METHOD AND SYSTEM FOR TREATING A FLOW BACK FLUID EXITING A WELL SITE

Abstract

The present invention relates to a method and system for treating a flow back fluid exiting a well site following stimulation of a subterranean formation. More specifically, the invention relates to processing the flow back fluid, and separating into a carbon dioxide rich stream and a carbon dioxide depleted stream, and continuing the separation until the carbon dioxide concentration in the flow back stream until the carbon dioxide concentration in the flow back gas diminishes to a point selected in a range of about 50-80 mol % in carbon dioxide concentration, after which the lower concentration carbon dioxide flow back stream continues to be separated into a carbon dioxide rich stream which is routed to waste or flare, and a hydrocarbon rich stream is formed.

Claims

1-10. (canceled)

11. A system for processing the flow back fluid from a well site following stimulation of a subterranean formation, comprises: a pretreatment unit to receive and process the flow back fluid from the well site and remove any one of water, solid particulates, liquid hydrocarbons, hydrogen sulfides or a combination thereof; a membrane unit downstream of the pretreatment unit to receive the pretreated flow back fluid therefrom and separate the pretreated flow back fluid into a carbon dioxide rich permeate stream and a carbon dioxide depleted retentate stream; a permeate cooling unit to receive the carbon dioxide rich permeate stream and reduce the temperature of the stream to a temperature ranging from about 40 to 20 F.; and a phase separator to receive the lower temperature carbon dioxide rich permeate stream from the permeate cooling unit and separate the stream into a first liquid stream of predominantly carbon dioxide and a first gaseous stream enriched in methane.

12. The system of claim 11, further comprising a second phase separator to receive the first liquid stream of predominantly carbon dioxide and further separate into a second liquid carbon dioxide product stream and a second gaseous phase stream enriched in methane.

13. The system of claim 11, further comprising a pressure reducing valve disposed between the cooling unit and the phase separator to reduce the pressure to a range of about 60-500 psig, and lower the temperature of the carbon dioxide rich permeate stream.

14. The system of claim 11, further comprising a chiller unit in communication with the cooling unit providing additional refrigeration to the cooling unit.

15. The system of claim 12, further comprising a heat exchanger disposed between the phase separators to warm the first liquid stream of predominantly carbon dioxide routed to the second phase separator.

16. The system of claim 11, further comprising a manifold where the carbon dioxide depleted retentate stream is mixed with the first and second gaseous phase streams to produce a cool process stream which is routed to the cooling unit.

17. The system of claim 11, wherein the entire system or parts of the system are mobile.

18. The system of claim 11, wherein the phase separator is selected from one that either has two or more separation stages, a vessel or column with trays and heaters, a distillation column with a reboiler or a combination thereof.

19. The system of claim 11, wherein the carbon dioxide depleted retentate stream, the first and second gaseous phase streams individually or in any combination is routed to the cooling unit.

20. The system of claim 17, wherein the parts of the system disposed downstream of the membrane are removed from operation upon switching to a second mode where the flow back fluid is separated into a carbon dioxide rich stream and the carbon dioxide rich stream is routed to waste or flared, while the carbon dioxide lean stream, rich in hydrocarbons is recovered as a product.

21. The system of claim 11, wherein the membrane unit is a polyether ether ketone (PEEK) membrane.

Description

BRIEF DESCRIPTION OF THE FIGURES

[0020] The objects and advantages of the invention will be better understood from the following detailed description of the preferred embodiments thereof in connection with the accompanying figures wherein like numbers denote same features throughout and wherein:

[0021] FIG. 1 illustrates a plot of CO.sub.2 concentration vs. time for flow back gas from a CO.sub.2-fractured well;

[0022] FIG. 2 is a schematic illustration of a system and associated process for treating the flow back fluid exiting a well site which is operated in a CO.sub.2 recovery mode;

[0023] FIG. 3 is a schematic illustration of a system and associated process for treating the flow back fluid exiting a well site which is operated in a CO.sub.2 rejection mode;

[0024] FIG. 4 is a detailed illustration of the system shown in FIG. 2;

[0025] FIG. 5 illustrates another embodiment of FIG. 2 with an alternate separation system;

[0026] FIG. 6 illustrates a further system of FIG. 2 with yet another separation system;

[0027] FIG. 7 depicts another exemplary system embodiment of the present invention;

[0028] FIG. 8 depicts the equipment for the process of FIG. 4 distributed among several mobile units.

[0029] FIG. 9 illustrates the use of one of the mobile units of FIG. 8, when operating in a CO.sub.2 rejection mode.

DETAILED DESCRIPTION OF THE INVENTION

[0030] The present invention provides a system for the treatment of a flow back fluid exiting a well site immediately following stimulation of a subterranean formation until the concentration approaches the natural CO.sub.2 concentration in the reservoir, irrespective of the type of formation. As explained below, the process commences immediately following stimulation, but the system may be employed at the well site for several months given that it is designed to ultimately switch to a CO.sub.2 rejection mode, where the hydrocarbon product is recovered and sent to a natural gas pipeline or processing plant.

