Hydrocarbon gas processing

Abstract

A process and an apparatus are disclosed for a compact processing assembly to improve the recovery of C.sub.2 (or C.sub.3) and heavier hydrocarbon components from a hydrocarbon gas stream. The preferred method of separating a hydrocarbon gas stream generally includes producing at least a substantially condensed first stream and a cooled second stream, expanding both streams to lower pressure, and supplying the streams to a fractionation tower. In the process and apparatus disclosed, the expanded first stream is heated to form a vapor fraction and a liquid fraction. The vapor fraction is combined with the tower overhead vapor, directed to a heat and mass transfer means inside a processing assembly, and cooled and partially condensed by the expanded first stream to form a residual vapor stream and a condensed stream. The condensed stream is combined with the liquid fraction and supplied to the tower at its top feed point.

Claims

1. In an apparatus for the separation of a gas stream containing methane, C.sub.2 components, C.sub.3 components, and heavier hydrocarbon components into a volatile residue gas fraction and a relatively less volatile fraction containing a major portion of said C.sub.2 components, C.sub.3 components, and heavier hydrocarbon components or said C.sub.3 components and heavier hydrocarbon components, in said apparatus there being (a) one or more heat exchange means and at least one dividing means to produce at least a first stream that has been cooled under pressure to condense substantially all of it, and at least a second stream that has been cooled under pressure; (b) a first expansion means connected to receive said substantially condensed first stream under pressure and expand it to a lower pressure, whereby said first stream is further cooled; (c) a distillation column connected to said first expansion means to receive said expanded further cooled first stream at a top feed position, with said distillation column producing at least an overhead vapor stream and a bottom liquid stream; (d) a second expansion means connected to receive said cooled second stream under pressure and expand it to said lower pressure; (e) said distillation column further connected to said second expansion means to receive said expanded second stream at a mid-column feed position; and (f) said distillation column adapted to fractionate at least said expanded further cooled first stream and said expanded second stream at said lower pressure whereby the components of said relatively less volatile fraction are recovered in said bottom liquid stream and said volatile residue gas fraction is discharged as said overhead vapor stream; the improvement wherein said apparatus includes (1) a heat and mass transfer means housed in a processing assembly and connected to said first expansion means to receive said expanded further cooled first stream and heat it, and thereafter discharging said heated expanded first stream as a vapor fraction and a liquid fraction; (2) a first combining means housed in said processing assembly connected to said distillation column and to said heat and mass transfer means to receive said overhead vapor stream and said vapor fraction and form a combined vapor stream; (3) said heat and mass transfer means being further connected to said first combining means to receive said combined vapor stream and cool it, thereby to supply at least a portion of the heating of element (1) while simultaneously condensing the less volatile components from said combined vapor stream, thereby forming a condensed stream and a residual vapor stream, and thereafter discharging said residual vapor stream from said processing assembly as said volatile residue gas fraction; (4) a second combining means housed in said processing assembly connected to said heat and mass transfer means to receive said condensed stream and said liquid fraction and form a combined liquid stream; (5) said second combining means being further connected to said distillation column to supply said combined liquid stream to said top feed position of said distillation column; and (6) control means adapted to regulate the quantity and temperature of said combined liquid stream to said distillation column to maintain the overhead temperature of said distillation column at a temperature whereby the major portions of the components in said relatively less volatile fraction are recovered in said bottom liquid stream.

2. The apparatus according to claim 1 wherein (1) a pumping means is connected to said second combining means to receive said combined liquid stream and pump it to higher pressure; and (2) said pumping means is further connected to said distillation column to supply said combined liquid stream to said top feed position of said distillation column.

3. The apparatus according to claim 2 wherein said pumping means is housed in said processing assembly.

4. The apparatus according to claim 1 wherein (1) said one or more heat exchange means is adapted to cool said gas stream under pressure sufficiently to partially condense it; (2) a separating means is connected to said one or more heat exchange means to receive said partially condensed gas stream and separate it into a vapor stream and at least one liquid stream; (3) said at least one dividing means is connected to said separating means and adapted to receive said vapor stream and divide it into at least a first vapor stream and a second vapor stream; (4) said one or more heat exchange means is connected to said at least one dividing means and adapted to receive said first vapor stream and cool it sufficiently to substantially condense it and thereby form said substantially condensed first stream; (5) said second expansion means is connected to said at least one dividing means and adapted to receive said second vapor stream and expand it to said lower pressure, thereby forming said expanded second stream; (6) a third expansion means is connected to said separating means to receive at least a portion of said at least one liquid stream and expand it to said lower pressure, said third expansion means being further connected to said distillation column to supply said expanded liquid stream to said distillation column at a lower mid-column feed position below said mid-column feed position; and (7) said distillation column is adapted to fractionate at least said expanded further cooled first stream, said expanded second stream, and said expanded liquid stream at said lower pressure whereby the components of said relatively less volatile fraction are recovered in said bottom liquid stream.

