquid product sfream 37 exits the bottom of the deethanizer at 195°F [9PC] and is cooled to 110°F [43°C] (sfream 37a) in heat exchanger 19 before flowing to storage.
• the operating pressure in deethanizer 17 is maintained slightly above the operating pressure of separator/absorber 15. This allows the deethanizer overhead vapor (sfream 36) to pressure flow through heat exchanger 20 and thence into the upper section of separator/absorber 15.
• the deethanizer overhead at -29°F [-34°C] is directed in heat exchange relation with the overhead (stream 34) from separator/absorber 15, cooling the stream to -108°F [-78°C] (stream 36a) and partially condensing it.
• the partially condensed stream is then supplied to the separator section in separator/absorber tower 15 where the condensed liquid is separated from the uncondensed vapor.
• the uncondensed vapor combines with the vapor rising from the lower absorbing section to form the cold distillation stream 34 leaving the upper region of separator/absorber 15.
• the condensed liquid is divided into two portions.
• sfream 39 is routed to the lower absorbing section of separator/absorber 15 as the cold liquid that contacts the vapors rising upward through the absorbing section.
• the other portion, sfream 38 is supplied to deethanizer 17 as reflux by pump 21, with reflux sfream 38a flowing to a top feed point on deethanizer 17 at -108°F [-78°C].
• -113°F [-81°C] is the cold residue gas sfream 34.
• the residue gas stream passes countercurrently to deethanizer overhead sfream 36 in heat exchanger 20 and is warmed to -95°F [-71°C] (stream 34a) as it provides cooling and partial condensation of the deethanizer overhead stream.
• the residue gas is further warmed to 75°F [24°C] (stream 34b) as it passes countercurrently to the incoming feed gas in heat exchanger 10.
• the residue gas is then re-compressed in two stages.
• the first stage is compressor 14 driven by expansion machine 13.
• the second stage is compressor 22 driven by a supplemental power source which compresses the residue gas (stream 34d) to sales line pressure.
• the residue gas product (sfream 34e) flows to the sales gas pipeline at 110°F [43 °C] and 613 psia [4,226 kPa(a)].
• FIG. 3 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. 3 are the same as those in FIG. 1. Accordingly, the FIG. 3 process can be compared with that of the FIG. 1 process to illusfrate the advantages of the present invention.
• inlet gas enters the plant at 80°F [27°C] and 1215 psia [8,377 kPa(a)] as sfream 31.
• the feed sfream 31 is cooled in exchanger 10 by heat exchange with cool residue gas at -56°F [-49°C] (stream 34a) and with separator/absorber liquids at -113°F [-80°C] (stream 35a).
• the cooled stream 31a (a dense-phase fluid at these conditions) is supplied directly to work expansion machine 13 at -35°F [-37°C].
• the work expansion machine 13 extracts mechanical energy from the high pressure feed by expanding the stream substantially isentropically from a pressure of about 1210 psia [8,343 kPa(a)] to a pressure of about 575 psia
• the operating pressure of deethanizer 17 is maintained below (not above) the operating pressure of separator/absorber 15. Consequently, a pump is not required for the combined liquid stream 35 from the bottom of the separator/absorber 15. Instead, the stream is flash expanded to slightly above the 450 psia [3,103 kPa(a)] operating pressure of deethanizer 17 by expansion valve 27, cooling stream 35 to -113°F [-80°C] (sfream 35a) before it provides cooling to the incoming feed gas as described earlier. Stream 35b, now at -73°F [-58°C], then enters deethanizer 17 at a mid-column feed point.
• stream 35b is stripped of its methane and C 2 components.
• the resulting liquid product sfream 37 exits the bottom of the deethanizer at 207°F [97°C] and is cooled to 110°F [43°C] (stream 37a) in heat exchanger 19 before flowing to storage.
• the deethamzer overhead vapor (stream 36) exits deethanizer 17 at -56°F [-49°C] and is warmed to 105°F [41°C] (sfream 36a) in heat exchanger 24 before entering compressor 25 (driven by a supplemental power source).
• Sfream 36b leaves compressor 25 at 600 psia [4,137 kPa(a)] and is cooled to 110°F [43 °C] (stream 36c) in heat exchanger 26.
• Sfream 36c is then directed in heat exchange relation with the deethanizer overhead vapor (sfream 36) in heat exchanger 24 to cool it (stream 36d) and conserve process cooling.
• stream 36d can now pressure flow through heat exchanger 20 and thence to the upper feed point of separator/absorber 15.
• the compressed deethanizer overhead at -41 °F [-40°C] is directed in heat exchange relation with the overhead (stream 34) from separator/absorber 15, cooling the sfream to -98°F [-72°C] (stream 36e) and partially condensing it.
• the partially condensed sfream is then supplied to the separator section in separator/absorber tower 15 where the condensed liquid is separated from the uncondensed vapor.
• the uncondensed vapor combines with the vapor rising from the lower absorbing section to form the cold distillation stream 34 leaving the upper region of separator/absorber 15.
• the condensed liquid is divided into two portions. One portion, sfream 39, is routed to the lower absorbing section of separator/absorber 15 as the cold liquid that contacts the vapors rising upward through the absorbing section.
• the other portion, sfream 38, is flash expanded to slightly above the operating pressure of deethanizer 17 by expansion valve 28 and the resulting sfream 38a is then supplied at -112°F [-80°C] to the separator section in deethanizer 17 where its condensed liquid is separated from its uncondensed vapor.
