Patent Application
20110067443
This invention relates to a process and an apparatus for the separation of a gas containing hydrocarbons. The applicants claim the benefits under Title 35, United States Code, Section 119(e) of prior U.S. Provisional Applications No. 61/244,181 which was filed on Sep. 21, 2009, No. 61/346,150 which was filed on May 19, 2010, and No. 61/351,045 which was filed on Jun. 3, 2010.
Ethylene, ethane, propylene, propane, and/or heavier hydrocarbons can be recovered from a variety of gases, such as natural gas, refinery gas, and synthetic gas streams obtained from other hydrocarbon materials such as coal, crude oil, naphtha, oil shale, tar sands, and lignite. Natural gas usually has a major proportion of methane and ethane, i.e., methane and ethane together comprise at least 50 mole percent of the gas. The gas also contains relatively lesser amounts of heavier hydrocarbons such as propane, butanes, pentanes, and the like, as well as hydrogen, nitrogen, carbon dioxide, and other gases.
The present invention is generally concerned with the recovery of ethylene, ethane, propylene, propane, and heavier hydrocarbons from such gas streams. A typical analysis of a gas stream to be processed in accordance with this invention would be, in approximate mole percent, 88.1% methane, 6.0% ethane and other C2 components, 2.5% propane and other C3 components, 0.2% iso-butane, 0.2% normal butane, and 0.5% pentanes plus, with the balance made up of nitrogen and carbon dioxide. Sulfur containing gases are also sometimes present.
The historically cyclic fluctuations in the prices of both natural gas and its natural gas liquid (NGL) constituents have at times reduced the incremental value of ethane, ethylene, propane, propylene, and heavier components as liquid products. This has resulted in a demand for processes that can provide more efficient recoveries of these products, for processes that can provide efficient recoveries with lower capital investment, and for processes that can be easily adapted or adjusted to vary the recovery of a specific component over a broad range. Available processes for separating these materials include those based upon cooling and refrigeration of gas, oil absorption, and refrigerated oil absorption. Additionally, cryogenic processes have become popular because of the availability of economical equipment that produces power while simultaneously expanding and extracting heat from the gas being processed. Depending upon the pressure of the gas source, the richness (ethane, ethylene, and heavier hydrocarbons content) of the gas, and the desired end products, each of these processes or a combination thereof may be employed.
The cryogenic expansion process is now generally preferred for natural gas liquids recovery because it provides maximum simplicity with ease of startup, operating flexibility, good efficiency, safety, and good reliability. U.S. Pat. Nos. 3,292,380; 4,061,481; 4,140,504; 4,157,904; 4,171,964; 4,185,978; 4,251,249; 4,278,457; 4,519,824; 4,617,039; 4,687,499; 4,689,063; 4,690,702; 4,854,955; 4,869,740; 4,889,545; 5,275,005; 5,555,748; 5,566,554; 5,568,737; 5,771,712; 5,799,507; 5,881,569; 5,890,378; 5,983,664; 6,182,469; 6,578,379; 6,712,880; 6,915,662; 7,191,617; 7,219,513; reissue U.S. Pat. No. 33,408; and co-pending application Ser. Nos. 11/430,412; 11/839,693; 11/971,491; 12/206,230; 12/689,616; 12/717,394; 12/750,862; 12/772,472; and 12/781,259 describe relevant processes (although the description of the present invention in some cases is based on different processing conditions than those described in the cited U.S. patents).
In a typical cryogenic expansion recovery process, a feed gas stream under pressure is cooled by heat exchange with other streams of the process and/or external sources of refrigeration such as a propane compression-refrigeration system. As the gas is cooled, liquids may be condensed and collected in one or more separators as high-pressure liquids containing some of the desired C2+ components. Depending on the richness of the gas and the amount of liquids formed, the high-pressure liquids may be expanded to a lower pressure and fractionated. The vaporization occurring during expansion of the liquids results in further cooling of the stream. Under some conditions, pre-cooling the high pressure liquids prior to the expansion may be desirable in order to further lower the temperature resulting from the expansion. The expanded stream, comprising a mixture of liquid and vapor, is fractionated in a distillation (demethanizer or deethanizer) column. In the column, the expansion cooled stream(s) is (are) distilled to separate residual methane, nitrogen, and other volatile gases as overhead vapor from the desired C2 components, C3 components, and heavier hydrocarbon components as bottom liquid product, or to separate residual methane, C2 components, nitrogen, and other volatile gases as overhead vapor from the desired C3 components and heavier hydrocarbon components as bottom liquid product.
