Technical Field

[0001] The present invention relates to a process for separating components of a sour natural gas.

Background Art

[0002] Many sources of natural gas contain sour gasses such as carbon dioxide and hydrogen sulphide. These sour gasses need to be removed before the natural gas (or components of the natural gas) can be used for many applications. However, separating carbon dioxide from natural gasses can be difficult because of the tendency for carbon dioxide to form an azeotropic mixture with ethane under typical distillation conditions (i.e. high pressure and low

temperature). Similarly, hydrogen sulphide can form an azeotropic mixture with ethane or propane under typical distillation conditions.

[0003] Many conventional techniques for breaking these azeotropes are undesirable because they are environmentally unsound (e.g. chemical absorption using amine solutions), capital intensive (e.g. adsorption using molecular sieves) and/or require complex, energy and resource intensive systems to operate (e.g. separation using absorption or extractive distillation techniques).

[0004] Azeotrope inhibiting agents (also referred to as entrainers in some azeotropic distillation systems) have been used in some separation processes. Addition of such an agent effectively prevents (or substantially reduces the extent of) azeotrope formation under the relevant conditions, which enables separation of components that would previously have been substantially inseparable. However, existing processes using azeotrope inhibiting agents typically require relatively large quantities of such agents, especially in extractive distillation techniques, which increases the cost and overall complexity of the systems.

Summary of Invention

[0005] In a first aspect, the present invention provides a process for separating components of a sour natural gas. The process comprises: operating a first distillation column into which the sour natural gas has been introduced at conditions whereby a methane enriched first overhead product and a first bottoms product which is substantially free of methane are produced; operating a second distillation column into which the first bottoms product has been introduced at conditions whereby a second overhead product which is substantially free of C3+ components (i.e. propane and heavier hydrocarbons) and a second bottoms product comprising C3+ components are produced; transferring the second overhead product to a third distillation column and adding an azeotrope inhibiting agent; and operating the third distillation column into which the second overhead product has been introduced at conditions whereby an ethane enriched third bottoms product is produced.

[0006] The present invention provides a distillation process via which methane, which forms a vast majority of many natural gases, is removed in a first step of the process. Advantageously, subsequent steps therefore typically involve a much smaller volume of gas, and therefore require less power for heating/cooling etc. and can be performed using smaller and less complicated processing equipment. Furthermore, as the methane and C3+ components of the natural gas are substantially removed in preliminary steps, the amount of azeotrope inhibiting agent required to prevent (or hinder) azeotrope formation between ethane and carbon dioxide or hydrogen sulphide is significantly reduced compared to conventional separation processes (the amount of the azeotrope inhibiting agent required will depend on the amount of CO2 in the gas feed, but may be as little as 10% of that required in other separation processes). Also advantageously, the present invention may be performed using only distillation columns for separation purposes, thereby avoiding the need for chemicals such as amines which have negative environmental effects and necessitate more complex processes and equipment.

[0007] Furthermore, some existing processes are not capable of separating methane, ethane and the C3+ components of a sour natural gas. For example, some processes may only be capable of producing mixtures of methane and ethane, which might require further processing, depending on the downstream application. In the process of the present invention, however, separate streams of relatively pure methane, ethane and C3+ components can be produced. In some embodiments of the present invention, described below, separate streams of relatively pure CO2 and H ₂ S can also be produced. [0008] The first distillation column is operated at conditions whereby a methane enriched first overhead product and a first bottoms product which is substantially free of methane are produced. The operating conditions and resultant products may be influenced by a number of factors, primarily the temperature and pressure conditions in the first distillation column, but other factors may be used to control the operating conditions and resultant products.

[0009] In some embodiments, for example, the operating conditions of the first distillation column may also be influenced by introducing a portion of the second bottoms product comprising C3+ components to the first distillation column at a location whereby sour species in the sour natural gas are prevented from freezing. In some embodiments, the location at which the sour natural gas is introduced into the distillation column may also influence the operating conditions of the first distillation column.

[0010] The second distillation column is operated at conditions whereby a second overhead product which is substantially free of C3+ components and a second bottoms product comprising C3+ components are produced. The operating conditions and resultant products may be influenced by a number of factors, primarily the temperature and pressure conditions in the second distillation column, but other factors may be used to control the operating conditions and resultant products.

