US20080078205A1 - Hydrocarbon Gas Processing - Google Patents

Hydrocarbon Gas Processing Download PDF

Info

Publication number
US20080078205A1
US20080078205A1 US11/839,693 US83969307A US2008078205A1 US 20080078205 A1 US20080078205 A1 US 20080078205A1 US 83969307 A US83969307 A US 83969307A US 2008078205 A1 US2008078205 A1 US 2008078205A1
Authority
US
United States
Prior art keywords
stream
contacting
components
expanded
vapor
Prior art date
Legal status (The legal status is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the status listed.)
Abandoned
Application number
US11/839,693
Other languages
English (en)
Inventor
Kyle T. Cuellar
Tony L. Martinez
John D. Wilkinson
Joe T. Lynch
Hank M. Hudson
Current Assignee (The listed assignees may be inaccurate. Google has not performed a legal analysis and makes no representation or warranty as to the accuracy of the list.)
Honeywell UOP LLC
Original Assignee
Ortloff Engineers Ltd
Priority date (The priority date is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the date listed.)
Filing date
Publication date
Family has litigation
First worldwide family litigation filed litigation Critical https://patents.darts-ip.com/?family=39259824&utm_source=google_patent&utm_medium=platform_link&utm_campaign=public_patent_search&patent=US20080078205(A1) "Global patent litigation dataset” by Darts-ip is licensed under a Creative Commons Attribution 4.0 International License.
Application filed by Ortloff Engineers Ltd filed Critical Ortloff Engineers Ltd
Priority to US11/839,693 priority Critical patent/US20080078205A1/en
Priority to CN2007800359949A priority patent/CN101517340B/zh
Priority to MYPI20091040A priority patent/MY157917A/en
Priority to BRPI0717082-3A priority patent/BRPI0717082B1/pt
Priority to MX2009002053A priority patent/MX2009002053A/es
Priority to AU2007305167A priority patent/AU2007305167B2/en
Priority to PCT/US2007/076199 priority patent/WO2008042509A2/en
Priority to CA2664224A priority patent/CA2664224C/en
Priority to CL2007002773A priority patent/CL2007002773A1/es
Priority to PE2007001299A priority patent/PE20080652A1/es
Priority to ARP070104325A priority patent/AR063066A1/es
Priority to SA7280532A priority patent/SA07280532B1/ar
Assigned to ORTLOFF ENGINEERS, LTD. reassignment ORTLOFF ENGINEERS, LTD. ASSIGNMENT OF ASSIGNORS INTEREST (SEE DOCUMENT FOR DETAILS). Assignors: CUELLAR, KYLE T., MR., HUDSON, HANK M., MR., LYNCH, JOE T., MR., MARTINEZ, TONY L., MR., WILKINSON, JOHN D., MR.
Publication of US20080078205A1 publication Critical patent/US20080078205A1/en
Priority to NO20090789A priority patent/NO20090789L/no
Priority to TN2009000078A priority patent/TN2009000078A1/fr
Assigned to UOP LLC reassignment UOP LLC ASSIGNMENT OF ASSIGNORS INTEREST (SEE DOCUMENT FOR DETAILS). Assignors: ORTLOFF ENGINEERS, LTD.
Abandoned legal-status Critical Current

