NO146554B - PROCEDURE AND APPARATUS FOR SEPARATING A SUPPLY GAS UNDER PRESSURE - Google Patents

Description

The present invention relates to the treatment of gases in containing hydrocarbons and other gases of similar volatility,

with a view to separating condensable fractions. In particular, the invention relates to the treatment of gases such as natural gas, synthesis gas and refined gases for the extraction of most of the propane and most of the ethane in the gas together with practically all heavier hydrocarbons in the gas, as well as an apparatus for this.

Gases containing hydrocarbons and other gases of similar volatility that can be treated according to the invention include natural gases, synthesis gases from other hydrocarbon-containing materials such as coal, crude oil, petrol, shale oil, tar sands and lignite. Natural gases usually have a predominant amount of methane and ethane (ie the combined and C^ fractions form at least 50% of the gas on a mole basis). There may also be smaller amounts of relatively heavier hydrocarbons such as propane,

butanes, pentanes and the like, as well as H^* N2' C02°s other gases.

A typical composition for natural gas that can be treated according to the invention is in approximate mol%: 80% methane, 10% ethane, 5% propane, 0.5% isobutane, 1.5% n-butane, 0.25/» isopentane, 0.25/» n-pentane, 0.5% hexane-+ and the rest nitrogen and carbon dioxide. Sulfur-containing gases are also often found in natural gas.

The later substantial increases of. the demand for ethane and propane from natural gas has led to a demand for a method that gives a high degree of recovery or yield of these products. Available processes for processing these gases are based on cooling and freezing the gas, oil absorption, absorption with cooled oil and the later cryogenic processes which use the principle of gas expansion through a mechanical device for the production of power while heat is simultaneously extracted from the system. Depending on the pressure of the gas source, the richness of the discovery (the content of ethane and heavier hydrocarbons) and the desired end products, these previously known methods or combinations of them can be used.

Recovery by cryogenic expansion is currently generally preferred for the recovery of ethane because it combines maximum simplicity with easy start-up, flexibility, good efficiency, safety and good reproducibility. US Patent Nos. 3,360,944, 3,292,380 and 3,292,381 describe relevant methods.

In extraction based on a typical cryogenic expansion process, an incoming gas stream (hereinafter also called charge or supply) is cooled under pressure by heat exchange with other material streams in the process and/or external cooling media such as e.g. from a propane compression refrigeration system. When the gas cools, some liquids are condensed and collected in one or more separators as high-pressure condensate containing most of the desired C, -, + components. The high-pressure condensate is then expanded to a lower pressure. The evaporation that takes place during the expansion of the liquid provides a further cooling of the remaining part of the liquid. The cooled stream that forms a mixture of steam and liquid is freed from methane, demethanized, in a demethanization tower. The demethanization tower is a distillation tower where the expansion-cooled stream is fractionated for the separation of remaining methane, nitrogen and other volatile gases such as top gas from the desired products consisting of ethane, propane and heavier components that appear as bottom fractions.

If the incoming charge does not condense completely, and most often it does not, the gases from this partial condensation are expanded to a lower pressure. More liquid is condensed due to the additional cooling during expansion. The pressure after the expansion is usually the same as the operating pressure of the methane removal column. The produced liquids are thus used as input charge to the methane removal column. Most often, residual gases and the top fraction from the demethanization column are combined as methane residual gas.

In ideal operation of such a separation process

the gases leaving the process will contain practically all the methane present in the charge and essentially no hydrocarbons equivalent to ethane or heavier constituents. The bottom fraction from the demethanization column will contain practically all heavier components and substantially no methane.

In practice, however, the ideal situation is not achieved for the reason that a normal methane removal plant is operated to a large extent as a distillation column. The methane product will therefore typically consist of gases that leave the top fraction stage in the column together with gases that have not undergone any distillation. Substantial losses of ethane occur because the gases emerging from the low-temperature separation contain ethane and heavier constituents which could be recovered if these gases could be further cooled or if they were brought into contact with a significant amount of relatively heavier hydrocarbons, -f .ex. and heavier, which could absorb the ethane.

As described in US Patent No. 4,171,964.

and in Norwegian application 772057, improved ethane yield is achieved by cooling the condensed high-pressure liquid before its expansion. Such cooling will reduce the temperature in the flash-expanded liquid charge that is fed to the methane deboiler (deboiler column) and thus improve the ethane yield. As further described in the aforementioned references, by pre-cooling the high-pressure charge - which is liquid - the temperature of the liquid can be reduced sufficiently so that it can be used as the input top charge on the methane reboiler column, while the expanded steam is led to the methane reboiler at an introduction level that lies between the top and the bottom of the column . This variation makes it possible to extract ethane which is present in the expanded gas and which would otherwise be lost.