[0011] The third distillation column is operated at conditions whereby an ethane enriched third bottoms product is produced. The operating conditions and resultant products may be influenced by a number of factors, primarily the temperature and pressure conditions in the third distillation column, but other factors may be used to control the operating conditions and resultant products.

[0012] In some embodiments, the azeotrope inhibitor agent also forms part of the third bottoms product. In such embodiments, the ethane enriched third bottoms product from the third distillation column may be transferred to a fourth distillation column, which is operated at conditions whereby the ethane and the azeotrope inhibitor agent are separated. That is, the third and fourth distillation columns may together provide a distillation system for separating a relatively pure stream of ethane from its azeotropic mixture with carbon dioxide (using the azeotrope inhibitor agent) and subsequently recovering the azeotrope inhibitor agent. In such embodiments, the azeotrope inhibitor agent separated from the ethane in the fourth distillation column may be recycled back to the third distillation column.

[0013] In some embodiments, products of the third distillation column other than the ethane enriched third bottoms product (i.e. an ethane-depleted third overhead product containing carbon dioxide and residual ethane) may be transferred from the third distillation column to a fifth distillation column and a second azeotrope inhibiting agent added. The fifth distillation column is operated at conditions whereby a carbon dioxide enriched product is produced.

[0014] In some embodiments, the carbon dioxide enriched product is separated together with the second azeotrope inhibiting agent. In such embodiments, the carbon dioxide enriched product from the fifth distillation column may be transferred to a sixth distillation column, which is operated at conditions whereby the carbon dioxide and the second azeotrope inhibitor agent are separated. That is, the fifth and sixth distillation columns may together provide a distillation system for separating a relatively pure stream of carbon dioxide from its azeotropic mixture with any residual ethane (using the second azeotrope inhibitor agent) and subsequently recovering the second azeotrope inhibitor agent. In such embodiments, the second azeotrope inhibitor agent separated from the carbon dioxide in the sixth distillation column may be recycled back to the fifth distillation column.

[0015] In some embodiments, the proportion of hydrogen sulphide in the sour natural gas may necessitate that it also be separated. Alternatively, it may be advantageous for other reasons to separate the hydrogen sulphide from the sour natural gas. In some embodiments, the hydrogen sulphide would either be separated with the second bottoms product comprising C3+

components or with the ethane enriched third bottoms product, depending on the composition of the sour natural gas and the operating conditions of the second and third distillation columns. In some embodiments, the hydrogen sulphide may be removed from the second overhead product before the second overhead product is introduced into the third distillation column. The second overhead product may, for example, be transferred to the third distillation column via a seventh distillation column, where a third azeotrope inhibiting agent is added and the seventh distillation column operated at conditions whereby an overhead product that is substantially free of hydrogen sulphide and the third azeotrope inhibiting agent is produced.

[0016] In some embodiments, a bottoms product of the seventh distillation column (primarily containing hydrogen sulphide and the third azeotrope inhibiting agent) is transferred to an eighth distillation column, the eighth distillation column being operated at conditions whereby hydrogen sulphide and the third azeotrope inhibitor agent are separated. That is, the seventh and eighth distillation columns may together provide a distillation system for separating a relatively pure stream of hydrogen sulphide from its azeotropic mixture with ethane (using the third azeotrope inhibitor agent) and subsequently recovering the third azeotrope inhibitor agent. In such embodiments, the third azeotrope inhibitor agent separated from the hydrogen sulphide in the eighth distillation column may be recycled back to the seventh distillation column. [0017] In a second aspect, the present invention provides a process for producing methane using the process of the first aspect of the present invention.

[0018] In a third aspect, the present invention provides a process for producing ethane using the process of the first aspect of the present invention.

[0019] In a fourth aspect, the present invention provides a process for producing propane and heavier hydrocarbons using the process of the first aspect of the present invention.

[0020] In a fifth aspect, the present invention provides a process for producing carbon dioxide using the process of the first aspect of the present invention.

[0021] In a sixth aspect, the present invention provides a process for producing hydrogen sulphide using the process of the first aspect of the present invention.

