Propylene, propane, and/or heavier hydrocarbons can be recovered from a variety of gases, such as natural gas, refinery gas, and synthetic gas streams obtained from other hydrocarbon materials such as coal, crude oil, naphtha, oil shale, tar sands, and lignite. Natural gas usually has a major proportion of methane and ethane, i. e., methane and ethane together comprise at least 50 mole percent of the gas.

The gas also contains relatively lesser amounts of heavier hydrocarbons such as propane, butanes, pentanes and the like, as well as hydrogen, nitrogen, carbon dioxide and other gases.

The present invention is generally concerned with the recovery of propylene, propane and heavier hydrocarbons from such gas streams. A typical analysis of a gas stream to be processed in accordance with this invention would be, in approximate mole percent, 92.6% methane, 4.7% ethane and other C2 components, 1.0% propane and other C3 components, 0.2% iso-butane, 0.2% normal butane, 0.1% pentanes plus, with the balance made up of nitrogen and carbon dioxide. Sulfur containing gases are also sometimes present.

The historically cyclic fluctuations in the prices of both natural gas and its natural gas liquid (NGL) constituents have reduced the incremental value of propylene, propane, and heavier components as liquid products. This has resulted in a demand for processes that can provide more efficient recoveries of these products.

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 (ethylene, ethane, 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 propylene and propane recovery because it provides maximum simplicity with ease of start up, operating flexibility, good efficiency, safety, and good reliability. U. S. Pat.

Nos. 4,157,904; 4,171,964; 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; 5,568,737; 5,555,748; 5,771,712; 5,799,507; 5,881,569; 5,890,378; and 6,182,468 B1 and reissue U. S. Pat. No. 33,408 describe relevant processes.

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 C3+ 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 (deethanizer) column. In the column, the expansion cooled stream (s) is (are) distilled 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 passed through a work expansion machine or engine, or an expansion valve, to a lower pressure at which additional

liquids are condensed as a result of further cooling of the stream. The pressure after expansion is slightly below the pressure at which the distillation column is operated.

The expanded stream then enters the lower section of an absorption column and is contacted with cold liquids to absorb the C3 components and heavier components from the vapor portion of the expanded stream. The liquids from the absorption column are then pumped into the deethanizer column at an upper column feed position, perhaps after heating to partially vaporize the stream.

The overhead distillation stream from the deethanizer passes in heat exchange relation with the residue gas from the absorber column and is cooled, condensing at least a portion of the distillation stream from the deethanizer. The cooled distillation stream then enters the upper section of the absorption column where the cold liquids contained in the stream can contact the vapor portion of the expanded stream as described earlier. Typically, the vapor portion (if any) of the cooled distillation stream and the absorber overhead vapor combine in an upper separator section in the absorber column as residual methane and C2 component product gas. Alternatively, the cooled distillation stream may be supplied to a separator to provide vapor and liquid streams. The vapor is combined with the absorber column overhead and the liquid is supplied to the absorber column as a top column feed. It may also be advantageous to supply a portion of the cold liquid condensate to the deethanizer tower to serve as reflux.

The separation that takes place in this process (producing a residue gas leaving the process which contains substantially all of the methane and C2 components in the feed gas with essentially none of the C3 components and heavier hydrocarbon components, and a bottoms fraction leaving the deethanizer which contains substantially all of the C3 components and heavier hydrocarbon components with essentially no methane, C2 components, or more volatile components) consumes energy for feed gas cooling, for reboiling the deethanizer, for refluxing the deethanizer, and/or for re-compressing the residue gas. The present invention provides a means for achieving this separation while substantially reducing the utility requirements (cooling, reboiling, refluxing, and/or re-compressing) needed for the recovery of the desired products.

In accordance with the present invention, it has been found that C3 recoveries in excess of 99 percent can be maintained while providing essentially complete rejection of C2 components to the residue gas stream. In addition, the present invention makes possible essentially 100 percent separation of C2 components and lighter components from the C3 components and heavier hydrocarbon components at reduced energy requirements. 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 column overhead temperatures of-50°F [-46°C] or colder.

For a better understanding of the present invention, reference is made to the following examples and drawings. Referring to the drawings: FIG. 1 is a flow diagram of a prior art cryogenic natural gas processing plant in accordance with United States Patent No. 5,771,712; FIG. 2 is a flow diagram of a prior art cryogenic natural gas processing plant of an alternative system in accordance with United States Patent No. 5,771,712; FIG. 3 is a flow diagram of a natural gas processing plant in accordance with the present invention; FIG. 4 is a flow diagram illustrating an alternative means of application of the present invention to a natural gas stream ; FIG. 5 is a flow diagram illustrating an alternative means of application of the present invention to a natural gas stream; FIG. 6 is a flow diagram illustrating an alternative means of application of the present invention to a natural gas stream; FIG. 7 is a flow diagram illustrating an alternative means of application of the present invention to a natural gas stream; FIG. 8 is a flow diagram illustrating an alternative means of application of the present invention to a natural gas stream; FIG. 9 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream; FIG. 10 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream;

FIG. 11 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream; FIG. 12 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream ; FIG. 13 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream ; FIG. 14 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream ; FIG. 15 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream ; FIG. 16 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream ; FIG. 17 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream ; and FIG. 18 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream.

In the following explanation of the above figures, tables are provided summarizing flow rates calculated for representative process conditions. In the tables appearing herein, the values for flow rates (in moles per hour) have been rounded to the nearest whole number for convenience. The total stream rates shown in the tables include all non-hydrocarbon components and hence are generally larger than the sum of the stream flow rates for the hydrocarbon components. Temperatures indicated are approximate values rounded to the nearest degree. It should also be noted that the process design calculations performed for the purpose of comparing the processes depicted in the figures are based on the assumption of no heat leak from (or to) the surroundings to (or from) the process. The quality of commercially available insulating materials makes this a very reasonable assumption and one that is typically made by those skilled in the art.

For convenience, process parameters are reported in both the traditional British units and in the units of the International System of Units (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 thousand British Thermal Units per hour (MBTU/hour) 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.

DESCRIPTION OF THE PRIOR ART Referring now to FIG. 1, in a simulation of prior art according to U. S.

Pat. No. 5,771,712, inlet gas enters the plant at 80°F [27°C] and 1215 psia [8,377 kPa (a)] as stream 31. If the inlet gas contains a concentration of sulfur compounds which would prevent the product streams from meeting specifications, the sulfur compounds are removed by appropriate pretreatment of the feed gas (not illustrated). In addition, the feed stream is usually dehydrated to prevent hydrate (ice) formation under cryogenic conditions. Solid desiccant has typically been used for this purpose.

The feed stream 31 is cooled in exchanger 10 by heat exchange with cool residue gas at-76°F [-60°C] (stream 34a). (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.). For the conditions stated, the feed stream pressure is above the cricondenbar, so no liquid will condense as the stream is cooled. Hence, the cooled stream 31 a (a dense-phase fluid at these conditions) is supplied directly to work expansion machine 13 at-14°F [-26°C]. (The cricondenbar is the maximum pressure at which a vapor phase can exist in a multi-phase fluid. At pressures below the cricondenbar, a separator or scrubber would typically be used to separate any condensed liquid contained in stream 31a from the vapor so that only the vapor is supplied to work expansion machine 13.) The work expansion machine 13 extracts mechanical energy from the high pressure feed by expanding the stream substantially isentropically from a pressure of about 1210 psia [8,343 kPa (a)] to a pressure of about 435 psia [2,999 kPa (a)] (the operating pressure of separator/absorber tower 15), with the work expansion cooling the expanded stream 31b to a temperature of approximately-104°F

[-76°C]. The expanded and partially condensed stream 31b is supplied to absorbing section 15b in a lower region of separator/absorber tower 15. The liquid portion of the expanded stream commingles with liquids falling downward from the absorbing section and the combined liquid stream 35 exits the bottom of separator/absorber tower 15 at-106°F [-77°C]. The vapor portion of the expanded stream rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C3 components and heavier components.

The separator/absorber tower 15 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. As is often the case in natural gas processing plants, the separator/absorber tower may consist of two sections. The upper section 15a is a separator wherein any vapor contained in the top feed is separated from its corresponding liquid portion, and wherein the vapor rising from the lower distillation or absorbing section 15b is combined with the vapor portion (if any) of the top feed to form the cold distillation stream 34 which exits the top of the tower. The lower, absorbing section 15b contains the trays and/or packing and provides the necessary contact between the liquids falling downward and the vapors rising upward to condense and absorb the C3 components and heavier components.