Links

Images

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
    • F25J3/0204—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
    • F25J3/0228—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
    • 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
    • F25J3/0238—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
    • 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
    • F25J2200/02—Processes or apparatus using separation by rectification in a single 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
    • F25J2200/30—Processes or apparatus using separation by rectification using a side column in a single pressure 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
    • F25J2200/78—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
    • 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
    • F25J2210/06—Splitting of the feed stream, e.g. for treating or cooling in different ways
    • 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
    • F25J2220/00—Processes or apparatus involving steps for the removal of impurities
    • F25J2220/60—Separating impurities from natural gas, e.g. mercury, cyclic hydrocarbons
    • F25J2220/66—Separating acid gases, e.g. CO2, SO2, H2S or RSH
    • 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
    • F25J2230/08—Cold compressor, i.e. suction of the gas at cryogenic temperature and generally without afterstage-cooler
    • 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
    • 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
    • F25J2290/00—Other details not covered by groups F25J2200/00 - F25J2280/00
    • 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 for the separation of a gas containing hydrocarbons.
  • the applicants claim the benefits under Title 35, United States Code, Section 119(e) of prior U.S. Provisional Application Nos. 60/848,299 which was filed on Sep. 28, 2006 and 60/897,683 which was filed on Jan. 25, 2007.
  • 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, 90.5% methane, 4.1% ethane and other C 2 components, 1.3% propane and other C 3 components, 0.4% iso-butane, 0.3% normal butane, 0.5% pentanes plus, and 2.6% carbon dioxide, with the balance made up of nitrogen. Sulfur containing gases are also sometimes present.
  • 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 + or 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 (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.
  • a portion of the vapor remaining from the partial condensation can be passed through a work expansion machine or engine, or an expansion valve, to a lower pressure at which additional liquids are condensed as a result of further cooling of the stream.
  • the pressure after expansion is essentially the same as the pressure at which the distillation column is operated.
  • the combined vapor-liquid phases resulting from the expansion are supplied as feed to the column.
  • 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 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 for two main reasons.
  • the first reason is that the conventional demethanizer is operated largely as a stripping column.
  • the methane product of the process therefore, typically comprises vapors leaving the top fractionation stage of the column, together with vapors not subjected to any rectification step.
  • the second reason that this ideal situation cannot be obtained is that carbon dioxide contained in the feed gas fractionates in the demethanizer and can build up to concentrations of as much as 5% to 10% or more in the tower even when the feed gas contains less than 1% carbon dioxide. At such high concentrations, formation of solid carbon dioxide can occur depending on temperatures, pressures, and the liquid solubility. It is well known that natural gas streams usually contain carbon dioxide, sometimes in substantial amounts. If the carbon dioxide concentration in the feed gas is high enough, it becomes impossible to process the feed gas as desired due to blockage of the process equipment with solid carbon dioxide (unless carbon dioxide removal equipment is added, which would increase capital cost substantially).
  • the present invention provides a means for generating a liquid reflux stream that will improve the recovery efficiency for the desired products while simultaneously substantially mitigating the problem of carbon dioxide icing.
  • 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, 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.
  • these processes require the use of a large amount of compression power to provide the motive force for recycling the reflux stream to the demethanizer, adding to both the capital cost and the operating cost of facilities using these processes.
  • the present invention also employs an upper rectification section (or a separate rectification column in some embodiments).
  • 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.
  • 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.
  • C 3 and C 4 + recoveries in excess of 99 percent can be obtained with no loss in C 2 + component recovery.
  • the present invention provides the further advantage of being able to maintain in excess of 99 percent 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 percent separation of methane and lighter components from the C 2 components and heavier components while maintaining the same recovery levels as the prior art and improving the safety factor with respect to the danger of carbon dioxide icing.
  • the present invention although applicable at lower pressures and warmer temperatures, is particularly advantageous when processing feed gases in the range of 400 to 1500 psia [2,758 to 10,342 kPa(a)] or higher under conditions requiring NGL recovery column overhead temperatures of ⁇ 50° F. [ ⁇ 46° C.] or colder.
  • FIG. 1 is a flow diagram of a prior art natural gas processing plant in accordance with U.S. Pat. No. 7,191,617;
  • FIG. 2 is a flow diagram of a natural gas processing plant in accordance with the present invention.
  • FIG. 3 is a concentration-temperature diagram for carbon dioxide showing the effect of 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 concentration-temperature diagram for carbon dioxide showing the effect of the present invention with respect to the process of FIG. 4 ;
  • FIGS. 6 through 9 are flow diagrams illustrating alternative means of application of the present invention to a natural gas stream.
  • FIG. 10 is a partial flow diagram illustrating alternative means of accomplishing the splitting of the vapor feed in accordance with the present invention.
  • 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 assignee's U.S. Pat. No. 7,191,617.
  • inlet gas enters the plant at 120° F. [49° C.] and 1040 psia [7,171 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 at ⁇ 28° F. [ ⁇ 33° C.] (stream 48 a ), demethanizer reboiler liquids at 35° F. [2° C.] (stream 41 ), demethanizer lower side reboiler liquids at ⁇ 10° F. [ ⁇ 23° C.] (stream 40 ), and demethanizer upper side reboiler liquids at ⁇ 79° F. [ ⁇ 62° C.] (stream 39 ).
  • the cooled stream 31 a enters separator 11 at ⁇ 15° F.
  • the vapor (stream 32 ) from separator 11 is divided into two streams, 35 and 36 .
  • Stream 35 containing about 36% of the total vapor, passes through heat exchanger 15 in heat exchange relation with the cold residue gas at ⁇ 127° F. [ ⁇ 88° C.] (stream 48 ) where it is cooled to substantial condensation.
  • the resulting substantially condensed stream 35 a at ⁇ 123° F. [ ⁇ 86° C.] is then flash expanded through expansion valve 16 to the operating pressure of fractionation tower 19 . During expansion a portion of the stream is vaporized, resulting in cooling of the total stream to ⁇ 134° F. [ ⁇ 92° C.].
  • the expanded stream 35 b is supplied to fractionation tower 19 at an upper mid-column feed point.
  • the remaining 64% of the vapor from separator 11 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 ⁇ 90° F. [ ⁇ 68° C.].
  • the typical commercially available expanders are capable of recovering on the order of 80-88% of the work theoretically available in an ideal isentropic expansion.
  • the work recovered is often used to drive a centrifugal compressor (such as item 18 ) that can be used to re-compress the residue gas (stream 48 b ), for example.
  • the partially condensed expanded stream 36 a is thereafter supplied as feed to fractionation tower 19 a second lower mid-column feed point.
  • the demethanizer in tower 19 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing.
  • the demethanizer tower consists of two sections: an upper absorbing (rectification) section 19 a that contains the trays and/or packing to provide the necessary contact between the vapor portion of the expanded streams 35 b 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 (demethanizing) section 19 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 19 a that contains the trays and/or packing to provide the necessary contact between the vapor portion of the expanded streams 35 b and 36 a rising upward and cold liquid falling downward to condense and absorb the C 2 components, C 3 components, and heavier components
  • the stripping section 19 b also includes reboilers (such as trim reboiler 20 and the reboiler and side reboilers described previously) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column to strip the liquid product, stream 42 , of methane and lighter components.
  • Stream 36 a enters demethanizer 19 at an intermediate feed position located in the lower region of absorbing section 19 a of demethanizer 19 .
  • the liquid portion of the expanded stream commingles with liquids falling downward from the absorbing section 19 a and the combined liquid continues downward into the stripping section 19 b of demethanizer 19 .
  • the vapor portion of the expanded stream rises upward through absorbing section 19 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 43 ) is withdrawn from the upper region of stripping section 19 b. This stream is then cooled from ⁇ 112° F. [ ⁇ 80° C.] to ⁇ 130° F. [ ⁇ 90° C.] and partially condensed (stream 43 a ) in heat exchanger 22 by heat exchange with the cold demethanizer overhead stream 38 exiting the top of demethanizer 19 at ⁇ 134° F. [ ⁇ 92° C].
  • the cold demethanizer overhead stream is warmed slightly to ⁇ 126° F. [ ⁇ 88° C.] (stream 38 a ) as it cools and condenses at least a portion of stream 43 .
  • the operating pressure in reflux separator 23 (428 psia [2,951 kPa(a)]) is maintained slightly below the operating pressure of demethanizer 19 .
  • This provides the driving force which causes distillation vapor stream 43 to flow through heat exchanger 22 and thence into the reflux separator 23 wherein the condensed liquid (stream 45 ) is separated from the uncondensed vapor (stream 44 ).
  • Stream 44 then combines with the warmed demethanizer overhead stream 38 a from heat exchanger 22 to form cold residue gas stream 48 at ⁇ 127° F. [ ⁇ 88° C.].
  • the liquid stream 45 from reflux separator 23 is pumped by pump 24 to a pressure slightly above the operating pressure of demethanizer 19 , and stream 45 a is then supplied as cold top column feed (reflux) to demethanizer 19 .
  • This cold liquid reflux absorbs and condenses the propane and heavier components rising in the upper rectification region of absorbing section 19 a of demethanizer 19 .
  • the feed streams are stripped of their methane and lighter components.
  • the resulting liquid product (stream 42 ) exits the bottom of tower 19 at 52° F. [11° C.], based on a typical specification of a methane to ethane ratio of 0.025:1 on a molar basis in the bottom product.
  • the distillation vapor stream forming the tower overhead (stream 38 ) is warmed in heat exchanger 22 as it provides cooling to distillation stream 43 as described previously, then combines with stream 44 to form the cold residue gas stream 48 .
  • the residue gas passes countercurrently to the incoming feed gas in heat exchanger 15 where it is heated to ⁇ 28° F.
  • stream 48 a [ ⁇ 33° C.] (stream 48 a ), and in heat exchanger 10 where it is heated to 107° F. [42° C.] (stream 48 b ) as it provides cooling as previously described.
  • the residue gas is then re-compressed in two stages, compressor 18 driven by expansion machine 17 and compressor 27 driven by a supplemental power source.
  • stream 48 d is cooled to 120° F. [49° C.] in discharge cooler 28
  • the residue gas product (stream 48 e ) flows to the sales gas pipeline at 1040 psia [7,171 kPa(a)].
  • FIG. 2 illustrates a flow diagram of a process in accordance with the present invention.