[0022] In a seventh aspect, the present invention provides methane, ethane, propane (and heavier hydrocarbons), carbon dioxide or hydrogen sulphide, produced using the process of the first aspect of the present invention.

Brief Description of the Drawings

[0023] The present invention will be described in further detail below with reference to the following figures, in which:

[0024] Figure 1 shows a process flow diagram in accordance with one embodiment of the present invention where components of a sour natural gas are separated;

[0025] Figure 2 shows a process flow diagram in accordance with another embodiment of the present invention where components of a sour natural gas are separated; and

[0026] Figure 3 shows a process flow diagram in accordance with yet another embodiment of the present invention where components of a sour natural gas are separated.

Description of Embodiments

[0027] As noted above, the present invention provides a process for separating components of a sour natural gas. The process comprises: operating a first distillation column into which the sour natural gas has been introduced at conditions whereby a methane enriched first overhead product and a first bottoms product which is substantially free of methane are produced; operating a second distillation column into which the first bottoms product has been introduced at conditions whereby a second overhead product which is substantially free of C3+ components (i.e. propane and heavier hydrocarbons) and a second bottoms product comprising C3+ components are produced; transferring the second overhead product to a third distillation column and adding an azeotrope inhibiting agent; and operating the third distillation column into which the second overhead product has been introduced at conditions whereby an ethane enriched third bottoms product is produced.

[0028] It will be understood by persons skilled in the art that, in the context of distillation separations, the use of the phrase "substantially free" does not preclude the presence of a residual amount of the recited component(s) in the relevant product. Indeed, it is typically not feasible using distillation techniques to completely separate components of a feed gas to obtain a completely pure product. Similarly, the term "enriched", in the context of the process of the present invention, will be understood as meaning that the relevant product contains a

significantly higher proportion of the recited component(s), but not precluding the presence of other components.

[0029] The present invention may be used to process a sour natural gas, that is, any natural gas containing one or more sour species. The composition of the gas stream may vary significantly, but the gas stream will generally contain methane, ethane, propane and higher hydrocarbons (C3+), water, as well as the sour species carbon dioxide and sometimes hydrogen sulphide.

[0030] Mole proportions of the components of typical natural gasses are set out below in Table

Table 1 - Typical composition of natural gas [0031] It should be noted that the proportion of carbon dioxide in some natural gasses can be as high as 60 or 70%. In such natural gasses, the proportions of the other components are relatively lower, but they will still tend to be in the same relative proportions to each other. The present invention is capable of processing natural gasses containing such high amounts of carbon dioxide.

[0032] The process of the present invention comprises operating a first distillation column into which the sour natural gas has been introduced at conditions whereby a methane enriched first overhead product and a first bottoms product which is substantially free of methane are produced. The operating conditions in the first distillation column may be influenced by a number of factors, the most important of which are the temperature and pressure conditions in the first distillation column.

[0033] The temperature and pressure conditions at which the first distillation column is operated will depend on the composition of the sour natural gas to be processed. The first distillation column may be operated at any temperature which (in combination with other factors which affect the conditions in the first distillation column) results in a methane enriched first overhead and a first bottoms product which is substantially free of methane being produced. Generally (depending on the composition of the natural gas and the operating pressure), the first distillation column may be operated with a top temperature in a range of about -90°C to about - 80°C, for example, at about, -80°C, -82°C, -85°C, -87°C, -or -90°C, and a bottom temperature in a range of about 15°C to about 50°C (e.g. 15°C to about 45°C or 40°C to about 50°C), for example, at about 15°C, 20°C, 25°C, 30°C, 35°C, 40°C, 42°C, 44°C, 46°C, 48°C or 50°C.

[0034] The first distillation column may be operated at any pressure which (in combination with other factors which affect the conditions in the first distillation column) results in a methane enriched first overhead and a first bottoms product which is substantially free of methane being produced. Generally, the first distillation column may be operated at a pressure of between about 40 and about 45 bar, for example, between about 43 and about 44 bar.

[0035] The first distillation column may, in some embodiments, contain a plurality of vapour- liquid contact devices such as trays or packing. The number of such contact devices required in the distillation column may, for example, depend on the required operating conditions and may be readily determined by a person skilled in the art. In some embodiments, the location of the distillation column at which the sour natural gas is introduced may also influence the operating conditions of the first distillation column. The first distillation column may, for example, contain 40-50 trays with the feed tray being located intermediate to the top section of the trays of the first distillation column.