If the feed gas is not totally condensed (typically it is not), the vapor remaining from the partial condensation can be split into two streams. One portion of the vapor is passed through a work expansion machine or engine, or an expansion valve, to a lower pressure at which additional liquids are condensed as a result of further cooling of the stream. The pressure after expansion is essentially the same as the pressure at which the distillation column is operated. The combined vapor-liquid phases resulting from the expansion are supplied as feed to the column.
The remaining portion of the vapor is cooled to substantial condensation by heat exchange with other process streams, e.g., the cold fractionation tower overhead. Some or all of the high-pressure liquid may be combined with this vapor portion prior to cooling. The resulting cooled stream is then expanded through an appropriate expansion device, such as an expansion valve, to the pressure at which the demethanizer is operated. During expansion, a portion of the liquid will vaporize, resulting in cooling of the total stream. The flash expanded stream is then supplied as top feed to the demethanizer. Typically, the vapor portion of the flash expanded stream and the demethanizer overhead vapor combine in an upper separator section in the fractionation tower as residual methane product gas. Alternatively, the cooled and expanded stream may be supplied to a separator to provide vapor and liquid streams. The vapor is combined with the tower overhead and the liquid is supplied to the column as a top column feed.
In the ideal operation of such a separation process, the residue gas leaving the process will contain substantially all of the methane in the feed gas with essentially none of the heavier hydrocarbon components, and the bottoms fraction leaving the demethanizer will contain substantially all of the heavier hydrocarbon components with essentially no methane or more volatile components. In practice, however, this ideal situation is not obtained because the conventional demethanizer is operated largely as a stripping column. The methane product of the process, therefore, typically comprises vapors leaving the top fractionation stage of the column, together with vapors not subjected to any rectification step. Considerable losses of C2, C3, and C4+ components occur because the top liquid feed contains substantial quantities of these components and heavier hydrocarbon components, resulting in corresponding equilibrium quantities of C2 components, C3 components, C4 components, and heavier hydrocarbon components in the vapors leaving the top fractionation stage of the demethanizer. The loss of these desirable components could be significantly reduced if the rising vapors could be brought into contact with a significant quantity of liquid (reflux) capable of absorbing the C2 components, C3 components, C4 components, and heavier hydrocarbon components from the vapors.
In recent years, the preferred processes for hydrocarbon separation use an upper absorber section to provide additional rectification of the rising vapors. The source of the reflux stream for the upper rectification section is typically a recycled stream of residue gas supplied under pressure. The recycled residue gas stream is usually cooled to substantial condensation by heat exchange with other process streams, e.g., the cold fractionation tower overhead. The resulting substantially condensed stream is then expanded through an appropriate expansion device, such as an expansion valve, to the pressure at which the demethanizer is operated. During expansion, a portion of the liquid will usually vaporize, resulting in cooling of the total stream. The flash expanded stream is then supplied as top feed to the demethanizer. Typically, the vapor portion of the expanded stream and the demethanizer overhead vapor combine in an upper separator section in the fractionation tower as residual methane product gas. Alternatively, the cooled and expanded stream may be supplied to a separator to provide vapor and liquid streams, so that thereafter the vapor is combined with the tower overhead and the liquid is supplied to the column as a top column feed. Typical process schemes of this type are disclosed in U.S. Pat. Nos. 4,889,545; 5,568,737; and 5,881,569; assignee's co-pending application Ser. No. 12/717,394; and in Mowrey, E. Ross, “Efficient, High Recovery of Liquids from Natural Gas Utilizing a High Pressure Absorber”, Proceedings of the Eighty-First Annual Convention of the Gas Processors Association, Dallas, Tex., Mar. 11-13, 2002. These processes use a compressor to provide the motive force for recycling the reflux stream to the demethanizer, adding to both the capital cost and the operating cost of facilities using these processes.