[0031] The system and process of the present invention, explained in detail below, operate in two modesCO.sub.2 recovery and CO.sub.2 rejection. During the first portion of the flow back, when CO.sub.2 concentration is relatively high, the process operates in CO.sub.2 recovery mode and produces a liquid product comprising mostly liquid CO.sub.2 with smaller amounts of hydrocarbons and nitrogen, suitable for use in subsequent CO.sub.2 fracturing operations or other uses. This mode also produces a hydrocarbon waste stream containing depleted amounts of CO.sub.2 that will typically be sent to a flare after it has been used in the downstream process for the production of liquid CO.sub.2 product.

[0032] The CO.sub.2 recovery process includes several unit operations including pretreatment, bulk gas separation, cooling, and phase separation/CO.sub.2 enrichment. In an exemplary embodiment a membrane is utilized for the separation following the pretreatment. The membrane preferentially permeates CO.sub.2 over methane, C.sub.2+ hydrocarbons and nitrogen and produces a permeate stream enriched in CO.sub.2. Cooling and partial condensation followed by phase separators achieve additional separation of methane and nitrogen from the CO.sub.2 enriched permeate to produce the CO.sub.2 rich product. During the second portion of the flow back, when CO.sub.2 concentration in flowback fluid drops below a certain level (i.e, a point selected in a range of 50-80 mol %), the process is reconfigured to perform CO.sub.2 rejection. This mode of operation is continued until CO.sub.2 concentration in the flowback fluid stabilizes to levels (e.g., 2-10 mol %) suitable for transport to a centralized gas processing plant or for direct addition to a natural gas pipeline. The products in CO.sub.2 rejection mode are gaseous and liquid hydrocarbon streams with the CO.sub.2 concentration controlled to a specific level, typically 2-10 mol % to meet the downstream requirements of the processing plant or a pipeline. The production of a liquid hydrocarbon stream is dependent upon the presence of C.sub.3+ hydrocarbons in the flowback fluid. During this mode of operation, a low pressure waste stream inclusive of a mixture of CO.sub.2 and hydrocarbons is produced and typically will be sent to a flare. CO.sub.2 is not readily recovered from the waste stream because of its low pressure and low CO.sub.2 concentration.

[0033] In the CO.sub.2 rejection mode, the same system as for CO.sub.2 recovery can be employed, but certain equipment is taken off line. For example, cooling and liquid CO.sub.2 phase separation unit operations are not needed for CO.sub.2 rejection. Thus, the entire system is modular and mobile in its entirety or only parts thereof, and can easily be moved from one well location to another.

[0034] There is a range of CO.sub.2 concentrations (50-80 mol %) in which the process could be operated in either the CO.sub.2 recovery mode or the CO.sub.2 rejection mode. Two operating modes (CO.sub.2 recovery and CO.sub.2 rejection) would always be carried out sequentially. Thus, switchover from CO.sub.2 recovery to CO.sub.2 rejection could occur at any point in the 50-75 mol % CO.sub.2 concentration range depending on the relative economic drivers for recovery of liquid CO.sub.2 product vs. recovery of hydrocarbon product streams.

[0035] With reference to FIG. 2, the system for treating the flow back fluid exiting the well site is illustrated. Flow back fluid 10 exiting the well site following stimulation of a particular formation enters a pretreatment unit 100, or alternatively bypasses the drier component (not shown) of the pretreatment unit 100 and then is routed to separator unit 200. Flow back fluid downstream of the well-head is typically at pressures ranging from about 1000-2000 psig and temperatures ranging from 80-130 F. The pretreatment unit 100 includes known processes for the removal of water (i.e., drier) and optionally H.sub.2S, solid particulates and/or liquid hydrocarbons from the flow back fluid 10. Stream 11 represents all contaminants removed by the pretreatment section. However, more than one such stream may exist depending on the configuration of the pretreatment unit 100. The pretreatment unit 100 may also include valves and instrumentation for controlling pressure and/or flow rate of flow back fluid to the downstream operational units. Liquid/gas phase separation may be required after any pressure reducing device such as a control valve. Flow back gas pressures above 500 psig are sufficient to achieve separation between CO.sub.2 and hydrocarbon. When flow back gas pressure is higher than 1000 psig, pressure difference between the feed and permeate side of the separation unit is set to meet the tolerance of the membrane utilized (e.g., typically up to 1000 psi).