  • the feed gas composition and conditions considered in the process presented in FIG. 2 are the same as those in FIG. 1 . Accordingly, the FIG. 2 process can be compared with that of the FIG. 1 process to illustrate the advantages of the present invention.
  • inlet gas enters the plant as stream 31 and is cooled in heat exchanger 10 by heat exchange with cool residue gas at ⁇ 66° F. [ ⁇ 54° C.] (stream 38 b ), demethanizer reboiler liquids at 48° F. [9° C.] (stream 41 ), demethanizer lower side reboiler liquids at 5° F. [ ⁇ 15° C.] (stream 40 ), and demethanizer upper side reboiler liquids at ⁇ 70° F. [ ⁇ 57° C.] (stream 39 ).
  • the cooled stream 31 a enters separator 11 at ⁇ 38° F.
  • the separator liquid (stream 33 ) may in some cases be divided into two streams, stream 47 and stream 37 .
  • all of the separator liquid in stream 33 is directed to stream 37 and is expanded to the operating pressure (approximately 470 psia [3,238 kPa(a)]) of fractionation tower 19 by expansion valve 12 , cooling stream 37 a to ⁇ 68° F. [ ⁇ 56° C.] before it is supplied to fractionation tower 19 at a lower mid-column feed point.
  • all of the separator liquid in stream 33 may be directed to stream 47 , or a portion of stream 33 may be directed to stream 37 with the remaining portion directed to stream 47 .
  • the vapor (stream 32 ) from separator 11 is divided into two streams, 34 and 36 .
  • Stream 34 containing about 22% of the total vapor, may in some embodiments be combined with a portion (stream 47 ) of separator liquid stream 33 to form combined stream 35 .
  • Stream 34 or 35 passes through heat exchanger 15 in heat exchange relation with the cold residue gas at ⁇ 105° F. [ ⁇ 76° C.] (stream 38 a ) where it is cooled to substantial condensation.
  • the resulting substantially condensed stream 35 a at ⁇ 101° F. [ ⁇ 74° C.] is then flash expanded through expansion valve 16 to the operating pressure of fractionation tower 19 .
  • the expanded stream 35 b leaving expansion valve 16 reaches a temperature of ⁇ 128° F. [ ⁇ 89° C.] and is supplied to fractionation tower 19 at an upper mid-column feed point.
  • the remaining 78% of the vapor from separator 11 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 ⁇ 102° F. [ ⁇ 74° C.].
  • the partially condensed expanded stream 36 a is thereafter supplied as feed to fractionation tower 19 a second lower mid-column feed point.
  • the demethanizer in tower 19 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing.
  • the demethanizer tower consists of two sections: an upper absorbing (rectification) section 19 a that contains the trays and/or packing to provide the necessary contact between the vapor portion of the expanded streams 35 b 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 (demethanizing) section 19 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 19 a that contains the trays and/or packing to provide the necessary contact between the vapor portion of the expanded streams 35 b and 36 a rising upward and cold liquid falling downward to condense and absorb the C 2 components, C 3 components, and heavier components
  • the stripping section 19 b also includes reboilers (such as trim reboiler 20 and the reboiler and side reboilers described previously) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column to strip the liquid product, stream 42 , of methane and lighter components.
  • Stream 36 a enters demethanizer 19 at an intermediate feed position located in the lower region of absorbing section 19 a of demethanizer 19 .
  • the liquid portion of the expanded stream commingles with liquids falling downward from the absorbing section 19 a and the combined liquid continues downward into the stripping section 19 b of demethanizer 19 .
  • the vapor portion of the expanded stream rises upward through absorbing section 19 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 43 ) is withdrawn from the upper region of stripping section 19 b at ⁇ 108° F. [ ⁇ 78° C.] below expanded stream 36 a and is compressed to approximately 609 psia [4,199 kPa(a)] by vapor compressor 21 .
  • the compressed stream 43 a is then cooled from ⁇ 78° F. [ ⁇ 61° C.] to ⁇ 125° F. [ ⁇ 87° C.] and substantially condensed (stream 43 b ) in heat exchanger 22 by heat exchange with the cold demethanizer overhead stream 38 exiting the top of demethanizer 19 at ⁇ 129° F. [ ⁇ 89° C.].
  • the cold demethanizer overhead stream is warmed to ⁇ 105° F. [ ⁇ 76° C.] (stream 38 a ) as it cools and condenses stream 43 a.
  • substantially condensed stream 43 b is at a pressure greater than the operating pressure of demethanizer 19 , it is flash expanded through expansion valve 25 to the operating pressure of fractionation tower 19 . During expansion a small portion of the stream is vaporized, resulting in cooling of the total stream to ⁇ 132° F. [ ⁇ 91° C.]. The expanded stream 43 c is then supplied as cold top column feed (reflux) to demethanizer 19 .
  • the vapor portion (if any) of stream 43 c combines with the distillation vapor rising from the upper fractionation stage to form residue gas stream 38 , while the cold liquid reflux portion absorbs and condenses the C 2 components, C 3 components, and heavier components rising in the upper rectification region of absorbing section 19 a of demethanizer 19 .
  • the feed streams are stripped of their methane and lighter components.
  • the resulting liquid product (stream 42 ) exits the bottom of tower 19 at 66° F. [19° C.].
  • the distillation vapor stream forming cold residue gas stream 38 is warmed in heat exchanger 22 as it provides cooling to compressed distillation stream 43 a as described previously.
  • the residue gas (stream 38 a ) passes countercurrently to the incoming feed gas in heat exchanger 15 where it is heated to ⁇ 66° F. [ ⁇ 54° C.] (stream 38 b ), and in heat exchanger 10 where it is heated to 110° F. [43° C.] (stream 38 c ) as it provides cooling as previously described.
  • stream 38 e is cooled to 120° F. [49° C.] in discharge cooler 28
  • the residue gas product flows to the sales gas pipeline at 1040 psia [7,171 kPa(a)].
  • Tables I and II show that, compared to the prior art, the present invention maintains essentially the same ethane recovery (83.05% versus 83.06%), but improves both the propane recovery (99.33% versus 98.50%) and butanes+ recovery (99.97% versus 99.94%). Comparison of Tables I and II further shows that these increased yields were achieved using less horsepower than the prior art (11,389 HP versus 12,464 HP, or more than 8% less).
  • the boost in pressure provided by vapor compressor 21 allows the column overhead (stream 38 ) to condense all of distillation vapor stream 43 , unlike the prior art process which can condense only a fraction of the stream.
  • the top reflux stream (stream 43 c ) for the present invention is more than 5 times greater than that of the prior art (stream 45 a ), providing much more efficient rectification in the upper region of absorbing section 19 a.
  • the quantity of secondary reflux stream 35 b can be correspondingly less without reducing the product yields.
  • FIG. 3 is a graph of the relation between carbon dioxide concentration and temperature.
  • Line 71 represents the equilibrium conditions for solid and liquid carbon dioxide in methane. (The liquid-solid equilibrium line in this graph is based on the data given in FIG. 16-33 on page 16-24 of the Engineering Data Book, Twelfth Edition, published in 2004 by the Gas Processors Suppliers Association, which is often used as a reference when checking for potential icing conditions.)
  • FIG. 3 Also plotted in FIG. 3 is a line representing the conditions for the liquids on the fractionation stages of demethanizer 19 in the prior art FIG. 1 process (line 72 ). As can be seen, a portion of this operating line lies above the liquid-solid equilibrium line, indicating that the prior art FIG. 1 process cannot be operated at these conditions without encountering carbon dioxide icing problems. As a result, it is not possible to use the FIG. 1 process under these conditions, so the prior art FIG. 1 process cannot actually achieve the recovery efficiencies stated in Table I in practice without removal of at least some of the carbon dioxide from the feed gas. This would, of course, substantially increase capital cost.
  • Line 73 in FIG. 3 represents the conditions for the liquids on the fractionation stages of demethanizer 19 in the present invention as depicted in FIG. 2 .
  • the present invention could tolerate a 20% higher concentration of carbon dioxide in its feed gas than the prior art FIG. 1 process could tolerate without risk of crossing the liquid-solid equilibrium line.
  • the prior art FIG. 1 process cannot be operated to achieve the recovery levels given in Table I because of icing, the present invention could in fact be operated at even higher recovery levels than those given in Table II without risk of icing.
  • the shift in the operating conditions of the FIG. 2 demethanizer as indicated by line 73 in FIG. 3 can be understood by comparing the distinguishing features of the present invention to the prior art process of FIG. 1 . While the shape of the operating line for the prior art FIG. 1 process (line 72 ) is similar to the shape of the operating line for the present invention (line 73 ), there is a key difference.
  • the operating temperatures of the critical upper fractionation stages in the demethanizer in the FIG. 2 process are warmer than those of the corresponding fractionation stages in the demethanizer in the prior art FIG. 1 process, effectively shifting the operating line of the FIG. 2 process away from the liquid-solid equilibrium line. The warmer temperatures of the fractionation stages in the FIG.
  • Another advantage of the present invention is a reduction in the amount of carbon dioxide leaving demethanizer 19 in liquid product stream 42 . Comparing stream 42 in Table I for the prior art FIG. 1 process to stream 42 in Table II for the FIG. 2 embodiment of the present invention reveals that there is nearly a 20% reduction in the quantity of carbon dioxide captured in stream 42 with the present invention. This generally reduces the product treating requirements by a corresponding amount, reducing both the capital cost and the operating cost of the treating system.
  • the reflux streams for absorbing section 19 a in demethanizer 19 of the prior art FIG. 1 process are streams 45 a and 35 b, while those for the present invention shown in the FIG. 2 process are streams 43 c and 35 b. Comparing these streams in Table I and Table II, note that the total amounts of C 2 components and carbon dioxide in the reflux streams in the prior art FIG. 1 process are 470 and 318 Lb. Moles/Hr [470 and 318 kg moles/Hr], respectively, versus 353 and 266 Lb. Moles/Hr [353 and 266 kg moles/Hr], respectively, for the reflux streams in the FIG. 2 process of the present invention.
  • FIG. 4 An alternative embodiment of the present invention is shown in FIG. 4 .
  • the feed gas composition and conditions considered in the process presented in FIG. 4 are the same as those in FIGS. 1 and 2 . Accordingly, FIG. 4 can be compared with the prior art FIG. 1 process to illustrate the advantages of the present invention, and can likewise be compared to the embodiment displayed in FIG. 2 .
  • inlet gas enters the plant as stream 31 and is cooled in heat exchanger 10 by heat exchange with cool residue gas at ⁇ 66° F. [ ⁇ 55° C.] (stream 38 b ), demethanizer reboiler liquids at 51° F. [11° C.] (stream 41 ), demethanizer lower side reboiler liquids at 10° F. [ ⁇ 12° C.] (stream 40 ), and demethanizer upper side reboiler liquids at ⁇ 65° F. [ ⁇ 54° C.] (stream 39 ).
  • the cooled stream 31 a enters separator 11 at ⁇ 38° F.
  • the separator liquid (stream 33 ) may in some cases be divided into two streams, stream 47 and stream 37 .
  • all of the separator liquid in stream 33 is directed to stream 37 and is expanded to the operating pressure (approximately 480 psia [3,309 kPa(a)]) of fractionation tower 19 by expansion valve 12 , cooling stream 37 a to ⁇ 67° F. [ ⁇ 55° C.] before it is supplied to fractionation tower 19 at a lower mid-column feed point.
  • all of the separator liquid in stream 33 may be directed to stream 47 , or a portion of stream 33 may be directed to stream 37 with the remaining portion directed to stream 47 .
  • the vapor (stream 32 ) from separator 11 is divided into two streams, 34 and 36 .
  • Stream 34 containing about 23% of the total vapor, may in some embodiments be combined with a portion (stream 47 ) of separator liquid stream 33 to form combined stream 35 .
  • Stream 34 or 35 passes through heat exchanger 15 in heat exchange relation with the cold residue gas at ⁇ 106° F. [ ⁇ 77° C.] (stream 38 a ) where it is cooled to substantial condensation.