The present invention also employs an upper rectification section (or a separate rectification column if plant size or other factors favor using separate rectification and stripping columns). However, the reflux stream for this rectification section is provided by using a side draw of the vapors rising in a lower portion of the tower combined with a portion of the column overhead vapor. Because of the relatively high concentration of C2 components in the vapors lower in the tower, a significant quantity of liquid can be condensed from this combined vapor stream with only a modest elevation in pressure, often using only the refrigeration available in the remaining portion of the cold overhead vapor leaving the upper rectification section of the column. This condensed liquid, which is predominantly liquid methane, can then be used to absorb C2 components, C3 components, C4 components, and heavier hydrocarbon components from the vapors rising through the upper rectification section and thereby capture these valuable components in the bottom liquid product from the demethanizer.
Heretofore, compressing either a portion of the cold overhead vapor stream or compressing a side draw vapor stream to provide reflux for the upper rectification section of the column has been employed in C2+ recovery systems, as illustrated in assignee's U.S. Pat. No. 4,889,545 and assignee's co-pending application Ser. No. 11/839,693, respectively. Surprisingly, applicants have found that combining a portion of the cold overhead vapor with the side draw vapor stream and then compressing the combined stream improves the system efficiency while reducing operating cost.
In accordance with the present invention, it has been found that C2 recovery in excess of 95% and C3 and C4+ recoveries in excess of 99% can be obtained. In addition, the present invention makes possible essentially 100% separation of methane and lighter components from the C2 components and heavier components at lower energy requirements compared to the prior art while maintaining the recovery levels. The present invention, although applicable at lower pressures and warmer temperatures, is particularly advantageous when processing feed gases in the range of 400 to 1500 psia [2,758 to 10,342 kPa(a)] or higher under conditions requiring NGL recovery column overhead temperatures of −50° F. [−46° C.] or colder.
For a better understanding of the present invention, reference is made to the following examples and drawings. Referring to the drawings:
In the following explanation of the above figures, tables are provided summarizing flow rates calculated for representative process conditions. In the tables appearing herein, the values for flow rates (in moles per hour) have been rounded to the nearest whole number for convenience. The total stream rates shown in the tables include all non-hydrocarbon components and hence are generally larger than the sum of the stream flow rates for the hydrocarbon components. Temperatures indicated are approximate values rounded to the nearest degree. It should also be noted that the process design calculations performed for the purpose of comparing the processes depicted in the figures are based on the assumption of no heat leak from (or to) the surroundings to (or from) the process. The quality of commercially available insulating materials makes this a very reasonable assumption and one that is typically made by those skilled in the art.
For convenience, process parameters are reported in both the traditional British units and in the units of the Système International d'Unités (SI). The molar flow rates given in the tables may be interpreted as either pound moles per hour or kilogram moles per hour. The energy consumptions reported as horsepower (HP) and/or thousand British Thermal Units per hour (MBTU/Hr) correspond to the stated molar flow rates in pound moles per hour. The energy consumptions reported as kilowatts (kW) correspond to the stated molar flow rates in kilogram moles per hour.
The feed stream 31 is cooled in heat exchanger 10 by heat exchange with cool residue gas (stream 43a), liquid product at 72° F. [22° C.] (stream 42a), demethanizer reboiler liquids at 52° F. [11° C.] (stream 41), and demethanizer side reboiler liquids at −20° F. [−29° C.] (stream 40). Note that in all cases exchanger 10 is representative of either a multitude of individual heat exchangers or a single multi-pass heat exchanger, or any combination thereof. (The decision as to whether to use more than one heat exchanger for the indicated cooling services will depend on a number of factors including, but not limited to, inlet gas flow rate, heat exchanger size, stream temperatures, etc.) The cooled stream 31a enters separator 11 at −18° F. [−28° C.] and 1025 psia [7,067 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33). The separator liquid (stream 33) is expanded to the operating pressure (approximately 392 psia [2,701 kPa(a)]) of fractionation tower 17 by expansion valve 16, cooling stream 33a to −53° F. [−47° C.] before it is supplied to fractionation tower 17 at a lower mid-column feed point.