[0036] In the exemplary embodiment of FIG. 2, pretreated flow back fluid 15 enters a membrane separation unit 200, where the pretreated flow back fluid 15 is separated to a carbon dioxide rich stream 20 and a carbon dioxide depleted stream 22. It will be recognized by those skilled in the art that other gas separation technologies, such as adsorption or absorption may be employed, although the membrane based system is preferred. In the event the flow back fluid has a CO.sub.2 concentration of 95 mol % or higher the flow back fluid can be routed to the cooling unit, discussed below, bypassing separation unit 200. Permeate 20 is higher in CO.sub.2 concentration (i.e., 100-83%) than the pretreated flow back fluid 15 (i.e., 100-50%) and the retentate 22 is lower (i.e., 35-40%) in CO.sub.2 concentration than the pretreated flow back fluid 15. Liquid hydrocarbons can form on the retentate side of the membrane due to Joule-Thomson cooling caused by the reduction in pressure of gas as it passes through the membrane wall. The tendency to form liquid hydrocarbons is determined by stream conditions and composition of the pretreated flowback fluid 15, the pressures of the permeate 20 and retentate 22, and the relative flow rates of permeate 20 and retentate 22. Suitable membrane separators for the situation when no liquid hydrocarbons are in contact with the membrane are commercially available from the likes of Natco Group, Inc., UOP, L.L.C., and Kvaerrner Process Systems US, Inc. For situations where liquid hydrocarbons may contact the membrane, the fluid separator for this application can be a separation unit having polyether ether ketone (PEEK) membranes. Suitable membrane separators to handle large fraction of C.sub.2+ components are commercially available from Porogen Corporation.

[0037] The pressure of carbon dioxide depleted stream (i.e., retentate) 22 will typically be 0.5-5 psi less than the feed pressure of pretreated flow back fluid 15, and the pressure of the carbon dioxide rich stream will typically be in the range of about 300-600 psig. In the case of a membrane, both streams 20 and 22 will typically exit the membrane at a lower temperature than the feed temperature, due to Joule-Thomson cooling associated with the pressure drop of the permeate across the membrane.

[0038] Permeate 20 (or carbon dioxide rich stream) is routed into a permeate cooling unit 300, where the permeate is cooled by indirect heat exchange with stream 42 from chiller unit 400 and a blend of cool process streams 80. Permeate 30 exiting cooling unit 300 is typically cooled to a temperature of 40 to 20 F. The blend of process streams 90 exits permeate cooling unit 300 and is typically sent to a flare. Naturally, the heat exchangers and Joule-Thomson valves employed, are known in the art, and are not described in any level of detail herein. Chiller 400 cools permeate 20 by heat exchange with either a refrigerant or a secondary heat transfer fluid 42. Refrigerant or secondary heat transfer fluid 44 is returned to chiller 400, where it is cooled by known processes and then it is recirculated as stream 42. The typical configuration for the chiller is a Carnot-cycle type (or derivative) mechanical refrigeration unit using a recirculating refrigerant. Such devices use a refrigerant compressor which may be driven by an electrical motor or, preferably, an engine typically fueled by natural gas, propane, gasoline or diesel fuel. If desired, the engine used to drive the refrigerant compressor may be a vehicle engine with a power take off mechanism. Alternative refrigeration processes may be used including heat driven absorption processes. Process stream 90 from the permeate cooling unit may be combusted to provide at least a portion of the heat needed for the heat driven absorption process.

[0039] Cooled permeate 30 is routed through a pressure reducing valve 901, where the pressure is reduced to the range of 60-500 psig, which further cools permeate (carbon dioxide rich stream) 32 to a temperature of about 70 to 20 F. The reduced pressure permeate 32 enters a first phase separator 500, where it is separated into a gaseous stream 52 enriched in the more volatile compounds contained in stream 32, such as methane and nitrogen, and a first liquid CO.sub.2 stream 50, which consists of predominantly CO.sub.2 and smaller amounts of methane, C.sub.2+ hydrocarbons and nitrogen. The first phase separator is typically operating at a pressure ranging from 60-500 psig, and preferably 265-340 psig.

[0040] Under some circumstances, first liquid CO.sub.2 stream 50 may be colder than the minimum allowable working temperature (MAWT) of liquid CO.sub.2 receiver tanks or transport tanks, which is typically 20 F. If this is the case, the first liquid CO.sub.2 stream 50 is warmed in a heat exchanger 600 to an acceptable temperature, typically warmer than about 20 F. The warmed liquid CO.sub.2 60 is sent to a second phase separator 700 where it is separated into a second gas stream 72 and a second liquid CO.sub.2 stream 70. Second liquid CO.sub.2 stream 70 is the desired product from the process and is sent to storage and/or transport. The process of warming liquid CO.sub.2 and sending it to a second phase separator causes the liquid CO.sub.2 from the second phase separator to have a lower concentration of methane than the liquid CO.sub.2 from the first phase separator. This results in lower methane concentration to occur in the headspace of the LCO.sub.2 storage tanks and reduces the tendency of the headspace gas to form gas mixtures that would be flammable when mixed with air. Optionally, a liquid CO.sub.2 pump can be used on stream 50 or stream 70.