The combined liquid stream 35 from the bottom of the separator/absorber tower 15 is supplied to deethanizer 17 by pump 16, entering at a mid-column feed point at-105°F [-76°C] as stream 35a to be stripped of its methane and C2 components. The deethanizer in tower 17, operating at about 450 psia [3,103 kPa (a)], is also a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. The deethanizer tower may also consist of two sections: an upper section 17a wherein any vapor contained in the top feed is separated from its corresponding liquid portion, and wherein the vapor rising from the lower distillation or deethanizing section 17b is combined with the vapor portion (if any) of the top feed to form distillation stream 36 which exits the top of the tower; and a lower, deethanizing section 17b that contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The deethanizing

section 17b also includes a reboiler 18 which heats and vaporizes a portion of the liquid at the bottom of the column to provide the stripping vapors which flow up the column to strip the liquid product, stream 37, of methane and C2 components. A typical specification for the bottom liquid product is to have an ethane to propane ratio of 0.02: 1 on a molar basis. The liquid product stream 37 exits the bottom of the deethanizer at 207°F [97°C] and is cooled to 110°F [43°C] (stream 37a) in heat exchanger 19 before flowing to storage.

The operating pressure in deethanizer 17 is maintained slightly above the operating pressure of separator/absorber tower 15. This allows the deethanizer overhead vapor (stream 36) to pressure flow through heat exchanger 20 and thence into the upper section of separator/absorber tower 15. In heat exchanger 20, the deethanizer overhead at-36°F [-38°C] is directed in heat exchange relation with the overhead (stream 34) from separator/absorber tower 15, cooling the stream to-107°F [-77°C] (stream 36a) and partially condensing it. The partially condensed stream is then supplied to the separator section in separator/absorber tower 15 where the condensed liquid is separated from the uncondensed vapor. The uncondensed vapor combines with the vapor rising from the lower absorbing section to form the cold distillation stream 34 leaving the upper region of separator/absorber tower 15. The condensed liquid is divided into two portions. One portion, stream 39, is routed to the lower absorbing section of separator/absorber tower 15 as the cold liquid that contacts the vapors rising upward through the absorbing section. The other portion, stream 38, is supplied to deethanizer 17 as reflux by pump 21, with reflux stream 38a flowing to a top feed point on deethanizer 17 at-107°F [-77°C].

The distillation stream leaving the top of separator/absorber tower 15 at-112°F [-80°C] is the cold residue gas stream 34. The residue gas stream passes countercurrently to deethanizer overhead stream 36 in heat exchanger 20 and is warmed to-76°F [-60°C] (stream 34a) as it provides cooling and partial condensation of the deethanizer overhead stream. The residue gas is further warmed to 54°F [12°C] (stream 34b) as it passes countercurrently to the incoming feed stream in heat exchanger 10. The residue gas is then re-compressed in two stages. The first stage is compressor 14 driven by expansion machine 13. The second stage is compressor 22

driven by a supplemental power source which compresses the residue gas (stream 34d) to sales line pressure. After cooling in discharge cooler 23, the residue gas product (stream 34e) flows to the sales gas pipeline at 110°F [43°C] and 1215 psia [8, 377 kPa (a)].

A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 1 is set forth in the following table: TABLE I (FIG. 1) Stream Flow Summary-Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 81,340 4,128 878 439 87, 840 35 5,023 2,108 880 439 8,558 36 7,473 3,432 6 0 11, 080 39 3,674 2,011 4 0 5,782 38 2,449 1,341 2 0 3,854 34 81,340 4,111 3 0 86,508 37 0 18 876 439 1,333 Recoveries* Propane 99.70% Butanes+ 100.00% Power Residue Gas Compression 37,593 HP [61,802 kW] Utility Heat Deethanizer Reboiler 53,478 MBTU/Hr [34,544 kW] * (Based on un-rounded flow rates)

In the prior art illustrated in FIG. 1, the refrigeration generated by work expansion machine 13 is not used efficiently in the process. This is evidenced by the relatively cool residue gas (stream 34b) temperature of 54°F [12°C] leaving heat exchanger 10 (compared to the temperature of the inlet stream 31), and by the fact that the process cooling available in the cold liquid (stream 35) leaving the bottom of separator/absorber 15 is not needed to provide a portion of the inlet gas cooling in heat exchanger 10. Ordinarily, this would indicate that less expansion ratio is needed across work expansion machine 13 to maintain the desired C3 component recovery efficiency, and that the operating pressure of separator/absorber 15 could be raised to reduce the external power requirements in compressor 22.

However, since the operating pressure of deethanizer 17 must of necessity be maintained somewhat higher than that of separator/absorber 15 so that its overhead stream 36 can pressure flow through heat exchanger 20 and into the separator section of separator/absorber 15, reducing the expansion ratio across work expansion machine 13 also means raising the operating pressure of deethanizer 17.

Unfortunately, this is not advisable in this instance because of the detrimental effect on distillation performance in deethanizer 17 that would result from the higher operating pressure. This effect is manifested by poor mass transfer in deethanizer 17 due to the phase behavior of its vapor and liquid streams. Of particular concern are the physical properties that affect the vapor-liquid separation efficiency, namely the liquid surface tension and the differential in the densities of the two phases. As a result, the operating pressure of deethanizer 17 should not be raised above the value shown in FIG. 1, so there is no means available to reduce the power consumption of compressor 22 using the prior art process.

FIG. 2 represents an alternative application of the prior art process in accordance with U. S. Pat. No. 5,771,712. The process of FIG. 2 has been applied to the same feed gas composition as described above for FIG. 1, but in this simulation of the process the inlet gas (stream 31) enters the plant at 80°F [27°C] and 580 psia [3,999 kPa (a)]. The feed stream 31 is cooled in exchanger 10 by heat exchange with cool residue gas at-95°F [-71°C] (stream 34a), with separator liquids at-92°F [-69°C] (stream 33a), and with separator/absorber liquids at-107°F [-77°C] (stream 35a). At

this operating pressure the feed stream is below the cricondenbar, so the cooled stream 31a enters separator 11 at-77°F [-60°C] and 570 psia [3,930 kPa (a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33).

The vapor (stream 32) from separator 11 enters work expansion machine 13 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 13 expands the vapor substantially isentropically from a pressure of about 570 psia [3,930 kPa (a)] to a pressure of about 380 psia [2,620 kPa (a)] (the operating pressure of separator/absorber 15), with the work expansion cooling the expanded stream 32a to a temperature of approximately-107°F [-77°C]. The expanded and partially condensed stream 32a enters the lower section of separator/absorber 15. The liquid portion of the expanded stream commingles with liquids falling downward from the absorbing section and the combined liquid stream 35 exits the bottom of separator/absorber 15 at-108°F [-78°C]. The vapor portion of the expanded stream rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C3 components and heavier components.

The combined liquid stream 35 from the bottom of the separator/absorber 15 is routed to heat exchanger 10 by pump 16 where it (stream 35a) is heated as it provides cooling of the incoming feed gas as described earlier.

The combined liquid stream is heated to-85°F [-65°C], partially vaporizing stream 35b before it is supplied as a mid-column feed to deethanizer 17. The separator liquid (stream 33) is flash expanded to slightly above the 395 psia [2,723 kPa (a)] operating pressure of deethanizer 17 by expansion valve 12, cooling stream 33 to-92°F [-69°C] (stream 33a) before it provides cooling to the incoming feed gas as described earlier.

Stream 33b, now at 65°F [18°C], then enters deethanizer 17 at a lower mid-column feed point. In the deethanizer, streams 35b and 33b are stripped of their methane and C2 components. The resulting liquid product stream 37 exits the bottom of the deethanizer at 195°F [91°C] and is cooled to 110°F [43°C] (stream 37a) in heat exchanger 19 before flowing to storage.

The operating pressure in deethanizer 17 is maintained slightly above the operating pressure of separator/absorber 15. This allows the deethanizer overhead

vapor (stream 36) to pressure flow through heat exchanger 20 and thence into the upper section of separator/absorber 15. In heat exchanger 20, the deethanizer overhead at-29°F [-34°C] is directed in heat exchange relation with the overhead (stream 34) from separator/absorber 15, cooling the stream to-108°F [-78°C] (stream 36a) and partially condensing it. The partially condensed stream is then supplied to the separator section in separator/absorber tower 15 where the condensed liquid is separated from the uncondensed vapor. The uncondensed vapor combines with the vapor rising from the lower absorbing section to form the cold distillation stream 34 leaving the upper region of separator/absorber 15. The condensed liquid is divided into two portions. One portion, stream 39, is routed to the lower absorbing section of separator/absorber 15 as the cold liquid that contacts the vapors rising upward through the absorbing section. The other portion, stream 38, is supplied to deethanizer 17 as reflux by pump 21, with reflux stream 38a flowing to a top feed point on deethanizer 17 at-108°F [-78°C].