The vapor (stream 32) from separator 11 is divided into two streams, 36 and 37. Stream 36, containing about 38% of the total vapor, passes through heat exchanger 12 in heat exchange relation with the cold residue gas (stream 43) where it is cooled to substantial condensation. The resulting substantially condensed stream 36a at −142° F. [−96° C.] is then flash expanded through expansion valve 13 to slightly above the operating pressure of fractionation tower 17. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in
The remaining 62% of the vapor from separator 11 (stream 37) enters a work expansion machine 14 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 14 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 37a to a temperature of approximately −94° F. [−70° C.]. The typical commercially available expanders are capable of recovering on the order of 80-85% of the work theoretically available in an ideal isentropic expansion. The work recovered is often used to drive a centrifugal compressor (such as item 15) that can be used to re-compress the residue gas (stream 43b), for example. The partially condensed expanded stream 37a is thereafter supplied as feed to fractionation tower 17 at a mid-column feed point.
The demethanizer in tower 17 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. The demethanizer tower consists of two sections: an upper absorbing (rectification) section 17a that contains the trays and/or packing to provide the necessary contact between the vapor portions of the expanded streams 36c and 37a rising upward and cold liquid falling downward to condense and absorb the C2 components, C3 components, and heavier components; and a lower, stripping 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 demethanizing section 17b also includes one or more reboilers (such as the reboiler and side reboiler 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 37a enters demethanizer 17 at an intermediate feed position located in the lower region of absorbing section 17a of demethanizer 17. The liquid portion of the expanded stream 37a commingles with liquids falling downward from absorbing section 17a and the combined liquid continues downward into stripping section 17b of demethanizer 17. The vapor portion of the expanded stream 37a rises upward through absorbing section 17a and is contacted with cold liquid falling downward to condense and absorb the C2 components, C3 components, and heavier components.
In stripping section 17b of demethanizer 17, the feed streams are stripped of their methane and lighter components. The resulting liquid product (stream 42) exits the bottom of tower 17 at 67° F. [19° C.] (based on a typical specification of a methane to ethane ratio of 0.015:1 on a volume basis in the bottom product) and is pumped to heat exchanger 10 by pump 20 to be heated to 116° F. [47° C.] as it provides cooling to the feed gas as described earlier.
Cold demethanizer overhead stream 39 exits the top of demethanizer 17 at −146° F. [−99° C.] and is divided into cold residue gas stream 43 and recycle stream 44. Recycle stream 44 is compressed to 492 psia [3,390 kPa(a)] by compressor 21 before entering heat exchanger 22. The compressed recycle stream 44a is cooled from −121° F. [−85° C.] to −140° F. [−96° C.] and substantially condensed by heat exchange with expanded substantially condensed stream 36b as described previously. The substantially condensed stream 44b is then expanded through an appropriate expansion device, such as expansion valve 23, to the demethanizer operating pressure, resulting in cooling of the total stream to −150° F. [−101° C.]. The expanded stream 44c is then supplied to fractionation tower 17 as the top column feed. The vapor portion of stream 44c combines with the vapors rising from the top fractionation stage of the column to form demethanizer overhead stream 39.
The cold residue gas stream 43 passes countercurrently to the incoming feed gas in heat exchanger 12 where it is heated to −26° F. [−32° C.] (stream 43a) and in heat exchanger 10 where it is heated to 98° F. [37° C.] (stream 43b). The residue gas is then re-compressed in two stages. The first stage is compressor 15 driven by expansion machine 14. The second stage is compressor 24 driven by a supplemental power source which compresses the residue gas (stream 43d) to sales line pressure. After cooling to 120° F. [49° C.] in discharge cooler 25, the residue gas product (stream 43e) flows to the sales gas pipeline at 1040 psia [7,171 kPa(a)], sufficient to meet line requirements (usually on the order of the inlet pressure).