[0036] In some embodiments, the operating conditions in the first distillation column may also be influenced by introducing a portion of the second bottoms product comprising C3+ components to the first distillation column at a location whereby sour species in the sour natural gas are prevented from freezing. Sour species such as carbon dioxide can sometimes freeze under the conditions encountered in the first distillation column, which is undesirable because it can prevent efficient separation and, in some circumstances, stop the process from operating. This may be a greater problem in natural gasses containing higher proportions of sour gasses (especially carbon dioxide), but can occur even when the proportion of carbon dioxide is relatively low. In such circumstances, recycling some (e.g. 5-10% of the total amount of the C3+ components in the sour natural gas) of the C3+ components back into the first distillation column can substantially prevent this from occurring. Typically, the C3+ components are recycled back into the first distillation column at a location above the input of the sour natural gas (e.g. at an allocation in the top 10 trays), but this may not be the case for all natural gasses. It is within the ability of a person skilled in the art to determine whether such recycling would be required for a given source of sour natural gas and, if so, an appropriate place to add the C3+ components into to the first distillation column.

[0037] Removing methane in the first step of the process of the present invention significantly reduces the volume of gas that needs to be processed in subsequent steps, thus simplifying those steps and reducing their energy requirements (as well as enabling a smaller amount of azeotrope inhibitor agent to be used, compared to some prior art processes, making the process cheaper, potentially safer and more environmentally friendly). Further, as methane is recovered in the first step, the ethane recovered in the later step (discussed below) is relatively pure (some prior art processes result in streams containing a mixture of methane and ethane).

[0038] The process of the present invention also comprises operating a second distillation column into which the first bottoms product has been introduced at conditions whereby a second overhead product which is substantially free of C3+ components (i.e. propane and heavier hydrocarbons) and a second bottoms product comprising C3+ components are produced. The operating conditions in the second distillation column may be influenced by a number of factors, the most important of which are the temperature and pressure conditions in the second distillation column. [0039] The temperature and pressure conditions at which the second distillation column is operated will again depend on the composition of the sour natural gas to be processed. The second distillation column may be operated at any temperature which (in combination with other factors which affect the conditions in the second distillation column) results in the production of a second overhead which is substantially free of C3+ components and a second bottoms product comprising C3+ components. Generally, the second distillation column may be operated with a top temperature in a range of about -5°C to about -25°C (e.g. about -15°C to about -25°C), for example, at about -5°C, -10°C, -12°C, -15°C, -18°C, -20°C,-22°C (e.g. - 22.6°C) or -25°C, and a bottom temperature in a range of about 85°C to about 110°C (e.g. about 85°C to about 95°C), for example, at about 85°C, 87°C, 90°C, 92°C, 93°C, 95°C, 98°C, 100°C, 105°C or 110°C.

[0040] The second distillation column may be operated at any pressure which (in combination with other factors which affect the conditions in the first distillation column) results in a second overhead which is substantially free of C3+ components and a second bottoms product comprising C3+ components. Generally, the second distillation column may be operated at a pressure of between about 20 and about 34 bar (e.g. about 20 and about 24 bar), for example, between about 21 and about 22 bar or about 20, 22, 24, 26, 28, 30, 32 or 34 bar.

[0041] The second distillation column may, in some embodiments, contain a plurality of vapour-liquid contact devices such as trays or packing. The number of such contact devices required in the distillation column may, for example, depend on the required operating conditions and may be readily determined by a person skilled in the art. In some embodiments, the location of the distillation column at which the methane-depleted sour natural gas (i.e. the first bottoms product) is introduced may also influence the operating conditions of the second distillation column. The second distillation column may, for example, contain 40-60 (e.g. 40- 50) trays with the feed tray being located intermediate a bottom and top tray of the second distillation column.

[0042] As noted above, as methane is removed from the sour natural gas in the first distillation column, up to about 70% of the volume of sour natural gas might have been removed, and the capacities of the second and downstream distillation columns can be reduced accordingly.