It is clear that the connection of external cooling on . this step of the process is difficult because of the very low temperatures at which it is operated. In typical methane reboiler operations, the expanded liquid and gas maintain temperatures in the range of -84°C to -123°C. Consequently, cooling of the condensed high-pressure liquid can best be achieved by heat exchange with currents diverted in the process as described in the above-mentioned patent.

One will understand from this that the high-pressure liquid that forms the incoming charge - generally contains volatile gases

(such as methane) as well as gases with lower volatility and that cooling of the incoming high-pressure supply during expansion is due to evaporation of part of the volatile gases. According to the present invention, the temperature drop achieved by expansion of the high-pressure supply can be increased by combining this supply with a stream from the process that has a bubble point lower than the high-pressure liquid supply's bubble point at the pressure to which the charge is expanded. Prior to expansion, the combined stream is therefore cooled to a temperature that is lower than the temperature of the high-pressure supply.

This can be done by cooling the liquid stream

of high-pressure charge or process gas (or both) before they are combined, or by cooling the combined stream if this is more appropriate. After expansion, the combined stream will reach a lower temperature due to the greater contained amount of the more volatile constituents which lower the bubble point of the combined stream and which evaporate at the lowest pressures and absorb increased amounts of vaporization heat.

It will be realized that in practice the bubble point temperature for incoming high-pressure liquid will be several degrees or more above the real treatment temperature due to the lack of complete equilibrium during condensation and separation of high-pressure liquid and gas. Such a condition also occurs when the high-pressure liquid is cooled as described in the aforementioned references. When the bubble point temperature significantly exceeds the actual treatment temperature for incoming high-pressure liquids, the temperature drop during expansion will be less than the temperature drop that could be achieved by expanding the high-pressure liquid at the bubble point.

According to the present invention, such a liquid can

charge of high-pressure liquid is combined with a more volatile process flow as described above and with moderate additional cooling give an improved operation and yield. This is because mixing the process gas into the high-pressure liquid charge will lead to the absorption of volatile gases until the actual bubble point of the high-pressure liquid can be reduced to the treatment temperature. Expansion of a liquid with such a reduced bubble point will produce colder cooling temperatures.

According to the present invention, there is thus provided a method and an apparatus for separating a supply gas under pressure. The essential features of the invention

signs appear from the following requirements.

For a better understanding of the invention, reference is made to

the attached forms and diagrams where:

fig. 1 shows a flow diagram of a prior art one-stage freezing plant for expansion cooling of natural gas, with a set of conditions applicable to a typical rich natural gas,

fig. 2 shows a similar flow diagram of a known type with a set of conditions for typical lean natural gas,

fig 3 shows a flow diagram from Norwegian

application no. 772057, and illustrates a technique by which high pressure input fluid can be precooled prior to expansion,

fig. 4 shows a diagram showing the application of the present invention to the pre-cooling of the incoming charge according to fig. 3, and

fig. 5 shows part of a flow chart with application of the present invention when precooling the incoming charge by flash expansion of a part of the charge.

Fig. 6A and 6B are curves of the carbon dioxide content versus temperature from an embodiment of the invention compared to the relationship in known embodiments.

In the following explanation of the attached diagrams, there are attached tables indicating flow quantities calculated for representative process conditions. The flow quantities are expressed in kg'mol/hour, which is rounded to the nearest whole number. The total flow rate shown in the tables includes all non-hydrocarbon components and is therefore usually greater than the sum of the hydrocarbon components' flow rate. The indicated temperatures are approximate, rounded to the nearest degree.

Reference is made to fig. 1 for a more complete description of a typical common ethane extraction process, where it is operated with an input gas from which carbon dioxide and sulfur compounds have been removed (if the concentration of such compounds in the input gas would cause the product not to meet specifications or cause icing of the equipment), and which is dehydrated, and this gas enters the process at 49°C and 64 kg/cm absolute, as stream 23. It is split into two parallel streams and cooled to -7°C by heat exchange with cold residual gas at -15°C in the heat exchanger 10, with product liquids (stream 26)

at 28°C in heat exchanger 11 and with the methane reboiler liquid at 12°C in methane reboiler 12. From these heat exchangers the flows are combined

again and then into the gas cooler, heat exchanger 13, where the combined stream is cooled to -12°C with propane refrigerant at -15°C. The cooled stream is split again into two parallel streams and cooled further with heat exchange with cold residual gas (stream 29) which maintains -77°C in heat exchanger 14, and with methane deboiler liquids at -62°C in methane sideboiler 15.