5. The apparatus according to claim 4 wherein (1) a third combining means is connected to said at least one dividing means and said separating means to receive said first vapor stream and at least a portion of said at least one liquid stream and form a combined stream; (2) said one or more heat exchange means is connected to said third combining means and adapted to receive said combined stream and cool it sufficiently to substantially condense it; (3) said first expansion means is connected to said one or more heat exchange means and adapted to receive said substantially condensed combined stream and expand it to said lower pressure whereby it is further cooled; (4) said heat and mass transfer means is adapted to receive said expanded further cooled combined stream and heat it, thereafter discharging said heated expanded combined stream as said vapor fraction and said liquid fraction; and (5) said third expansion means is adapted to receive any remaining portion of said at least one liquid stream and expand it to said lower pressure, whereupon said expanded liquid stream is supplied to said distillation column at said lower mid-column feed position.

6. The apparatus according to claim 2 wherein (1) said one or more heat exchange means is adapted to cool said gas stream under pressure sufficiently to partially condense it; (2) a separating means is connected to said one or more heat exchange means to receive said partially condensed gas stream and separate it into a vapor stream and at least one liquid stream; (3) said at least one dividing means is connected to said separating means and adapted to receive said vapor stream and divide it into at least a first vapor stream and a second vapor stream; (4) said one or more heat exchange means is connected to said at least one dividing means and adapted to receive said first vapor stream and cool it sufficiently to substantially condense it and thereby form said substantially condensed first stream; (5) said second expansion means is connected to said at least one dividing means and adapted to receive said second vapor stream and expand it to said lower pressure, thereby forming said expanded second stream; (6) a third expansion means is connected to said separating means to receive at least a portion of said at least one liquid stream and expand it to said lower pressure, said third expansion means being further connected to said distillation column to supply said expanded liquid stream to said distillation column at a lower mid-column feed position below said mid-column feed position; and (7) said distillation column is adapted to fractionate at least said expanded further cooled first stream, said expanded second stream, and said expanded liquid stream at said lower pressure whereby the components of said relatively less volatile fraction are recovered in said bottom liquid stream.

7. The apparatus according to claim 3 wherein (1) said one or more heat exchange means is adapted to cool said gas stream under pressure sufficiently to partially condense it; (2) a separating means is connected to said one or more heat exchange means to receive said partially condensed gas stream and separate it into a vapor stream and at least one liquid stream; (3) said at least one dividing means is connected to said separating means and adapted to receive said vapor stream and divide it into at least a first vapor stream and a second vapor stream; (4) said one or more heat exchange means is connected to said at least one dividing means and adapted to receive said first vapor stream and cool it sufficiently to substantially condense it and thereby form said substantially condensed first stream; (5) said second expansion means is connected to said at least one dividing means and adapted to receive said second vapor stream and expand it to said lower pressure, thereby forming said expanded second stream; (6) a third expansion means is connected to said separating means to receive at least a portion of said at least one liquid stream and expand it to said lower pressure, said third expansion means being further connected to said distillation column to supply said expanded liquid stream to said distillation column at a lower mid-column feed position below said mid-column feed position; and (7) said distillation column is adapted to fractionate at least said expanded further cooled first stream, said expanded second stream, and said expanded liquid stream at said lower pressure whereby the components of said relatively less volatile fraction are recovered in said bottom liquid stream.

Description

(1) FIGS. 1, 2, and 3 are flow diagrams of prior art natural gas processing plants in accordance with U.S. Pat. No. 4,157,904 or 4,278,457;

(2) FIGS. 4, 5, and 6 are flow diagrams of natural gas processing plants adapted to use the present invention; and

(3) FIGS. 7 and 8 are flow diagrams illustrating alternative means of application of the present invention to a natural gas processing plant.

(4) 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.

(5) For convenience, process parameters are reported in both the traditional British units and in the units of the Syst?me International d'Unit?s (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.

DESCRIPTION OF THE PRIOR ART

(6) FIG. 1 is a process flow diagram showing the design of a processing plant to recover C.sub.2+ components from natural gas using prior art according to U.S. Pat. No. 4,157,904 or 4,278,457. In this simulation of the process, inlet gas enters the plant at 100? F. [38? C.] and 915 psia [6,307 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.

