Patent Application
20180347899
This invention relates to a process and apparatus for improving the separation of a gas containing hydrocarbons. Assignees S.M.E. Products LP and Ortloff Engineers, Ltd. were parties to a joint research agreement that was in effect before the invention of this application was made. The applicants claim the benefits under Title 35, United States Code, Section 119(e) of prior U.S. Provisional Application No. 62/513,851 which was filed on Jun. 1, 2017 and prior U.S. Provisional Application No. 62/667,833 which was filed on May 7, 2018.
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/or other gases.
The present invention is generally concerned with improving 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, 78.6% methane, 12.5% ethane and other C2 components, 4.9% propane and other C3 components, 0.6% iso-butane, 1.4% normal butane, and 1.1% pentanes plus, with the balance made up of nitrogen and carbon dioxide. Sulfur containing gases are also sometimes present.
The historically cyclic fluctuations in the prices of both natural gas and its natural gas liquid (NGL) constituents have at times reduced the incremental value of ethane, ethylene, propane, propylene, and heavier components as liquid products. This has resulted in a demand for processes that can provide more efficient recoveries of these products, for processes that can provide efficient recoveries with lower capital investment, and for processes that can be easily adapted or adjusted to vary the recovery of a specific component over a broad range. Available processes for separating these materials include those based upon cooling and refrigeration of gas, oil absorption, and refrigerated oil absorption. Additionally, cryogenic processes have become popular because of the availability of economical equipment that produces power while simultaneously expanding and extracting heat from the gas being processed. Depending upon the pressure of the gas source, the richness (ethane, ethylene, and heavier hydrocarbons content) of the gas, and the desired end products, each of these processes or a combination thereof may be employed.
The cryogenic expansion process is now generally preferred for natural gas liquids recovery because it provides maximum simplicity with ease of startup, operating flexibility, good efficiency, safety, and good reliability. U.S. Pat. Nos. 3,292,380; 4,061,481; 4,140,504; 4,157,904; 4,171,964; 4,185,978; 4,251,249; 4,278,457; 4,519,824; 4,617,039; 4,687,499; 4,689,063; 4,690,702; 4,854,955; 4,869,740; 4,889,545; 5,275,005; 5,555,748; 5,566,554; 5,568,737; 5,771,712; 5,799,507; 5,881,569; 5,890,378; 5,983,664; 6,182,469; 6,578,379; 6,712,880; 6,915,662; 7,191,617; 7,219,513; 8,590,340; 8,881,549; 8,919,148; 9,021,831; 9,021,832; 9,052,136; 9,052,137; 9,057,558; 9,068,774; 9,074,814; 9,080,810; 9,080,811; 9,476,639; 9,637,428; 9,783,470; 9,927,171; 9,933,207; and 9,939,195; reissue U.S. Pat. No. 33,408; and co-pending application Ser. Nos. 11/839,693; 12/868,993; 12/869,139; 14/714,912; 14/828,093; 15/259,891; 15/332,670; 15/332,706; 15/332,723; and 15/668,139 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 and co-pending applications).
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. For many of these processes, the source of the reflux stream for the upper rectification section is 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. Typical process schemes of this type are disclosed in U.S. Pat. Nos. 4,889,545; 5,568,737; 5,881,569; 9,052,137; and 9,080,811 and in Mowrey, E. Ross, “Efficient, High Recovery of Liquids from Natural Gas Utilizing a High Pressure Absorber”, Proceedings of the Eighty-First Annual Convention of the Gas Processors Association, Dallas, Tex., Mar. 11-13, 2002. Unfortunately, in addition to the additional rectification section in the demethanizer, these processes also require surplus compression capacity 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.
Another means of providing a reflux stream for the upper rectification section is to withdraw a distillation vapor stream from a lower location on the tower (and perhaps combine it with a portion of the tower overhead vapor). This vapor (or combined vapor) stream is compressed to higher pressure, then cooled to substantial condensation, expanded to the tower operating pressure, and supplied as top feed to the tower. Typical process schemes of this type are disclosed in U.S. Pat. No. 9,476,639 and co-pending application Ser. Nos. 11/839,693; 12/869,139; and Ser. No. 15/259,891. These also require an additional rectification section in the demethanizer, plus a compressor to provide motive force for recycling the reflux stream to the demethanizer, again adding to both the capital cost and the operating cost of facilities using these processes.