A summary of stream flow rates and energy consumption for the process illustrated in
In the simulation of the
The vapor (stream 32) from separator 11 is divided into two streams, 34 and 37. Stream 34, containing about 28% of the total vapor, passes through heat exchanger 12 in heat exchange relation with the cold residue gas (stream 43) where it is cooled to substantial condensation. The resulting substantially condensed stream 36a at −140° F. [−96° C.] is then flash expanded through expansion valve 13 to the operating pressure of fractionation tower 17. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in
The remaining 72% of the vapor from separator 11 (stream 37) enters a work expansion machine 14 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 14 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 37a to a temperature of approximately −97° F. [−72° C.]. The partially condensed expanded stream 37a is thereafter supplied as feed to fractionation tower 17 at a mid-column feed point (located below the feed point of stream 36b).
The demethanizer in tower 17 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. The demethanizer tower consists of two sections: an upper absorbing (rectification) section 17a that contains the trays and/or packing to provide the necessary contact between the vapor portion of the expanded streams 36b and 37a rising upward and cold liquid falling downward to condense and absorb the C2 components, C3 components, and heavier components from the vapors rising upward; and a lower, stripping 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 demethanizing section 17b also includes one or more reboilers (such as the reboiler and side reboiler 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 37a enters demethanizer 17 at an intermediate feed position located in the lower region of absorbing section 17a of demethanizer 17. The liquid portion of the expanded stream 37a commingles with liquids falling downward from absorbing section 17a and the combined liquid continues downward into stripping section 17b of demethanizer 17. The vapor portion of the expanded stream 37a rises upward through absorbing section 17a and is contacted with cold liquid falling downward to condense and absorb the C2 components, C3 components, and heavier components.
A portion of the distillation vapor (stream 45) is withdrawn from the upper region of absorbing section 17a in fractionation column 17, above the feed position of expanded stream 36b in the middle region of absorbing section 17a. The distillation vapor stream 45 at −142° F. [−96° C.] is combined with a first portion (stream 44) of overhead vapor stream 39 at −144° F. [−98° C.] to form combined vapor stream 46 at −144° F. [−98° C.]. The combined vapor stream 46 is compressed to 686 psia [4,728 kPa(a)] by reflux compressor 21, then cooled from −84° F. [−65° C.] to −140° F. [−96° C.] and substantially condensed (stream 46b) in heat exchanger 12 by heat exchange with cold residue gas stream 43, the remaining second portion of demethanizer overhead stream 39 exiting the top of demethanizer 17.
The substantially condensed stream 46b is flash expanded to the operating pressure of demethanizer 17 by expansion valve 23. A portion of the stream is vaporized, further cooling stream 46c to −149° F. [−101° C.] before it is supplied as cold top column feed (reflux) to demethanizer 17. This cold liquid reflux absorbs and condenses the C2 components, C3 components, and heavier components rising in the upper rectification region of absorbing section 17a of demethanizer 17.
In stripping section 17b of demethanizer 17, the feed streams are stripped of their methane and lighter components. The resulting liquid product (stream 42) exits the bottom of tower 17 at 69° F. [21° C.] (based on a typical specification of a methane to ethane ratio of 0.015:1 on a volume basis in the bottom product) and is pumped to heat exchanger 10 by pump 20 to be heated to 116° F. [47° C.] as it provides cooling to the feed gas as described earlier. The cold residue gas stream 43 passes countercurrently to the incoming feed gas and compressed combined vapor stream in heat exchanger 12 where it is heated to −37° F. [−39° C.] (stream 43a), and countercurrently to the incoming feed gas in heat exchanger 10 where it is heated to 97° F. [36° C.] (stream 43b) as it provides cooling as previously described. The residue gas is then re-compressed in two stages, compressor 15 driven by expansion machine 14 and compressor 24 driven by a supplemental power source. After stream 43d is cooled to 120° F. [49° C.] in discharge cooler 25, the residue gas product (stream 43e) flows to the sales gas pipeline at 1040 psia [7,171 kPa(a)], sufficient to meet line requirements (usually on the order of the inlet pressure).