The distillation stream leaving the top of separator/absorber 15 at -113°F [-81°C] is the cold residue gas stream 34. The residue gas stream passes countercurrently to deethanizer overhead stream 36 in heat exchanger 20 and is warmed to-95°F [-71°C] (stream 34a) as it provides cooling and partial condensation of the deethanizer overhead stream. The residue gas is further warmed to 75°F [24°C] (stream 34b) as it passes countercurrently to the incoming feed gas in heat exchanger 10. The residue gas is then re-compressed in two stages. The first stage is compressor 14 driven by expansion machine 13. The second stage is compressor 22 driven by a supplemental power source which compresses the residue gas (stream 34d) to sales line pressure. After cooling in discharge cooler 23, the residue gas product (stream 34e) flows to the sales gas pipeline at 110°F [43°C] and 613 psia [4,226 kPa (a)].

A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 2 is set forth in the table below:

TABLE II (FIG. 2) Stream Flow Summary-Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 81,340 4,128 878 439 87, 840 32 80, 218 3,702 573 126 85,651 33 1,122 427 306 313 2,189 35 1,876 964 547 126 3,556 36 3,451 1,700 76 0 5,306 39 1,816 1,300 61 0 3,229 38 454 325 15 0 807 34 81,340 4,113 87 0 86,594 37 0 16 791 439 1,246 Recoveries* Propane 90.09% Butanes+ 99.99% Power Residue Gas Compression 18,911 HP [31,089 kW] Utility Heat Deethanizer Reboiler 17,844 MBTU/Hr [11,526 kW] * (Based on un-rounded flow rates) In the prior art illustrated in FIG. 2, the much lower feed gas pressure results in much less refrigeration (compared to FIG. 1) being generated by work expansion machine 13. Consequently, much better process heat integration is required to achieve the desired C3 component recovery efficiency than was the case for the FIG. 1 processing conditions. The process cooling available from the residue gas (stream 34a) in heat exchanger 10 must be supplemented by the process cooling

available in the cold separator liquids (stream 33) and the cold liquid (stream 35) leaving the bottom of separator/absorber 15 in order to accomplish the necessary inlet gas cooling. In fact, the process heat integration is so complete that the only means available to increase the C3 component recovery with the prior art process for these processing conditions would be to increase the expansion ratio across work expansion machine 13 (by reducing the operating pressure of separator/absorber 15) to increase the refrigeration generated by the machine, and/or to add external refrigeration. Of course, this would have the undesired consequence of increasing the external power requirements, either in compressor 22, in the external refrigeration system, or both.

DESCRIPTION OF THE INVENTION Example 1 FIG. 3 illustrates a flow diagram of a process in accordance with the present invention. The feed gas composition and conditions considered in the process presented in FIG. 3 are the same as those in FIG. 1. Accordingly, the FIG. 3 process can be compared with that of the FIG. 1 process to illustrate the advantages of the present invention.

In the simulation of the FIG. 3 process, inlet gas enters the plant at 80°F [27°C] and 1215 psia [8,377 kPa (a)] as stream 31. The feed stream 31 is cooled in exchanger 10 by heat exchange with cool residue gas at-56°F [-49°C] (stream 34a) and with separator/absorber liquids at-113°F [-80°C] (stream 35a). The cooled stream 31 a (a dense-phase fluid at these conditions) is supplied directly to work expansion machine 13 at-35°F [-37°C].

The work expansion machine 13 extracts mechanical energy from the high pressure feed by expanding the stream substantially isentropically from a pressure of about 1210 psia [8, 343 kPa (a)] to a pressure of about 575 psia [3,964 kPa (a)] (the operating pressure of separator/absorber tower 15), with the work expansion cooling the expanded stream 31b to a temperature of approximately-98°F [-72°C]. The expanded and partially condensed stream 31b enters the lower section of separator/absorber 15. The liquid portion of the expanded stream commingles with liquids falling downward from the absorbing section and the combined liquid stream

35 exits the bottom of separator/absorber 15 at-100°F [-73°C]. The vapor portion of the expanded stream rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C3 components and heavier components.

Unlike the prior art process illustrated in FIG. 1, in the present invention the operating pressure of deethanizer 17 is maintained below (not above) the operating pressure of separator/absorber 15. Consequently, a pump is not required for the combined liquid stream 35 from the bottom of the separator/absorber 15.

Instead, the stream is flash expanded to slightly above the 450 psia [3,103 kPa (a)] operating pressure of deethanizer 17 by expansion valve 27, cooling stream 35 to -113°F [-80°C] (stream 35a) before it provides cooling to the incoming feed gas as described earlier. Stream 35b, now at-73°F [-58°C], then enters deethanizer 17 at a mid-column feed point. In the deethanizer, stream 35b is stripped of its methane and C2 components. The resulting liquid product stream 37 exits the bottom of the deethanizer at 207°F [97°C] and is cooled to 110°F [43°C] (stream 37a) in heat exchanger 19 before flowing to storage.

The deethanizer overhead vapor (stream 36) exits deethanizer 17 at -56°F [-49°C] and is warmed to 105°F [41°C] (stream 36a) in heat exchanger 24 before entering compressor 25 (driven by a supplemental power source). Stream 36b leaves compressor 25 at 600 psia [4,137 kPa (a)] and is cooled to 110°F [43°C] (stream 36c) in heat exchanger 26. Stream 36c is then directed in heat exchange relation with the deethanizer overhead vapor (stream 36) in heat exchanger 24 to cool it (stream 36d) and conserve process cooling.

With the increase in pressure provided by compressor 25, stream 36d can now pressure flow through heat exchanger 20 and thence to the upper feed point of separator/absorber 15. In heat exchanger 20, the compressed deethanizer overhead at-41°F [-40°C] is directed in heat exchange relation with the overhead (stream 34) from separator/absorber 15, cooling the stream to-98°F [-72°C] (stream 36e) and partially condensing it. The partially condensed stream is then supplied to the separator section in separator/absorber tower 15 where the condensed liquid is separated from the uncondensed vapor. The uncondensed vapor combines with the

vapor rising from the lower absorbing section to form the cold distillation stream 34 leaving the upper region of separator/absorber 15. The condensed liquid is divided into two portions. One portion, stream 39, is routed to the lower absorbing section of separator/absorber 15 as the cold liquid that contacts the vapors rising upward through the absorbing section. The other portion, stream 38, is flash expanded to slightly above the operating pressure of deethanizer 17 by expansion valve 28 and the resulting stream 38a is then supplied at-112°F [-80°C] to the separator section in deethanizer 17 where its condensed liquid is separated from its uncondensed vapor.

The uncondensed vapor combines with the vapor rising from the lower distillation section to form the deethanizer overhead stream 36 leaving the upper region of deethanizer 17, while the condensed liquid is routed to the lower distillation section of deethanizer 17 as reflux for the vapors rising upward through the distillation section.

The distillation stream leaving the top of separator/absorber 15 at -103°F [-75°C] is the cold residue gas stream 34. The residue gas stream passes countercurrently to compressed deethanizer overhead stream 36d in heat exchanger 20 and is warmed to-56°F [-49°C] (stream 34a) as it provides cooling and partial condensation of the compressed deethanizer overhead stream. The residue gas is further warmed to 75°F [24°C] (stream 34b) as it passes countercurrently to the incoming feed gas in heat exchanger 10. The residue gas is then re-compressed in two stages. The first stage is compressor 14 driven by expansion machine 13. The second stage is compressor 22 driven by a supplemental power source which compresses the residue gas (stream 34d) to sales line pressure. After cooling in discharge cooler 23, the residue gas product (stream 34e) flows to the sales gas pipeline at 110°F [43°C] and 1215 psia [8, 377 kPa (a)].

A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 3 is set forth in the table below:

TABLE III (FIG. 3) Stream Flow Summary-Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 81,340 4,128 878 439 87,840 35 11, 825 2,920 879 439 16,261 36 17,501 4,827 6 0 22,642 39 7,839 2,658 3 0 10,653 38 5,676 1,925 2 0 7,714 34 81, 340 4,111 3 0 86, 506 37 0 18 876 439 1,333 Recoveries* Propane 99.70% Butanes+ 100.00% Power Residue Gas Compression 28,422 HP [46,725 kW] Overhead Vapor Compression 3,810 HP [6,264 kW] Total Compression 32,232 HP [52, 989 kW] UtilityHeat Deethanizer Reboiler 51,073 MBTU/Hr [32, 990 kW] * (Based on un-rounded flow rates) Comparison of the utility consumptions of the prior art process displayed in Table I with the utility consumptions of the present invention displayed in Table III shows that the present invention maintains the desired C3 component recovery while reducing the utility heat requirement and substantially reducing the compression horsepower. The utility heat requirement is more than four percent lower than the prior art process, while the compression horsepower is more than

fourteen percent lower than the prior art process.

