DE10226596A1 - Process for liquefying a hydrocarbon-rich stream with simultaneous recovery of a C3 + -rich fraction with high yield - Google Patents

Abstract

It becomes a process for liquefying a hydrocarbon-rich stream, in particular a natural gas stream, with simultaneous recovery of a C¶3 + ¶-rich fraction with high yield, the liquefaction of the hydrocarbon-rich stream taking place in heat exchange with a refrigerant mixture cycle cascade. DOLLAR A According to the invention, DOLLAR A a) the hydrocarbon-rich stream (1, 1 ') is expanded before its separation into a C¶2-¶-rich and into a C¶3 + ¶-rich fraction (X), DOLLAR A b ) the relaxed hydrocarbon-rich stream (2) is fed to a C¶3 ¶ absorption process (T1) and in this into the C¶2-¶ rich fraction (3) to be supplied to the liquefaction (E2, E3) and into a first condensate fraction (8) separated, DOLLAR A c) the C¶2-¶-rich fraction (3) to be supplied to the liquefaction condenses (V), DOLLAR A d) before the liquefaction (E2, E3), the first condensate fraction (8) warmed (E ) fed to a C¶2¶ stripping process (T2), DOLLAR A e) from the bottom of the C¶2¶ stripping process (T2) a second, high C¶3 + ¶ condensate fraction (11) Yield obtained, and DOLLAR A f) at the head of the C¶2¶ stripping process (T2) withdrawn a C¶3-¶-rich gas fraction, partially condensed (E) and the C¶3¶ absorption process (T1) as Detergent added (10th ).

Description

The invention relates to a process for liquefying a hydrocarbon-rich stream, in particular a natural gas stream, with simultaneous recovery of a C ₃₊ -rich fraction with high yield, the liquefaction of the hydrocarbon-rich stream in heat exchange with the refrigerants being one of at least three different refrigerant compositions having existing refrigerant mixture circuits, and wherein the hydrocarbon-rich stream to be liquefied is separated into a C _2– -rich fraction which is subjected to the liquefaction and into a C ₃₊ -rich fraction.

Gas liquefaction plants are either as so-called LNG Baseload Plants Liquefaction plants of natural gas to supply natural gas as primary energy - or as so-called peak shaving Plants - well Liquefaction plants of natural gas to meet peak demand - designed.

LNG Baseload Plants are usually included Refrigeration cycles operated, which consist of hydrocarbon mixtures. These are mixed cycles energetically more efficient than expander circuits and allow for the large liquefaction capacities the Baseload Plants correspondingly relatively low energy consumption.

A generic method for liquefying a hydrocarbon-rich stream is known for example from German Offenlegungsschrift 197 16 415. Here, the hydrocarbon-rich stream is liquefied against a refrigerant mixture circuit cascade, which consists of three refrigerant mixture circuits having different refrigerant compositions. The first of the three refrigerant mixture circuits is used for pre-cooling the hydrocarbon-rich stream to be liquefied, the second refrigerant mixture circuit is for actual liquefaction, and the third refrigerant mixture circuit is for sub-cooling the liquefied hydrocarbon-rich stream. Based on 1 Such a method for liquefying a hydrocarbon-rich stream will be explained in more detail.

According to the in the 1 The method shown is an optionally pretreated natural gas stream, which has a temperature between 10 and 50 ° C. and a pressure between 30 and 80 bar, is fed via line 10 to a first heat exchanger E1.

The pretreatment steps that may be necessary, such as drying, CO ₂ removal, sulfur removal, etc., are not discussed in more detail below; the usual procedures are known to the person skilled in the art.

The natural gas flow is in the heat exchanger E1 against the refrigerant mixture expanded in a P13 expansion valve of the first or PRC (Precooling Refrigerant Cycle) refrigerant mixture circuit precooled in line P14 to a temperature between -35 and -55 ° C.

The third refrigerant mixture or SRC (Subcooling Refrigerant Cycle) refrigerant mixture circuit becomes the heat exchanger E1 via line S5 with a temperature between 10 and 50 ° C and a pressure between 30 and 60 bar supplied and in the heat exchanger E1 against the already mentioned Refrigerant mixture cooled in line P14 and partially condensed with the refrigerant mixture in line P14 evaporates at a pressure between 1.5 and 6 bar. The refrigerant mixture of the SRC refrigerant mixture circuit leaves the heat exchanger E1 over Line S6 with a temperature between -35 and -55 ° C.