These streams are recombined and enter a high-pressure separator 16 at -43°C and 63.3 kg/cm<2> as stream 23a. The condensed liquid (stream 24) is separated and led to the methane reboiler 19 through the expansion valve 30. An expansion motor can be used instead of the expansion valve 30 if desired.

The cooled gas from the high-pressure separator 16

passes through the expansion machine 17 where the gas is expanded. from 63.3 kg/cm 2 to 20.4 kg/cm ?. The working expansion cools the gas to -87°C. The expansion machine 17 is preferably a turbine expander with a compressor 21 mounted on the expansion machine's shaft. In certain previously known embodiments, the expansion machine 17 is replaced by a normal expansion valve.

The liquid which is condensed during the expansion is flushed out in the low-pressure separator 18. The liquid goes by level control through line 25 to the methane reboiler column 19 at the top and flows from a pipe distillation bowl (not shown) as a top charge to the column 19

It is noted that in certain embodiments the low pressure separator 18 can be inserted as part of the methane reboiler 19 and then occupy the top of the column. In this case, the outlet stream from the expander enters over a pipe distillation bowl at the bottom of the separator section at the top of the column. The liquid then flows from the distillation plate (bowl) as top charge to the methane reboiler section of the column.

Liquid charge to the methane reboiler 19 goes down through the column and comes into contact with gases which distil methane from the liquid and produce a methane-free liquid product at the bottom. The heat required for the production of distillation gases comes from the heat exchangers 12 and 15-

Gases freed from condensed liquid in the methane reboiler 19 pass through line 27 together with cold outgoing gas from separator 18 through line 28. The combined gas flow then returns through line 29 to heat exchangers 14 and 10. After these heat exchangers, the gas flows through the compressor 21 which is operated by the expansion machine 17 and is directly connected to this. The compressor 21 compresses the gas to an output pressure of approx. 21.4 kg/cm . The gas then enters compressor 22 and is compressed to the final output pressure of

63.3 kg/cm .

Flow rates for incoming and outgoing liquid components, degree of recovery for outgoing liquid and compression effect for this previously known method which is shown in fig. 1 appears from the following table.

In fig. 2, a typical lean natural gas is treated and it is cooled according to a previously known method of the same construction as shown in fig. 1. Incoming gas stream 33 is cooled to ~55°C and flows to high-pressure separator 16 as stream 33a where the contained liquid is separated and fed via level regulator through line 34 and the expansion valve 30 to methane reboiler 19 in the middle of the reboiler column.

Cold gas from separator 16 passes through the expansion machine 17 where, due to expansion work from 63.3 kg/cm<2> to 17.6 kg/cm2, it is cooled to -103°C. The liquid that condenses during the expansion is separated in low-pressure separator 18 and fed via level control through line 35 to methane reboiler 19 as top charge to the column.

Data for this case appears in the following table:

In the previous method shown in fig. 1 and 2, the electron recovery rate is 73% for the charge that constitutes rich gas and 79% for the lean gas charge. It is known that a certain improvement in yield can be achieved by introducing one or more cooling steps followed by one or more separation steps or by changing the temperature in separator 16 or the pressure in separator 18. Recovery rates for ethane and propane produced on this the way, even if they are possibly improved in relation to the cases illustrated in fig. 1 and fig. 2, is, however, considerably lower than the yields that can be obtained according to the method of the present invention.

As an illustration, the process conditions in fig. 2 is changed by reducing the column pressure to 15.8 kg/cm o. At this lower pressure, the ethane and propane yields will be somewhat increased (to 82.96% and 98.66%, respectively), however, the lower operating pressure requires a significant increase of the power requirement of 2983 BHK (brake horsepower).

Fig. 3 shows a method, as described in

the previously mentioned US patent no. 4,171,964 for pre-cooling incoming high-pressure charge in liquid form. According to the procedure in fig. 3, the partially condensed gas charge 33a is obtained at -55°C

and 63.3 kg/cm o in the same way as appears from fig. 2. Incoming gas was here assumed to be a lean gas with a composition like stream 33 in fig. 2. The partially condensed gas 33a enters the high-pressure separator 16 where liquid and gas are separated.