  • the resulting substantially condensed stream 35 a at ⁇ 102° F. [ ⁇ 74° C.] is then flash expanded through expansion valve 16 to the operating pressure of fractionation tower 19 .
  • the expanded stream 35 b leaving expansion valve 16 reaches a temperature of ⁇ 127° F. [ ⁇ 88° C.] and is supplied to fractionation tower 19 at an upper mid-column feed point.
  • the remaining 77% of the vapor from separator 11 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 ⁇ 101° F. [ ⁇ 74° C.].
  • the partially condensed expanded stream 36 a is thereafter supplied as feed to fractionation tower 19 a second lower mid-column feed point.
  • a portion of the distillation vapor (stream 43 ) is withdrawn from the lower region of absorbing section 19 a of demethanizer 19 at ⁇ 113° F. [ ⁇ 81° C.] above expanded stream 36 a and is compressed to approximately 619 psia [4,266 kPa(a)] by vapor compressor 21 .
  • the compressed stream 43 a is then cooled from ⁇ 84° F. [ ⁇ 65° C.] to ⁇ 124° F. [ ⁇ 87° C.] and substantially condensed (stream 43 b ) in heat exchanger 22 by heat exchange with the cold demethanizer overhead stream 38 exiting the top of demethanizer 19 at ⁇ 128° F. [ ⁇ 89° C.].
  • the cold demethanizer overhead stream is warmed to ⁇ 106° F. [ ⁇ 77° C.] (stream 38 a ) as it cools and condenses stream 43 a.
  • substantially condensed stream 43 b is at a pressure greater than the operating pressure of demethanizer 19 , it is flash expanded through expansion valve 25 to the operating pressure of fractionation tower 19 . During expansion a small portion of the stream is vaporized, resulting in cooling of the total stream to ⁇ 131° F. [ ⁇ 91° C.]. The expanded stream 43 c is then supplied as cold top column feed (reflux) to demethanizer 19 .
  • the vapor portion (if any) of stream 43 c combines with the distillation vapor rising from the upper fractionation stage to form residue gas stream 38 , while the cold liquid reflux portion absorbs and condenses the C 2 components, C 3 components, and heavier components rising in the upper rectification region of absorbing section 19 a of demethanizer 19 .
  • the feed streams are stripped of their methane and lighter components.
  • the resulting liquid product (stream 42 ) exits the bottom of tower 19 at 70° F. [21° C.].
  • the distillation vapor stream forming cold residue gas stream 38 is warmed in heat exchanger 22 as it provides cooling to compressed distillation stream 43 a as described previously.
  • the residue gas (stream 38 a ) passes countercurrently to the incoming feed gas in heat exchanger 15 where it is heated to ⁇ 66° F. [ ⁇ 55° C.] (stream 38 b ), and in heat exchanger 10 where it is heated to 110° F. [43° C.] (stream 38 c ) as it provides cooling as previously described.
  • stream 38 e is cooled to 120° F. [49° C.] in discharge cooler 28
  • the residue gas product flows to the sales gas pipeline at 1040 psia [7,171 kPa(a)].
  • FIG. 4 embodiment maintains the same ethane recovery while improving the propane recovery (99.50% versus 99.33%) and butanes+ recovery (99.98% versus 99.97%) slightly.
  • comparison of Tables II and III further shows that these yields were achieved using about 3% less horsepower than that required by the FIG. 2 embodiment of the present invention.
  • the drop in the power requirements for the FIG. 4 embodiment is mainly due to the lower content of C 2 + components in top reflux stream 43 c, which provides more efficient rectification in the upper region of absorbing section 19 a so that demethanizer 19 can be operated at a slightly higher operating pressure (thereby reducing compression requirements) without reducing product yields.
  • the concentrations of C 2 components and particularly the C 3 + components in stream 43 of the FIG. 4 embodiment are significantly lower, so that higher product yields are achieved using less power than the FIG. 2 embodiment.
  • the lower concentrations of C 2 components and C 3 + components in stream 43 of the FIG. 4 embodiment are the result of withdrawing the distillation vapor from the lower region of absorbing section 19 a rather than from the upper region of stripping section 19 b as in the FIG. 2 embodiment.
  • FIG. 5 is another graph of the relation between carbon dioxide concentration and temperature, with line 71 as before representing the equilibrium conditions for solid and liquid carbon dioxide in methane and line 72 representing the conditions for the liquids on the fractionation stages of demethanizer 19 in the prior art process of FIG. 1 .
  • Line 74 in FIG. 5 represents the conditions for the liquids on the fractionation stages of demethanizer 19 in the present invention as depicted in FIG. 4 , and shows a safety factor of 1.2 between the anticipated operating conditions and the icing conditions for the FIG. 4 process.
  • this embodiment of the present invention could also tolerate an increase of 20 percent in the concentration of carbon dioxide without risk of icing.
  • this improvement in the icing safety factor could be used to advantage by operating the demethanizer at lower pressure (i.e., with colder temperatures on the fractionation stages) to raise the C 2 + component recovery levels without encountering icing problems.
  • the shape of line 74 in FIG. 5 for the FIG. 4 embodiment is very similar to that of line 73 in FIG. 3 for the FIG. 2 embodiment.
  • the primary difference is the significantly lower carbon dioxide concentrations of the liquids on the fractionation stages in the lower section of the FIG. 4 demethanizer due to withdrawing the distillation vapor stream at a higher location on the column in this embodiment.
  • 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 one theoretical stage, and it is believed that even the equivalent of a fractional theoretical stage may allow achieving these benefits.
  • all or a part of the expanded substantially condensed distillation stream 43 c from expansion valve 25 , all or a part of the expanded substantially condensed stream 35 b from expansion valve 16 , and all or a part of the expanded stream 36 a from work expansion machine 17 can be combined (such as in the piping joining the expansion valve to the demethanizer) and if thoroughly intermingled, the vapors and liquids will mix together and separate in accordance with the relative volatilities of the various components of the total combined streams.
  • Such commingling of the three streams shall be considered for the purposes of this invention as constituting an absorbing section.
  • the substantially condensed distillation stream 43 b may be advantageous to split into at least two streams as shown in FIGS. 6 through 9 .
  • This allows a portion (stream 51 ) to be supplied above the location where vapor distillation stream 43 is withdrawn (and perhaps also above the feed location of expanded stream 36 a ), either lower in the absorbing section of fractionation tower 19 ( FIGS. 6 and 7 ) or lower on absorber column 19 ( FIGS. 8 and 9 ), to increase the liquid flow in that part of the distillation system and improve the rectification of stream 43 .
  • expansion valve 26 is used to expand stream 51 to the column operating pressure (forming stream 51 a ), while expansion valve 25 is used to expand the remaining portion (stream 50 ) to the column operating pressure so that the resulting stream 50 a can then be supplied to the top of the absorbing section in demethanizer 19 ( FIGS. 6 and 7 ) or to the top of absorber column 19 ( FIGS. 8 and 9 ).
  • FIGS. 8 and 9 depict a fractionation tower constructed in two vessels, absorber (rectifier) column 19 (a contacting and separating device) and stripper column 29 (a distillation column).
  • the overhead vapor (stream 46 ) from stripper column 29 is split into two portions.
  • One portion (stream 43 ) is routed to compressor 21 and thence to heat exchanger 22 to generate reflux for absorber column 19 as described earlier.
  • the remaining portion (stream 49 ) flows to the lower section of absorber column 19 to be contacted by expanded substantially condensed stream 35 b and the expanded substantially condensed distillation stream (either stream 50 a, or streams 50 a and 51 a ).
  • Pump 30 is used to route the liquids (stream 52 ) from the bottom of absorber column 19 to the top of stripper column 29 so that the two towers effectively function as one distillation system.
  • all of the overhead vapor (stream 46 ) flows to the lower section of absorber column 19 , and distillation vapor stream 43 is withdrawn from a location higher in absorber column 19 , above the feed location of expanded stream 36 a.
  • the decision whether to construct the fractionation tower as a single vessel (such as demethanizer 19 in FIGS. 2 , 4 , 6 , and 7 ) or multiple vessels will depend on a number of factors such as plant size, the distance to fabrication facilities, etc.
  • 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 ) and/or the substantially condensed distillation stream (stream 43 b ).
  • distillation stream 43 is substantially condensed and the resulting condensate used to absorb valuable C 2 components, C 3 components, and heavier components from the vapors rising through the upper region of absorbing section 19 a of demethanizer 19 ( FIGS. 2 , 4 , 6 , and 7 ) or absorber column 19 ( FIGS. 8 and 9 ).
  • 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 19 a of demethanizer 19 ( FIGS.
  • distillation stream 43 a in heat exchanger 22 may favor partial condensation, rather than total condensation, of distillation stream 43 a in heat exchanger 22 .
  • Other circumstances may favor that distillation stream 43 be a total vapor side draw from fractionation column 19 rather than a partial vapor side draw.
  • it may be advantageous to use external refrigeration to provide some portion of the cooling of distillation stream 43 a in heat exchanger 22 .
  • the splitting of the vapor feed may be accomplished in several ways.
  • vapor splitting may be effected in a separator.
  • the splitting of the vapor occurs following cooling, and perhaps after separation of any liquids which may have been formed.
  • the high pressure gas may be split, however, prior to any cooling of the inlet gas as shown in FIG. 10 .
  • Streams 35 b, 36 a, and 37 a in FIG. 10 may all be fed to a distillation column (such as demethanizer 19 in FIGS.
  • streams 35 b and 36 a may be fed to a contacting and separating device and stream 37 a may be fed to a distillation column (such as absorber column 19 and stripper column 29 , respectively, in FIGS. 8 and 9 ).
  • the cooling of stream 53 in heat exchanger 10 in FIG. 10 may be accomplished or supplemented by additional process streams (such as streams 39 , 40 , and 41 in FIGS. 2 , 4 , and 6 through 9 ) and/or external refrigeration.
  • separator 11 in FIGS. 2 , 4 , and 6 through 10 may not be needed.
  • the cooled feed stream 31 a leaving heat exchanger 10 in FIGS. 2 , 4 , and 6 through 9 or the cooled stream 53 a leaving heat exchanger 10 in FIG. 10 may not contain any liquid (because it is above its dewpoint, or because it is above its cricondenbar), so that separator 11 shown in FIGS. 2 , 4 , and 6 through 10 is not required.
  • the high pressure liquid (stream 33 ) in FIGS. 2 , 4 , and 6 through 9 need not be expanded and fed to a mid-column feed point on the distillation column. Instead, all or a portion of it (dashed stream 47 ) may be combined with the portion of the separator vapor (stream 34 ) to form combined stream 35 that flows to heat exchanger 15 . Any remaining portion of the liquid (dashed stream 37 ) may be expanded through an appropriate expansion device, such as expansion valve 12 , to form stream 37 a which is then fed to a mid-column feed point on distillation column 19 ( FIGS. 2 , 4 , 6 , and 7 ) or stripper column 29 ( FIGS. 8 and 9 ). Stream 33 in FIGS. 2 , 4 , and 6 through 9 and/or stream 37 in FIGS. 2 , 4 , and 6 through 10 may also be used for inlet gas cooling or other heat exchange service before or after the expansion step prior to flowing to the demethanizer.
  • the use of external refrigeration to supplement the cooling available to the inlet gas and/or the distillation stream from other process streams may be employed, particularly in the case of a rich inlet gas.
  • the use and distribution of separator liquids and demethanizer side draw liquids for process heat exchange, and the particular arrangement of heat exchangers for inlet gas cooling must be evaluated for each particular application, as well as the choice of process streams for specific heat exchange services.
  • 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 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 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.