Comparing the present invention to the prior art process displayed in FIG. 1, note that while the operating pressure of deethanizer 17 is the same in both cases, the operating pressure of separator/absorber 15 in the present invention is significantly higher than in the FIG. 1 process, 575 psia [3,964 kPa (a)] versus 435 psia [2,999 kPa (a)]. Accordingly, the residue gas enters compressor 14 at a higher pressure in the FIG. 3 process and less compression horsepower is therefore needed to deliver the residue gas to pipeline pressure. Further, with separator/absorber 15 operating at a higher pressure than deethanizer 17, it is no longer necessary to pump the absorber bottom liquid (stream 35) and the reflux stream (stream 38) to feed deethanizer 17, eliminating the capital and operating cost of pumps 16 and 21 in the FIG. 1 process.

In essence, work expansion machine 13 and compressors 14 and 22 represent an open cycle mechanical-compression refrigeration loop that provides the process cooling in the prior art process of FIG. 1, with a working fluid (streams 31 and 34) that is predominantly methane. In the present invention illustrated in FIG. 3, the refrigeration provided by this cycle has been reduced by the addition of a second open cycle refrigeration loop powered by compressor 25. Examination of Table III shows that the working fluid for this second cycle (the deethanizer overhead, stream 36) has a substantially lower concentration of methane and a substantially higher concentration of ethane than the working fluid in the first cycle. In general, the efficiency of mechanical-compression refrigeration cycles improves as the molecular weight of the working fluid increases. This effect, together with the much lower flow rate of stream 36 compared to streams 31/34 and the lower compression ratio needed from compressor 25 compared to compressors 14/22, accounts for most of the improvement in energy efficiency of the present invention relative to the prior art of FIG. 1. As a measure of this increase in efficiency, note that the total reflux generated for the two columns in the present invention shown in FIG. 3 (the sum of streams 38 and 39) is nearly twice that for the prior art of FIG. 1, and yet this is accomplished using 14% less power in the mechanical-refrigeration cycles.

With compressor 25 supplying the motive force to cause the deethanizer overhead (stream 36 in FIG. 3) to flow through heat exchanger 20 and thence to separator/absorber 15, the operating pressures of separator/absorber 15 and deethanizer 17 are no longer coupled together as they are in the prior art process.

Instead, the operating pressures of the two columns can be optimized independently.

In the case of deethanizer 17, the pressure can be selected to insure good distillation characteristics, while for separator/absorber 15 the pressure can be selected to optimize the process cooling versus the residue gas compression requirements.

Example 2 FIG. 3 represents the preferred embodiment of the present invention for the temperature and pressure conditions shown because it typically provides the simplest plant arrangement for a given C3 component recovery level. A slightly more complex design that maintains the same C3 component recovery with lower utility consumption can be achieved using another embodiment of the present invention as illustrated in the FIG. 4 process. The feed gas composition and conditions considered in the process presented in FIG. 4 are the same as those in FIGS. 1 and 3.

Accordingly, FIG. 4 can be compared with the FIG. 1 process to illustrate the advantages of the present invention, and can likewise be compared to the embodiment displayed in FIG. 3.

In the simulation of the FIG. 4 process, the feed gas cooling and expansion scheme is much the same as that used in FIG. 3. The difference lies in the manner in which the vapor distillation stream 36 leaving the overhead of deethanizer 17 is used to generate reflux for deethanizer 17 and separator/absorber 15. Referring to FIG. 4, the deethanizer overhead vapor (stream 36) exits deethanizer 17 at-39°F [-39°C] and is warmed to 105°F [41°C] (stream 36a) in heat exchanger 24 before entering compressor 25 (driven by a supplemental power source). Stream 36b leaves compressor 25 at 600 psia [4,137 kPa (a)] and is cooled to 110°F [43°C] (stream 36c) in heat exchanger 26. Stream 36c is then directed in heat exchange relation with the deethanizer overhead vapor (stream 36) in heat exchanger 24 to cool it to-24°F [-31°C] (stream 36d) and conserve process cooling.

In heat exchanger 20, the compressed deethanizer overhead (stream 36d) is directed in heat exchange relation with the overhead (stream 34) from separator/absorber 15, cooling the stream to-50°F [-46°C] (stream 36e) and partially condensing it before it is withdrawn. The partially condensed stream 36e enters separator 30 where the condensed liquid is separated from the uncondensed vapor.

The condensed liquid (stream 38) from separator 30 is flash expanded to slightly above the operating pressure of deethanizer 17 by expansion valve 28 (stream 38a), which partially vaporizes the stream and cools it further to-63°F [-53°C]. It is then supplied to the separator section in deethanizer 17 where the liquid is separated from the flash vapor. The flash vapor combines with the vapor rising from the lower distillation section to form the deethanizer overhead stream 36 leaving the upper region of deethanizer 17, while the liquid is routed to the lower distillation section of deethanizer 17 as reflux for the vapors rising upward through the distillation section.

The uncondensed vapor (stream 39) from separator 30 is routed back to heat exchanger 20 to also direct it in heat exchange relation with the overhead (stream 34) from separator/absorber 15, cooling the stream to-98°F [-72°C] (stream 39a) and partially condensing it. The partially condensed stream is then supplied to the separator section in separator/absorber tower 15 where its condensed liquid is separated from its uncondensed vapor. The uncondensed vapor combines with the vapor rising from the lower absorbing section to form the cold distillation stream 34 leaving the upper region of separator/absorber 15, while the condensed liquid is routed to the lower absorbing section of separator/absorber 15 as the cold liquid that contacts the vapors rising upward through the absorbing section.

A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 4 is set forth in the table below:

TABLE IV (FIG. 4) Stream Flow Summary-Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 81,340 4,128 878 439 87, 840 35 11, 237 2,788 878 439 15,530 36 13,548 5,755 10 0 19,583 39 11, 237 2,770 2 0 14,198 38 2,311 2,985 8 0 5,385 34 81,340 4, 111 3 0 86,507 37 0 18 876 439 1,333 Recoveries* Propane 99.70% Butanes+ 100.00% Power Residue Gas Compression 28,405 HP [46,697 kW] Overhead Vapor Compression 3,246 HP [5,336 kW] 31,651 HP [52, 034 kW] Utility Heat Deethanizer Reboiler 51,255 MBTU/Hr [33, 108 kW] * (Based on un-rounded flow rates) Comparison of the utility consumptions of the prior art process displayed in Table I with the utility consumptions of the present invention displayed in Table IV shows that this embodiment of the present invention also maintains the desired C3 component recovery while reducing the utility heat requirement and substantially reducing the compression horsepower. The utility heat requirement is more than four percent lower than the prior art process, while the compression

horsepower is more than fifteen percent lower than the prior art process.

Comparison of the utility consumptions displayed in Tables III and IV for the FIG. 3 and FIG. 4 processes shows that the FIG. 4 embodiment of the present invention requires slightly less compression horsepower (about 2 percent) than the FIG. 3 embodiment, but uses slightly more utility heat for the deethanizer reboiler (less than 1 percent), with the total utility requirements being about 1 percent lower for the FIG. 4 embodiment. The improvement in efficiency can be understood by comparing the reflux stream for deethanizer 17 (stream 38) in the FIG. 4 embodiment of the present invention with the corresponding stream in the FIG. 3 embodiment.

Whereas stream 38 in FIG. 3 is predominantly methane, stream 38 in FIG. 4 is predominantly ethane because it is withdrawn after only partial cooling in heat exchanger 20 so that proportionally less of the more volatile methane has been condensed. Not only is ethane a more effective reflux liquid than methane for rectifying the C3 and heavier components from the vapors rising upward in deethanizer 17 (as reflected by the much lower flow rate of stream 38 in the FIG. 4 embodiment), the deethanizer overhead (stream 36) has a lower concentration of methane (because less methane enters deethanizer 17 in the reflux) so that the mechanical-compression refrigeration efficiency of compressor 25 is improved.

Although this embodiment of the present invention is more efficient than the FIG. 3 embodiment, the choice of whether to include the additional equipment that the FIG. 4 process requires will generally depend on factors which include plant size and available equipment, as well as the relative costs of compression horsepower and utility heat.

Example 3 A third embodiment of the present invention is shown in FIG. 5, wherein a different method of implementing the second mechanical-compression refrigeration cycle is applied to the present invention. The feed gas composition and conditions considered in the process illustrated in FIG. 5 are the same as those in FIGS. 1,3, and 4. Accordingly, FIG. 5 can be compared with the FIG. 1 process to illustrate the advantages of the present invention, and can likewise be compared to the

embodiments displayed in FIGS. 3 and 4.