• the uncondensed vapor combines with the vapor rising from the lower distillation section to form the deethanizer overhead stream 36 leaving the upper region of deethanizer 17, while the condensed liquid is routed to the lower distillation section of deethanizer 17 as reflux for the vapors rising upward through the distillation section.
• the distillation sfream leaving the top of separator/absorber 15 at -103°F [-75°C] is the cold residue gas stream 34.
• the residue gas sfream passes countercurrently to compressed deethanizer overhead stream 36d in heat exchanger 20 and is warmed to -56°F [-49°C] (stream 34a) as it provides cooling and partial condensation of the compressed deethanizer overhead stream.
• the residue gas is further warmed to 75 °F [24°C] (stream 34b) as it passes countercurrently to the incoming feed gas in heat exchanger 10.
• the residue gas is then re-compressed in two stages.
• the first stage is compressor 14 driven by expansion machine 13.
• the second stage is compressor 22 driven by a supplemental power source which compresses the residue gas (stream 34d) to sales line pressure.
• the residue gas product (sfream 34e) flows to the sales gas pipeline at 110°F [43°C] and 1215 psia [8,377 kPa(a)].
• a summary of sfream flow rates and energy consumptions for the process illustrated in FIG. 3 is set forth in the table below: TABLE ⁇ i
• FIG. 3 (FIG. 3)
• separator/absorber 15 operating at a higher pressure than deethanizer 17, it is no longer necessary to pump the absorber bottom liquid (stream 35) and the reflux stream (sfream 38) to feed deethanizer 17, eliminating the capital and operating cost of pumps 16 and 21 in the FIG. 1 process.
• work expansion machine 13 and compressors 14 and 22 represent an open cycle mechanical-compression refrigeration loop that provides the process cooling in the prior art process of FIG. 1, with a working fluid (streams 31 and 34) that is predominantly methane.
• the refrigeration provided by this cycle has been reduced by the addition of a second open cycle refrigeration loop powered by compressor 25.
• Table III shows that the working fluid for this second cycle (the deethanizer overhead, sfream 36) has a substantially lower concenfration of methane and a substantially higher concenfration of ethane than the working fluid in the first cycle.
• the efficiency of mechanical-compression refrigeration cycles improves as the molecular weight of the working fluid increases.
• the operating pressures of separator/absorber 15 and deethanizer 17 are no longer coupled together as they are in the prior art process. Instead, the operating pressures of the two columns can be optimized independently. In the case of deethanizer 17, the pressure can be selected to insure good distillation characteristics, while for separator/absorber 15 the pressure can be selected to optimize the process cooling versus the residue gas compression requirements.
• FIG. 3 represents the preferred embodiment of the present invention for the temperature and pressure conditions shown because it typically provides the simplest plant arrangement for a given C 3 component recovery level.
• a slightly more complex design that maintains the same C 3 component recovery with lower utility consumption can be achieved using another embodiment of the present invention as illustrated in the FIG. 4 process.
• the feed gas composition and conditions considered in the process presented in FIG. 4 are the same as those in FIGS. 1 and 3. Accordingly, FIG. 4 can be compared with the FIG. 1 process to illustrate the advantages of the present invention, and can likewise be compared to the embodiment displayed in FIG. 3. In the simulation of the FIG. 4 process, the feed gas cooling and expansion scheme is much the same as that used in FIG. 3.
• the deethanizer overhead vapor (stream 36) exits deethanizer 17 at -39°F [-39°C] and is warmed to 105°F [41°C] (stream 36a) in heat exchanger 24 before entering compressor 25 (driven by a supplemental power source).
• Stream 36b leaves compressor 25 at 600 psia [4,137 kPa(a)] and is cooled to 110°F [43°C] (sfream 36c) in heat exchanger 26.
• Stream 36c is then directed in heat exchange relation with the deethanizer overhead vapor (stream 36) in heat exchanger 24 to cool it to -24°F [-31°C] (sfream 36d) and conserve process cooling.
• the compressed deethanizer overhead (sfream 36d) is directed in heat exchange relation with the overhead (stream 34) from separator/absorber 15, cooling the sfream to -50°F [-46°C] (sfream 36e) and partially condensing it before it is withdrawn.
• the partially condensed stream 36e enters separator 30 where the condensed liquid is separated from the uncondensed vapor.
• the condensed liquid (stream 38) from separator 30 is flash expanded to slightly above the operating pressure of deethanizer 17 by expansion valve 28 (stream 38a), which partially vaporizes the sfream and cools it further to -63 °F [-53 °C]. It is then supplied to the separator section in deethanizer 17 where the liquid is separated from the flash vapor.
• the flash vapor combines with the vapor rising from the lower distillation section to form the deethanizer overhead sfream 36 leaving the upper region of deethanizer 17, while the liquid is routed to the lower distillation section of deethanizer 17 as reflux for the vapors rising upward through the distillation section.
• the uncondensed vapor (sfream 39) from separator 30 is routed back to heat exchanger 20 to also direct it in heat exchange relation with the overhead (sfream 34) from separator/absorber 15, cooling the sfream to -98°F [-72°C] (stream 39a) and partially condensing it.