Landscapes

  • Engineering & Computer Science (AREA)
  • Physics & Mathematics (AREA)
  • Mechanical Engineering (AREA)
  • Thermal Sciences (AREA)
  • General Engineering & Computer Science (AREA)
  • Chemical & Material Sciences (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • General Chemical & Material Sciences (AREA)
  • Oil, Petroleum & Natural Gas (AREA)
  • Separation By Low-Temperature Treatments (AREA)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
US11/839,693 2006-09-28 2007-08-16 Hydrocarbon Gas Processing Abandoned US20080078205A1 (en)

Priority Applications (14)

Application Number Priority Date Filing Date Title
US11/839,693 US20080078205A1 (en) 2006-09-28 2007-08-16 Hydrocarbon Gas Processing
CN2007800359949A CN101517340B (zh) 2006-09-28 2007-08-17 烃类气体处理
MYPI20091040A MY157917A (en) 2006-09-28 2007-08-17 Hydrocarbon gas processing
BRPI0717082-3A BRPI0717082B1 (pt) 2006-09-28 2007-08-17 Processo para a separação de um fluxo de gás
MX2009002053A MX2009002053A (es) 2006-09-28 2007-08-17 Procesamiento de gas hidrocarbonado.
AU2007305167A AU2007305167B2 (en) 2006-09-28 2007-08-17 Hydrocarbon gas processing
PCT/US2007/076199 WO2008042509A2 (en) 2006-09-28 2007-08-17 Hydrocarbon gas processing
CA2664224A CA2664224C (en) 2006-09-28 2007-08-17 Hydrocarbon gas processing
PE2007001299A PE20080652A1 (es) 2006-09-28 2007-09-26 Proceso y aparato para la recuperacion de etano, etileno, propano, propileno e hidrocarburos mas pesados de una corriente de gas de hidrocarburos
CL2007002773A CL2007002773A1 (es) 2006-09-28 2007-09-26 Proceso y aparato para la recuperacion de etano, e hidrocarburos superiores de una corriente de gas, que comprende etapas de enfriamiento, division de corrientes, expansion, franccionamiento e intercambio de calor entre corriente de destilacion de vapor comprimida y corriente de vapor de cabeza de destilacion.
ARP070104325A AR063066A1 (es) 2006-09-28 2007-09-28 Procesamiento de gas de hidrocarburos
SA7280532A SA07280532B1 (ar) 2006-09-28 2007-09-29 عملية معالجة غاز الهيدروكربون
NO20090789A NO20090789L (no) 2006-09-28 2009-02-19 Hydrokarbongassprosessering
TN2009000078A TN2009000078A1 (en) 2006-09-28 2009-03-10 Hydrocarbon gas processing