[0041] Retentate (carbon dioxide depleted stream) 22 and phase separator gas streams 52 and 72 are routed through pressure reducing valves 902, 903 and 904, respectively. The reduced pressure streams 24, 54 and 74 are blended into stream 80 and employed for cooling by indirect heat exchange in the permeate cooling unit 300, as discussed above. Although, a particular system and process configuration is shown in this FIG. 2, where the retentate and phase separator gases are blended prior to being routed to the permeate cooling unit 300, others are contemplated. For example, in another embodiment, these streams are not blended or only part of the retentate and/or phase separator gases are utilized in process cooling unit 300. Likewise, process configurations in which all the cooling is provided by chiller 400 may be employed or ones in which the membrane feed and/or retentate are cooled by chiller 400 and/or by blended stream 80 may be used.

[0042] An alternative process configuration is to employ purge gas 90 to cool the membrane feed 15 and/or flow back fluid 10 or to employ the mechanical chiller and/or cool process streams to cool the membrane feed. Cooling the membrane feed has the benefit of lowering the temperature of the membrane material and increasing the selectivity of the membrane for CO.sub.2. Another potential benefit of cooling prior to the membrane is the potential to separate hydrocarbon liquids by phase separation.

[0043] Once the concentration of carbon dioxide in the flow back fluid 10 diminishes to a range of about 50-80 mol %, the lower concentration flow back fluid 10 continues to be separated, but the system switches to a CO.sub.2 rejection mode. With reference to FIG. 3, the process for rejecting the CO.sub.2 from flow back fluid recovered from wells fractured with high pressure CO.sub.2 is shown. Flow back fluid 10, enters a pretreatment unit 100. Some or all pretreatment steps may be bypassed when operating in CO.sub.2 rejection mode. Only those pretreatment steps are employed which are needed for protecting the membrane or for producing a hydrocarbon rich product stream (i.e., pipeline ready natural gas). For example, water removal may not need to be done in the pretreatment process unit 100, as sufficient drying of the hydrocarbon rich product stream 22 will likely be performed by the membrane unit 20. Contaminants and/or other components that are removed in the pretreatment unit 100, exit as stream 11. Depending on how the pretreatment unit 100 is configured, there may be more than one such stream.

[0044] Pretreated flow back fluid 15 enters the membrane unit 200, which produces a permeate 20 and a retentate 22, which is sent to a phase separator 800, if needed. If liquid hydrocarbons exist in the retentate 22, a liquid hydrocarbon stream 16 is recovered separately from the gaseous stream 28. The liquid hydrocarbon stream is either mixed with the oil produced from the well or further processed and sold separately as natural gas liquids. CO.sub.2 concentration of the retentate gas 28 is reduced to a specified concentration and the retentate gas 28 is sent to downstream processing plant or pipeline as product. The concentration of CO.sub.2 in the retentate is generally in the range of 2-10 mol %. Permeate 20 contains mostly CO.sub.2 and some hydrocarbons and is typically sent to flare as a waste gas 94. The permeate stream is typically set at a low pressure in the range of 5-50 psig. Membrane feed pressure is typically controlled to a pressure such that the pressure difference between feed and permeate does not cause the membrane material to rupture, but which is high enough to send the retentate 28 as a product without the need for a retentate compressor. Meanwhile, permeate cooling unit 300, chiller 400, phase separators 500 and 700 and liquid CO.sub.2 heater 600 are not needed when operating in CO.sub.2 rejection mode. As discussed above, these elements of the system are modular, and mobile. Thus, they can be removed and employed at the next well site where the subterranean formation is or about to be stimulated.

[0045] FIG. 4, depicts a particular arrangement of the system embodied in FIG. 2. This arrangement, provides the system particulars within the permeate cooling process 300. This system, as discussed above, is capable of operating in either CO.sub.2 recovery mode or CO.sub.2 rejection mode.