[0043] The process of the present invention also comprises transferring the second overhead product (depleted in methane and C3+ components) to a third distillation column, adding an azeotrope inhibiting agent, and operating the third distillation column into which the second overhead has been introduced at conditions whereby an ethane enriched third bottoms product is produced. The operating conditions in the third distillation column may be influenced by a number of factors, the most important of which are the temperature and pressure conditions in the third distillation column.

[0044] The temperature and pressure conditions in the third distillation column will depend on the composition of the remaining components of the sour natural gas to be processed. The third distillation column may be operated at any temperature which (in combination with other factors which affect the conditions in the third distillation column) results in an ethane enriched third bottoms product being produced (typically, the azeotrope inhibitor agent will also be present in the ethane enriched third bottoms product). Generally, the third distillation column may be operated with a top temperature in a range of about 5°C to about -35°C (e.g. about-5°C to about -10°C), for example, at about -5°C, -7°C, -9°C, -10°C, -15°C, -20°C, -25°C, -30°C or -35°C, and a bottom temperature in a range of about 20°C to about 55°C (e.g. about 50°C to about 55°C), for example, at about 20°C, 25°C, 30°C, 35°C, 40°C, 45°C, 50°C, 52°C, 52.5°C, 54°C or 55°C.

[0045] The third distillation column may be operated at any pressure which (in combination with other factors which affect the conditions in the third distillation column) results in an ethane enriched third bottoms product being produced. Generally, the third distillation column may be operated at a pressure of between about 15 and about 50 bar (e.g. about 45 and about 50 bar), for example, between about 47 and about 49 bar or about 15, 20, 25, 30, 35, 40, 45 or 50 bar.

[0046] The third distillation column may, in some embodiments, contain a plurality of vapour- liquid contact devices such as trays or packing. The number of such contact devices required in the distillation column may, for example, depend on the required operating conditions and may be readily determined by a person skilled in the art. In some embodiments, the location of the distillation column at which the second overhead is introduced may also influence the operating conditions of the third distillation column. The third distillation column may, for example, contain 40-60 (e.g. 40-50) trays with the feed tray being located intermediate a bottom and top tray of the third distillation column.

[0047] The azeotrope inhibiting agent prevents formation of an azeo tropic mixture (e.g. an ethane-carbon dioxide azeotrope or an ethane-hydrogen sulphide azeotrope) or breaks an azeotrope already formed under the conditions experienced in the third distillation column. The formation of such an azeotrope would prevent (or greatly complicate) effective separation of the components of the azeotropic mixture using distillation techniques. However, the presence of the azeotrope inhibiting agent enables these components to be separated. Whilst the azeotrope between the ethane and sour gas gets broken, it is typically replaced by an azeotrope between the azeotrope inhibiting agent and either the sour gas or ethane. However, once the sour gas or ethane is separated via distillation from the azeotropic mixture of the other of the sour gas/ethane and the azeotrope inhibiting agent, the newly formed azeotropic mixture itself can be separated, for example in another distillation column (under different conditions).

[0048] Any species which is capable of inhibiting or breaking an azeotrope between

components of the sour natural gas present in the third distillation column may be used in the process of the present invention. Suitable azeotrope inhibiting agents for use in the third distillation column can be determined by considering the relative boiling points or volatilities of the components of the azeotrope present under the conditions within the third distillation column. Suitable azeotrope inhibiting agents for breaking azeotropes between ethane and carbon dioxide which form under the conditions in the third distillation column include, for example, ethanol, acetone, sulfolane and mixtures thereof.

[0049] The amount of the azeotrope inhibiting agent that needs to be added will depend on factors such as the volume of gas in the second overhead, the relative proportions of the components in the second overhead and the operating conditions in the third distillation column. The amount of the azeotrope inhibiting agent could readily be determined for any given process by a person skilled in the art. The amount of the azeotrope inhibiting agent added may, for example, be between about 5 and 10 mol% (with respect to the total feed to the third distillation column).