By first following the gases 113 that leave the separator 16, these gases flow into an expansion engine 17 where mechanical energy is drained from the gas. As the gas is expanded from a pressure of approx. 63.3 kg/cm 0 to 17.6 kg/cm 2 the expansion work will cool the gas 113 to approx. -103°C. The expanded and partially condensed gas 113 is fed as input charge to the methane reboiler 19, where the gas portion rises and forms part of the methane reboiler top extraction 117. Top extractor 117 from the methane reboiler runs at a temperature of -104°C together with the gases 116 from the flash the evaporation which is described later and forms the residual gas 118. The combined cold residual gas flow 118 passes through the heat exchanger 119- The heated residual gas leaves the -87°C heat exchanger 119 and is fed back to the preceding cooling stage, e.g. illustrated in fig. 2, where further cooling effect from the still cold gas is utilized and the gas is compressed in compressor 21 (see fig. 2) which is driven by the expansion motor 17 and compressed further to the line pressure of 63.3 kg/cm o at the final compressor 22.

One now looks at the liquid 34 which comes from the separator 16 and then passes through the heat exchanger 119 in heat exchange with the cold residual gas 118. This results in a pre-cooling of the liquid part of the partially condensed high-pressure charge in gaseous form.

The subcooled liquid is expanded through a suitable expansion device such as e.g. an expansion valve 120 to a pressure of

about. 17.6 kg/cm . During the expansion, part of the charge will evaporate and cool the remaining part. In the process illustrated in fig. 3, the expanded stream leaves the expansion valve 120 at a temperature of -106°C and enters the separator. The liquid part is separated and fed as stream 115 to distillation column 19 as top charge. It may be noted that

in comparison with fig. 2, the expanded liquids through line J>H into the methane reboiler column only reach a temperature of -92°C. Because the current 115 in fig. 3 is significantly colder, it can be used as a top charge on the methane reboiler column for extraction of ethane in the stream 113- The extracted ethane is tapped off with the methane reboiler bottom product 125. The bottom product 125 is heat exchanged with the incoming charge to utilize the cooling effect as generally shown in fig. 1 and 2.

In connection with fig. 3, it is mentioned that for heat-economic reasons, one or more methane reboiler-sideboilers will be arranged as heat exchangers with a cold incoming charge (not shown in fig. 3), but as generally illustrated in fig. 1 and 2. With regard to the numbers and calculations made in connection with fig. 3 and which can be found in the following table, two such side boilers have been assumed as shown in fig. 2. These side boilers are important for the overall heat economy in the process. Subcooling of the liquid stream 34 with the tppp gas 118 reduces the available cooling effect from the stream 118. However, the increased load on the methane reboiler 19 with the liquid stream 115 which is cooled as shown in fig. 3 provide additional available cooling capacity in the side reboilers. Thus, the process's total heat balance remains essentially unchanged.

Incoming and outgoing flow quantities, yields and expansion/compression requirements in the process as shown in fig. 3 appears from the table:

For further comparison with the method according to the present invention which follows in the later examples, a second base case was calculated according to the flow diagram in fig. 3 from the same lean gas. With this modified flow plan, the incoming gas charge in the process at 49°C and 64 kg/cm<2> was cooled to -56°C in precoolers (e.g. heat exchangers 10, 11, 12, 14 and 15 in Fig. 2 ) instead of -55°C and the column was run at a slightly lower pressure of 16.9 kg/cm p instead of 17.6 kg/cm 2. The result was a slightly increased yield of ethane and propane and increased horsepower requirements in the process. A summary of the changed flow conditions and quantities for this alternative base case appears in table IV:

Example 1

The present invention is illustrated by this example, which is to be seen together with fig. 4. Fig. 4 shows part of a flow diagram where a lean gas 33a holding 63.3 kg/cm<2> is cooled to -55°C and enters separator 16. The gas charge is cooled and partially condensed by heat exchange with different process streams (these heat exchangers are not shown), including side reboilers on the methane reboiler column 19 (side reboilers not shown), and by heat exchange with methane reboiler bottom product and product gas as described in fig. 2. If necessary and as indicated in fig. 1 and 2, external cooling can also be used. The process conditions as described in fig. 4 and the flow quantities that appear in table V which follows correspond to the treatment of a lean gas charge with a composition as in table II and fig. 2.

According to the method shown in fig. 4, the partially condensed gas 33a, which contains a liquid part and a gas part, is led into the high-pressure separator 16 where the liquid part. separated from. The liquid from separator 16 (stream 34) is combined with part of the gas from separator 16 (stream 169). The combined stream then passes through heat exchanger 154 in heat exchange with the top gas 158 from the methane reboiler, which provides cooling and condensation of the combined stream. The cooled stream at -102°C is then expanded through a suitable expansion device such as e.g. an expansion valve 155 to a pressure of approx. 17.6 kg/cm . During expansion, part of the guide will evaporate and cool the remaining part. In the process of fig. 4, the expanded stream 157 that leaves the expansion valve 155 reaches a temperature of -108°C and goes to the top of the distillation column 19 as top charge.