In the simulation of the FIG. 5 process, inlet gas enters the plant at 80°F [27°C] and 1215 psia [8,377 kPa (a)] as stream 31. The feed stream 31 is cooled in exchanger 10 by heat exchange with cool residue gas at-70°F [-57°C] (stream 34a), with cool vapor at-49°F [-45°C] (stream 41 a), and with separator/absorber liquids at-112°F [-80°C] (stream 35a). The cooled stream 31a (a dense-phase fluid at these conditions) is supplied directly to work expansion machine 13 at-32°F [-36°C].

The work expansion machine 13 extracts mechanical energy from the high pressure feed by expanding the stream substantially isentropically from a pressure of about 1210 psia [8, 343 kPa (a)] to a pressure of about 515 psia [3,551 kPa (a)] (the operating pressure of separator/absorber tower 15), with the work expansion cooling the expanded stream 31b to a temperature of approximately-104°F [-75°C]. The expanded and partially condensed stream 31b enters the lower section of separator/absorber 15. The liquid portion of the expanded stream commingles with liquids falling downward from the absorbing section and the combined liquid stream 35 exits the bottom of separator/absorber 15 at-104°F [-76°C]. The vapor portion of the expanded stream rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C3 components and heavier components.

The combined liquid stream 35 is flash expanded to slightly above the 450 psia [3,103 kPa (a)] operating pressure of deethanizer 17 by expansion valve 27, cooling stream 35 to-112°F [-80°C] (stream 35a) before it provides cooling to the incoming feed gas as described earlier. Stream 35b, now at-92°F [-69°C], then enters deethanizer 17 at a mid-column feed point. In the deethanizer, stream 35b is stripped of its methane and C2 components. The resulting liquid product stream 37 exits the bottom of the deethanizer at 207°F [97°C] and is cooled to 110°F [43°C] (stream 37a) in heat exchanger 19 before flowing to storage.

The deethanizer overhead vapor (stream 36) exits deethanizer 17 at -44°F [-42°C] and flows through heat exchanger 20. In heat exchanger 20, the deethanizer overhead is directed in heat exchange relation with the overhead (stream 34) from separator/absorber 15 and the uncondensed vapor (stream 41) from separator

30, cooling the stream to-102°F [-74°C] (stream 36a) and partially condensing it.

The partially condensed stream is then supplied to separator 30 where the condensed liquid (stream 40) is separated from the uncondensed vapor (stream 41).

In this embodiment of the present invention, the liquid condensed from the deethanizer overhead (stream 40) is at a lower pressure than the two columns (separator/absorber 15 and deethanizer 17), so it is pumped by pump 21 so that it can be used as reflux. After pumping, stream 40a is then divided into two portions. One portion, stream 39, is supplied by control valve 29 to the separator section in separator/absorber tower 15 at-100°F [-74°C] (stream 39a) where its liquid is separated from any vapor that forms. (As the stream is at elevated pressure relative to the pressure at which it condensed, it is unlikely that any vapor will form.) Any vapor that may form combines with the vapor rising from the lower absorbing section to form the cold distillation stream 34 leaving the upper region of separator/absorber 15, while the condensed liquid is routed to the lower absorbing section of separator/absorber 15 as the cold liquid that contacts the vapors rising upward through the absorbing section. The other portion of the pumped liquid (stream 38) is flash expanded to slightly above the operating pressure of deethanizer 17 by expansion valve 28 (stream 38a). It is then supplied at-101°F [-74°C] to the separator section in deethanizer 17 where its liquid is separated from any flash vapor that forms. Any flash vapor combines with the vapor rising from the lower distillation section to form the deethanizer overhead stream 36 leaving the upper region of deethanizer 17, while the condensed liquid is routed to the lower distillation section of deethanizer 17 as reflux for the vapors rising upward through the distillation section.

The distillation stream leaving the top of separator/absorber 15 at -107°F [-77°C] is the cold absorber overhead stream 34. The absorber overhead stream passes countercurrently to deethanizer overhead stream 36 in heat exchanger 20 and is warmed to-70°F [-57°C] (stream 34a) as it provides cooling and partial condensation of the deethanizer overhead stream. The absorber overhead stream is further warmed to 75°F [24°C] (stream 34b) as it passes countercurrently to the incoming feed gas in heat exchanger 10. The uncondensed vapor (stream 41) leaves separator 30 at-102°F [-74°C] and also passes countercurrently to deethanizer

overhead stream 36 in heat exchanger 20 and is warmed to-49°F [-45°C] (stream 41 a) as it too provides cooling and partial condensation of the deethanizer overhead stream. The vapor stream is further warmed to 65°F [18°C] (stream 41b) as it passes countercurrently to the incoming feed gas in heat exchanger 10.

The warm absorber overhead stream 34b and the warm vapor stream 41b are then re-compressed in two stages. The first stage for the absorber overhead stream is compressor 14 driven by expansion machine 13, while the first stage for the vapor stream is compressor 25 driven by a supplemental power source. The two partially compressed streams (streams 34c and 41c, respectively) combine to form the residue gas, stream 42. The combined residue gas stream then enters compressor 22 driven by a supplemental power source, which provides the second stage of compression to raise the residue gas (stream 42a) to sales line pressure. After cooling in discharge cooler 23, the residue gas product (stream 42b) flows to the sales gas pipeline at 110°F [43°C] and 1215 psia [8, 377 kPa (a)].

A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 5 is set forth in the table below:

TABLE V (FIG. 5) Stream Flow Summary-Lb. Moles/Hr [kg moleslHr] Stream Methane Ethane Propane Butanes+ Total 31 81,340 4,128 878 439 87,840 35 9,637 2,685 879 439 13,813 36 12,609 4,631 6 0 17,508 39 3,352 2,213 3 0 5,669 38 2,972 1,963 3 0 5,027 41 6,285 454 0 0 6,811 34 75,055 3,657 3 0 79,696 42 81, 340 4,111 3 0 86,507 37 0 18 876 439 1,332 Recoveries* Propane 99.70% Butanes+ 100.00% Power Residue Gas Compression 32,712 HP [53,778 kW] Vapor Compression 1,413 HP [2,323 kW] 34,125 HP [56,101 kW] UtilityHeat Deethanizer Reboiler 56,696 MBTU/Hr [36,623 kW] * (Based on un-rounded flow rates) Comparison of the utility consumptions of the prior art process displayed in Table I with the utility consumptions of the present invention displayed in Table V shows that this embodiment of the present invention also maintains the

desired C3 component recovery while substantially reducing the compression horsepower. Although the utility heat requirement is about six percent higher than the prior art process, the compression horsepower is more than nine percent lower than the prior art process, so the total utility requirements is about four percent lower than the prior art.

Comparison of the utility consumptions displayed in Tables III, IV, and V for the FIG. 3, FIG. 4, and FIG. 5 embodiments of the present invention shows that the FIG. 5 embodiment requires slightly more compression horsepower and utility heating than either the FIG. 3 or the FIG. 4 embodiment. However, if multiple stage compression or multi-wheel centrifugal compression is used to compress the residue gas stream 42, it may be possible to compress the vapor stream 41b using an intermediate stage or wheel, eliminating the need for a separate compressor like compressor 25. Thus, factors such as plant size and available equipment will determine whether the FIG. 5 embodiment would be preferable for a specific circumstance.

Example 4 A slightly more complex design than the FIG. 5 embodiment that maintains the same C3 component recovery with lower utility consumption can be achieved using another embodiment of the present invention as illustrated in the FIG. 6 process. The feed gas composition and conditions considered in the process presented in FIG. 6 are the same as those in FIGS. 1 and 5. Accordingly, FIG. 6 can be compared with the FIG. 1 process to illustrate the advantages of the present invention, and can likewise be compared to the embodiment displayed in FIG. 5.

In the simulation of the FIG. 6 process, the feed gas cooling and expansion scheme is much the same as that used in FIG. 5. The difference lies in the manner in which the vapor distillation stream 36 leaving the overhead of deethanizer 17 is used to generate reflux for deethanizer 17 and separator/absorber 15. Referring to FIG. 6, the deethanizer overhead vapor (stream 36) exits deethanizer 17 at-39°F [-40°C] and flows through heat exchanger 20. In heat exchanger 20, the deethanizer overhead is directed in heat exchange relation with the overhead (stream 34) from

separator/absorber 15 and the uncondensed vapor (stream 41) from separator 30, cooling the stream to-60°F [-51°C] (stream 36a) and partially condensing it before it is withdrawn. The partially condensed stream 36a enters separator 29 where the condensed liquid (stream 38) is separated from the uncondensed vapor (stream 40).