• the partially condensed stream is then supplied to the separator section in separator/absorber tower 15 where its condensed liquid is separated from its uncondensed vapor.
• the uncondensed vapor combines with the vapor rising from the lower absorbing section to form the cold distillation sfream 34 leaving the upper region of separator/absorber 15, while the condensed liquid is routed to the lower absorbing section of separator/absorber 15 as the cold liquid that contacts the vapors rising upward through the absorbing section.
• FIG. 4 (FIG. 4)
• FIG. 4 embodiment of the present invention requires slightly less compression horsepower (about 2 percent) than the FIG. 3 embodiment, but uses slightly more utility heat for the deethanizer reboiler (less than 1 percent), with the total utility requirements being about 1 percent lower for the FIG. 4 embodiment.
• the improvement in efficiency can be understood by comparing the reflux stream for deethanizer 17 (stream 38) in the FIG. 4 embodiment of the present invention with the corresponding sfream in the FIG. 3 embodiment. Whereas stream 38 in FIG. 3 is predominantly methane, stream 38 in FIG.
• FIG. 4 process requires will generally depend on factors which include plant size and available equipment, as well as the relative costs of compression horsepower and utility heat.
• Example 3 A third embodiment of the present invention is shown in FIG. 5, wherein a different method of implementing the second mechanical-compression refrigeration cycle is applied to the present invention.
• the feed gas composition and conditions considered in the process illustrated in FIG. 5 are the same as those in FIGS. 1, 3, and 4. Accordingly, FIG. 5 can be compared with the FIG. 1 process to illustrate the advantages of the present invention, and can likewise be compared to the embodiments displayed in FIGS. 3 and 4.
• inlet gas enters the plant at 80°F [27°C] and 1215 psia [8,377 kPa(a)] as stream 31.
• the feed sfream 31 is cooled in exchanger 10 by heat exchange with cool residue gas at -70°F [-57°C] (sfream 34a), with cool vapor at -49°F [-45°C] (stream 41a), and with separator/absorber liquids at -112°F [-80°C] (stream 35a).
• the cooled stream 31a (a dense-phase fluid at these conditions) is supplied directly to work expansion machine 13 at -32°F [-36°C].
• the work expansion machine 13 extracts mechanical energy from the high pressure feed by expanding the sfream substantially isentropically from a pressure of about 1210 psia [8,343 kPa(a)] to apressure of about 515 psia
• the deethanizer overhead vapor (stream 36) exits deethanizer 17 at -44°F [-42°C] and flows through heat exchanger 20.
• the deethanizer overhead is directed in heat exchange relation with the overhead (stream 34) from separator/absorber 15 and the uncondensed vapor (stream 41) from separator 30, cooling the sfream to -102°F [-74°C] (sfream 36a) and partially condensing it.
• the partially condensed sfream is then supplied to separator 30 where the condensed liquid (stream 40) is separated from the uncondensed vapor (stream 41).
• the liquid condensed from the deethanizer overhead (sfream 40) is at a lower pressure than the two columns
• sfream 40a is then divided into two portions. One portion, stream 39, is supplied by control valve 29 to the separator section in separator/absorber tower 15 at -100°F [-74°C] (stream 39a) where its liquid is separated from any vapor that forms.
• any vapor that may form combines with the vapor rising from the lower absorbing section to form the cold distillation stream 34 leaving the upper region of separator/absorber 15, while the condensed liquid is routed to the lower absorbing section of separator/absorber 15 as the cold liquid that contacts the vapors rising upward through the absorbing section.
• the other portion of the pumped liquid (stream 38) is flash expanded to slightly above the operating pressure of deethanizer 17 by expansion valve 28 (stream 38a). It is then supplied at -101°F [-74°C] to the separator section in deethanizer 17 where its liquid is separated from any flash vapor that forms.
• Any flash vapor combines with the vapor rising from the lower distillation section to form the deethamzer overhead sfream 36 leaving the upper region of deethanizer 17, while the condensed liquid is routed to the lower distillation section of deethanizer 17 as reflux for the vapors rising upward through the distillation section.
• the distillation sfream leaving the top of separator/absorber 15 at -107°F [-77°C] is the cold absorber overhead stream 34.
• the absorber overhead sfream passes countercurrently to deethanizer overhead stream 36 in heat exchanger 20 and is warmed to -70°F [-57°C] (sfream 34a) as it provides cooling and partial condensation of the deethanizer overhead stream.
• the absorber overhead stream is further warmed to 75°F [24°C] (stream 34b) as it passes countercurrently to the incoming feed gas in heat exchanger 10.
• the uncondensed vapor (stream 41) leaves separator 30 at -102°F [-74°C] and also passes countercurrently to deethanizer overhead stream 36 in heat exchanger 20 and is warmed to -49°F [-45°C] (stream 41a) as it too provides cooling and partial condensation of the deethanizer overhead stream.
• the vapor stream is further warmed to 65°F [18°C] (sfream 41b) as it passes countercurrently to the incoming feed gas in heat exchanger 10.