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables I and II shows that the present invention maintains essentially the same recoveries as the prior art. However, further comparison of Tables I and II shows that the product yields were achieved using significantly less power than the prior art. In terms of the recovery efficiency (defined by the quantity of ethane recovered per unit of power), the present invention represents more than a 4% improvement over the prior art of the
Like the prior art of the
In accordance with this invention, it is generally advantageous to design the absorbing (rectification) section of the demethanizer to contain multiple theoretical separation stages. However, the benefits of the present invention can be achieved with as few as two theoretical stages. For instance, all or a part of the expanded reflux stream (stream 46c) leaving expansion valve 23 and all or a part of the expanded substantially condensed stream 36b from expansion valve 13 can be combined (such as in the piping joining the expansion valves 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 two streams, combined with contacting at least a portion of expanded stream 37a, shall be considered for the purposes of this invention as constituting an absorbing section.
Some circumstances may favor withdrawing the distillation vapor stream 45 in
As described earlier, the compressed combined vapor stream 46a is substantially condensed and the resulting condensate used to absorb valuable C2 components, C3 components, and heavier components from the vapors rising through absorbing section 17a of demethanizer 17 or through absorber column 17. However, the present invention is not limited to this embodiment. It may be advantageous, for instance, to treat only a portion of these vapors in this manner, or to use only a portion of the condensate as an absorbent, in cases where other design considerations indicate portions of the vapors or the condensate should bypass absorbing section 17a of demethanizer 17 or absorber column 17. Some circumstances may favor partial condensation, rather than substantial condensation, of compressed combined vapor stream 46a in heat exchanger 12. Other circumstances may favor that distillation vapor stream 45 be a total vapor side draw from fractionation column 17 or absorber column 17 rather than a partial vapor side draw. It should also be noted that, depending on the composition of the feed gas stream, it may be advantageous to use external refrigeration to provide partial cooling of compressed combined vapor stream 46a in heat exchanger 12.
Feed gas conditions, plant size, available equipment, or other factors may indicate that elimination of work expansion machine 14, or replacement with an alternate expansion device (such as an expansion valve), is feasible. Although individual stream expansion is depicted in particular expansion devices, alternative expansion means may be employed where appropriate. For example, conditions may warrant work expansion of the substantially condensed portion of the feed stream (stream 36a) or the substantially condensed reflux stream (stream 46b) leaving heat exchanger 12.
Depending on the quantity of heavier hydrocarbons in the feed gas and the feed gas pressure, the cooled feed stream 31a leaving heat exchanger 10 in
The high pressure liquid (stream 33 in
In accordance with the present invention, the use of external refrigeration to supplement the cooling available to the inlet gas from other process streams may be employed, particularly in the case of a rich inlet gas. The use and distribution of separator liquids and demethanizer side draw liquids for process heat exchange, and the particular arrangement of heat exchangers for inlet gas cooling must be evaluated for each particular application, as well as the choice of process streams for specific heat exchange services.
In accordance with the present invention, the splitting of the vapor feed may be accomplished in several ways. In the processes of
It will also be recognized that the relative amount of feed found in each branch of the split vapor feed will depend on several factors, including gas pressure, feed gas composition, the amount of heat which can economically be extracted from the feed, and the quantity of horsepower available. More feed to the top of the column may increase recovery while decreasing power recovered from the expander thereby increasing the recompression horsepower requirements. Increasing feed lower in the column reduces the horsepower consumption but may also reduce product recovery. The relative locations of the mid-column feeds may vary depending on inlet composition or other factors such as desired recovery levels and amount of liquid formed during inlet gas cooling. Moreover, two or more of the feed streams, or portions thereof, may be combined depending on the relative temperatures and quantities of individual streams, and the combined stream then fed to a mid-column feed position.
The present invention provides improved recovery of C2 components, C3 components, and heavier hydrocarbon components or of C3 components and heavier hydrocarbon components per amount of utility consumption required to operate the process. An improvement in utility consumption required for operating the demethanizer or deethanizer process may appear in the form of reduced power requirements for compression or re-compression, reduced power requirements for external refrigeration, reduced energy requirements for tower reboilers, or a combination thereof.
While there have been described what are believed to be preferred embodiments of the invention, those skilled in the art will recognize that other and further modifications may be made thereto, e.g. to adapt the invention to various conditions, types of feed, or other requirements without departing from the spirit of the present invention as defined by the following claims.
| Number | Date | Country | |
|---|---|---|---|
| 61244181 | Sep 2009 | US | |
| 61346150 | May 2010 | US | |
| 61351045 | Jun 2010 | US |