The remaining steam from separator 16 (stream 170) is fed into expansion motor 17 where mechanical energy is drained from this part of the charge. When the gas is expanded from the pressure of 63.3 kg/cm 2 and to approx. 17.6 kg/cm 2, the working expansion will cool the gas to approx. -103°C. The expanded and partially condensed gas 153 is fed to the methane reboiler 19 in an intermediate step.

In comparison with the first basic case in fig.

3, it will be seen that the liquid 115 in fig. 3 which enters the methane reboiler column reaches a temperature of -106°C. In order to reach a lower temperature of -107°C at the top of the column in the alternative base case, a reduced column pressure was necessary.

The reduced column pressure increased the horsepower requirement, but improved yield only slightly. In fig. 4, as a result of the combination of the liquid 34 from the separator 16 with part of the high-pressure charge 169 before subcooling in the heat exchanger 154,

achieve that the incoming charge on the methane reboiler top maintains -108°C without lowering the methane reboiler pressure.

Incoming and outgoing flow quantities, yields and expansion/compression requirements for example 1 can be found in table

V.

Comparison between the ethane and propane yield for Tables III and V shows that without enrichment of the liquids from separator 16, the ethane yield is 89.1% and the propane yield is 97.7%. Enrichment of the separator liquids according to example 1 (see fig. 4) increases the ethane and propane yields to 92.2% and 98.3% respectively.

Comparison between Tables IV and V shows further

that the improvement of yields according to the present process is not simply due to increased effect. On the contrary, Table IV shows that even when the process conditions in the base case were changed so that the methane digester was operated at a lower pressure, which increased the horsepower requirement to 2773 BHK, the ethane and propane yields were only increased to 90.66% and 98.08%, respectively. When using the present invention as in example 1, the yield of ethane and propane was increased compared to the alternative base case,

although somewhat less effect was included.

From a practical point of view and particularly in connection with lean gases, the entire amount of liquid from separator 16 will be combined with part of the gas from separator 16. The combined stream will then be cooled and expanded as described. The amount of gas used in the combined stream is sufficient to achieve the desired cooling and temperature regulation for the top charge of the methane still. When added to the gas entering the top of the column, the liquids from separator 16 will increase the entering charge surface tension under column conditions and thereby reduce the formation of small liquid particles that are difficult to separate from in the top gas.

For richer gases and when there is more liquid

from separator 16 than is allowed to maintain the conditions at the top of the column, it may be more economical from a practical point of view to divide the liquid from separator 16 and expand a part directly to the tower or possibly after some subcooling. This can lead to savings in the heat exchange requirement and higher yields.

As indicated in the above-mentioned references, a number of modified flow schemes are characterized by subcooling part or all of the liquid quantity from separator 16, with which the present invention can be used. Two or more of these techniques can be used simultaneously. Among these flow schemes are the following: 1. Non-condensed gases from separator 16 can be expanded, e.g. in an expansion engine for the production of a cold partially condensed liquid and gas. The liquid is separated and fed to the methane reboiler column. All or part of the liquid that is thereby separated can be used as a cooling source for subcooling liquid condensate 34 from separator 16. Alternatively, all or part of the expanded gas can be used. In addition, side reboilers can be used for cooling the condensate 34 from separator 16. According to the present invention, the flow diagrams can be modified by combining liquid condensate from separator 16 with part of the gas from that separator before subcooling and flashing the liquid condensate. 2. Condensate from separator 16 can be led to sub-cooling by heat exchange and then to an expansion valve where it is expanded from line pressure (63> 3 kg/cm in fig. 1-4)

to the operating pressure in the methane reboiler column. This will give a gas-liquid mixture that can be separated either in a low-pressure separator or introduced directly to the methane reboiler column with internal separation devices for separating gas and liquid in the column. The flash expansion provides further cooling of the charge on the column. Part of the cooled external liquid obtained in this way is used as a cooling medium for heat exchange with the high-pressure condensate from separator 16 and is then fed to the methane reboiler column as a second charge on a middle stage in the column. According to the present invention, such a method can

is improved by enriching the liquid condensate leaving separator 16

with part of the gases from this separator before subcooling and flashing the liquid condensate.