The second refrigerant mixture or LRC (Liquefaction Refrigerant Cycle) refrigerant mixture circuit becomes the heat exchanger E1 via line L5 with a temperature between 10 and 50 ° C and a pressure between 15 and 40 bar supplied and in the heat exchanger E1 against the refrigerant mixture of the PRC mixed refrigerant circuit condensed in line P14. The refrigerant mixture of the LRC-refrigerant mixture circuit becomes from the heat exchanger Subtract E1 at a temperature between -35 and -55 ° C.

That evaporated and overheated Refrigerant mixture of the PRC mixed refrigerant circuit contains in line P14 preferably 0 to 70 mol% of ethylene or ethane, 30 to 70 mol% Propane and 0 to 30 mol% butane. This refrigerant mixture will separate P1 supplied with a pressure of 1.5 to 6 bar. That on the head of the separator P1 via line P2 withdrawn gaseous Refrigerant mixture is in the compressor P3 to a pressure between 6 and 15 bar compacted. Subsequently takes place, preferably against sea water, against air or against a appropriate cooling medium, a cool down of the compressed refrigerant mixture in cooler P4 to a temperature between 10 and 50 ° C. Partial condensation can occur of the refrigerant mixture occur.

The refrigerant mixture is then fed to a further separator P6 via line P5. The gaseous fraction of the refrigerant mixture obtained at the top of the separator P6 is fed to the second compressor stage P8 and compressed therein to a pressure between 10 and 30 bar. If present, the liquid fraction from the separation P6 is pumped to a pressure between 10 and 30 bar by means of the pump P7 - this is preferably a centrifugal pump - and then with that in the compressor P8 compressed refrigerant mixture flow merged.

The compression of the refrigerant mixture first or PRC mixed refrigerant circuit preferably in a two-stage, single or multi-housing centrifugal compression device, which both the cooler P4 as well as the deposit P6 contains. In the case of very large quantities instead of the centrifugal compression device also an axial compression device be provided.

The compressed refrigerant mixture of the PRC-refrigerant mixture circuit is in the cooler P9, preferably against sea water or an appropriate cooling medium condensed at a temperature in the range of 10 to 50 ° C and possibly slight supercooled. Subsequently is the refrigerant mixture over the P10 line to the heat exchanger E1 fed and subcooled to a temperature between -35 and -55 ° C against itself.

The evaporation temperature after the Joule-Thomson relaxation in the relief valve P13 - or alternatively in addition in a relaxation turbine - can essentially be achieved depends on the degree of hypothermia before expansion and from the evaporation pressure in the temperature range between -38 and from -58 ° C.

The second or LRC mixed refrigerant circuit serves, as already mentioned, partial or complete liquefaction of the pre-cooled natural gas flow in line 20. The refrigerant mixture this LRC refrigerant mixture circuit preferably consists of a mixture of 0 to 20 mol% of methane, 35 to 90 mole% ethylene or ethane and 0 to 30 mole% propane. The chilled Natural gas flow becomes the heat exchanger E2 over Line 20 fed, in this cooled down to a temperature between -70 and -100 ° C and then via line 30 from the heat exchanger E2 deducted.

The third refrigerant mixture or SRC refrigerant mixture circuit becomes the heat exchanger E2 over Line S6 with a temperature between -35 and -55 ° C fed and against the refrigerant of the LRC mixed refrigerant circuit condensed in line L10. The refrigerant mixture in the line L10 evaporates at a pressure level between 1.5 and 6 bar. The cooled Refrigerant mixture of the SRC mixed refrigerant circuit with a temperature between -70 and –100 ° C via line S7 from the heat exchanger E2 deducted.

That evaporated and overheated Refrigerant mixture of the LRC mixed refrigerant circuit in line L10 the separator L1 is at a pressure between 1.5 and 6 bar supplied. The gaseous refrigerant mixture obtained at the top of the separator L1 is piped L2 fed to the compressor L3 and compressed to a pressure between 10 and 40 bar. The compressor L3 is preferably designed as a single-case axial or centrifugal compressor. Such cold suction compressors have the advantage that the medium to be sucked in before suction does not need to be warmed up to ambient temperature, which saves heating space and thus the heat exchanger smaller dimensioned and can be produced cheaper.