In this embodiment of the present invention, the liquid condensed from the deethanizer overhead (stream 38) is at a lower pressure than deethanizer 17, so it is pumped by pump 28 so that it can be used as reflux. After pumping, stream 38a is supplied at-60°F [-51°C] to the separator section in deethanizer 17 where the liquid is routed to the lower distillation section of deethanizer 17 as reflux for the vapors rising upward through the distillation section.

The uncondensed vapor (stream 40) from separator 29 is routed back to heat exchanger 20 to also direct it in heat exchange relation with the overhead (stream 34) from separator/absorber 15 and the uncondensed vapor (stream 41) from separator 30, cooling the stream to-102°F [-74°C] (stream 40a) and partially condensing it.

The partially condensed stream is then supplied to separator 30 where the condensed liquid (stream 39) is separated from the uncondensed vapor (stream 41). Since the operating pressure of separator 30 is lower than the operating pressure of separator/absorber 15, pump 21 is used to direct the condensed liquid (stream 39a) at -100°F [-73°C] to the separator section in separator/absorber tower 15, where the condensed liquid is routed to the lower absorbing section of separator/absorber 15 as the cold liquid that contacts the vapors rising upward through the absorbing section.

The distillation stream leaving the top of separator/absorber 15 at -107°F [-77°C] is the cold absorber overhead stream 34. The absorber overhead stream passes countercurrently to deethanizer overhead stream 36 and vapor stream 40 in heat exchanger 20 and is warmed to-74°F [-59°C] (stream 34a) as it provides cooling and partial condensation of the deethanizer overhead stream and the vapor stream. The absorber overhead stream is further warmed to 75°F [24°C] (stream 34b) as it passes countercurrently to the incoming feed gas in heat exchanger 10. The uncondensed vapor (stream 41) leaves separator 30 at-102°F [-74°C] and also passes countercurrently to deethanizer overhead stream 36 and vapor stream 40 in heat exchanger 20 and is warmed to-44°F [-42°C] (stream 41a) as it too provides cooling

and partial condensation of the streams. The vapor stream is further warmed to 65°F [18°C] (stream 41b) as it passes countercurrently to the incoming feed gas in heat exchanger 10.

The warm absorber overhead stream 34b and the warm vapor stream 41b are then re-compressed in two stages. The first stage for the absorber overhead stream is compressor 14 driven by expansion machine 13, while the first stage for the vapor stream is compressor 25 driven by a supplemental power source. The two partially compressed streams (streams 34c and 41c, respectively) combine to form the residue gas, stream 42. The combined residue gas stream then enters compressor 22 driven by a supplemental power source, which provides the second stage of compression to raise the residue gas (stream 42a) to sales line pressure. After cooling in discharge cooler 23, the residue gas product (stream 42b) flows to the sales gas pipeline at 110°F [43°C] and 1215 psia [8,377 kPa (a)].

A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 6 is set forth in the table below:

TABLE VI (FIG. 6) Stream Flow Summary-Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 81, 340 4,128 878 439 87, 840 35 11,196 2,812 879 439 15,518 36 12,511 5,182 13 0 17,956 40 11,195 2,795 3 0 14,185 38 1,316 2,388 10 0 3,771 41 7,955 585 0 0 8,632 39 3,240 2,210 3 0 5,553 34 73,384 3,526 2 0 77,875 42 81,340 4,111 3 0 86,507 37 0 18 876 439 1,333 Recoveries* Propane 99.71% Butanes+ 100.00% Power Residue Gas Compression 31,592 HP [51,937 kW] Vapor Compression 1,940 HP [3,189 kW] 33,532 HP [55, 126 kW] Utility Heat Deethanizer Reboiler 54,144 MBTU/Hr [34, 974 kW] * (Based on un-rounded flow rates) Comparison of the utility consumptions of the prior art process displayed in Table I with the utility consumptions of the present invention displayed

in Table VI shows that this embodiment of the present invention also maintains the desired C3 component recovery while substantially reducing the compression horsepower. Although the utility heat requirement is about one percent higher than the prior art process, the compression horsepower is more than ten percent lower than the prior art process, so the total utility requirements is about six percent lower than the prior art.

Comparison of the utility consumptions displayed in Tables V and VI for the FIG. 5 and FIG. 6 processes shows that the FIG. 6 embodiment of the present invention requires slightly less compression horsepower (about 2 percent) than the FIG. 5 embodiment, and uses slightly less utility heat for the deethanizer reboiler (about 4 percent), with the total utility requirements being about 3 percent lower for the FIG. 6 embodiment. The improvement in efficiency can be understood by comparing the reflux stream for deethanizer 17 (stream 38) in the FIG. 6 embodiment of the present invention with the corresponding stream in the FIG. 5 embodiment.

Whereas stream 38 in FIG. 5 is predominantly methane, stream 38 in FIG. 6 is predominantly ethane because it is withdrawn after only partial cooling in heat exchanger 20 so that proportionally less of the more volatile methane has been condensed. Not only is ethane a more effective reflux liquid than methane for rectifying the C3 and heavier components from the vapors rising upward in deethanizer 17 (as reflected by the much lower flow rate of stream 38 in the FIG. 6 embodiment), the deethanizer overhead (stream 36) has a lower concentration of methane (because less methane enters deethanizer 17 in the reflux) so that the mechanical-compression refrigeration efficiency of compressor 25 is improved.

Although this embodiment of the present invention is more efficient than the FIG. 5 embodiment, the choice of whether to include the additional equipment that the FIG. 6 process requires will generally depend on factors which include plant size and available equipment, as well as the relative costs of compression horsepower and utility heat.

Example 5 FIG. 7 illustrates a flow diagram of a process in accordance with the present invention when applied to the feed gas composition and conditions considered in the process presented in FIG. 2. Accordingly, the FIG. 7 process can be compared with that of the FIG. 2 process to illustrate the advantages of the present invention.

In the simulation of the FIG. 7 process, inlet gas enters the plant at 80°F [27°C] and 580 psia [3,999 kPa (a)] as stream 31. The feed stream 31 is cooled in exchanger 10 by heat exchange with cool residue gas at-93°F [-70°C] (stream 34a), with separator liquids at-110°F [-79°C] (stream 33a), and with separator/absorber liquids at-121°F [-85°C] (stream 35a). At this operating pressure the feed stream is below the cricondenbar, so the cooled stream 31 a enters separator 11 at-80°F [-62°C] and 570 psia [3,930 kPa (a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33).

The vapor (stream 32) from separator 11 enters work expansion machine 13 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 13 expands the vapor substantially isentropically from a pressure of about 570 psia [3,930 kPa (a)] to a pressure of about 410 psia [2,827 kPa (a)] (the operating pressure of separator/absorber 15), with the work expansion cooling the expanded stream 32a to a temperature of approximately-104°F [-76°C]. The expanded and partially condensed stream 32a enters the lower section of separator/absorber 15. The liquid portion of the expanded stream commingles with liquids falling downward from the absorbing section and the combined liquid stream 35 exits the bottom of separator/absorber 15 at-106°F [-76°C]. The vapor portion of the expanded stream rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C3 components and heavier components.

In the present invention, separator/absorber 15 operates at a higher pressure than deethanizer 17, so the combined liquid stream 35 from the bottom of the separator/absorber 15 is flash expanded to slightly above the 290 psia [1,999 kPa (a)] operating pressure of deethanizer 17 by expansion valve 27, cooling stream 35 to -121°F [-85°C] (stream 35a) before it provides cooling to the incoming feed gas as

described earlier. The combined liquid stream is heated to-85°F [-65°C], partially vaporizing stream 35b before it is supplied as a mid-column feed to deethanizer 17.

The separator liquid (stream 33) is flash expanded to slightly above the operating pressure of deethanizer 17 by expansion valve 12, cooling stream 33 to-110°F [-79°C] (stream 33a) before it provides cooling to the incoming feed gas as described earlier. Stream 33b, now at 65°F [18°C], then enters deethanizer 17 at a lower mid-column feed point. In the deethanizer, streams 35b and 33b are stripped of their methane and C2 components. The resulting liquid product stream 37 exits the bottom of the deethanizer at 164°F [73°C] and is cooled to 110°F [43°C] (stream 37a) in heat exchanger 19 before flowing to storage.

The deethanizer overhead vapor (stream 36) exits deethanizer 17 at -47°F [-44°C] and is warmed to 105°F [41°C] (stream 36a) in heat exchanger 24 before entering compressor 25 (driven by a supplemental power source). Stream 36b leaves compressor 25 at 435 psia [2,999 kPa (a)] and is cooled to 110°F [43°C] (stream 36c) in heat exchanger 26. Stream 36c is then directed in heat exchange relation with the deethanizer overhead vapor (stream 36) in heat exchanger 24 to cool it (stream 36d) and conserve process cooling.