3- The non-condensed gas leaving separator 16 can be expanded in a working expander from a high pressure such as 63.3 kg/cm 2 in fig. 1-4 and to the operating pressure in the methane still,

and the entire cooled gas-liquid mixture from the expansion can be used for subcooling the condensate from separator 16. The subcooled condensate is then flash-expanded and led to the methane reboiler as input charge. This performance can be improved by enriching the liquid condensate from separator 16 with a part of the gases leaving the separator before subcooling and flashing the condensate.

Instead of or in addition to these methods, external cooling can be used if an increased yield is desired, however, one of the advantages of the invention is described in the aforementioned applications and consists in the fact that when the condensate is subcooled, an improved yield can be achieved without the need for increased power.

Yet another embodiment of the invention can be seen from the example below which is linked together with fig. 5:

Example 2

Fig. 5 shows part of a flow core for extracting ethane and heavier components from a hydrocarbon mixture containing methane, ethane and heavier hydrocarbons.

As shown in fig. 5, a partially condensed high-pressure gas 174 is introduced

to separator 16 with -48°C and 63.3 kg/cm<2>. Cooling the gas charge to -48°C can, for example, shown in fig. 1 and 2 occur by heat exchange with methane residual gas or other process streams flowing from side reboilers and bottom streams (these heat exchangers are not shown) and, if necessary, with suitable external cooling. The calculations on which the present example is based are based on two methane side digesters. In contrast to fig. 1 and 2, however, with the present calculations (applicable temperature, pressure and flow rates) it is based on the incoming gas having a composition between lean and rich gas as respectively used in fig. l..and 2' and tables I and II.

As indicated in fig. 5. liquid and gas are separated from each other in separator 16. The gas from separator 16 is divided into two parts. The first part 176 passes through expander 17 where, due to work expansion from 63.3 kg/cm to 20.4 kg/cm, it is cooled to approx. ~92°C. From expander 17, the cooled gas goes to methane reboiler 19 as middle charge. The second gas 177 is combined with part of the subcooled liquid from heat exchanger 184 when it goes to heat exchanger 185.

The liquid 175 from separator 16 passes through heat exchanger 184 where it is subcooled to -90°C by heat exchange with the cold flow from the expansion valve 182. The subcooled liquid is then divided into two parts. The first part 178 passes through the expansion valve 182 where it is expanded and flash-evaporated by pressure reduction from 63.3 to 17.6 kg/cm . The cold flow from the expansion valve 182 then flows through heat exchanger 184 where it is used to subcool the liquids from separator 16. From heat exchanger 184 the flow goes to methane reboiler 19 as the lowest charge with -55°C.

The remaining liquid part 179 from heat exchanger 184, still at high pressure, is combined with part 177 of the gas from separator 16. The combined flow passes through heat exchanger 185 which is cooled to approx. -96°C with the top extract from column l80. At this temperature, the combined stream is substantially condensed. The condensed stream enters the expansion valve 184 where it is expanded and flash-evaporated by pressure reduction from 63 kg/cm to 17.6 kg/cm. From the expansion valve 183, the cold flow continues to the methane reboiler 19 as top charge.

The relevant data appears in the table below:

It is noted that in addition to the method indicated in fig. 5 for treating the cooled liquid 175 from separator 16, other methods can in some situations be used with advantage. Such an alternative method consists in passing a part of the cooled liquid 179 directly from the separator through another expansion valve to the methane reboiler column 19

at an intermediate level.

According to another method, the liquid 175 from separator 16 can be subcooled with residual gas instead of self-cooling as shown in fig. 5. With such an alternative, the high-pressure condensate can be cooled in two successive heat exchangers, each of which uses residual gas as a cooling medium. After passing through the first heat exchanger, the partially cooled high-pressure charge is divided into two parts. The first part is expanded through an expansion valve and fed to the methane reboiler column as the middle feed. The second part of the partially condensed condensate continues through the second heat exchanger where it is further cooled and combined with gas from separator 16. The combined stream is thereby further cooled and expanded and is fed to column 19 as top charge. Alternatively, gas from separator 16 could be added to the second part of the partially cooled stream before it entered other heat exchangers, thereby avoiding subsequent cooling of the combined stream.

In yet another modification of the present invention, the flash-expanded current, for example 186

on fig. 5j is led to heat exchange with the work-expanded stream I87 for cooling this and heating stream 186. If stream 187 thereby cools sufficiently, it can advantageously be used as top charge to the methane reboiler and stream 186 can be used as intermediate charge on the column since stream 186 as is clear from the plan in fig. 5 is richer in heavier constituents such as C2+ and stream 187 contains more lighter constituents, i.e. methane and non-condensable gases.