(7) The feed stream 31 is cooled in heat exchanger 10 by heat exchange with cool residue gas (stream 39a), demethanizer reboiler liquids at 44? F. [7? C.] (stream 41), and demethanizer side reboiler liquids at ?49? F. [?45? C.] (stream 40). (In some cases, the use of one or more supplemental external refrigeration streams may be advantageous as shown by the dashed line.) Stream 31a then enters separator 11 at ?24? F. [?31? C.] and 900 psia [6,203 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33).

(8) 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 FIG. 1 is according to U.S. Pat. No. 4,157,904. Otherwise, the process of FIG. 1 is according to U.S. Pat. No. 4,278,457.) For the process illustrated in FIG. 1, stream 35 contains 100% of the total separator liquid. Stream 34, containing about 31% of the total separator vapor, is combined with stream 35 and the combined stream 36 passes through heat exchanger 12 in heat exchange relation with the cold residue gas (stream 39) where it is cooled to substantial condensation. The resulting substantially condensed stream 36a at ?134? F. [?92? C.] is then flash expanded through expansion valve 13 to the operating pressure (approximately 395 psia [2,721 kPa(a)]) of fractionation tower 17. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in FIG. 1, the expanded stream 36b leaving expansion valve 13 reaches a temperature of ?140? F. [?96? C.] and is supplied to separator section 17a in the upper region of fractionation tower 17. The liquids separated therein become the top feed to demethanizing section 17b.

(9) The remaining 69% 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 ?95? F. [?70? C.]. The typical commercially available expanders are capable of recovering on the order of 80-85% of the work theoretically available in an ideal isentropic expansion. The work recovered is often used to drive a centrifugal compressor (such as item 15) that can be used to re-compress the residue gas (stream 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 before it is supplied to fractionation tower 17 at a lower mid-column feed point.

(10) 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.

(11) The liquid product stream 42 exits the bottom of the tower at 67? F. [19? C.], based on a typical specification of a methane to ethane ratio of 0.010:1 on a molar basis in the bottom product. The residue gas (demethanizer overhead vapor stream 39) passes countercurrently to the incoming feed gas in heat exchanger 12 where it is heated from ?139? F. [?95? C.] to ?37? F. [?38? C.] (stream 39a) and in heat exchanger 10 where it is heated to 91? F. [33? 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 110? F. [43? C.] in discharge cooler 20, the residue gas product (stream 39e) flows to the sales gas pipeline at 915 psia [6,307 kPa(a)], sufficient to meet line requirements (usually on the order of the inlet pressure).

(12) A summary of stream flow rates and energy consumption for the process illustrated in FIG. 1 is set forth in the following table:

(13) TABLE-US-00001 TABLE I (FIG. 1) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 12,398 546 233 229 13,726 32 12,202 504 183 82 13,281 33 196 42 50 147 445 34 3,909 161 59 26 4,255 36 4,105 203 109 173 4,700 37 8,293 343 124 56 9,026 39 12,393 55 5 1 12,636 42 5 491 228 228 1,090 Recoveries* Ethane 89.85% Propane 98.05% Butanes+ 99.71% Power Residue Gas Compression 5,569 HP [9,155 kW] *(Based on un-rounded flow rates)

(14) FIG. 2 is a process flow diagram showing one manner in which the design of the processing plant in FIG. 1 can be adjusted to operate at a lower C.sub.2 component recovery level. This is a common requirement when the relative values of natural gas and liquid hydrocarbons are variable, causing recovery of the C.sub.2 components to be unprofitable at times. The process of FIG. 2 has been applied to the same feed gas composition and conditions as described previously for FIG. 1. However, in the simulation of the process of FIG. 2, the process operating conditions have been adjusted to reject nearly all of C.sub.2 components to the residue gas rather than recovering them in the bottom liquid product from the fractionation tower.

(15) In this simulation of the process, inlet gas enters the plant at 100? F. [38? C.] and 915 psia [6,307 kPa(a)] as stream 31 and is cooled in heat exchanger 10 by heat exchange with cool residue gas stream 39a. (One consequence of operating the FIG. 2 process to reject nearly all of the C.sub.2 components to the residue gas is that the temperatures of the liquids flowing down fractionation tower 17 are much warmer, to the point that side reboiler stream 40 and reboiler stream 41 can no longer be used to cool the inlet gas and all of the column reboil heat must be supplied by supplemental reboiler 18.) Cooled stream 31a enters separator 11 at 2? F. [?17? C.] and 900 psia [6,203 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33).