• the first stage for the absorber overhead stream is compressor 14 driven by expansion machine 13, while the first stage for the vapor sfream is compressor 25 driven by a supplemental power source.
• the two partially compressed streams combine to form the residue gas, sfream 42.
• the combined residue gas stream then enters compressor 22 driven by a supplemental power source, which provides the second stage of compression to raise the residue gas (sfream 42a) to sales line pressure.
• the residue gas product (stream 42b) flows to the sales gas pipeline at 110°F [43°C] and 1215 psia [8,377 kPa(a)].
• FIG. 5 (FIG. 5)
• FIG. 5 embodiment requires slightly more compression horsepower and utility heating than either the FIG. 3 or the FIG. 4 embodiment.
• multiple stage compression or multi-wheel centrifugal compression is used to compress the residue gas sfream 42, it may be possible to compress the vapor sfream 41b using an intermediate stage or wheel, eliminating the need for a separate compressor like compressor 25.
• factors such as plant size and available equipment will determine whether the FIG. 5 embodiment would be preferable for a specific circumstance.
• Example 4 A slightly more complex design than the FIG. 5 embodiment that maintains the same C 3 component recovery with lower utility consumption can be achieved using another embodiment of the present invention as illustrated in the FIG. 6 process.
• the feed gas composition and conditions considered in the process presented in FIG. 6 are the same as those in FIGS. 1 and 5. Accordingly, FIG. 6 can be compared with the FIG. 1 process to illustrate the advantages of the present invention, and can likewise be compared to the embodiment displayed in FIG. 5.
• the feed gas cooling and expansion scheme is much the same as that used in FIG. 5. The difference lies in the manner in which the vapor distillation stream 36 leaving the overhead of deethanizer 17 is used to generate reflux for deethanizer 17 and separator/absorber 15. Referring to FIG.
• the deethanizer overhead vapor (stream 36) exits deethamzer 17 at -39°F [-40°C] and flows through heat exchanger 20.
• the deethanizer overhead is directed in heat exchange relation with the overhead (stream 34) from separator/absorber 15 and the uncondensed vapor (sfream 41) from separator 30, cooling the sfream to -60°F [-51°C] (stream 36a) and partially condensing it before it is withdrawn.
• the partially condensed stream 36a enters separator 29 where the condensed liquid (sfream 38) is separated from the uncondensed vapor (sfream 40).
• the liquid condensed from the deethanizer overhead (stream 38) is at a lower pressure than deethanizer 17, so it is pumped by pump 28 so that it can be used as reflux.
• stream 38a is supplied at -60°F [-51°C] to the separator section in deethanizer 17 where the liquid is routed to the lower distillation section of deethanizer 17 as reflux for the vapors rising upward through the distillation section.
• the uncondensed vapor (stream 40) from separator 29 is routed back to heat exchanger 20 to also direct it in heat exchange relation with the overhead (stream 34) from separator/absorber 15 and the uncondensed vapor (sfream 41) from separator 30, cooling the stream to -102°F [-74°C] (stream 40a) and partially condensing it.
• the partially condensed sfream is then supplied to separator 30 where the condensed liquid (stream 39) is separated from the uncondensed vapor (stream 41).
• pump 21 is used to direct the condensed liquid (stream 39a) at -100°F [-73°C] to the separator section in separator/absorber tower 15, where the condensed liquid is routed to the lower absorbing section of separator/absorber 15 as the cold liquid that contacts the vapors rising upward through the absorbing section.
• the distillation stream leaving the top of separator/absorber 15 at -107°F [-77°C] is the cold absorber overhead sfream 34.
• the absorber overhead stream passes countercurrently to deethanizer overhead stream 36 and vapor stream 40 in heat exchanger 20 and is warmed to -74°F [-59°C] (stream 34a) as it provides cooling and partial condensation of the deethanizer overhead stream and the vapor stream.
• the absorber overhead sfream is further warmed to 75°F [24°C] (stream 34b) as it passes countercurrently to the incoming feed gas in heat exchanger 10.
• the uncondensed vapor (stream 41) leaves separator 30 at -102°F [-74°C] and also passes countercurrently to deethanizer overhead stream 36 and vapor sfream 40 in heat exchanger 20 and is warmed to -44°F [-42°C] (stream 41a) as it too provides cooling and partial condensation of the sfreams.
• the vapor sfream is further warmed to 65°F [18°C] (stream 41b) as it passes countercurrently to the incoming feed gas in heat exchanger 10.
• the warm absorber overhead stream 34b and the warm vapor stream 41b are then re-compressed in two stages.
• the first stage for the absorber overhead stream is compressor 14 driven by expansion machine 13, while the first stage for the vapor stream is compressor 25 driven by a supplemental power source.
• the two partially compressed streams (sfreams 34c and 41c, respectively) combine to form the residue gas, stream 42.
• the combined residue gas sfream then enters compressor 22 driven by a supplemental power source, which provides the second stage of compression to raise the residue gas (stream 42a) to sales line pressure.
• the residue gas product (stream 42b) flows to the sales gas pipeline at 110°F [43°C] and 1215 psia [8,377 kPa(a)].
• FIG. 6 (FIG. 6)