However, there are many gas processing plants that have been built in the U.S. and other countries according to U.S. Pat. Nos. 4,157,904 and 4,278,457 (as well as other processes) that have no upper absorber section to provide additional rectification of the rising vapors and cannot be easily modified to add this feature. Also, these plants do not usually have surplus compression capacity to allow recycling a reflux stream. As a result, these plants are not as efficient when operated to recover C2 components and heavier components from the gas (commonly referred to as “ethane recovery”), and are particularly inefficient when operated to recover only the C3 components and heavier components from the gas (commonly referred to as “ethane rejection”).
The present invention is a novel means of providing additional rectification that can be easily added to existing gas processing plants to increase the recovery of the desired C2 components and/or C3 components without requiring additional residue gas compression. The incremental value of this increased recovery is often substantial. For the Examples given later, the incremental income from the additional recovery capability over that of the prior art is in the range of US$710,000 to US$4,720,000 [€590,000 to €3,930,000] per year using an average incremental value US$0.10-0.58 per gallon [€22-129 per m3] for hydrocarbon liquids compared to the corresponding hydrocarbon gases.
The present invention also combines what heretofore have been individual equipment items into a common housing, thereby reducing both the plot space requirements and the capital cost of the addition. Surprisingly, applicants have found that the more compact arrangement also significantly increases the product recovery at a given power consumption, thereby increasing the process efficiency and reducing the operating cost of the facility. In addition, the more compact arrangement also eliminates much of the piping used to interconnect the individual equipment items in traditional plant designs, further reducing capital cost and also eliminating the associated flanged piping connections. Since piping flanges are a potential leak source for hydrocarbons (which are volatile organic compounds, VOCs, that contribute to greenhouse gases and may also be precursors to atmospheric ozone formation), eliminating these flanges reduces the potential for atmospheric emissions that may damage the environment.
In accordance with the present invention, it has been found that C2 recoveries in excess of 99% can be obtained. Similarly, in those instances where recovery of C2 components is not desired, C3 recoveries in excess of 96% can be maintained. 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 39a), pumped liquid product at 20° F. [−7° C.] (stream 42a), demethanizer reboiler liquids at 0° F. [−18° C.] (stream 41), demethanizer side reboiler liquids at −45° F. [−43° C.] (stream 40), and propane refrigerant. Stream 31a then enters separator 11 at −29° F. [−34° C.] and 795 psia [5,479 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33).
The vapor (stream 32) from separator 11 is divided into two streams, 34 and 37. The liquid (stream 33) from separator 11 is optionally divided into two streams, 35 and 38. (Stream 35 may contain from 0% to 100% of the separator liquid in stream 33. If stream 35 contains any portion of the separator liquid, then the process of
The remaining 70% 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 −126° F. [−88° 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 39b), for example. The partially condensed expanded stream 37a is thereafter supplied as feed to fractionation tower 17 at an upper mid-column feed point. The remaining separator liquid in stream 38 (if any) is expanded to the operating pressure of fractionation tower 17 by expansion valve 16, cooling stream 38a to −85° F. [−65° C.] before it is supplied to fractionation tower 17 at a lower 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. As is often the case in natural gas processing plants, the fractionation tower may consist of two sections. The upper section 17a is a separator wherein the partially vaporized top feed is divided into its respective vapor and liquid portions, and wherein the vapor rising from the lower distillation or demethanizing section 17b is combined with the vapor portion of the top feed to form the cold demethanizer overhead vapor (stream 39) which exits the top of the tower. The lower, demethanizing section 17b contains the trays and/or packing and provides the necessary contact between the liquids falling downward and the vapors rising upward. The demethanizing section 17b also includes reboilers (such as the reboiler and the side reboiler described previously and supplemental reboiler 18) 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.