[0050] The azeotrope inhibiting agent may be added at any convenient time. For example, the azeotrope inhibiting agent may be added to or mixed with the second overhead before it is transferred to (e.g. before it is introduced into or it enters) the third distillation column (e.g. to ensure good mixing occurs so that the azeotrope either does not form inside the column or, if it does form, it is broken relatively quickly). In contrast, entrainers used in extractive distillations are often added to the top of the tower while the feed is added to the bottom or the middle to allow the entrainer to fall through the feed in the tower and hence change the azeotropic properties of the mixture and also extract the required components from it. However, the azeotrope inhibiting agent could, if appropriate, be added in the process of the present invention by injecting it into the third distillation column, either on the same tray or a different tray. In embodiments (discussed below) where hydrogen sulphide is removed from the second overhead product before it is introduced into the third distillation column, the azeotrope inhibiting agent would be added after the hydrogen sulphide removal step had been carried out. [0051] The carbon dioxide in the sour natural gas will almost always be separated with the ethane (because of its tendency to form azeotropes with ethane). When the sour natural gas also contains hydrogen sulphide, this would either typically be separated with the second bottoms product comprising C3+ components or with the ethane enriched product, depending on the composition of the sour natural gas and the operating conditions of the second and third distillation columns. In such cases, it is within the ability of those skilled in the art to determine if and when the hydrogen sulphide should be removed from any given sour natural gas and operating conditions.

[0052] In embodiments where the hydrogen sulphide is separated with the second bottoms product comprising C3+ components, the process of the present invention may include a further step in which the hydrogen sulphide is separated from the C3+ components. It is within the ability of those skilled in the art to determine the most effective technique to perform such a separation, using any state of the art H ₂ S removal system. In embodiments where the hydrogen sulphide is separated with the ethane enriched product, the process of the present invention may include a further step in which the hydrogen sulphide is separated from the ethane. Again, it is within the ability of those skilled in the art to determine the most effective technique to perform such a separation.

[0053] As will be described in further detail below, a convenient time to remove hydrogen sulphide is to remove it from the second overhead product before it is introduced into the third distillation column. The second overhead product is depleted in methane and C3+ components and hence has a significantly reduced volume.

[0054] The process described above results in relatively clean streams of methane and C3+ components and, depending on the proportion of sour gasses in the natural gas, ethane.

However, in some embodiments of the present invention, it may be desirable to improve the purity of the ethane product, recover the azeo trope inhibiting agent and/or recover useable carbon dioxide and/or hydrogen sulphide as well. In such embodiments, one or more of the additional distillation systems described below may be used in the process.

[0055] Typically, the azeotrope inhibitor agent also forms part of the third bottoms product. Accordingly, in some embodiments, the ethane enriched third bottoms product from the third distillation column may be transferred to a fourth distillation column, which is operated at conditions whereby the ethane and the azeotrope inhibitor agent are separated. The fourth distillation column can be used to produce a stream of ethane (typically as the overhead product of the fourth distillation column) that is substantially free of the azeotrope inhibiting agent (typically formed as the bottoms product) and other contaminants (e.g. carbon dioxide, which was separated from the ethane in the third distillation column). The azeotrope inhibitor agent separated from the ethane in the fourth distillation column may also be recycled back to the third distillation column in order to even further improve the efficiency of the process of the present invention.

[0056] The operating conditions in the fourth distillation column may be influenced by a number of factors, the most important of which are the temperature and pressure conditions in the fourth distillation column. The temperature and pressure conditions in the fourth distillation column will depend on the composition of the ethane enriched product from the third distillation column.

[0057] The fourth distillation column may be operated at any temperature which (in

combination with other factors which affect the conditions in the fourth distillation column) results in substantial separation of ethane and the azeotrope inhibitor. Generally, the fourth distillation column may be operated with a top temperature in a range of about -50°C to about - 55°C, for example at about -50°C, -52°C or -55°C, and a bottom temperature in a range of about 130°C to about 140°C, for example, at about 130°C, 133°C, 135°C, 137°C or 140 °C.

[0058] The fourth distillation column may be operated at any pressure which (in combination with other factors which affect the conditions in the fourth distillation column) results in substantial separation of ethane and the azeotrope inhibitor. Generally, the fourth distillation column may be operated at a pressure of between about 5 and about 9 bar (e.g. about 5 and about 7 bar), for example, at about 5, 6, 7, 8 or 9 bar.