Applications Claiming Priority (3)

Application Number Priority Date Filing Date Title
US84829906P 2006-09-28 2006-09-28
US89768307P 2007-01-25 2007-01-25
US11/839,693 US20080078205A1 (en) 2006-09-28 2007-08-16 Hydrocarbon Gas Processing

Publications (1)

Publication Number Publication Date
US20080078205A1 true US20080078205A1 (en) 2008-04-03

Family

ID=39259824

Family Applications (1)

Application Number Title Priority Date Filing Date
US11/839,693 Abandoned US20080078205A1 (en) 2006-09-28 2007-08-16 Hydrocarbon Gas Processing

Country Status (14)

Country Link
US (1) US20080078205A1 (pt)
CN (1) CN101517340B (pt)
AR (1) AR063066A1 (pt)
AU (1) AU2007305167B2 (pt)
BR (1) BRPI0717082B1 (pt)
CA (1) CA2664224C (pt)
CL (1) CL2007002773A1 (pt)
MX (1) MX2009002053A (pt)
MY (1) MY157917A (pt)
NO (1) NO20090789L (pt)
PE (1) PE20080652A1 (pt)
SA (1) SA07280532B1 (pt)
TN (1) TN2009000078A1 (pt)
WO (1) WO2008042509A2 (pt)

Cited By (38)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US20080190136A1 (en) * 2007-02-09 2008-08-14 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing
US20080293086A1 (en) * 2006-09-18 2008-11-27 Cobalt Technologies, Inc. A Delaware Corporation Real time monitoring of microbial enzymatic pathways
US20090081715A1 (en) * 2007-09-07 2009-03-26 Cobalt Technologies, Inc., A Delaware Corporation Engineered Light-Emitting Reporter Genes
US20090100862A1 (en) * 2007-10-18 2009-04-23 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing
WO2010017061A1 (en) * 2008-08-06 2010-02-11 Ortloff Engineers, Ltd. Liquefied natural gas production
WO2010096223A1 (en) * 2009-02-17 2010-08-26 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US20100251764A1 (en) * 2009-02-17 2010-10-07 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing
US20100258401A1 (en) * 2007-01-10 2010-10-14 Pilot Energy Solutions, Llc Carbon Dioxide Fractionalization Process
US20100275647A1 (en) * 2009-02-17 2010-11-04 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing
US20100287983A1 (en) * 2009-02-17 2010-11-18 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing
US20100287984A1 (en) * 2009-02-17 2010-11-18 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US20100330633A1 (en) * 2009-06-26 2010-12-30 Cobalt Technologies, Inc. Integrated System and Process for Bioproduct Production
US20100326134A1 (en) * 2009-02-17 2010-12-30 Ortloff Engineers Ltd. Hydrocarbon Gas Processing
US20110067442A1 (en) * 2009-09-21 2011-03-24 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing
US20110226013A1 (en) * 2010-03-31 2011-09-22 S.M.E. Products Lp Hydrocarbon Gas Processing
US20110226011A1 (en) * 2010-03-31 2011-09-22 S.M.E. Products Lp Hydrocarbon Gas Processing
US20110226014A1 (en) * 2010-03-31 2011-09-22 S.M.E. Products Lp Hydrocarbon Gas Processing
US20110232328A1 (en) * 2010-03-31 2011-09-29 S.M.E. Products Lp Hydrocarbon Gas Processing
WO2011123289A1 (en) * 2010-03-31 2011-10-06 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US8460906B2 (en) 2008-04-09 2013-06-11 Cobalt Technologies, Inc. Immobilized product tolerant microorganisms
US20140060113A1 (en) * 2012-09-04 2014-03-06 Linde Aktiengesellschaft Method for separating c2+-hydrocarbons or c3+-hydrocarbons from a hydrocarbon-rich fraction
US20140060112A1 (en) * 2012-09-04 2014-03-06 Linde Aktiengesellschaft Method for separating c2+-hydrocarbons from a hydrocarbon-rich fraction
US8667812B2 (en) 2010-06-03 2014-03-11 Ordoff Engineers, Ltd. Hydrocabon gas processing
US8794030B2 (en) 2009-05-15 2014-08-05 Ortloff Engineers, Ltd. Liquefied natural gas and hydrocarbon gas processing
US8850849B2 (en) 2008-05-16 2014-10-07 Ortloff Engineers, Ltd. Liquefied natural gas and hydrocarbon gas processing
AU2011233590B2 (en) * 2010-03-31 2015-02-26 Uop Llc Hydrocarbon gas processing
US9021832B2 (en) 2010-01-14 2015-05-05 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US9052137B2 (en) 2009-02-17 2015-06-09 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US9637428B2 (en) 2013-09-11 2017-05-02 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US9783470B2 (en) 2013-09-11 2017-10-10 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US9790147B2 (en) 2013-09-11 2017-10-17 Ortloff Engineers, Ltd. Hydrocarbon processing
US10533794B2 (en) 2016-08-26 2020-01-14 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US10551118B2 (en) 2016-08-26 2020-02-04 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US10551119B2 (en) 2016-08-26 2020-02-04 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US11428465B2 (en) 2017-06-01 2022-08-30 Uop Llc Hydrocarbon gas processing
US11543180B2 (en) 2017-06-01 2023-01-03 Uop Llc Hydrocarbon gas processing
US11578915B2 (en) * 2019-03-11 2023-02-14 Uop Llc Hydrocarbon gas processing
US11643604B2 (en) 2019-10-18 2023-05-09 Uop Llc Hydrocarbon gas processing

Families Citing this family (3)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN104263402A (zh) * 2014-09-19 2015-01-07 华南理工大学 一种利用能量集成高效回收管输天然气中轻烃的方法
KR20180064471A (ko) * 2015-10-06 2018-06-14 엑손모빌 업스트림 리서치 캄파니 탄화수소 처리 플랜트에 있는 통합된 냉각 및 액화 모듈
CN110118468B (zh) * 2019-05-10 2020-02-11 西南石油大学 一种带自冷循环适用于富气的乙烷回收方法

Citations (12)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US6250105B1 (en) * 1998-12-18 2001-06-26 Exxonmobil Upstream Research Company Dual multi-component refrigeration cycles for liquefaction of natural gas
US6269655B1 (en) * 1998-12-09 2001-08-07 Mark Julian Roberts Dual mixed refrigerant cycle for gas liquefaction
US6272882B1 (en) * 1997-12-12 2001-08-14 Shell Research Limited Process of liquefying a gaseous, methane-rich feed to obtain liquefied natural gas
US6308531B1 (en) * 1999-10-12 2001-10-30 Air Products And Chemicals, Inc. Hybrid cycle for the production of liquefied natural gas
US6336344B1 (en) * 1999-05-26 2002-01-08 Chart, Inc. Dephlegmator process with liquid additive
US6347532B1 (en) * 1999-10-12 2002-02-19 Air Products And Chemicals, Inc. Gas liquefaction process with partial condensation of mixed refrigerant at intermediate temperatures
US6363744B2 (en) * 2000-01-07 2002-04-02 Costain Oil Gas & Process Limited Hydrocarbon separation process and apparatus
US6526777B1 (en) * 2001-04-20 2003-03-04 Elcor Corporation LNG production in cryogenic natural gas processing plants
US6578379B2 (en) * 2000-12-13 2003-06-17 Technip-Coflexip Process and installation for separation of a gas mixture containing methane by distillation
US20030158458A1 (en) * 2002-02-20 2003-08-21 Eric Prim System and method for recovery of C2+ hydrocarbons contained in liquefied natural gas
US20060032269A1 (en) * 2003-02-25 2006-02-16 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US20080190136A1 (en) * 2007-02-09 2008-08-14 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing

Family Cites Families (3)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US20020166336A1 (en) * 2000-08-15 2002-11-14 Wilkinson John D. Hydrocarbon gas processing
US6742358B2 (en) * 2001-06-08 2004-06-01 Elkcorp Natural gas liquefaction
US6945075B2 (en) * 2002-10-23 2005-09-20 Elkcorp Natural gas liquefaction