(16) 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 FIG. 2, stream 35 contains 100% of the total separator liquid. Stream 34, containing about 24% of the total separator vapor, is combined with stream 35 and the combined stream 36 passes through heat exchanger 12 in heat exchange relation with the cold residue gas (stream 39) where it is cooled to substantial condensation. The resulting substantially condensed stream 36a at ?102? F. [?75? C.] is then flash expanded through expansion valve 13 to the operating pressure (approximately 398 psia [2,747 kPa(a)]) of fractionation tower 17. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in FIG. 2, the expanded stream 36b leaving expansion valve 13 reaches a temperature of ?137? F. [?94? C.] and is supplied to fractionation tower 17 at the top feed point.

(17) The remaining 76% 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 ?71? F. [?57? 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 the operating pressure of fractionation tower 17 by expansion valve 16, cooling stream 38a before it is supplied to fractionation tower 17 at a lower mid-column feed point.

(18) Note that when fractionation tower 17 is operated to reject the C.sub.2 components to the residue gas product as shown in FIG. 2, the column is typically referred to as a deethanizer and its lower section 17b is called a deethanizing section. The liquid product stream 42 exits the bottom of deethanizer 17 at 230? F. [110? C.], based on a typical specification of an ethane to propane ratio of 0.020:1 on a molar basis in the bottom product. The residue gas (deethanizer overhead vapor stream 39) passes countercurrently to the incoming feed gas in heat exchanger 12 where it is heated from ?108? F. [?78? C.] to ?36? F. [?38? C.] (stream 39a) and in heat exchanger 10 where it is heated to 99? F. [37? C.] (stream 39b) as it provides cooling as previously described. The residue gas is then re-compressed in two stages, compressor 15 driven by expansion machine 14 and compressor 19 driven by a supplemental power source. After stream 39d is cooled to 110? F. [43? C.] in discharge cooler 20, the residue gas product (stream 39e) flows to the sales gas pipeline at 915 psia [6,307 kPa(a)].

(19) A summary of stream flow rates and energy consumption for the process illustrated in FIG. 2 is set forth in the following table:

(20) TABLE-US-00002 TABLE II (FIG. 2) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 12,398 546 233 229 13,726 32 12,304 526 208 117 13,470 33 94 20 25 112 256 34 3,040 130 51 29 3,328 36 3,134 150 76 141 3,584 37 9,264 396 157 88 10,142 39 12,398 542 15 2 13,276 42 0 4 218 227 450 Recoveries* Propane 93.60% Butanes+ 99.12% Power Residue Gas Compression 5,565 HP [9,149 kW] *(Based on un-rounded flow rates)

(21) Product economics sometimes favor rejecting only a portion of the C.sub.2 components to the residue gas product. FIG. 3 is a process flow diagram showing one manner in which the design of the processing plant in FIG. 1 can be adjusted to operate at an intermediate C.sub.2 component recovery level. The process of FIG. 3 has been applied to the same feed gas composition and conditions as described previously for FIGS. 1 and 2. However, in the simulation of the process of FIG. 3, the process operating conditions have been adjusted to recover about half as much of the C.sub.2 components in the bottom liquid product from the fractionation tower compared to the quantity of C.sub.2 components recovered by the FIG. 1 process.

(22) In this simulation of the process, inlet gas enters the plant at 100? F. [38? C.] and 915 psia [6,307 kPa(a)] as stream 31 and is cooled in heat exchanger 10 by heat exchange with cool residue gas stream 39a and demethanizer side reboiler liquids at 63? F. [17? C.] (stream 40). (At the C.sub.2 component recovery level of the FIG. 3 process, the side reboiler stream 40 is still cool enough to be used to cool the inlet gas, reducing the amount of column reboil heat that must be supplied by supplemental reboiler 18.) Cooled stream 31a enters separator 11 at 8? F. [?14? C.] and 900 psia [6,203 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33).

(23) 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 FIG. 3, stream 35 contains 100% of the total separator liquid. Stream 34, containing about 27% of the total separator vapor, is combined with stream 35 and the combined stream 36 passes through heat exchanger 12 in heat exchange relation with the cold residue gas (stream 39) where it is cooled to substantial condensation. The resulting substantially condensed stream 36a at ?120? F. [?85? C.] is then flash expanded through expansion valve 13 to the operating pressure (approximately 384 psia [2,644 kPa(a)]) of fractionation tower 17. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in FIG. 3, the expanded stream 36b leaving expansion valve 13 reaches a temperature of ?141? F. [?96? C.] and is supplied to fractionation tower 17 at the top feed point.