• FIG. 6 embodiment of the present invention requires slightly less compression horsepower (about 2 percent) than the FIG. 5 embodiment, and uses slightly less utility heat for the deethanizer reboiler (about 4 percent), with the total utility requirements being about 3 percent lower for the FIG. 6 embodiment.
• the improvement in efficiency can be understood by comparing the reflux stream for deethanizer 17 (sfream 38) in the FIG. 6 embodiment of the present invention with the corresponding stream in the FIG. 5 embodiment. Whereas sfream 38 in FIG. 5 is predominantly methane, stream 38 in FIG.
• FIG. 6 process requires will generally depend on factors which include plant size and available equipment, as well as the relative costs of compression horsepower and utility heat.
• Example 5 FIG. 7 illustrates a flow diagram of a process in accordance with the present invention when applied to the feed gas composition and conditions considered in the process presented in FIG. 2. Accordingly, the FIG. 7 process can be compared with that of the FIG. 2 process to illustrate the advantages of the present invention.
• inlet gas enters the plant at 80°F [27°C] and 580 psia [3,999 kPa(a)] as stream 31.
• the feed stream 31 is cooled in exchanger 10 by heat exchange with cool residue gas at -93 °F [-70°C] (stream 34a), with separator liquids at -110°F [-79°C] (sfream 33 a), and with separator/absorber liquids at -121°F [-85°C] (stream 35a).
• the feed stream is below the cricondenbar, so the cooled sfream 31a enters separator 11 at -80°F [-62°C] and 570 psia [3,930 kPa(a)] where the vapor (sfream 32) is separated from the condensed liquid (sfream 33).
• the vapor (sfream 32) from separator 11 enters work expansion machine 13 in which mechanical energy is extracted from this portion of the high pressure feed.
• the machine 13 expands the vapor substantially isenfropically from a pressure of about 570 psia [3,930 kPa(a)] to a pressure of about 410 psia [2,827 kPa(a)] (the operating pressure of separator/absorber 15), with the work expansion cooling the expanded stream 32a to a temperature of approximately -104°F [-76°C].
• the expanded and partially condensed stream 32a enters the lower section of separator/absorber 15.
• the liquid portion of the expanded sfream commingles with liquids falling downward from the absorbing section and the combined liquid sfream 35 exits the bottom of separator/absorber 15 at -106°F [-76°C].
• the vapor portion of the expanded stream rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C 3 components and heavier components.
• separator/absorber 15 operates at a higher pressure than deethanizer 17, so the combined liquid stream 35 from the bottom of the separator/absorber 15 is flash expanded to slightly above the 290 psia [1,999 kPa(a)] operating pressure of deethanizer 17 by expansion valve 27, cooling stream 35 to -121°F [-85°C] (stream 35a) before it provides cooling to the incoming feed gas as described earlier.
• the combined liquid stream is heated to -85°F [-65°C], partially vaporizing stream 35b before it is supplied as a mid-column feed to deethanizer 17.
• the separator liquid (sfream 33) is flash expanded to slightly above the operating pressure of deethanizer 17 by expansion valve 12, cooling sfream 33 to -110°F [-79°C] (sfream 33a) before it provides cooling to the incoming feed gas as described earlier.
• Sfream 33b now at 65°F [18°C]
• sfreams 35b and 33b are stripped of their methane and C 2 components.
• the resulting liquid product sfream 37 exits the bottom of the deethanizer at 164°F [73°C] and is cooled to 110°F [43°C] (stream 37a) in heat exchanger 19 before flowing to storage.
• the deethanizer overhead vapor (sfream 36) exits deethanizer 17 at -47°F [-44°C] and is warmed to 105°F [41°C] (sfream 36a) in heat exchanger 24 before entering compressor 25 (driven by a supplemental power source).
• Stream 36b leaves compressor 25 at 435 psia [2,999 kPa(a)] and is cooled to 110°F [43°C] (stream 36c) in heat exchanger 26.
• Sfream 36c is then directed in heat exchange relation with the deethanizer overhead vapor (stream 36) in heat exchanger 24 to cool it (stream 36d) and conserve process cooling.
• sfream 36d can now pressure flow through heat exchanger 20 and thence to the upper feed point of separator/absorber 15.
• the compressed deethanizer overhead at -31°F [-35°C] is directed in heat exchange relation with the overhead (sfream 34) from separator/absorber 15, cooling the stream to -106°F [-77°C] (stream 36e) and partially condensing it.
• the partially condensed sfream is then supplied to the separator section in separator/absorber tower 15 where the condensed liquid is separated from the uncondensed vapor.
• the uncondensed vapor combines with the vapor rising from the lower absorbing section to form the cold distillation stream 34 leaving the upper region of separator/absorber 15.
• the condensed liquid is divided into two portions. One portion, sfream 39, is routed to the lower absorbing section of separator/absorber 15 as the cold liquid that contacts the vapors rising upward through the absorbing section.
• the other portion, sfream 38 is flash expanded to slightly above the operating pressure of deethanizer 17 by expansion valve 28 (stream 38a). It is then supplied at -124°F [-87°C] to the separator section in deethanizer 17 where its condensed liquid is separated from its uncondensed vapor.
• the uncondensed vapor combines with the vapor rising from the lower distillation section to form the deethanizer overhead sfream 36 leaving the upper region of deethanizer 17, while the condensed liquid is routed to the lower distillation section of deethanizer 17 as reflux for the vapors rising upward through the distillation section.