The liquid product stream 42 exits the bottom of the tower at 7° F. [−14° C.], based on a typical specification of a methane concentration of 0.5% on a volume basis in the bottom product. It is pumped to higher pressure by pump 21 (stream 42a) and then heated to 95° F. [35° C.] (stream 42b) as it provides cooling of the feed gas in heat exchanger 10 as described earlier. The residue gas (demethanizer overhead vapor stream 39) passes countercurrently to the incoming feed gas in heat exchanger 12 where it is heated from −176° F. [−115° C.] to −47° F. [−44° C.] (stream 39a) and in heat exchanger 10 where it is heated to 113° F. [45° C.] (stream 39b). 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 19 driven by a supplemental power source which compresses the residue gas (stream 39d) to sales line pressure. After cooling to 120° F. [49° C.] in discharge cooler 20, the residue gas product (stream 39e) flows to the sales gas pipeline at 765 psia [5,272 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 this simulation of the process, inlet gas enters the plant at 120° F. [49° C.] and 815 psia [5,617 kPa(a)] as stream 31 and is cooled in heat exchanger 10 by heat exchange with cool residue gas stream 39a and flashed separator liquids (stream 38a). (One consequence of operating the
The vapor (stream 32) from separator 11 is divided into two streams, 34 and 37, and the liquid (stream 33) is optionally divided into two streams, 35 and 38. For the process illustrated in
The remaining 67% 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 −103° F. [−75° C.] before it is supplied as feed to fractionation tower 17 at an upper mid-column feed point. The remaining separator liquid in stream 38 (if any) is expanded to slightly above the operating pressure of fractionation tower 17 by expansion valve 16, cooling stream 38a to −61° F. [−51° C.] before it is heated to 103° F. [39° C.] in heat exchanger 10 as described previously, with heated stream 40a then supplied to fractionation tower 17 at a lower mid-column feed point.
Note that when fractionation tower 17 is operated to reject the C2 components to the residue gas product as shown in
A summary of stream flow rates and energy consumption for the process illustrated in
Co-pending application Ser. No. 15/332,723 describes one means of improving the performance of the
Most of the process conditions shown for the
Substantially condensed stream 151b at −171° F. [−113° C.] is then flash expanded through expansion valve 23 to slightly above the operating pressure of fractionation tower 17. During expansion a portion of the stream may be vaporized, resulting in cooling of the total stream. In the process illustrated in
The flash expanded stream 151c is further vaporized as it provides cooling and partial condensation of the partially rectified vapor stream, and exits the heat and mass transfer means in rectifying section 117b at −178° F. [−117° C.]. The heated flash expanded stream discharges into separator section 117d of processing assembly 117 and is separated into its respective vapor and liquid phases. The vapor phase combines with the remaining portion (stream 152) of overhead vapor stream 39 to form a combined vapor stream that enters a mass transfer means in absorbing section 117c of processing assembly 117. The mass transfer means may consist of a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing, but could also be comprised of a non-heat transfer zone in a fin and tube type heat exchanger, a plate type heat exchanger, a brazed aluminum type heat exchanger, or other type of heat transfer device, including multi-pass and/or multi-service heat exchangers. The mass transfer means is configured to provide contact between the cold condensed liquid leaving the bottom of the heat and mass transfer means in rectifying section 117b and the combined vapor stream arising from separator section 117d. As the combined vapor stream rises upward through absorbing section 117c, it is contacted with the cold liquid falling downward to condense and absorb C2 components, C3 components, and heavier components from the combined vapor stream. The resulting partially rectified vapor stream is then directed to the heat and mass transfer means in rectifying section 117b of processing assembly 117 for further rectification as described previously.