The compressed refrigerant mixture of the LRC-refrigerant mixture circuit is in the cooler L4, preferably against sea water or an appropriate cooling medium, cooled down to a temperature between 10 and 50 ° C. That from the cooler L4 via line L5 refrigerant mixture drawn off as already mentioned, in the heat exchanger E1 liquefied, via line L6 the heat exchanger E2 fed and subcooled to a temperature between -70 and -100 ° C against itself. The Evaporation temperature of the refrigerant mixture after the Joule-Thomson relaxation in the relief valve L9 - or alternatively in an expansion turbine - lies between –72 and –112 ° C.

The third or SRC mixed refrigerant circuit serves the complete if necessary Liquefaction and hypothermia of the liquefied Hydrocarbon-rich electricity or natural gas electricity. This hypothermia is sensible or necessary so that no more than the required amount of the flash gas after the expansion of the liquefied hydrocarbon-rich Current in a downstream arranged nitrogen removal unit accrues.

The third refrigerant mixture or SRC refrigerant mixture circuit preferably consists essentially of a mixture of 0 to 15 Mol% nitrogen, 30 to 65 mol% methane and 0 to 45 mol% ethylene or ethan.

The line 30 to the heat exchanger E3 fed liquefied Hydrocarbon-rich electricity is heated to a temperature in the heat exchanger E3 from -145 subcooled to –160 ° C. After this hypothermia is the hydrocarbon-rich or natural gas flow via line 40 from the heat exchanger E3 subtracted and essentially to atmospheric pressure using a Joule-Thomson relaxation in the relief valve 50 - or alternatively in a relaxation turbine - relaxed.

The refrigerant mixture of the third or SRC refrigerant mixture circuit fed to the heat exchanger E3 via line S7 is subcooled in the heat exchanger E3 and then also subjected to a Joule-Thomson expansion in the expansion valve S10. Instead of the expansion valve S10, an expansion turbine can in turn be provided. The pressure relief valve S10 relaxes at a pressure level between 1.5 and 6 bar. The evaporation of the refrigerant mixture in the heat exchanger E3 serves Both the subcooling of the already liquefied hydrocarbon-rich stream and the self-subcooling of the refrigerant mixture of the SRC / refrigerant mixture circuit, which has not yet expanded.

That evaporated and overheated Refrigerant mixture of the SRC refrigerant mixture circuit is via line S11 fed to a separator S1. The gaseous refrigerant mixture obtained at the top of the separator S1 is piped S2 fed to a compressor S3. The refrigerant mixture is compressed in the compressor S3 a pressure between 30 and 60 bar. The one coming out of compressor S3 Refrigerant mixture subsequently in the cooler S4, preferably against sea water or an appropriate cooling medium, cooled.

The compression of the refrigerant mixture SRC refrigerant mixture circuit is preferably carried out in a single or multi-stage, single or multi-housing centrifugal compression device S3. In the case of very large ones Quantities can also be used instead of the centrifugal compression device an axial compression device can be provided.

Each of the three mixed refrigerant circuits has advantageously downstream of the respective expansion valve P13, L9 or S10 a separator / storage tank P11, L7 or S8. In principle can these separators / storage tanks too provided at any other suitable location of the refrigerant mixture circuits become.

From these separators / storage tanks P11, L7 and S8 becomes the liquid fraction over the Lines P16, L12 and S13 withdrawn and the vaporous top fraction (Flash gas) of the refrigerant mixture fed. This procedure ensures a good distribution of liquid and Gas and therefore good heat transfer in the heat exchangers E1, E2 and E3, especially if they are so-called plate-fin type heat exchangers acts, guaranteed.

In lines P16, L12 and S13 control valves P15, L11 and S12 are provided. These control valves serve to control the fluid level to be regulated within the separators / storage tanks P11, L7 or S8.

In the event of a system shutdown, the control valves P15, L11 and S12 are closed, so that the separators / storage tanks P11, L7 and S8 are filled with the refrigerant mixture of the respective refrigerant mixture circuit; for this purpose it makes sense that the separators / storage tanks P11, L7 and S8 in the 1 Shut-off valves, not shown, are provided. This enables the refrigerant mixture to be stored at the coldest point of the respective refrigerant mixture circuit, which speeds up the start-up procedure when restarting. The separators / storage tanks P11, L7 and S8 should preferably be dimensioned such that they can store the total amount of mixed refrigerant in a mixed refrigerant circuit.

As a rule, the compressors P8, P3, L3 and S3 powered by own gas turbines. But it can also several or even all compressors together from several gas turbines or a gas turbine G are driven - represented by the dash-dotted line Line.