(24) The remaining 73% 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 ?69? F. [?56? 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 the operating pressure of fractionation tower 17 by expansion valve 16, cooling stream 38a before it is supplied to fractionation tower 17 at a lower mid-column feed point.

(25) The liquid product stream 42 exits the bottom of the tower at 130? F. [54? C.]. The residue gas (deethanizer overhead vapor stream 39) passes countercurrently to the incoming feed gas in heat exchanger 12 where it is heated from ?122? F. [?86? C.] to ?29? F. [?34? C.] (stream 39a) and in heat exchanger 10 where it is heated to 86? F. [30? C.] (stream 39b) as it provides cooling as previously described. The residue gas is then re-compressed in two stages, compressor 15 driven by expansion machine 14 and compressor 19 driven by a supplemental power source. After stream 39d is cooled to 110? F. [43? C.] in discharge cooler 20, the residue gas product (stream 39e) flows to the sales gas pipeline at 915 psia [6,307 kPa(a)].

(26) A summary of stream flow rates and energy consumption for the process illustrated in FIG. 3 is set forth in the following table:

(27) TABLE-US-00003 TABLE III (FIG. 3) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 12,398 546 233 229 13,726 32 12,316 529 211 124 13,496 33 82 17 22 105 230 34 3,351 144 57 34 3,671 36 3,433 161 79 139 3,901 37 8,965 385 154 90 9,825 39 12,398 300 8 1 13,025 42 0 246 225 228 701 Recoveries* Ethane 45.00% Propane 96.51% Butanes+ 99.56% Power Residue Gas Compression 5,564 HP [9,147 kW] *(Based on un-rounded flow rates)

DESCRIPTION OF THE INVENTION

Example 1

(28) In those cases where the C.sub.2 component recovery level in the liquid product must be reduced (as in the FIG. 2 prior art process described previously, for instance), the present invention offers significant efficiency advantages over the prior art process depicted in FIG. 2. FIG. 4 illustrates a flow diagram of the FIG. 2 prior art process that has been adapted to use the present invention. The operating conditions of the FIG. 4 process have been adjusted as shown to reduce the ethane content of the liquid product to the same level as that of the FIG. 2 prior art process. The feed gas composition and conditions considered in the process presented in FIG. 4 are the same as those in FIG. 2. Accordingly, the FIG. 4 process can be compared with that of the FIG. 2 process to illustrate the advantages of the present invention.

(29) Most of the process conditions shown for the FIG. 4 process are much the same as the corresponding process conditions for the FIG. 2 process. The main differences are the disposition of flash expanded substantially condensed stream 36b and column overhead vapor stream 39. In the FIG. 4 process, substantially condensed stream 36a is flash expanded through expansion valve 13 to slightly above the operating pressure (approximately 402 psia [2,774 kPa(a)]) 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 FIG. 4, the expanded stream 36b leaving expansion valve 13 reaches a temperature of ?138? F. [?94? C.] before it is directed into a heat and mass transfer means inside rectifying section 117a of processing assembly 117. This heat and mass transfer means may be comprised of 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 heat and mass transfer means is configured to provide heat exchange between a combined vapor stream flowing upward through one pass of the heat and mass transfer means, and the flash expanded substantially condensed stream 36b flowing downward, so that the combined vapor stream is cooled while heating the expanded stream. As the combined vapor stream is cooled, a portion of it is condensed and falls downward while the remaining combined vapor stream continues flowing upward through the heat and mass transfer means. The heat and mass transfer means provides continuous contact between the condensed liquid and the combined vapor stream so that it also functions to provide mass transfer between the vapor and liquid phases, thereby providing rectification of the combined vapor stream. The condensed liquid from the bottom of the heat and mass transfer means is directed to separator section 117b of processing assembly 117.

(30) 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 117a at ?105? F. [?76? C.]. The heated flash expanded stream discharges into separator section 117b 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 heat and mass transfer means in rectifying section 117a as previously described, and the liquid phase combines with the condensed liquid from the bottom of the heat and mass transfer means to form combined liquid stream 152. Combined liquid stream 152 leaves the bottom of processing assembly 117 and is pumped to higher pressure by pump 21 so that stream 152a at ?102? F. [?75? C.] can enter fractionation column 17 at the top feed point. The vapor remaining from the cooled combined vapor stream leaves the heat and mass transfer means inside rectifying section 117a of processing assembly 117 at ?117? F. [?83? C.] as cold residue gas stream 151, which is then heated and compressed as described previously for stream 39 in the FIG. 2 process.