The liquid phase (if any) from the heated flash expanded stream leaving rectifying section 117b of processing assembly 117 that is separated in separator section 117d combines with the distillation liquid leaving the bottom of the mass transfer means in absorbing section 117c of processing assembly 117 to form combined liquid stream 154. Combined liquid stream 154 leaves the bottom of processing assembly 117 and is pumped to higher pressure by pump 24 (stream 154a at −170° F. [−112° C.]). Further cooled stream 36b at −169° F. [−112° C.] is flash expanded through expansion valve 13 to the operating pressure of fractionation tower 17. During expansion a portion of the stream may be vaporized, resulting in cooling of the total stream to −177° F. [−116° C.]. Flash expanded stream 36c then joins with pumped stream 154a to form combined feed stream 155, which then enters fractionation column 17 at the top feed point at −176° F. [−116° C.].
The further rectified vapor stream leaves the heat and mass transfer means in rectifying section 117b of processing assembly 117 at −182° F. [−119° C.] and enters the heat exchange means in cooling section 117a of processing assembly 117. The vapor is heated to −96° F. [−71° C.] as it provides cooling to streams 36a and 151a as described previously. The heated vapor is then discharged from processing assembly 117 as cool residue gas stream 153, which is heated and compressed as described previously for stream 39 in the
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables I and III shows that, compared to the
The process of co-pending application Ser. No. 15/332,723 can also be operated to reject nearly all of the C2 components to the residue gas rather than recovering them in the liquid product. The operating conditions of the
Most of the process conditions shown for the
The flash expanded stream 36b is further vaporized as it provides cooling and partial condensation of the combined vapor stream, and exits the heat and mass transfer means in rectifying section 117b at −83° F. [−64° C.]. The heated flash expanded stream discharges into separator section 117d of processing assembly 117 and is separated into its respective vapor and liquid phases. The vapor phase combines with overhead vapor stream 39 to form the combined vapor stream that enters the mass transfer means in absorbing section 117c as described previously, and the liquid phase combines with the condensed liquid from the bottom of the mass transfer means in absorbing section 117c to form combined liquid stream 154. Combined liquid stream 154 leaves the bottom of processing assembly 117 and is pumped to higher pressure by pump 24 so that stream 154a at −73° F. [−58° C.] can enter fractionation column 17 at the top feed point. The further rectified vapor stream leaves the heat and mass transfer means in rectifying section 117b and discharges from processing assembly 117 at −104° F. [−76° C.] as cold residue gas stream 153, which is then heated and compressed as described previously for stream 39 in the
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables II and IV shows that, compared to the
In those cases where it is desirable to maximize the recovery of C2 components in the liquid product (as in the
Most of the process conditions shown for the
Absorbing section 117c inside processing assembly 117 contains a mass transfer means. This mass transfer means may consist of a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing, but could also be comprised of a non-heat transfer zone in a fin and tube type heat exchanger, a plate type heat exchanger, a brazed aluminum type heat exchanger, or other type of heat transfer device, including multi-pass and/or multi-service heat exchangers. The mass transfer means is configured to provide contact between cold condensed liquid leaving the bottom of a heat and mass transfer means in rectifying section 117b inside processing assembly 117 and column overhead vapor stream 39 arising from separator section 117d inside processing assembly 117. As the column overhead vapor stream rises upward through absorbing section 117c, it is contacted with the cold liquid falling downward to condense and absorb C2 components, C3 components, and heavier components from the vapor stream. The resulting partially rectified vapor stream is then directed to the heat and mass transfer means in rectifying section 117b inside processing assembly 117 for further rectification.
Substantially condensed stream 151e at −178° F. [−117° C.] is flash expanded through expansion valve 23 to slightly above the operating pressure of fractionation tower 17. During expansion a portion of the stream may be vaporized, resulting in cooling of the total stream. In the process illustrated in
The flash expanded stream 151f is further vaporized as it provides cooling and partial condensation of the partially rectified vapor stream, and exits the heat and mass transfer means in rectifying section 117b inside processing assembly 117 at −182° F. [−119° C.]. The heated flash expanded stream then mixes with the further rectified vapor stream to form a combined stream at −181° F. [−119° C.] that is directed to the heat exchange means in cooling section 117a inside processing assembly 117. The combined stream is heated as it provides cooling to streams 151d and 36a as described previously.