If the hydrocarbon-rich stream to be liquefied has a certain proportion of higher or heavy hydrocarbons, these must be removed from the hydrocarbon-rich stream before the actual liquefaction process, since they would otherwise freeze out in the liquefaction part and lead to relocations within the liquefaction process. This separation of heavy hydrocarbons usually takes place by providing a so-called HHC (Heavy Hydrocarbon) column, which is used to separate the heavy hydrocarbons and benzene from the hydrocarbon-rich stream to be liquefied. Such a procedure is also in the aforementioned German Offenlegungsschrift 197 16 415 described; see for example 2 as well as the associated figure description.

Basically, it should be noted that with liquefaction a hydrocarbon-rich stream is usually attempted, the pressure at which the hydrocarbon-rich stream is present to use and maintain the liquefaction in the heat exchangers in one if possible to realize warm temperature range. This procedure results in a comparatively small outlay on apparatus, in particular heat exchangers, as well as low operating costs. Usually natural gas is under a pressure of at least 50 bar, often also under a pressure of 70 bar and above.

The maximum usable pressure of the liquefying However, hydrocarbon-rich electricity is limited. One of the Causes lies in that the rectification separation of heavy hydrocarbons from the hydrocarbon rich to be liquefied Current and thus the setting of the maximum permissible calorific value of the LNG product flow through the rapprochement at the critical pressure and reducing the density difference of vapor and liquid is difficult or limited in the HHC column. Far away come with conventional ones liquefaction plants for a variety of reasons Plate heat exchanger for use; however, these "work" when reached or more uneconomical above certain design pressures.

The article "High Efficiency 6 MTPA LNG Train Design Via Two Different Mixed Refrigerant Processes" (AIChE Spring National Meeting, Manch 10–14, 2002, New Orleans, Louisiana) is over Scheme 3.2 shows a liquefaction process for natural gas, in which the natural gas stream to be liquefied relaxes before a C ₃ / C ₄ and a C ₅₊ fraction is separated, and the natural gas stream withdrawn from the separation is then compressed again before being fed into the liquefaction part , A C ₃ / C ₄ and a C ₅₊ fraction are separated by rectification in conventional rectification columns, in which part of the steam rising in the columns is condensed by a refrigerant and then used as a liquid reflux.

Using this procedure, however, it is only possible with a very high expenditure of energy to obtain a C ₃₊ -rich fraction with a high yield. The term "high yield" should be understood to mean yields of at least 60%.

The object of the present invention is to provide a generic method which, in addition to the liquefaction of a hydrocarbon-rich stream, in particular a natural gas stream, makes it possible at the same time to obtain a C ₃ + -rich fraction with high yield.

• a) der Kohlenwasserstoff-reiche Strom vor seiner Auftrennung in eine C_2–-reiche und in eine C₃₊-reiche Fraktion entspannt,
• b) der entspannte Kohlenwasserstoff-reiche Strom einem C₃-Absorptionsprozess zugeführt und in diesem in die der Verflüssigung zuzuführende C_{2 –}-reiche Fraktion und in eine erste Kondensatfraktion aufgetrennt,
• c) die der Verflüssigung zuzuführende C_2–-reiche Fraktion vor der Verflüssigung verdichtet,
• d) die erste Kondensatfraktion angewärmt einem C₂-Stripp-Prozess zugeführt,
• e) aus dem Sumpf des C₂-Stripp-Prozesses eine zweite, C₃₊-reiche Kondensatfraktion mit hoher Ausbeute gewonnen, und
• f) am Kopf des C₂-Stripp-Prozesses eine C_3–-reiche Gasfraktion abgezogen, partiell kondensiert und dem C₃-Absorptionsprozess als Waschmittel zugeführt wird.

To solve this task, it is proposed that

• a) the hydrocarbon-rich stream relaxes before it is separated into a C _2– -rich and a C ₃₊ -rich fraction,
• b) the relaxed hydrocarbon-rich stream is fed to a C ₃ absorption process and separated therein into the C _{2 -} -rich fraction to be fed to the liquefaction and into a first condensate fraction,
• c) the C _2– -rich fraction to be supplied to the liquefaction compresses before the liquefaction,
• d) the first condensate fraction is heated and fed to a C ₂ stripping process,
• e) a second, C ₃₊ -rich condensate fraction is obtained in high yield from the bottom of the C ₂ stripping process, and
• f) a C ₃ - -rich gas fraction is drawn off at the head of the C ₂ stripping process, partially condensed and fed to the C ₃ absorption process as a detergent.