(31) A summary of stream flow rates and energy consumption for the process illustrated in FIG. 4 is set forth in the following table:

(32) TABLE-US-00004 TABLE IV (FIG. 4) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 12,398 546 233 229 13,726 32 12,318 529 212 125 13,499 33 80 17 21 104 227 34 3,570 153 61 36 3,912 36 3,650 170 82 140 4,139 37 8,748 376 151 89 9,587 39 9,525 856 31 4 10,699 152 777 485 112 144 1,578 151 12,398 541 1 0 13,260 42 0 5 232 229 466 Recoveries* Propane 99.65% Butanes+ 100.00% Power Residue Gas Compression 5,565 HP [9,149 kW] *(Based on un-rounded flow rates)

(33) A comparison of Tables II and IV shows that, compared to the prior art, the present invention improves propane recovery from 93.60% to 99.65% and butane+ recovery from 99.12% to 100.00%. The economic impact of these improved recoveries is significant. Using an average incremental value $1.08/gallon [214/m.sup.3] for hydrocarbon liquids compared to the corresponding hydrocarbon gases, the improved recoveries represent more than US$1,120,000 [835,000] of additional annual revenue for the plant operator. Comparison of Tables II and IV further shows that these increased product yields were achieved using the same power as the prior art. In terms of the recovery efficiency (defined by the quantity of C.sub.3 components and heavier components recovered per unit of power), the present invention represents more than a 3% improvement over the prior art of the FIG. 2 process.

(34) The improvement in recovery efficiency provided by the present invention over that of the prior art of the FIG. 2 process is primarily due to the indirect cooling of the column vapor provided by flash expanded stream 36b in rectifying section 117a of processing assembly 117, rather than the direct-contact cooling that characterizes stream 36b in the prior art process of FIG. 2. Although stream 36b is quite cold, it is not an ideal reflux stream because it contains significant concentrations of the C.sub.3 components and C.sub.4+ components that deethanizer 17 is supposed to capture, resulting in losses of these desirable components due to equilibrium effects at the top of column 17 for the prior art process of FIG. 2. For the present invention shown in FIG. 4, however, there are no equilibrium effects to overcome because there is no direct contact between flash expanded stream 36b and the combined vapor stream to be rectified.

(35) The present invention has the further advantage of using the heat and mass transfer means in rectifying section 117a to simultaneously cool the combined 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 C.sub.3 components and heavier hydrocarbon components can be removed from the combined vapor stream using the refrigeration available in expanded stream 36b than is possible using conventional mass transfer equipment and conventional heat transfer equipment.

(36) 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 replaces three separate equipment items in the prior art of U.S. Pat. No. 4,854,955 (heat exchanger 23, the upper absorbing section in the top of distillation column 24, and reflux drum 26 in FIG. 4 of U.S. Pat. No. 4,854,955) with a single equipment item (processing assembly 117 in FIG. 4 of the present invention). This reduces the plot space requirements and eliminates the interconnecting piping, reducing the capital cost of modifying a process plant to use the present invention. Second, elimination of the interconnecting piping means that a processing plant modified to use the present invention has far fewer flanged connections compared to the prior art of U.S. Pat. No. 4,854,955, reducing the number of potential leak sources in the plant. Hydrocarbons are volatile organic compounds (VOCs), some of which are classified as greenhouse gases and some of which may be precursors to atmospheric ozone formation, which means the present invention reduces the potential for atmospheric releases that may damage the environment.

(37) 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 FIG. 4, only two connections (commonly referred to as tie-ins) to the existing plant are needed: for flash expanded substantially condensed stream 36b (represented by the dashed line between stream 36b and stream 152a that is removed from service), and for column overhead vapor stream 39 (represented by the dashed line between stream 39 and stream 151 that is removed from service). The existing plant can continue to operate while the new processing assembly 117 is installed near fractionation tower 17, with just a short plant shutdown when installation is complete to make the new tie-ins to these two existing lines. The plant can then be restarted, with all of the existing equipment remaining in service and operating exactly as before, except that the product recovery is now higher with no increase in residue gas compression power.

Example 2

(38) The present invention also offers advantages when product economics favor rejecting only a portion of the C.sub.2 components to the residue gas product. The operating conditions of the FIG. 4 process can be altered as illustrated in FIG. 5 to increase the ethane content of the liquid product to the same level as that of the FIG. 3 prior art process. The feed gas composition and conditions considered in the process presented in FIG. 5 are the same as those in FIG. 3. Accordingly, the FIG. 5 process can be compared with that of the FIG. 3 process to further illustrate the advantages of the present invention.