• the distillation sfream leaving the top of separator/absorber 15 at -111°F [-79°C] is the cold residue gas sfream 34.
• the residue gas sfream passes countercurrently to compressed deethanizer overhead stream 36 in heat exchanger 20 and is warmed to -93°F [-70°C] (stream 34a) as it provides cooling and partial condensation of the compressed deethanizer overhead stream.
• the residue gas is further warmed to 74°F [23 °C] (stream 34b) as it passes countercurrently to the incoming feed gas in heat exchanger 10.
• the residue gas is then re-compressed in two stages.
• the first stage is compressor 14 driven by expansion machine 13.
• the second stage is compressor 22 driven by a supplemental power source which compresses the residue gas (stream 34d) to sales line pressure.
• the residue gas product (stream 34e) flows to the sales gas pipeline at 110°F [43°C] and 613 psia [4,226 kPa(a)].
• FIG. 7 (FIG. 7)
• FIG. 2 note that the operating pressure of deethanizer 17 is significantly lower in the present invention than in the FIG. 2 process, 290 psia [1,999 kPa(a)] versus 395 psia [2,723 kPa(a)], and the operating pressure of separator/absorber 15 is significantly higher in the present invention than in the FIG. 2 process, 410 psia [2,827 kPa(a)] versus 380 psia [2,620 kPa(a)]. Accordingly, the residue gas enters compressor 14 at a higher pressure in the FIG. 7 process and less compression horsepower is therefore needed to deliver the residue gas to pipeline pressure.
• separator/absorber 15 operating at a higher pressure than deethanizer 17, it is no longer necessary to pump the absorber bottom liquid (stream 35) and the reflux sfream (sfream 38) to feed deethanizer 17, eliminating the capital and operating cost of pumps 16 and 21 in the FIG. 2 process.
• the deethanizer overhead (stream 36) in the FIG. 7 process provides a more efficient working fluid for a mechanical-compression refrigeration cycle than the inlet gas (stream 31) and residue gas (sfream 34) which are predominantly methane, so that the refiigeration provided to the process by the cycle including compressor 25 not only reduces the refrigeration required from the cycle using compressors 14 and 22, but reduces the total refrigeration energy consumption as well.
• the liquid streams used to provide part of the feed gas cooling, the cold separator liquids (stream 33) and the cold liquid (sfream 35) leaving the bottom of separator/absorber 15, are cooled by flash expansion (streams 33a and 35a, respectively) before entering heat exchanger 10.
• FIG. 7 process versus 195°F [91°C] for sfream 37 in the FIG. 2 process).
• much less sensible heating is required for the column liquids in deethanizer 17, reducing the heating load in reboiler 18 accordingly.
• Total energy consumption for the FIG. 7 embodiment of the present invention is only 85 percent of that required for the prior art of FIG. 2.
• Example 6 In the embodiment of the present invention shown in FIG. 7, the process was operated to achieve the same C 3 component recovery level as the prior art process shown in FIG. 2, with the resulting reduction in the utility consumption due to the better efficiency of the present invention.
• FIG. 8 shows the present invention when applied to match the compression power used by the prior art FIG. 2 process.
• the feed gas composition and conditions considered in the process presented in FIG. 8 are the same as those in FIG. 2. Accordingly, the FIG. 8 process can be compared with that of the FIG. 2 process to illusfrate the advantages of the present invention.
• the feed gas cooling and expansion scheme, the deethanizer overhead compression and cooling scheme, and the tower reflux schemes are essentially the same as those used in FIG. 7.
• the only difference for the FIG. 8 embodiment of the present invention is that the operating pressures of separator/absorber 15 and deethanizer 17 have been adjusted to increase the recovery level for the C 3 components, with the corresponding drops in the process operating temperatures that result from the increase in process cooling (due primarily to the increase in expansion ratio across work expansion machine 13). Note that relative to the FIG. 7 embodiment, in the FIG.
• the operating pressure of separator/absorber 15 has been lowered from 410 psia [2,827 kPa(a)] to 395 psia [2,723 kPa(a)], and the operating pressure of deethanizer 17 has been lowered from 290 psia [1,999 kPa(a)] to 285 psia [1,965 kPa(a)].
• FIG. 8 (FIG. 8)
• the vapor-liquid mixture from heat exchanger 20 can be used without separation, or the liquid portion thereof may be separated. Such commingling of the two streams shall be considered for the purposes of this invention as constituting a contacting device.
• the partially condensed stream from heat exchanger 20 can be separated (using separator 30 as shown in FIG. 9, for instance), and then all or a part of the separated liquid supplied to the separator/absorber or mixed with the vapors fed thereto (with any remaining portion of the separated liquid supplied to the deethanizer).
• the overhead vapors are partially condensed and used to absorb valuable C 3 components and heavier components from the vapors leaving the work expansion machine.
• the present invention is not limited to this embodiment. It may be advantageous, for instance, to treat only a portion of the outlet vapor from the work expansion machine in this manner, or to use only a portion of the overhead condensate as an absorbent, in cases where other design considerations indicate portions of the expansion machine outlet or overhead condensate should bypass the separator/absorber.