Patent Citations (12)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US6272882B1 (en) * 1997-12-12 2001-08-14 Shell Research Limited Process of liquefying a gaseous, methane-rich feed to obtain liquefied natural gas
US6269655B1 (en) * 1998-12-09 2001-08-07 Mark Julian Roberts Dual mixed refrigerant cycle for gas liquefaction
US6250105B1 (en) * 1998-12-18 2001-06-26 Exxonmobil Upstream Research Company Dual multi-component refrigeration cycles for liquefaction of natural gas
US6336344B1 (en) * 1999-05-26 2002-01-08 Chart, Inc. Dephlegmator process with liquid additive
US6308531B1 (en) * 1999-10-12 2001-10-30 Air Products And Chemicals, Inc. Hybrid cycle for the production of liquefied natural gas
US6347532B1 (en) * 1999-10-12 2002-02-19 Air Products And Chemicals, Inc. Gas liquefaction process with partial condensation of mixed refrigerant at intermediate temperatures
US6363744B2 (en) * 2000-01-07 2002-04-02 Costain Oil Gas & Process Limited Hydrocarbon separation process and apparatus
US6578379B2 (en) * 2000-12-13 2003-06-17 Technip-Coflexip Process and installation for separation of a gas mixture containing methane by distillation
US6526777B1 (en) * 2001-04-20 2003-03-04 Elcor Corporation LNG production in cryogenic natural gas processing plants
US20030158458A1 (en) * 2002-02-20 2003-08-21 Eric Prim System and method for recovery of C2+ hydrocarbons contained in liquefied natural gas
US20060032269A1 (en) * 2003-02-25 2006-02-16 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US20080190136A1 (en) * 2007-02-09 2008-08-14 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing

Cited By (87)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US20080293086A1 (en) * 2006-09-18 2008-11-27 Cobalt Technologies, Inc. A Delaware Corporation Real time monitoring of microbial enzymatic pathways
US9481834B2 (en) 2007-01-10 2016-11-01 Pilot Energy Solutions, Llc Carbon dioxide fractionalization process
US10316260B2 (en) 2007-01-10 2019-06-11 Pilot Energy Solutions, Llc Carbon dioxide fractionalization process
USRE44462E1 (en) 2007-01-10 2013-08-27 Pilot Energy Solutions, Llc Carbon dioxide fractionalization process
US8709215B2 (en) 2007-01-10 2014-04-29 Pilot Energy Solutions, Llc Carbon dioxide fractionalization process
US20100258401A1 (en) * 2007-01-10 2010-10-14 Pilot Energy Solutions, Llc Carbon Dioxide Fractionalization Process
US8590340B2 (en) 2007-02-09 2013-11-26 Ortoff Engineers, Ltd. Hydrocarbon gas processing
US20080190136A1 (en) * 2007-02-09 2008-08-14 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing
US20090081715A1 (en) * 2007-09-07 2009-03-26 Cobalt Technologies, Inc., A Delaware Corporation Engineered Light-Emitting Reporter Genes
US8919148B2 (en) 2007-10-18 2014-12-30 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US20090100862A1 (en) * 2007-10-18 2009-04-23 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing
US8460906B2 (en) 2008-04-09 2013-06-11 Cobalt Technologies, Inc. Immobilized product tolerant microorganisms
US8850849B2 (en) 2008-05-16 2014-10-07 Ortloff Engineers, Ltd. Liquefied natural gas and hydrocarbon gas processing
WO2010017061A1 (en) * 2008-08-06 2010-02-11 Ortloff Engineers, Ltd. Liquefied natural gas production
US20110120183A9 (en) * 2008-08-06 2011-05-26 Ortloff Engineers, Ltd. Liquefied natural gas and hydrocarbon gas processing
US8584488B2 (en) 2008-08-06 2013-11-19 Ortloff Engineers, Ltd. Liquefied natural gas production
EA018269B1 (ru) * 2008-08-06 2013-06-28 Ортлофф Инджинирс, Лтд. Получение сжиженного природного газа
US20100031700A1 (en) * 2008-08-06 2010-02-11 Ortloff Engineers, Ltd. Liquefied natural gas and hydrocarbon gas processing
CN102112829A (zh) * 2008-08-06 2011-06-29 奥特洛夫工程有限公司 液化天然气生产
US9052137B2 (en) 2009-02-17 2015-06-09 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US8881549B2 (en) 2009-02-17 2014-11-11 Ortloff Engineers, Ltd. Hydrocarbon gas processing
KR101619563B1 (ko) * 2009-02-17 2016-05-10 오르트로프 엔지니어스, 리미티드 탄화수소 가스 처리
EA022672B1 (ru) * 2009-02-17 2016-02-29 Ортлофф Инджинирс, Лтд. Обработка углеводородного газа
US9080811B2 (en) 2009-02-17 2015-07-14 Ortloff Engineers, Ltd Hydrocarbon gas processing
US20100326134A1 (en) * 2009-02-17 2010-12-30 Ortloff Engineers Ltd. Hydrocarbon Gas Processing
US9021831B2 (en) 2009-02-17 2015-05-05 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US20100275647A1 (en) * 2009-02-17 2010-11-04 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing
CN102317725A (zh) * 2009-02-17 2012-01-11 奥特洛夫工程有限公司 烃气体加工
US20100287983A1 (en) * 2009-02-17 2010-11-18 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing
US20100251764A1 (en) * 2009-02-17 2010-10-07 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing
US20100236285A1 (en) * 2009-02-17 2010-09-23 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing
WO2010096223A1 (en) * 2009-02-17 2010-08-26 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US20100287984A1 (en) * 2009-02-17 2010-11-18 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US9939195B2 (en) 2009-02-17 2018-04-10 Ortloff Engineers, Ltd. Hydrocarbon gas processing including a single equipment item processing assembly
US9939196B2 (en) 2009-02-17 2018-04-10 Ortloff Engineers, Ltd. Hydrocarbon gas processing including a single equipment item processing assembly
US9933207B2 (en) 2009-02-17 2018-04-03 Ortloff Engineers, Ltd. Hydrocarbon gas processing
AU2010216329B2 (en) * 2009-02-17 2013-11-14 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US8794030B2 (en) 2009-05-15 2014-08-05 Ortloff Engineers, Ltd. Liquefied natural gas and hydrocarbon gas processing
US9074173B2 (en) 2009-06-26 2015-07-07 Cobalt Technologies Inc. Integrated system and process for bioproduct production
US20100330633A1 (en) * 2009-06-26 2010-12-30 Cobalt Technologies, Inc. Integrated System and Process for Bioproduct Production
US8497105B2 (en) 2009-06-26 2013-07-30 Cobalt Technologies, Inc. Integrated system and process for bioproduct production
WO2011049672A1 (en) * 2009-09-21 2011-04-28 Ortloff Engineer, Ltd. Hydrocarbon gas processing
US20110067443A1 (en) * 2009-09-21 2011-03-24 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing
JP2013505421A (ja) * 2009-09-21 2013-02-14 オートロフ・エンジニアーズ・リミテッド 炭化水素ガス処理
US20110067442A1 (en) * 2009-09-21 2011-03-24 Ortloff Engineers, Ltd. Hydrocarbon Gas Processing
CN102498360A (zh) * 2009-09-21 2012-06-13 奥特洛夫工程有限公司 碳氢化合物气体处理
WO2011034709A1 (en) * 2009-09-21 2011-03-24 Ortloff Engineers, Ltd. Hydrocarbon gas processing
WO2011034710A1 (en) * 2009-09-21 2011-03-24 Ortloff Engineers, Ltd. Hydrocarbon gas processing
KR101619568B1 (ko) 2009-09-21 2016-05-10 오르트로프 엔지니어스, 리미티드 탄화수소 가스 처리공정
EA028835B1 (ru) * 2009-09-21 2018-01-31 Ортлофф Инджинирс, Лтд. Переработка углеводородного газа
AU2010308519B2 (en) * 2009-09-21 2015-05-07 Uop Llc Hydrocarbon gas processing
EA021947B1 (ru) * 2009-09-21 2015-10-30 Ортлофф Инджинирс, Лтд. Переработка углеводородного газа
US9476639B2 (en) 2009-09-21 2016-10-25 Ortloff Engineers, Ltd. Hydrocarbon gas processing featuring a compressed reflux stream formed by combining a portion of column residue gas with a distillation vapor stream withdrawn from the side of the column
JP2013505422A (ja) * 2009-09-21 2013-02-14 オートロフ・エンジニアーズ・リミテッド 炭化水素ガス処理
AU2010295869B2 (en) * 2009-09-21 2015-07-09 Ortloff Engineers, Ltd. Hydrocarbon gas processing
EA024075B1 (ru) * 2009-09-21 2016-08-31 Ортлофф Инджинирс, Лтд. Переработка углеводородного газа
US9021832B2 (en) 2010-01-14 2015-05-05 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US9057558B2 (en) * 2010-03-31 2015-06-16 Ortloff Engineers, Ltd. Hydrocarbon gas processing including a single equipment item processing assembly
US20110232328A1 (en) * 2010-03-31 2011-09-29 S.M.E. Products Lp Hydrocarbon Gas Processing
EA023957B1 (ru) * 2010-03-31 2016-07-29 Ортлофф Инджинирс, Лтд. Переработка углеводородного газа
US20110226013A1 (en) * 2010-03-31 2011-09-22 S.M.E. Products Lp Hydrocarbon Gas Processing
US9074814B2 (en) 2010-03-31 2015-07-07 Ortloff Engineers, Ltd. Hydrocarbon gas processing
KR20120139656A (ko) * 2010-03-31 2012-12-27 에스.엠.이. 프로덕츠 엘피 탄화수소 가스 처리방법
WO2011123289A1 (en) * 2010-03-31 2011-10-06 Ortloff Engineers, Ltd. Hydrocarbon gas processing
AU2011233590B2 (en) * 2010-03-31 2015-02-26 Uop Llc Hydrocarbon gas processing
US9068774B2 (en) * 2010-03-31 2015-06-30 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US20110226014A1 (en) * 2010-03-31 2011-09-22 S.M.E. Products Lp Hydrocarbon Gas Processing
US9052136B2 (en) 2010-03-31 2015-06-09 Ortloff Engineers, Ltd. Hydrocarbon gas processing
KR101714102B1 (ko) 2010-03-31 2017-03-08 오르트로프 엔지니어스, 리미티드 탄화수소 가스 처리방법
US20110226011A1 (en) * 2010-03-31 2011-09-22 S.M.E. Products Lp Hydrocarbon Gas Processing
US8667812B2 (en) 2010-06-03 2014-03-11 Ordoff Engineers, Ltd. Hydrocabon gas processing
US9389015B2 (en) * 2012-09-04 2016-07-12 Linde Aktiengesellschaft Method for separating C2+-hydrocarbons or C3+-hydrocarbons from a hydrocarbon-rich fraction
US20140060113A1 (en) * 2012-09-04 2014-03-06 Linde Aktiengesellschaft Method for separating c2+-hydrocarbons or c3+-hydrocarbons from a hydrocarbon-rich fraction
US20140060112A1 (en) * 2012-09-04 2014-03-06 Linde Aktiengesellschaft Method for separating c2+-hydrocarbons from a hydrocarbon-rich fraction
US10793492B2 (en) 2013-09-11 2020-10-06 Ortloff Engineers, Ltd. Hydrocarbon processing
US9927171B2 (en) 2013-09-11 2018-03-27 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US10227273B2 (en) 2013-09-11 2019-03-12 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US9637428B2 (en) 2013-09-11 2017-05-02 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US9783470B2 (en) 2013-09-11 2017-10-10 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US9790147B2 (en) 2013-09-11 2017-10-17 Ortloff Engineers, Ltd. Hydrocarbon processing
US10533794B2 (en) 2016-08-26 2020-01-14 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US10551118B2 (en) 2016-08-26 2020-02-04 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US10551119B2 (en) 2016-08-26 2020-02-04 Ortloff Engineers, Ltd. Hydrocarbon gas processing
US11428465B2 (en) 2017-06-01 2022-08-30 Uop Llc Hydrocarbon gas processing
US11543180B2 (en) 2017-06-01 2023-01-03 Uop Llc Hydrocarbon gas processing
US11578915B2 (en) * 2019-03-11 2023-02-14 Uop Llc Hydrocarbon gas processing
US11643604B2 (en) 2019-10-18 2023-05-09 Uop Llc Hydrocarbon gas processing