The process according to the invention now makes it possible to obtain a fraction rich in C ₃₊ with a high yield and a greatly reduced energy expenditure. This is achieved in that, on the one hand, the pressure of the hydrocarbon-rich stream to be liquefied is reduced before the C ₃₊ -rich fraction to be recovered and is only increased again before being fed into the liquefaction part, and on the other hand the separation process of the C _{3+ to be} recovered -rich fraction consists of an innovative combination of a C ₃ absorption process with a C ₂ stripping process. The C ₂ stripping process follows the C ₃ absorption process.

Further advantageous configurations of the method according to the invention are objects the dependent Claims.

According to an advantageous embodiment of the method according to the invention, the pressure of the hydrocarbon-rich stream is reduced in the expansion by 10 to 60. Likewise, the pressure of the C ₂ - -rich fraction subjected to the compression is preferably increased by 20 to 100%.

Further developing the method according to the invention it is also proposed that the relaxation and / or compression are carried out in several stages.

Advantageously, the the relaxation of the liquefied Hydrocarbon-rich electricity generated energy to drive the compressor or compressors used.

It is also advantageous if the C ₂ stripping process is operated at a slightly higher pressure, preferably at a pressure which is 1 to 5 bar higher than the C ₃ absorption process. In this case, the first condensate fraction drawn off from the C ₃ absorption process is preferably pumped to the pressure prevailing in the C ₂ stripping process.

The method according to the invention and further refinements of the same, which are the subject of the dependent claims, are described below with reference to 2 illustrated embodiment explained in more detail.

In the 2 are already based on the in the 1 The procedure described in detail, three refrigerant mixture cycles only partially shown. Their mode of action will no longer be discussed below.

The hydrocarbon rich to be liquefied Electricity or natural gas flow is supplied to the heat exchanger E1 via a line 1 fed. On an appropriately chosen one The temperature level becomes the hydrocarbon rich to be liquefied Electricity from the heat exchanger E1 over Line 1 'withdrawn and fed to the separator D. The separator D serves to separate off the partial condensation in the heat exchanger E1 resulting liquid portion in the liquefied Hydrocarbon-rich electricity.

This liquid portion is withdrawn from the bottom of the separator D via line 9, in which a relief valve a can be provided, heated in the heat exchanger E and partially evaporated and then fed to the C ₂ stripper T2 to be described. The procedure described above relating to the liquid fraction drawn off from the bottom of the separator D is optional.

At the top of the separator D, the hydrocarbon-rich stream to be liquefied is drawn off, expanded in one or more stages in the expander X and then fed via line 2 to the C ₃ absorber T1.

A first condensate fraction is drawn off from the sump of the C ₃ absorber T1 via line 8, pumped to the pressure prevailing in the C ₂ stripper T2 by means of the pump P and, after heating in the heat exchanger E, is fed to the C ₂ stripper T2 on its head , The C ₂ stripper T2 has a sump heater R1.

The provision of the pump P is necessary at least when - according to an advantageous embodiment of the method according to the invention - the C ₂ stripping process T2 at a slightly higher pressure, preferably at a pressure 1 to 5 bar higher than the C ₃ absorption process T1 is operated.

From the bottom of the C ₂ stripper T2, a second condensate fraction, which represents the C ₃₊ -rich fraction to be obtained, is withdrawn via line 11 and, as a rule, fed to further processing. The yield of the C ₃₊ -rich fraction drawn off via line 11 is at least 60%. In principle, with appropriate process control, any high yields can be achieved.

At the head of the C ₂ -Strippers T2 a gaseous, essentially C ₃ -rich fraction is withdrawn, partially condensed in the heat exchanger E and then passed via line 10 to the head of the C ₃ absorber T1.

At the top of the C ₃ absorber T1, a C _{2 -} -rich fraction is drawn off via line 3 and subjected to the single-stage or multi-stage compression V. The compressor V is preferably driven by the expander X - represented by the dash-dotted connecting line.

The compressed C _{2 -} -rich fraction is then fed via line 3 'to a branch point and divided into two partial streams - lines 4 and 5.