[0046] Permeate 20 from the membrane is optionally split into a first permeate stream 21a which enters a first permeate cooler 310 and a second permeate stream 21b which enters the LCO.sub.2 heater 600. Cooled permeate 23 from first permeate cooler 310 is blended with cooled permeate 27 from the LCO.sub.2 heater 600 to form a blended stream of cooled permeate 25 which enters second permeate cooler 320 and is further cooled by heat exchange with refrigerant 42 from the chiller 400. Refrigerant 44 is returned to the chiller 400 from second permeate cooler 320. Further cooled permeate 26 is additionally cooled in a third permeate cooler 330 by the blend of process streams 54, 74, and 24 to form stream 80. Additionally cooled permeate 30 exits the permeate cooling process 300 and is sent to J-T valve 901. Stream 82 exits third permeate cooler 330, passes through J-T valve 905, which reduces the pressure and temperature of the blend of process gases 84. Low pressure process stream 84 provides cooling to first permeate cooler 310. The waste gas 90 from permeate cooler 310 is sent to a flare. Stream temperatures within the permeate cooling process 300 will vary as the composition and stream conditions of the flowback gas 10 changes over time. At permeate pressures of about 400 psig, the following temperature ranges will occur: Permeate streams 20, 21a and 21b will generally be in the range of 10 to 100 F. Cooled permeate 23 from the first permeate cooler 310 will generally be 0-25 F. The temperature of permeate 26 from the second permeate cooler will generally be 40 to +5 F. The temperature of permeate 30 from the third permeate cooler 330 will generally be 1-10 F. colder than the temperature of permeate 26 from the second permeate cooler.

[0047] Other equipment items, including the pretreatment unit 100, membrane unit 200, chiller 400, first phase separator 500, LCO.sub.2 heater 600, second phase separator 700, and J-T valves 901, 902, 903 and 904 are similar to the ones discussed with respect to FIG. 2, above, and the retentate phase separator 800 is similar the one depicted in FIG. 3.

[0048] The arrangement of permeate cooling heat exchangers of FIG. 4 provides several additional benefits. Process streams 22, 52 and 72 are pressurized fluids normally containing a high concentration of CO.sub.2. The ultimate destination for these streams is a flare with an outlet at atmospheric pressure. If these streams are depressurized and sent directly to flare, without being heated, several unfavorable conditions will occur. Temperature of the vent stream can become lower than minimum allowable working temperature of carbon steel and result in the need for more expensive materials of construction, such as 300 series stainless steel, to be used for the waste gas piping and flare stack. Solid CO.sub.2 can form resulting in potential blockages. Condensation of hydrocarbon components can also occur, resulting in potential pooling of liquid hydrocarbons within the flare stack. To counter these unfavorable conditions, first permeate cooler 310 warms low pressure waste gas to temperatures greater than the minimum allowable working temperature of carbon steel. Permeate cooler 310 also causes any condensed hydrocarbons in stream 84 to vaporize. Formation of solid CO.sub.2 may generally be prevented by operating all streams containing significant amounts of CO.sub.2 at temperatures warmer than the triple point of pure CO.sub.2 (69.7 F.). This is accomplished in FIG. 4 by depressurizing and warming stream 80 in stages. Formation of solid CO.sub.2 in low pressure waste streams 84 and 90 is prevented by careful selection of the pressure of stream 80. Stream 80 pressure is generally in the range of 15-165 psig and is set such that stream 80 is warmer than the CO.sub.2 triple point temperature (69.7 F.), yet cold enough that significant heat transfer occurs in third permeate cooler 330 and prevents solid CO.sub.2 from forming in stream 84. In the exemplary systems shown in FIG. 2 and FIG. 4, receiver tanks or transport tanks may be utilized as phase separator 700. In such configurations, stream 60 would be sent from the process to a receiver tank or transport tank. Stream 72 would return vapor from the receiver tank or transport tank to the process. Stream 70 would accumulate in the receiver tank or transport tank. The advantage of such a configuration is the elimination of the need for a second phase separator and associated liquid level control valve in the CO.sub.2 recovery process.

[0049] With reference to FIG. 4, the preferred heat source for the LCO.sub.2 heater 600 is stream 21b, which is a portion of stream 20. Other sources of heat may be used, including ambient heat, refrigerant, engine coolant or oil, or compressor oil. Alternate process streams that can be used to heat stream 50 may include streams 10, 15, 20, 22, 23, 26 and 30. In the exemplary system of FIG. 4, stream 27 from the LCO.sub.2 heater 600 returns to the process by mixing with stream 23, at the outlet of the first permeate cooler 310. Alternative return locations for stream 27 are stream 26 (outlet of second permeate cooler), stream 30 (outlet of third permeate cooler) or stream 32 (outlet of valve 901). The purpose of the LCO.sub.2 heater and second phase separator is to provide a liquid CO.sub.2 product that is not colder than the minimum allowable working temperature of the receiver. Normally, the receiver vessel will have a MAWT of 20 F., and so the temperature of liquid CO.sub.2 product from the process needs to warmer than 20 F. at the receiver pressure. One alternate configuration is to not include the use of the LCO.sub.2 heater and second phase separator. Instead the process may be stopped once the liquid product temperature becomes colder than 20 F. This configuration takes advantage of the fact that recovered LCO.sub.2 from the first phase separator may be warmer than 20 F. at the start of the flowback and generally becomes colder as the flowback proceeds. However, the alternative configuration will usually result in less CO.sub.2 recovery than the process which includes two phase separators and a liquid CO.sub.2 heater.