• Feed gas conditions, plant size, available equipment, or other factors may indicate that elimination of work expansion machine 13, or replacement with an alternate expansion device (such as an expansion valve), is feasible, or that total (rather than partial) condensation of the overhead stream in heat exchanger 20 is possible or is preferred.
• the separator/absorber can be constructed either as a separate vessel or as a section of the deethanizer column. The use and distribution of the separator liquids, the separator/absorber liquids, and the reflux liquids for process heat exchange, the particular arrangement of heat exchangers for feed gas cooling, and the choice of process streams for specific heat exchange services must be evaluated for each particular application. For instance, as shown in FIG.
• separator liquids may be routed directly to deethanizer 17 via an expansion device (such as expansion valve 12a shown in FIG. 10), with part or none of the liquid used for process cooling in heat exchanger 10.
• separator/absorber liquids may be routed directly to deethanizer 17 via an expansion device (such as expansion valve 27a shown in FIG. 10), with part or none of the liquid used for process cooling in heat exchanger 10.
• the condensed liquid that serves as reflux for deethanizer 17 can be used for process cooling before being supplied to the column. As shown in FIG.
• this liquid may be let down to slightly above the operating pressure of deethanizer 17 (using a device such as expansion valve 28) and used for process cooling (such as in heat exchanger 20 as shown) before being routed to deethanizer 17, with part or none of the liquid routed directly to deethanizer 17 (via expansion valve 28a, for example).
• a device such as expansion valve 28
• process cooling such as in heat exchanger 20 as shown
• External refrigeration may also be employed to generate some or all of the reflux for the deethanizer as illustrated in FIG. 11.
• all of the condensed liquid contained in the partially condensed stream leaving heat exchanger 20 might be directed only to the separator/absorber rather than a portion feeding the deethanizer.
• the heated flash expanded liquid (stream 35b) from the bottom of the separator/absorber could be used to cool a distillation sfream (stream 40) from the deethanizer in heat exchanger 50 to partially condense the distillation stream (stream 40a), whereupon the condensed liquid (stream 38) is separated from the uncondensed vapor (stream 36) in separator 51.
• Reflux pump 52 could then direct the condensed liquid (sfream 38a) to deethanizer 17 to serve as its reflux, with the further heated sfream 35c from heat exchanger 50 feeding deethanizer 17 at a mid-column feed point.
• the heated flash expanded liquid (sfream 35b) from the bottom of the separator/absorber may contain an adequate quantity of liquid to serve as the reflux for the deethanizer, as shown in FIG. 13 and by the dashed lines in FIGS. 15 and 17. Further, as shown in FIGS.
• the relative amount of feed found in each branch of the condensed liquid contained in stream 36e that is split between the two towers in FIGS. 3, 7, and 8 will depend on several factors, including gas pressure, feed gas composition and the quantity of horsepower available.
• the relative amount of condensation in separator 30 in FIG. 4 the relative amount of feed contained in stream 40a that is split between the two towers in FIG. 5, and the relative amount of condensation in separators 29 and 30 in FIG. 6 will also depend on factors such as these.
• the optimum split or distribution generally cannot be predicted without evaluating the particular circumstances for a specific application of the present invention.
• the mid-column feed positions depicted in FIGS. 3 through 8 are the preferred feed locations for the process operating conditions described.
• the relative locations of the mid-column feeds may vary depending on inlet composition or other factors such as desired recovery levels, etc.
• two or more of the feed streams, or portions thereof, may be combined depending on the relative temperatures and quantities of individual sfreams, and the combined stream then fed to a mid-column feed position.
• FIGS. 3 through 8 are the preferred embodiments for the compositions and pressure conditions shown. Although individual stream expansion is depicted in particular expansion devices, alternative expansion means may be employed where appropriate. For example, conditions may warrant work expansion of some or all of the liquid streams (such as streams 33, 35, and/or 38 in FIG. 7).
• FIGS. 3 through 14, 17, and 18 depict using a supplemental power source for compressor 25 to compress this stream, while compressing the residue gas (stream 34b in FIGS. 3, 4, 7, and 8) or absorber overhead (sfream 34b in FIGS. 5 and 6) using compressor 14 driven by expansion machine 13.
• An improvement in utility consumption required for operating the 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.
• increased C 3 component recovery can be obtained for a fixed utility consumption.
• complete rejection of the C 2 components to the residue gas is not required by the present invention. If the project economics favor recovery of the C 2 components in the liquid product (sfream 37), the process operating conditions can be altered to recover in the liquid product a significant portion of the C 2 components present in the feed gas. Preliminary calculations indicate that perhaps 40% of the C 2 components can be recovered in this fashion.