(39) Most of the process conditions shown for the FIG. 5 process are much the same as the corresponding process conditions for the FIG. 3 process. The main differences are again the disposition of flash expanded substantially condensed stream 36b and column overhead vapor stream 39. In the FIG. 5 process, substantially condensed stream 36a is flash expanded through expansion valve 13 to slightly above the operating pressure (approximately 390 psia [2,691 kPa(a)]) 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 FIG. 5, the expanded stream 36b leaving expansion valve 13 reaches a temperature of ?141? F. [?96? C.] before it is directed into the heat and mass transfer means inside rectifying section 117a of processing assembly 117. As the combined vapor stream flows upward through one pass of the heat and mass transfer means and is cooled, a portion of it is condensed and falls downward while the remaining combined vapor stream continues flowing upward. The heat and mass transfer means provides continuous contact between the condensed liquid and the combined vapor stream so that it also functions to provide mass transfer between the vapor and liquid phases, thereby providing rectification of the combined vapor stream. The condensed liquid from the bottom of the heat and mass transfer means is directed to separator section 117b of processing assembly 117.

(40) 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 117a at ?136? F. [?93? C.]. The heated flash expanded stream discharges into separator section 117b 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 heat and mass transfer means in rectifying section 117a as previously described, and the liquid phase combines with the condensed liquid from the bottom of the heat and mass transfer means to form combined liquid stream 152. Combined liquid stream 152 leaves the bottom of processing assembly 117 and is pumped to higher pressure by pump 21 so that stream 152a at ?133? F. [?92? C.] can enter fractionation column 17 at the top feed point. The vapor remaining from the cooled combined vapor stream leaves the heat and mass transfer means inside rectifying section 117a of processing assembly 117 at ?128? F. [?89? C.] as cold residue gas stream 151, which is then heated and compressed as described previously for stream 39 in the FIG. 3 process.

(41) A summary of stream flow rates and energy consumption for the process illustrated in FIG. 5 is set forth in the following table:

(42) TABLE-US-00005 TABLE V (FIG. 5) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 12,398 546 233 229 13,726 32 12,317 529 212 124 13,497 33 81 17 21 105 229 34 3,632 156 63 37 3,980 36 3,713 173 84 142 4,209 37 8,685 373 149 87 9,517 39 10,689 425 12 1 11,435 152 2,004 298 95 143 2,627 151 12,398 300 1 0 13,017 42 0 246 232 229 709 Recoveries* Ethane 45.00% Propane 99.65% Butanes+ 100.00% Power Residue Gas Compression 5,565 HP [9,149 kW] *(Based on un-rounded flow rates)

(43) A comparison of Tables III and V shows that, compared to the prior art, the present invention improves propane recovery from 96.51% to 99.65% and butane+ recovery from 99.56% to 100.00%. The economic impact of these improved recoveries is significant. Using an average incremental value $0.74/gallon [145/m.sup.3] for hydrocarbon liquids compared to the corresponding hydrocarbon gases, the improved recoveries represent more than US$575,000 [430,000] of additional annual revenue for the plant operator. Comparison of Tables III and V further shows that these increased product yields were achieved using the same power as the prior art. In terms of the recovery efficiency (defined by the quantity of C.sub.3 components and heavier components recovered per unit of power), the present invention represents nearly a 2% improvement over the prior art of the FIG. 3 process.

(44) The FIG. 5 embodiment of the present invention provides the same advantages related to processing efficiency and the compact arrangement of processing assembly 117 as the FIG. 4 embodiment. The FIG. 5 embodiment of the present invention overcomes the equilibrium limitations associated with expanded stream 36b in the prior art FIG. 3 process to remove the heavier components from the column overhead vapor via indirect cooling and simultaneous mass transfer in rectifying section 117a of process assembly 117. The FIG. 5 embodiment of the present invention also replaces the three separate equipment items in the prior art of the U.S. Pat. No. 4,854,955 process with a single equipment item (processing assembly 117 in FIG. 5). This reduces the plot space requirements and eliminates the interconnecting piping, reducing the capital cost to modify an existing process plant to use this embodiment of the present invention while also reducing the potential for atmospheric releases of hydrocarbons that may damage the environment.

Example 3

(45) The present invention can also be operated to recover the maximum amount of C.sub.2 components in the liquid product. The operating conditions of the FIG. 4 process can be altered as illustrated in FIG. 6 to increase the ethane content of the liquid product to the same level as that of the FIG. 1 prior art process. The feed gas composition and conditions considered in the process presented in FIG. 6 are the same as those in FIG. 1. Accordingly, the FIG. 6 process can be compared with that of the FIG. 1.