[0050] In another alternative embodiment, and as illustrated in FIG. 5, a distillation column 500 in conjunction with a reboiler 650 is employed to adjust the temperature of the liquid CO.sub.2 product. Such a distillation column may be configured a number of ways including the use of trays or packing, internal or external reboiler, and internal or external or no overhead condenser. The distillation column can be employed as a replacement for the two pressure vessels (phase separators), and the numerous distillation stages of separation yields a higher concentration CO.sub.2 stream 50 (and 70). In a configuration of the present invention, and as shown in FIG. 6, the distillation column 500 is a phase separator with trays and an internal heater.

[0051] In another exemplary embodiment of the invention, and with reference to FIG. 7, an expansion device, such as a gas turbine 110 can be disposed downstream of the pretreatment unit 100, but upstream of the membrane unit 200. Refrigeration produced when employing such an expansion device can be increased by integrating a compressor with the expander. This system has the benefits of not requiring a utility to drive a refrigerant compressor and requiring less space for process equipment. Pretreated flow back fluid 9 is routed to compressor/expander 110, where it is compressed to a pressure of 1300-2000 psia. The compressed flow back fluid 12 is cooled in heat exchanger 120, illustrated here as an air cooled heat exchanger. Flow back fluid 13, now cooled and elevated in pressure is sent to a pressure reducing valve 900. The resulting flowback fluid 15 is sent to the membrane unit 200, which produces permeate 20 and retentate 22. Permeate 20 is higher in CO.sub.2 concentration than retentate 22. Permeate 20 is cooled in heat exchanger 300, illustrated here as a multi-stream heat exchanger. The function of heat exchanger 300 could also be performed by multiple dual stream heat exchangers as shown in previous embodiments. Cooled permeate 30 enters a pressure reducing valve 901. The reduced pressure permeate 32 comprises both liquid and vapor, and is sent to phase separator 500, where it is separated into a gaseous stream 52, enriched in methane and a liquid stream 50 which consists of predominantly liquid CO.sub.2 and smaller amounts of methane, C.sub.2+ hydrocarbon and nitrogen. Liquid stream 50 may be sent as product to storage or may be further processed by heating and additional phase separation as shown in previous embodiments.

[0052] Retentate 22 and phase separated gas streams 52 and 72 enter J-T valves 902, 903 and 904 respectively. The reduced pressure streams 24, 54 and 74 are blended into stream 80 and employed for cooling by indirect heat exchange in heat exchanger 300. The resulting blended gas stream 81 is sent to compressor/expander 110 where it is reduced in pressure and provides the driving energy for the compressor. Blended gas stream 82 from the expander provides cooling in heat exchanger 300, passes through pressure reducing valve 905 and is sent again to heat exchanger 300 to provide cooling. The resulting waste gas stream 90 is sent to flare or vented.

[0053] As illustrated in FIG. 8 an exemplary distribution system for the process equipment shown in the embodiment of FIG. 2 is shown among multiple mobile units. These multiple mobile units are useful when the footprint and/or weight of the process equipment is greater than may practically be carried on a single mobile unit. In this example, the dryer 100 A is mounted on one mobile unit. The remainder of the pretreatment equipment 100 B, the membrane unit 200, and the retentate phase separator 800 are mounted another mobile unit. The permeate cooling process equipment 300, chiller 400, first and second CO.sub.2 phase separators 500 and 700, and liquid CO.sub.2 heater 600 are mounted on a third mobile unit.

[0054] FIG. 9 depicts the equipment distribution of FIG. 8, but with the elimination of mobile units that are needed when operating in CO.sub.2 recovery mode, but not needed when operating in CO.sub.2 rejection mode. In this example, Mobile Units A and C are not shown, as all the equipment needed for operating in CO.sub.2 rejection mode is mounted on Mobile Unit B. This distribution of process equipment enables Mobile Units A and C to be disconnected and used at a different location, with a different Mobile Unit B, once CO.sub.2 recovery mode has ended.

[0055] The invention is further explained through the following Example, which is based on various embodiments of the system, but it is in no way to be construed as limiting the present invention.

Example

[0056] Performance of the process was evaluated through simulation of the system shown in the embodiment of FIG. 4. Assumed process conditions of the flowback gas are shown in Table 1, and are assumed to be constant while the flowback process is operating.

TABLE-US-00001 TABLE 1 Flowback Gas Process Conditions Flow Rate, MMSCFD 5 Temperature, F. 120 Pressure, psia 1215

[0057] Operating conditions for the CO.sub.2 recovery process are shown in Table 2.