The distillation liquid leaving the bottom of the mass transfer means in absorbing section 117c discharges from the bottom of processing assembly 117 (stream 154) and is pumped to higher pressure by pump 24 (stream 154a at −172° F. [−113° C.]). Further cooled substantially condensed stream 36b at −160° F. [−107° C.] is flash expanded through expansion valve 13 to the operating pressure of fractionation tower 17. During expansion a portion of the stream may be vaporized, resulting in cooling of the total stream to −172° F. [−114° C.]. Flash expanded stream 36c then joins with pumped stream 154a to form combined feed stream 155, which enters fractionation column 17 at the top feed point at −172° F. [−114° C.].
The heated combined stream 152 is discharged from the heat exchange means in cooling section 117a inside processing assembly 117 at −80° F. [−62° C.]. It is divided into the previously described stream 151, and into cool residue gas stream 153 which is then heated and compressed as described previously for stream 39 in the
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables I and V shows that, compared to the prior art of
The improvement in recovery efficiency provided by the present invention over that of the prior art of the
The present invention has the further advantage of using the heat and mass transfer means in rectifying section 117b to simultaneously cool the column overhead vapor stream and condense the heavier hydrocarbon components from it, providing more efficient rectification than using reflux in a conventional distillation column. As a result, more of the C2 components, C3 components, and heavier hydrocarbon components can be removed from the column overhead vapor stream using the refrigeration available in flash expanded stream 151f than is possible using conventional mass transfer equipment and conventional heat transfer equipment.
The present invention offers two other advantages over the prior art in addition to the increase in processing efficiency. First, the compact arrangement of processing assembly 117 of the present invention incorporates what would normally be three separate equipment items (the heat exchange means in cooling section 117a, the heat and mass transfer means in rectifying section 117b, and the mass transfer means in absorbing section 117c) into a single equipment item (processing assembly 117 in
One additional advantage of the present invention is how easily it can be incorporated into an existing gas processing plant to effect the superior performance described above. As shown in
The main reason the present invention is more efficient than our co-pending application Ser. No. 15/332,723 depicted in
The present invention also offers advantages when product economics favor rejecting the C2 components to the residue gas product. The present invention can be easily reconfigured to operate in a manner similar to that of our U.S. Pat. Nos. 9,637,428 and 9,927,171 as shown in
When operating the present invention in this manner, many of the process conditions shown for the
For the operating conditions shown in
The flash expanded stream 36b is further vaporized as it provides cooling and partial condensation of the partially rectified vapor stream, and exits the heat and mass transfer means in rectifying section 117b inside processing assembly 117 at −83° F. [−64° C.]. The heated flash expanded stream 36c is then mixed with pumped liquid stream 154a to form combined feed stream 155, which enters fractionation column 17 at the top feed point at −82° F. [−64° C.].
The further rectified vapor stream leaves the heat and mass transfer means in rectifying section 117b inside processing assembly 117 at −104° F. [−76° C.]. Since the heat exchange means in cooling section 117a inside processing assembly 117 has been idled, the vapor simply discharges from processing assembly 117 as cool residue gas stream 153, which is heated and compressed as described previously for stream 39 in the
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables II and VI shows that, compared to the prior art, the
In the embodiment of the present invention shown in
Some circumstances may favor mounting the liquid pump inside the processing assembly to further reduce the number of equipment items and the plot space requirements. Such embodiments are shown in
Some circumstances may favor locating the processing assembly at a higher elevation than the top feed point on fractionation column 17. In such cases, it may be possible for distillation liquid stream 154 to flow by gravity head and combine with stream 36c so that the resulting combined feed stream 155 then flows to the top feed point on fractionation column 17 as shown in
Some circumstances may favor eliminating cooling section 117a from processing assembly 117, and using a heat exchange means external to the processing assembly for feed cooling, such as heat exchanger 27 shown in
The present invention provides improved recovery of C2 components, 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 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 supplemental heating, 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 | |
|---|---|---|---|
| 62667833 | May 2018 | US | |
| 62513851 | Jun 2017 | US |