Landscapes

• Engineering & Computer Science (AREA)
• Chemical & Material Sciences (AREA)
• Oil, Petroleum & Natural Gas (AREA)
• Organic Chemistry (AREA)
• Thermal Sciences (AREA)
• General Engineering & Computer Science (AREA)
• Physics & Mathematics (AREA)
• Mechanical Engineering (AREA)
• Chemical Kinetics & Catalysis (AREA)
• General Chemical & Material Sciences (AREA)
• Analytical Chemistry (AREA)
• Water Supply & Treatment (AREA)
• Vaporization, Distillation, Condensation, Sublimation, And Cold Traps (AREA)
• Separation By Low-Temperature Treatments (AREA)
• Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)

Priority Applications (1)

Applications Claiming Priority (4)

Publications (1)

Family

ID=26919445

Family Applications (1)

Country Status (5)

Cited By (7)

Families Citing this family (61)

Citations (2)

• 2001
  • 2001-07-26 US US09/915,904 patent/US20020166336A1/en not_active Abandoned
  • 2001-08-13 PE PE2001000806A patent/PE20020320A1/es not_active Application Discontinuation
  • 2001-08-14 WO PCT/US2001/025208 patent/WO2002014242A1/en not_active Ceased
  • 2001-08-14 AU AU2001283307A patent/AU2001283307A1/en not_active Abandoned
  • 2001-08-15 AR ARP010103907A patent/AR030455A1/es unknown

Patent Citations (3)

Also Published As

Similar Documents

Legal Events