Other alternative methods for forming the cooled liquid 175 are described in the aforementioned patent applications. These alternative processes can be used in different combinations. These different schemes can also be used instead of or simultaneously with cooling with residual gas of the enriched stream before it is used as top charge on column 19-

These alternative methods are particularly useful when the cooling capacity of the top draw 180 from the column due to rich charge into the process is insufficient to cool the entire volume of liquid passing through line 175 to the desired low temperature.

As is known, natural gas can contain relatively large amounts of carbon dioxide. Carbon dioxide in the methane reboiler can lead to icing of the internal parts of the column under the cooling conditions used. Even when the incoming charge gas contains less than 1% carbon dioxide, it will fractionate in the methane reboiler and build up to concentrations of up to 5 or 10% or more. At such high concentrations, the carbon dioxide can freeze out depending on temperature, pressure, whether the carbon dioxide is in liquid or gas phase and the solubility of the liquid phase.

In the present invention, it has been found that when the gas from the high-pressure separator is expanded and fed to the methane reboiler lower than the top level, the problem of carbon dioxide icing can essentially be avoided. The gas from the high-pressure separator typically contains a large amount of methane compared to ethane and carbon dioxide. When the gas is thus fed to a middle level of the column, separator gas from the high-pressure separator will dilute the carbon dioxide concentration and prevent it from increasing to icing level.

The advantage of the present invention can be easily seen by setting up the carbon dioxide concentration and temperature for the different distillation stages in the methane still. To illustrate the production of such a diagram, the flow process illustrated in fig. 4 is used for treating a gas charge with the following composition:

The same incoming gas was treated according to fig. 2. However, in order to achieve a more efficient utilization of the cooling effect, the cooling-heat exchangers upstream of the high-pressure separator were completely relocated. The main conditions under which the process operated were the following:

Curves were set up for each of these cases with regard to CC^ concentration as a function of temperature in the methane reboiler, shown in fig. 6A and 6B. In fig. one also finds the liquid-solid and gas-solid equilibria. Equilibrium data indicated in fig. 6A and 6B apply to the methane-carbon dioxide system. These data are generally assumed to be representative of methane and ethane systems. If the CO^ concentration at a particular temperature in the column is at or higher than the equilibrium line at this temperature, it can be assumed that icing is occurring. For practical purposes, the designer usually uses a safety margin, ie the real concentration should be less than the "icing" concentration by a suitable safety factor.

In fig. 6A the carbon dioxide concentration rises

in the methane reboiler above the limit level. Such a gas could not be used in a normal process and therefore could not be used without pre-treatment to remove a significant amount of the carbon dioxide. When the expanded gases are used as middle charge on the column

according to the present invention, however, the CC^ concentration will be reduced in the methane reboiler to a point well below the "icing" level.

When constructing methane reboiler columns for

use in the present invention respectively, the designer will routinely ensure that icing does not occur in the column.

Even when the gas is fed at a middle stage in the column, it is possible that icing may occur if the process is set up for the highest possible ethane cabin. Such an arrangement will usually require the coldest possible temperature at the top of the column. This will cause the carbon dioxide concentration to go to the right

the curves in fig. 6A and 6B. Depending on the arrangement, the result may be an unacceptably high concentration of carbon dioxide near the top of the column. In such cases, it may be necessary to accept a somewhat lower ethane yield to avoid icing or to pretreat the incoming gas to reduce the carbon dioxide content to a point where the concentration does not presents problems in the methane reboiler. As an alternative, icing in such cases can be avoided by other modifications of the conditions. For example, it may be possible to run the high-pressure separator at a higher temperature, increase the relative amount of gas from the high-pressure separator that is expanded through expander 17 or expand part of the gases from the high-pressure separator through an isenthalpic expansion valve. If such changes can be made within the heat balance constraints of the process, icing can be avoided without reduced ethane yield".

In connection with the above method, it should be noted that in some cases the charge introduced at the top of the methane reboiler is a liquid that has been expanded from a high pressure to the pressure of the methane reboiler (see, for example, Fig. 4 and 5)-

In such cases, it may be beneficial to self-cool the liquid feed at the top of the column. This happens by dividing the liquid stream at the top of the column into two streams, either before or after expansion. (Both streams are expanded if the top charge of the column is split before expansion). One of the two expanded streams is fed directly to heat exchange with the top charge before expansion.