TABLE-US-00002 TABLE 2 CO.sub.2 Recovery Operating Conditions Permeate Pressure, psia 415 Chiller Outlet Temperature, F. 0 LCO2 Product Pressure, psia 265 LCO2 Product Temperature, F. 20

[0058] Performance of a CO.sub.2 recovery process, is summarized in Table 3, below. The purpose of the process is to recover and liquefy CO.sub.2 from the flowback gas. The first column with the heading Elapsed Time, indicates the time from the start of the flowback gas flow, in approximately equal time periods. The 2.sup.nd-8.sup.th columns indicate the assumed dry basis composition of the flowback gas. Also shown in the table are the amount of CO.sub.2 contained in the flowback gas, the amount of CO.sub.2 recovered as liquid, and the purity of the recovered liquid CO.sub.2. Over the periods shown in Table 3, in the aggregate, CO.sub.2 concentration in the flowback decreases from 95.60% to 54.92%. The mass rate of CO.sub.2 contained in the flowback gas is proportional to the CO.sub.2 concentration and declines with time. Initially, the effectiveness of the CO.sub.2 recovery process (i.e. the fraction of CO.sub.2 in the flowback gas that is recovered as liquid) is about 93%. However, the recovery effectiveness declines with CO.sub.2 concentration, so that by Period 10, when CO.sub.2 concentration in the flow back gas is 54.9%, only about 23% of the contained CO.sub.2 is recovered.

TABLE-US-00003 TABLE 3 CO.sub.2 Recovery Process Performance CO2 CO2 LCO2 Elapsed AVERAGE FEED COMPOSITION (Dry Basis) Contained in Recovered as Product Time CO2 N2 C1 C2 C3 nC4 nC5 Flowback Gas Liquid Product Purity Period mol % mol % mol % mol % mol % mol % mol % tpd tpd mol % 1 95.6% 0.1% 2.9% 0.4% 0.5% 0.3% 0.3% 277 258 96.9% 2 91.1% 0.1% 5.8% 0.7% 1.0% 0.7% 0.6% 264 228 97.8% 3 86.6% 0.2% 8.7% 1.1% 1.5% 1.0% 0.9% 251 200 97.6% 4 82.0% 0.3% 11.7% 1.4% 2.0% 1.3% 1.3% 238 173 97.5% 5 77.5% 0.3% 14.6% 1.8% 2.5% 1.7% 1.6% 225 147 97.2% 6 73.0% 0.4% 17.6% 2.2% 3.0% 2.0% 1.9% 212 122 97.0% 7 68.5% 0.5% 20.5% 2.5% 3.5% 2.4% 2.2% 199 99 96.7% 8 64.0% 0.5% 23.4% 2.9% 4.0% 2.7% 2.5% 185 77 96.3% 9 59.4% 0.6% 26.4% 3.2% 4.5% 3.0% 2.8% 172 56 95.9% 10 54.9% 0.7% 29.3% 3.6% 5.0% 3.4% 3.2% 159 37 95.5%

[0059] Table 3 also indicates the changes in liquid CO.sub.2 product purity that occur as the composition of the flow back gas changes. CO.sub.2 is separated and purified in several steps. The pretreatment unit 100, removes contaminants such as water, solid particulates, liquid hydrocarbons, or hydrogen sulfide. The membrane unit 200 removes some of the methane, as well as most of the N.sub.2 and heavier hydrocarbons. The flash tanks 500 and 700 remove additional methane. Most of the C.sub.2+ hydrocarbons contained in the permeate will accumulate in the liquid CO.sub.2 product. Thus, as C.sub.2+ hydrocarbon concentration in the flowback gas increases over time, the purity of liquid CO.sub.2 product decreases due to the corresponding increasing concentration of C.sub.2+ in the permeate. The exception to the trend of decreasing CO.sub.2 purity is when feed is over 95% CO.sub.2. When this occurs, the CO.sub.2 concentration is high enough that the membrane unit is not needed. Therefore, it is bypassed. The CO.sub.2 purity for Period 1 is lower than for Period 2 because on Period 1, hydrocarbons are not removed by the membrane.

[0060] Although Table 3 indicates the performance of the CO.sub.2 recovery process for flow back CO.sub.2 concentrations down to 54%, there is considerable flexibility of the process regarding when the CO.sub.2 recovery process may be ended and when the CO.sub.2 rejection process started. If it is desired to produce more liquid CO.sub.2 at the expense of producing natural gas and natural gas liquids, the CO.sub.2 recovery process may be extended. If it is desired to produce more natural gas and hydrocarbon condensates and less liquid CO.sub.2 product, the CO.sub.2 recovery process may be ended at an earlier point. Generally, the change in operating modes will take when the CO.sub.2 concentration of the flowback is in the range of 50-80 mol %.

[0061] Operating parameters of the CO.sub.2 rejection process, which produces natural gas and hydrocarbon condensates from flowback gas, is simulated based on the embodiment shown in FIG. 3, and are shown in Table 4, below. The pressure of flow back gas to the membrane is controlled at 915 psia and the permeate pressure is set at 30 psia. Natural gas liquids are produced both at the pressure reduction step in the pretreatment unit 100 and in the retentate stream 22. Phase separators are used at both of these locations to separate the hydrocarbon condensates.