The line 4 through the heat exchanger E2 and E3 guided Partial flow is partially or completely liquefied in this (heat exchanger E2) and possibly completely liquefied and supercooled (Heat exchanger E3). The hypothermic liquefied hydrocarbon-rich Electricity is then on the line 4 'of the separation column Fed T3, previously for the purpose of heating the reboiler R2 before Relaxation in the expansion valve b through the bottom of the separation column T3 led becomes.

The separation column T3 is used for the separation of nitrogen, a stream rich in nitrogen and methane being drawn off at the top of the separation column T3 via line 6. This nitrogen and methane-rich stream drawn off via line 6 - the so-called tail gas stream - is heated in the heat exchanger E4 against the second partial stream of the compressed C _{2 -} -rich fraction carried in line 5. The liquefied C _{2 -} - rich fraction is then passed via line 5 'and expansion valve c also to the separation column T3 - either on the same tray or any tray below the feed point of the hydrocarbon-rich stream in the line 4'.

That from the bottom of the separation column T3 withdrawn, liquefied and hypothermic Natural gas is piped by means of the pump P ' 7 fed to storage and / or further processing.

Of course you can on the in the 2 separation column T3 shown are dispensed with; in this case the entire, compressed C _{2 -} -rich fraction would be passed through the heat exchangers E2 and E3.

After the pressure in the process according to the invention is reduced only for the purpose of obtaining a C ₃₊ -rich fraction upstream of the liquefaction process and increased again before the actual liquefaction process, the advantages associated with a high pressure can be realized in the liquefaction part while the production of the C _{3 +} -rich fraction is facilitated due to the reduced pressure.

Claims (9)

Process for liquefying a hydrocarbon-rich stream, in particular a natural gas stream, with simultaneous recovery of a C ₃₊ -rich fraction with high yield, the liquefaction of the hydrocarbon-rich stream in the heat exchange with the refrigerants of a refrigerant mixture circuit consisting of at least three refrigerant mixture circuits having different refrigerant compositions takes place, and wherein the rich hydrocarbon to be liquefied stream to a C _{2 -} -rich fraction, which is subjected to liquefaction is separated and into a C3 ₊ -rich fraction, characterized in that a) the hydrocarbon-rich stream ( 1, 1 ') before its separation into a C _{2 -} -rich and into a C ₃₊ -rich fraction relaxes (X), b) the relaxed hydrocarbon-rich stream (2) is fed to a C ₃ absorption process (T1) and in this into the C _{2 -} -rich F to be fed to the liquefaction (E2, E3) reaction (3) and separated into a first condensate fraction (8), c) the C _Z = rich fraction (3) to be supplied to the liquefaction condenses (V) before the liquefaction (E2, E3), d) the first condensate fraction (8) is warmed (E) fed to a C ₂ stripping process (T2), e) a second, C ₃ + -rich condensate fraction (11) was obtained from the bottom of the C ₂ stripping process (T2) with high yield, and f) at the head of the C ₂ stripping process (T2), a C ₃ - -rich gas fraction is drawn off, partially condensed (E) and fed to the C ₃ absorption process (T1) as a detergent (10). A method according to claim 1, characterized in that the hydrocarbon-rich stream (1, 1 ') is partially condensed before it relaxes (X) and the liquid portion which forms is separated off (D) and fed to the C ₂ stripping process (T2) (9). A method according to claim 2, characterized in that the separated (D) and the C ₂ stripping process (T2) supplied (9) liquid portion is relaxed (a) and / or heated (E) before being fed. Method according to one of the preceding claims 1 to 3, characterized in that the pressure of the hydrocarbon-rich Current (1, 1 ') in the relaxation (X) is reduced by 10 to 60%. Method according to one of the preceding claims 1 to 4, characterized in that the pressure of the C ₂ - - rich fraction (7) subjected to the compression (V) in the compression (V) is increased by 20 to 100%. Method according to one of the preceding claims 1 to 5, characterized in that the relaxation (X) and / or the Compression (V) are carried out in several stages. Method according to one of the preceding claims 1 to 6, characterized in that the one obtained during relaxation (X) Energy is used to drive the compressor or compressors (V). Method according to one of the preceding claims 1 to 7, characterized in that the C ₂ stripping process (T2) is operated at a slightly higher pressure, preferably at a pressure 1 to 5 bar higher than the C ₃ absorption process (T1) becomes. A method according to claim 8, characterized in that the first condensate fraction (8) is pumped (P) to the pressure prevailing in the C ₂ stripping process (T2).