With the increase in pressure provided by compressor 25, stream 36d can now pressure flow through heat exchanger 20 and thence to the upper feed point of separator/absorber 15. In heat exchanger 20, the compressed deethanizer overhead at-31°F [-35°C] is directed in heat exchange relation with the overhead (stream 34) from separator/absorber 15, cooling the stream to-106°F [-77°C] (stream 36e) and partially condensing it. The partially condensed stream is then supplied to the separator section in separator/absorber tower 15 where the condensed liquid is separated from the uncondensed vapor. The uncondensed vapor combines with the vapor rising from the lower absorbing section to form the cold distillation stream 34 leaving the upper region of separator/absorber 15. The condensed liquid is divided into two portions. One portion, stream 39, is routed to the lower absorbing section of separator/absorber 15 as the cold liquid that contacts the vapors rising upward through the absorbing section. The other portion, stream 38, is flash expanded to slightly above the operating pressure of deethanizer 17 by expansion valve 28 (stream 38a). It

is then supplied at-124°F [-87°C] to the separator section in deethanizer 17 where its condensed liquid is separated from its uncondensed vapor. The uncondensed vapor combines with the vapor rising from the lower distillation section to form the deethanizer overhead stream 36 leaving the upper region of deethanizer 17, while the condensed liquid is routed to the lower distillation section of deethanizer 17 as reflux for the vapors rising upward through the distillation section.

The distillation stream leaving the top of separator/absorber 15 at -111°F [-79°C] is the cold residue gas stream 34. The residue gas stream passes countercurrently to compressed deethanizer overhead stream 36 in heat exchanger 20 and is warmed to-93°F [-70°C] (stream 34a) as it provides cooling and partial condensation of the compressed deethanizer overhead stream. The residue gas is further warmed to 74°F [23'C] (stream 34b) as it passes countercurrently to the incoming feed gas in heat exchanger 10. The residue gas is then re-compressed in two stages. The first stage is compressor 14 driven by expansion machine 13. The second stage is compressor 22 driven by a supplemental power source which compresses the residue gas (stream 34d) to sales line pressure. After cooling in discharge cooler 23, the residue gas product (stream 34e) flows to the sales gas pipeline at 110°F [43°C] and 613 psia [4,226 kPa (a)].

A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 7 is set forth in the table below:

TABLE VII (FIG. 7) Stream Flow Summary-Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 81,340 4,128 878 439 87,840 32 79,899 3,596 521 106 85,149 33 1,441 533 357 333 2,691 35 1,906 891 490 106 3,435 36 3,860 1,739 69 0 5,753 39 2,051 1,325 55 0 3,487 38 513 331 14 0 872 34 81,340 4,113 87 0 86,594 37 0 16 791 439 1,246 Recoveries* Propane 90.05% Butanes+ 100.00% Power Residue Gas Compression 15,990 HP [26,287 kW] Overhead Vapor Compression 1,414 HP [2,325 kW] 17,404 HP [28, 612 kW] Utility Heat DeethanizerReboiler 11, 515 MBTU/Hr [7, 438 kW] * (Based on un-rounded flow rates) Comparison of the utility consumptions of the prior art process displayed in Table II with the utility consumptions of the present invention displayed in Table VII shows that the present invention maintains the desired C3 component

recovery while substantially reducing the utility heat requirement and reducing the compression horsepower. The utility heat requirement is more than thirty-five percent lower than the prior art process, while the compression horsepower is eight percent lower than the prior art process.

Comparing the present invention to the prior art process displayed in FIG. 2, note that the operating pressure of deethanizer 17 is significantly lower in the present invention than in the FIG. 2 process, 290 psia [1,999 kPa (a)] versus 395 psia [2,723 kPa (a)], and the operating pressure of separator/absorber 15 is significantly higher in the present invention than in the FIG. 2 process, 410 psia [2,827 kPa (a)] versus 380 psia [2,620 kPa (a)]. Accordingly, the residue gas enters compressor 14 at a higher pressure in the FIG. 7 process and less compression horsepower is therefore needed to deliver the residue gas to pipeline pressure. Further, with separator/absorber 15 operating at a higher pressure than deethanizer 17, it is no longer necessary to pump the absorber bottom liquid (stream 35) and the reflux stream (stream 38) to feed deethanizer 17, eliminating the capital and operating cost of pumps 16 and 21 in the FIG. 2 process.

As described earlier, the deethanizer overhead (stream 36) in the FIG. 7 process provides a more efficient working fluid for a mechanical-compression refrigeration cycle than the inlet gas (stream 31) and residue gas (stream 34) which are predominantly methane, so that the refrigeration provided to the process by the cycle including compressor 25 not only reduces the refrigeration required from the cycle using compressors 14 and 22, but reduces the total refrigeration energy consumption as well. Also, note that the liquid streams used to provide part of the feed gas cooling, the cold separator liquids (stream 33) and the cold liquid (stream 35) leaving the bottom of separator/absorber 15, are cooled by flash expansion (streams 33a and 35a, respectively) before entering heat exchanger 10. As a result, these streams are considerably colder than the corresponding streams in the FIG. 2 process, allowing better heat integration and more efficient process cooling than the prior art process can provide. This is a consequence of operating deethanizer 17 at a lower pressure than separator/absorber 15 in the present invention, which is not possible with the prior art process.

The substantial reduction in the utility heat required for deethanizer reboiler 18 for the present invention is a consequence of the lower operating pressure that is possible for deethanizer 17 in the FIG. 7 process. With the pressure lower in deethanizer 17, the bubble point temperatures of all the liquid streams in the column are lower, including the bottom liquid product (164°F [73°C] for stream 37 in the FIG. 7 process, versus 195°F [91°C] for stream 37 in the FIG. 2 process). Thus, much less sensible heating is required for the column liquids in deethanizer 17, reducing the heating load in reboiler 18 accordingly. Total energy consumption for the FIG. 7 embodiment of the present invention is only 85 percent of that required for the prior art of FIG. 2.

Example 6 In the embodiment of the present invention shown in FIG. 7, the process was operated to achieve the same C3 component recovery level as the prior art process shown in FIG. 2, with the resulting reduction in the utility consumption due to the better efficiency of the present invention. Alternatively, it is also possible to adjust the operating conditions of the present invention to increase the C3 component recovery level while keeping the utility consumption the same as the prior art process, or to provide some combination of better recovery and lower utility consumption. For example, FIG. 8 shows the present invention when applied to match the compression power used by the prior art FIG. 2 process. The feed gas composition and conditions considered in the process presented in FIG. 8 are the same as those in FIG. 2.

Accordingly, the FIG. 8 process can be compared with that of the FIG. 2 process to illustrate the advantages of the present invention.

In the simulation of the FIG. 8 process, the feed gas cooling and expansion scheme, the deethanizer overhead compression and cooling scheme, and the tower reflux schemes are essentially the same as those used in FIG. 7. The only difference for the FIG. 8 embodiment of the present invention is that the operating pressures of separator/absorber 15 and deethanizer 17 have been adjusted to increase the recovery level for the C3 components, with the corresponding drops in the process operating temperatures that result from the increase in process cooling (due primarily

to the increase in expansion ratio across work expansion machine 13). Note that relative to the FIG. 7 embodiment, in the FIG. 8 embodiment the operating pressure of separator/absorber 15 has been lowered from 410 psia [2,827 kPa (a)] to 395 psia [2,723 kPa (a)], and the operating pressure of deethanizer 17 has been lowered from 290 psia [1,999 kPa (a)] to 285 psia [1,965 kPa (a)].

A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 8 is set forth in the table below: TABLE VIII (FIG. 8) Stream Flow Summary-Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 81, 340 4,128 878 439 87,840 32 79,899 3,596 521 106 85,149 33 1,441 533 357 333 2,691 35 2,186 1,055 531 106 3,928 36 4,209 1,944 53 0 6,299 39 2,328 1,493 42 0 3,926 38 582 373 11 0 982 34 81,340 4,112 33 0 86,538 37 0 17 846 439 1,302

Recoveries* Propane 96.30% Butanes+ 100.00% Power Residue Gas Compression 17,428 HP [28, 651 kW] Overhead Vapor Compression 1,483 HP [2, 43 8 kW] 18, 911 HP [31, 089 kW] Utility Heat Deethanizer Reboiler 12,909 MBTU/Hr [8, 339 kW] * (Based on un-rounded flow rates) Comparison of the utility consumptions of the prior art process displayed in Table II with the utility consumptions of the present invention displayed in Table VIII shows that the present invention uses the same amount of external power for compression as the prior art process while increasing the C3 component recovery and substantially reducing the utility heat requirement. The C3 component recovery increases from 90.09% in the prior art FIG. 2 process to 96.30% in the present invention, an increase of over six percentage points. The utility heat requirement for the present invention is more than twenty-eight percent lower than the prior art process. The choice of whether to apply the present invention to increase the C3 component recovery level, to reduce the utility consumptions, or to provide some combination of increased recovery and reduced utility consumption will normally be governed by the specific circumstances of each application, as the optimum will depend on such factors as plant size, available equipment, and the relative values of the recovered liquid product components and the utilities consumed.