TABLE-US-00004 TABLE 4 CO.sub.2 Rejection Operating Conditions Membrane Feed Pressure, psia 915 Permeate Pressure, psia 30 Retentate CO.sub.2 Concentration, mol % 5%

[0062] Performance of a CO.sub.2 rejection process, is shown in Table 5, below. This CO.sub.2 rejection process uses the same pretreatment unit and membrane unit as the CO.sub.2 recapture unit. The permeate cooling process 300, chiller 400, phase separators 500 and 700, and LCO.sub.2 heater 600 are not used by the CO.sub.2 rejection process, and may be removed and used at another location once the CO.sub.2 recapture process has ended and the CO.sub.2 rejection process has started.

[0063] The first columns of Table 5 are similar to Table 3, indicating the same elapsed time from the start of flowback and the same composition of the flowback gas. The time period in Table 5 is from Period 5 to Period 28, while the time period in Table 3 is from Period 1 to Period 10. The overlap in time periods is shown to illustrate that the process can be used to reject CO.sub.2 for producing a product natural gas stream when CO.sub.2 recovery is still an option. At 77.5% CO.sub.2 in the feed, the product rates from the CO.sub.2 rejection process are 0.23 MMSCFD of natural gas and 8200 gpd of hydrocarbon condensate. At 8.2% CO.sub.2 in the feed, the product rates have increased to 3.52 MMSCFD of natural gas and 13400 gpd of hydrocarbon condensate.

TABLE-US-00005 TABLE 5 CO.sub.2 Rejection Process Performance Elapsed AVERAGE FEED COMPOSITION (Dry Basis) Recovered Recovered HC Time CO2 N2 C1 C2 C3 nC4 nC5 Natural Gas Condensates Period mol % mol % mol % mol % mol % mol % mol % MMSCFD gpd 5 77.5% 0.3% 14.6% 1.8% 2.5% 1.7% 1.6% 0.23 8200 6 73.0% 0.4% 17.6% 2.2% 3.0% 2.0% 1.9% 0.34 9200 7 68.5% 0.5% 20.5% 2.5% 3.5% 2.4% 2.2% 0.47 10100 8 64.0% 0.5% 23.4% 2.9% 4.0% 2.7% 2.5% 0.62 10900 9 59.4% 0.6% 26.4% 3.2% 4.5% 3.0% 2.8% 0.77 11500 10 54.9% 0.7% 29.3% 3.6% 5.0% 3.4% 3.2% 0.94 12100 11 50.4% 0.7% 32.2% 4.0% 5.5% 3.7% 3.5% 1.13 12500 12 45.9% 0.8% 35.2% 4.3% 6.0% 4.1% 3.8% 1.33 12800 13 41.4% 0.9% 38.1% 4.7% 6.5% 4.4% 4.1% 1.54 13100 14 37.0% 0.9% 40.9% 5.0% 6.9% 4.7% 4.4% 1.76 13300 15 32.8% 1.0% 43.7% 5.4% 7.4% 5.0% 4.7% 1.98 13400 16 28.8% 1.1% 46.3% 5.7% 7.8% 5.3% 5.0% 2.20 13500 17 25.4% 1.1% 48.5% 6.0% 8.2% 5.6% 5.2% 2.40 13600 18 22.4% 1.2% 50.4% 6.2% 8.5% 5.8% 5.4% 2.58 13700 19 19.9% 1.2% 52.1% 6.4% 8.8% 6.0% 5.6% 2.74 13700 20 17.7% 1.2% 53.5% 6.6% 9.1% 6.2% 5.8% 2.88 13800 21 15.8% 1.3% 54.7% 6.7% 9.3% 6.3% 5.9% 3.00 13800 22 14.2% 1.3% 55.8% 6.9% 9.4% 6.4% 6.0% 3.11 13800 23 12.8% 1.3% 56.7% 7.0% 9.6% 6.5% 6.1% 3.20 13900 24 11.6% 1.3% 57.5% 7.1% 9.7% 6.6% 6.2% 3.28 13900 25 10.6% 1.3% 58.1% 7.2% 9.8% 6.7% 6.3% 3.35 13900 26 9.7% 1.4% 58.7% 7.2% 9.9% 6.8% 6.3% 3.42 13900 27 8.9% 1.4% 59.2% 7.3% 10.0% 6.8% 6.4% 3.47 13600 28 8.3% 1.4% 59.6% 7.3% 10.1% 6.9% 6.4% 3.52 13400

[0064] While the invention has been described in detail with reference to specific embodiments thereof, it will become apparent to one skilled in the art that various changes and modifications can be made, and equivalents employed, without departing from the scope of the appended claims.