Also Published As

Publication number Publication date
CL2007002773A1 (es) 2008-01-18
CN101517340B (zh) 2011-07-27
MY157917A (en) 2016-08-15
AU2007305167A1 (en) 2008-04-10
CA2664224C (en) 2015-02-24
BRPI0717082A2 (pt) 2013-10-29
MX2009002053A (es) 2009-03-09
BRPI0717082B1 (pt) 2020-02-18
CA2664224A1 (en) 2008-04-10
WO2008042509A2 (en) 2008-04-10
TN2009000078A1 (en) 2010-08-19
CN101517340A (zh) 2009-08-26
SA07280532B1 (ar) 2009-12-22
AU2007305167B2 (en) 2012-02-23
WO2008042509A3 (en) 2008-07-03
AR063066A1 (es) 2008-12-23
PE20080652A1 (es) 2008-06-14
NO20090789L (no) 2009-06-22

Similar Documents

Publication Publication Date Title
CA2664224C (en) Hydrocarbon gas processing
US10753678B2 (en) Hydrocarbon gas processing
US7191617B2 (en) Hydrocarbon gas processing
US9476639B2 (en) Hydrocarbon gas processing featuring a compressed reflux stream formed by combining a portion of column residue gas with a distillation vapor stream withdrawn from the side of the column
US8919148B2 (en) Hydrocarbon gas processing
US8590340B2 (en) Hydrocarbon gas processing
US9021832B2 (en) Hydrocarbon gas processing
US20190170435A1 (en) Hydrocarbon Gas Processing
US11578915B2 (en) Hydrocarbon gas processing
US11643604B2 (en) Hydrocarbon gas processing
US20210116174A1 (en) Hydrocarbon gas processing

Legal Events

Date Code Title Description
AS Assignment

Owner name: ORTLOFF ENGINEERS, LTD., TEXAS

Free format text: ASSIGNMENT OF ASSIGNORS INTEREST;ASSIGNORS:CUELLAR, KYLE T., MR.;MARTINEZ, TONY L., MR.;WILKINSON, JOHN D., MR.;AND OTHERS;REEL/FRAME:019939/0685;SIGNING DATES FROM 20070928 TO 20071002

STPP Information on status: patent application and granting procedure in general

Free format text: RESPONSE TO NON-FINAL OFFICE ACTION ENTERED AND FORWARDED TO EXAMINER

STPP Information on status: patent application and granting procedure in general

Free format text: FINAL REJECTION MAILED

STPP Information on status: patent application and granting procedure in general

Free format text: RESPONSE AFTER FINAL ACTION FORWARDED TO EXAMINER

STPP Information on status: patent application and granting procedure in general

Free format text: DOCKETED NEW CASE - READY FOR EXAMINATION

STPP Information on status: patent application and granting procedure in general

Free format text: NON FINAL ACTION MAILED

STCV Information on status: appeal procedure

Free format text: NOTICE OF APPEAL FILED

AS Assignment

Owner name: UOP LLC, ILLINOIS

Free format text: ASSIGNMENT OF ASSIGNORS INTEREST;ASSIGNOR:ORTLOFF ENGINEERS, LTD.;REEL/FRAME:054188/0807

Effective date: 20200918

STCB Information on status: application discontinuation

Free format text: ABANDONED -- FAILURE TO RESPOND TO AN OFFICE ACTION