Other Embodiments In accordance with this invention, it is generally advantageous to design the separator/absorber to provide a contacting device composed of multiple

theoretical separation stages. However, 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. For instance, all or a part of the partially condensed stream leaving heat exchanger 20 and all or a part of the partially condensed stream from work expansion machine 13 can be combined (such as in the piping joining the expansion machine to the separator/absorber) 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. In such an embodiment, the vapor-liquid mixture from heat exchanger 20 can be used without separation, or the liquid portion thereof may be separated. Such commingling of the two streams shall be considered for the purposes of this invention as constituting a contacting device. In another variation of the foregoing, the partially condensed stream from heat exchanger 20 can be separated (using separator 30 as shown in FIG. 9, for instance), and then all or a part of the separated liquid supplied to the separator/absorber or mixed with the vapors fed thereto (with any remaining portion of the separated liquid supplied to the deethanizer).

As described earlier in the preferred embodiment, the overhead vapors are partially condensed and used to absorb valuable C3 components and heavier components from the vapors leaving the work expansion machine. However, the present invention is not limited to this embodiment. It may be advantageous, for instance, to treat only a portion of the outlet vapor from the work expansion machine in this manner, or to use only a portion of the overhead condensate as an absorbent, in cases where other design considerations indicate portions of the expansion machine outlet or overhead condensate should bypass the separator/absorber. Feed gas conditions, plant size, available equipment, or other factors may indicate that elimination of work expansion machine 13, or replacement with an alternate expansion device (such as an expansion valve), is feasible, or that total (rather than partial) condensation of the overhead stream in heat exchanger 20 is possible or is preferred. It should also be noted that the separator/absorber can be constructed either as a separate vessel or as a section of the deethanizer column.

The use and distribution of the separator liquids, the separator/absorber liquids, and the reflux liquids for process heat exchange, the particular arrangement of heat exchangers for feed gas cooling, and the choice of process streams for specific heat exchange services must be evaluated for each particular application. For instance, as shown in FIG. 10, all or a part of the separator liquids (stream 33) may be routed directly to deethanizer 17 via an expansion device (such as expansion valve 12a shown in FIG. 10), with part or none of the liquid used for process cooling in heat exchanger 10. Similarly, all or a part of the separator/absorber liquids (stream 35) may be routed directly to deethanizer 17 via an expansion device (such as expansion valve 27a shown in FIG. 10), with part or none of the liquid used for process cooling in heat exchanger 10. Additionally, the condensed liquid that serves as reflux for deethanizer 17 (stream 38 in FIG. 10) can be used for process cooling before being supplied to the column. As shown in FIG. 10, all or a part of this liquid may be let down to slightly above the operating pressure of deethanizer 17 (using a device such as expansion valve 28) and used for process cooling (such as in heat exchanger 20 as shown) before being routed to deethanizer 17, with part or none of the liquid routed directly to deethanizer 17 (via expansion valve 28a, for example).

Moreover, the use of external refrigeration to supplement the cooling available to the feed gas from other process streams may be employed as illustrated in FIG. 11, particularly in the case of an inlet gas richer than that used in Example 1.

External refrigeration may also be employed to generate some or all of the reflux for the deethanizer as illustrated in FIG. 11. In such cases, all of the condensed liquid contained in the partially condensed stream leaving heat exchanger 20 (stream 36e in FIG. 11) might be directed only to the separator/absorber rather than a portion feeding the deethanizer. Note also in FIG. 11 (and as was shown previously in FIG. 10) that other process streams and/or external refrigeration may be used to supplement the cooling provided to the deethanizer overhead by the separator/absorber overhead (stream 34) in heat exchanger 20, such as the flash expanded liquid (stream 35a in FIG. 11) from the bottom of the separator/absorber.

Still other alternative means for generating the reflux stream for the deethanizer may be advantageous, depending of the particular application of the

present invention. For example, as shown in FIG. 12, the heated flash expanded liquid (stream 35b) from the bottom of the separator/absorber could be used to cool a distillation stream (stream 40) from the deethanizer in heat exchanger 50 to partially condense the distillation stream (stream 40a), whereupon the condensed liquid (stream 38) is separated from the uncondensed vapor (stream 36) in separator 51.

Reflux pump 52 could then direct the condensed liquid (stream 38a) to deethanizer 17 to serve as its reflux, with the further heated stream 35c from heat exchanger 50 feeding deethanizer 17 at a mid-column feed point. Depending on the particular circumstances, the heated flash expanded liquid (stream 35b) from the bottom of the separator/absorber may contain an adequate quantity of liquid to serve as the reflux for the deethanizer, as shown in FIG. 13 and by the dashed lines in FIGS. 15 and 17.

Further, as shown in FIGS. 13 through 18, it may be advantageous to direct some or all of the flash expanded separator liquid (stream 33a) to separator/absorber 15 rather than to deethanizer 17, either to a separate fractionation zone in separator/absorber 15 or to the same fractionation zone as the outlet (stream 32a) from work expansion machine 13.

It will also be recognized that the relative amount of feed found in each branch of the condensed liquid contained in stream 36e that is split between the two towers in FIGS. 3,7, and 8 will depend on several factors, including gas pressure, feed gas composition and the quantity of horsepower available. Similarly, the relative amount of condensation in separator 30 in FIG. 4, the relative amount of feed contained in stream 40a that is split between the two towers in FIG. 5, and the relative amount of condensation in separators 29 and 30 in FIG. 6 will also depend on factors such as these. The optimum split or distribution generally cannot be predicted without evaluating the particular circumstances for a specific application of the present invention. The mid-column feed positions depicted in FIGS. 3 through 8 are the preferred feed locations for the process operating conditions described. However, the relative locations of the mid-column feeds may vary depending on inlet composition or other factors such as desired recovery levels, etc. 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. FIGS. 3 through 8 are the preferred embodiments for the compositions and pressure conditions shown. Although individual stream expansion is depicted in particular expansion devices, alternative expansion means may be employed where appropriate. For example, conditions may warrant work expansion of some or all of the liquid streams (such as streams 33,35, and/or 38 in FIG. 7).

It will also be recognized that the manner in which the deethanizer overhead stream (stream 36 in FIGS. 3,4,7 through 14,17, and 18) or the vapor stream (stream 41 in FIGS. 5 and 6) is compressed can be accomplished in a variety of ways. FIGS. 3 through 14,17, and 18 depict using a supplemental power source for compressor 25 to compress this stream, while compressing the residue gas (stream 34b in FIGS. 3, 4,7, and 8) or absorber overhead (stream 34b in FIGS. 5 and 6) using compressor 14 driven by expansion machine 13. In other circumstances, for instance, it may be desirable to drive compressor 25 with expansion machine 13 and use a supplemental power source for compressor 14. It may also be desirable to combine the services of compressor 25 and compressor 14 into a compound machine driven by expansion machine 13. Further, as shown by the dashed equipment in FIGS. 3,4,7 through 14,17, and 18, some circumstances may favor reducing the capital cost of the facility by eliminating heat exchanger 24 and/or heat exchanger 26 (at the expense of increasing the cooling load on heat exchanger 20 and either reducing the product recoveries or increasing the power consumption of compressor 22). Choices such as these must generally be evaluated for each application, as factors such as gas composition, plant size, desired recovery level, and available equipment must all be considered.

The present invention provides improved recovery of C3 components per amount of utility consumption required to operate the process. An improvement in utility consumption required for operating the deethanizer process may appear in the form of reduced power requirements for compression or re-compression, reduced power requirements for external refrigeration, reduced energy requirements for tower reboilers, or a combination thereof. Alternatively, if desired, increased C3 component recovery can be obtained for a fixed utility consumption.

It should also be noted that complete rejection of the C2 components to the residue gas is not required by the present invention. If the project economics favor recovery of the C2 components in the liquid product (stream 37), the process operating conditions can be altered to recover in the liquid product a significant portion of the C2 components present in the feed gas. Preliminary calculations indicate that perhaps 40% of the C2 components can be recovered in this fashion.

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.