Claims (10)

1. Method for separating a feed gas under pressure into a relatively volatile residual gas and a relatively less volatile fraction, the feed gas containing hydrocarbons, methane and ethane which together constitute a major proportion of said feed gas, where (a) the feed gas under pressure is cooled sufficiently to partially condensing the gas to form a pressurized liquid portion and a pressurized vapor portion, (b) at least some of the pressurized liquid portion is expanded in an expansion device to a lower pressure to form an expanded liquid portion, wherein part of the liquid portion part during the expansion step. evaporating so that the expanded liquid portion is cooled to a refrigeration temperature, whereby the remainder of said liquid portion remains as liquid, and (c) at least some of the expanded liquid portion is then processed in a fractionating column to separate the relatively less volatile fraction, characterized in that one (1) combines at least a portion of the pressurized liquid resulting from step (a) with a process stream having a bubble point below the bubble point of said pressurized liquid portion to thereby form a combined stream, (2) prior to combining the liquid part under pressure and the process stream according to step (1), cools either one or both of said liquid part or said process stream, or after said combination cools the combined stream, whereby a cooled combined stream is obtained which has a temperature below the bubble point of said liquid part under pressure, (3) supplies the cooled combined stream to the expansion device used in step (b ), wherein the combined stream is expanded to said lower pressure to form an expanded combined stream, and the cooling temperature obtained in expansion step (b) is reduced, (4) then feeding at least some of the expanded combined stream to said fractionating column at a first charge point , and (5) expands at least a part of said steam part under pressure in a working expansion machine with said lower pressure and supplies the thus obtained expanded part to the fractionation column at another charge point which is located lower on the column than said first charge point. 2. Method according to claim 1, characterized in that the process stream which has a bubble point below the bubble point for said liquid part obtained in step (a) is part of said steam part under pressure. 3. Method according to claim 1 or 2, characterized in that at least 25% of the steam portion under pressure obtained in step (a) is expanded to said lower pressure in step (5). 4. Method according to claim 1, 2 or 3, characterized in that the amount of said steam part under pressure which is expanded in step (5) is sufficient to reduce the risk of ice formation of carbon dioxide in the fractionation column. 5. Method according to claims 1-4, characterized in that the first charge point in step (4) is the top charge in the fractionation column. 6. Apparatus for carrying out the method according to claim 1, for separating a feed gas under pressure into a volatile residual gas and a relatively less volatile fraction, where the feed gas under pressure contains hydrocarbons, methane and ethane comprising the main part of said feed gas, where the device has : (a) cooling device (10,11,12,13,14 and 15) for cooling the supply gas under pressure sufficient to partially condense the gas thereby forming a liquid part under pressure and a vapor part under pressure, (b) an expansion device ( 155, 182,183) connected to said cooling device (a) for receiving the liquid part under pressure and for expanding it to a lower pressure whereby it is further cooled, (c) fractionation devices (19) for receiving said expanded liquid part from the expansion device ( b) and to separate the relatively less volatile fraction, characterized by (i) devices for receiving at least part of the liquid part und is pressure obtained from the cooling device (a) and for combining this part with a process stream having a bubble point below the bubble point of the liquid part under pressure to thereby form a combined stream, (ii) cooling devices (154,184,185) for cooling the liquid part under pressure, the process stream or the combined stream, in any combination, being sufficient for the combined stream to have a temperature below the bubble point of the pressurized liquid portion obtained from cooling device (a), (iii) devices connected to the expansion device (b ) for receiving the combined stream at a temperature below the bubble point of the liquid portion under pressure so that the combined stream is expanded in the expansion device (b) to said lower pressure and thereby forms an expanded combined stream, (iv) devices (157,186) connecting the expansion device (b) with the fractionation column for supplying at least part of the expanded combine t stream as a charge to the fractionation column at a first charge point, and (v) work expansion device (17) connected to the cooling device (a) for receiving at least some of the vapor portion under pressure and expanding this to said lower pressure, forming an expanded vapor portion , the working expansion device being further connected for supplying the expanded steam part to the fractionation column (19) at another charge point which is located lower on the column than the first charge point. 7. Apparatus according to claim 6, characterized in that the devices (i) for receiving the liquid part under pressure for combining this with a process stream that has a bubble point below the bubble point of the liquid part under pressure, is connected to receive part of the steam part under pressure obtained from cooling device (a), said steam part under pressure thereby providing the process stream which has a bubble point below the bubble point of the liquid part under pressure. 8. Apparatus according to claim 6 or 7, characterized in that the work expansion devices (v) are arranged to receive at least 25% of the steam portion under pressure obtained in step (a). 9. Apparatus according to claims 6-8, characterized in that the quantity of the steam part under pressure which is supplied to the working expansion devices (v) is sufficient to reduce the risk of carbon dioxide ice formation in the fractionation column. 10. Apparatus according to claims 6-9, characterized in that the first charge point associated with the devices (iv) represents the top charge of the fractionation column.