(46) Most of the process conditions shown for the FIG. 6 process are much the same as the corresponding process conditions for the FIG. 1 process. The main differences are again the disposition of flash expanded substantially condensed stream 36b and column overhead vapor stream 39. In the FIG. 6 process, substantially condensed stream 36a is flash expanded through expansion valve 13 to slightly above the operating pressure (approximately 396 psia [2,731 kPa(a)]) 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 FIG. 6, the expanded stream 36b leaving expansion valve 13 reaches a temperature of ?140? F. [?96? C.] before it is directed into the heat and mass transfer means inside rectifying section 117a of processing assembly 117. As the combined vapor stream flows upward through one pass of the heat and mass transfer means and is cooled, a portion of it is condensed and falls downward while the remaining combined vapor stream continues flowing upward. The heat and mass transfer means provides continuous contact between the condensed liquid and the combined vapor stream so that it also functions to provide mass transfer between the vapor and liquid phases, thereby providing rectification of the combined vapor stream. The condensed liquid from the bottom of the heat and mass transfer means is directed to separator section 117b of processing assembly 117.

(47) 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 117a at ?141? F. [?96? C.]. (Note that the temperature of stream 36b drops slightly as it is heated, due to the pressure drop through the heat and mass transfer means and the resulting vaporization of some of the liquid methane contained in the stream.) The heated flash expanded stream discharges into separator section 117b 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 heat and mass transfer means in rectifying section 117a as previously described, and the liquid phase combines with the condensed liquid from the bottom of the heat and mass transfer means to form combined liquid stream 152. Combined liquid stream 152 leaves the bottom of processing assembly 117 and is pumped to higher pressure by pump 21 so that stream 152a at ?141? F. [?96? C.] can enter fractionation column 17 at the top feed point. The vapor remaining from the cooled combined vapor stream leaves the heat and mass transfer means inside rectifying section 117a of processing assembly 117 at ?139? F. [?95? C.] as cold residue gas stream 151, which is then heated and compressed as described previously for stream 39 in the FIG. 1 process.

(48) A summary of stream flow rates and energy consumption for the process illustrated in FIG. 6 is set forth in the following table:

(49) TABLE-US-00006 TABLE VI (FIG. 6) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31 12,398 546 233 229 13,726 32 12,200 503 183 82 13,278 33 198 43 50 147 448 34 3,784 156 57 25 4,118 36 3,982 199 107 172 4,566 37 8,416 347 126 57 9,160 39 12,265 55 5 1 12,508 152 3,854 198 107 172 4,432 151 12,393 56 5 1 12,642 42 5 490 228 228 1,084 Recoveries* Ethane 89.79% Propane 98.03% Butanes+ 99.71% Power Residue Gas Compression 5,569 HP [9,155 kW] *(Based on un-rounded flow rates)

(50) A comparison of Tables I and VI shows that the present invention achieves essentially the same recovery levels as the prior art when the process is operated to recover the maximum amount of C.sub.2 components. When operated in this manner, the temperature driving force for indirect cooling and simultaneous mass transfer in rectifying section 117a of process assembly 117 is very low because the temperature of column overhead stream 39 is almost the same as the temperature of flash expanded stream 36b, reducing the effectiveness of rectifying section 117a. Although there is no improvement in the component recoveries compared to the prior art when the present invention is operated in this manner, there is no decline either. This means there is no penalty when economics favor operating the plant to recover the maximum amount of C.sub.2 components in the liquid product, but the plant has all the advantages described previously for Examples 1 and 2 when economics favor operating the plant to reject some or all of the C.sub.2 components to the residue gas product.

Other Embodiments

(51) Some circumstances may favor also mounting the liquid pump inside the processing assembly to further reduce the number of equipment items and the plot space requirements. Such an embodiment is shown in FIG. 7, with pump 121 mounted inside processing assembly 117 as shown to send the combined liquid stream from separator section 117b to the top feed point of column 17 via conduit 152. The pump and its driver may both be mounted inside the processing assembly if a submerged pump or canned motor pump is used, or just the pump itself may be mounted inside the processing assembly (using a magnetically-coupled drive for the pump, for instance). For either option, the potential for atmospheric releases of hydrocarbons that may damage the environment is reduced still further.

(52) 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 combined liquid stream 152 to flow to the top feed point on fractionation column 17 by gravity head as shown in FIG. 8, eliminating the need for pump 21/121 shown in the FIGS. 4 through 7 embodiments.

(53) The present invention provides improved recovery of C.sub.3 components and heavier hydrocarbon components per amount of utility consumption required to operate the process. An improvement in utility consumption required for operating the 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.

(54) 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.