[0059] The fourth distillation column may, in some embodiments, contain a plurality of vapour- liquid contact devices such as trays or packing. The number of such contact devices required in the distillation column may, for example, depend on the required operating conditions and may be readily determined by a person skilled in the art. In some embodiments, the location of the distillation column at which the ethane enriched product from the third distillation column is introduced may also influence the operating conditions of the fourth distillation column. The fourth distillation column may, for example, contain 25-35 trays with the feed tray being located intermediate a bottom and top tray of the fourth distillation column.

[0060] In some embodiments, products from the third distillation column other than the ethane enriched third bottoms product (i.e. an ethane depleted third overhead product, including C0 ₂ and residual ethane) may be transferred to a fifth distillation column and a second azeotrope inhibiting agent added, the fifth distillation column being operated at conditions whereby the ethane-CC azeotrope is broken and a carbon dioxide enriched product is produced. Typically, the carbon dioxide enriched product is separated together with the second azeotrope inhibitor agent as bottom product.

[0061] The operating conditions in the fifth distillation column may be influenced by a number of factors, the most important of which are the temperature and pressure conditions in the fifth distillation column. The temperature and pressure conditions in the fifth distillation column will depend on the composition of product received from the third distillation column.

[0062] The fifth distillation column may be operated at any temperature which (in combination with other factors which affect the conditions in the fifth distillation column) results in production of a carbon dioxide enriched product. Generally, the fifth distillation column may be operated with a top temperature in a range of about -15°C to about -55°C (e.g. -25°C to about - 55°C or about -15°C to about -25°C), for example, at about -15°C, -17°C, -20°C, -23°C, - 25°C, -30°C, -35°C, -40°C, -45°C, -50°C or -55°C and a bottom temperature in a range of about 0°C to about 32°C (e.g. about 0°C to about 5°C), for example, at about 0°C, 1.5°C, 3°C, 5°C, 10°C, 15°C, 20°C, 25°C, 30°C or 32°C.

[0063] The fifth distillation column may be operated at any pressure which (in combination with other factors which affect the conditions in the fifth distillation column) results in production of a carbon dioxide enriched product. Generally, the fifth distillation column may be operated at a pressure of between about 29 and about 35 bar (e.g. about 29 and about 32 bar), for example, at about 29, 30, 31, 32, 33, 34 or 35 bar.

[0064] The fifth distillation column may, in some embodiments, contain a plurality of vapour- liquid contact devices such as trays or packing. The number of such contact devices required in the distillation column may, for example, depend on the required operating conditions and may be readily determined by a person skilled in the art. In some embodiments, the location of the distillation column at which the product received from the third distillation column is introduced may also influence the operating conditions of the fifth distillation column. The fifth distillation column may, for example, contain 30-40 trays with the feed tray being located intermediate a bottom and top tray of the fifth distillation column.

[0065] The second azeotrope inhibiting agent acts to substantially prevent formation of (or break) an azeotropic mixture (e.g. an azeotrope between carbon dioxide and the residual ethane) under the conditions experienced in the fifth distillation column. Such an azeotrope would prevent effective separation of the C0 ₂ and residual ethane in the distillation column. The presence of the second azeotrope inhibiting agent in the column enables a vast majority of the carbon dioxide to be recovered.

[0066] Typically, the second azeotrope inhibitor agent will be different to the azeotrope inhibitor agent (i.e. that used to break the carbon dioxide/ethane azeotrope in the third column) because each agent changes the equilibrium diagram differently, and due to the different relative proportions of ethane and carbon dioxide in the gasses in each distillation column. In some cases, however, it is possible that the second azeotrope inhibitor agent will be the same as the azeotrope inhibitor agent.

[0067] Any azeotrope inhibiting agent which is capable of inhibiting an azeotrope between components present in the fifth distillation column may be used in the process of the present invention. Suitable azeotrope inhibiting agents include, for example, propargylOl (propargyl alcohol), ammonia and mixtures thereof.

[0068] The amount of the second azeotrope inhibiting agent that needs to be added will depend on factors such as the volume of gas in the fifth distillation column, the relative proportions of the components in the gas and the operating conditions in the fifth distillation column. The amount of the second azeotrope inhibiting agent could readily be determined for any given process by a person skilled in the art. The amount of the second azeotrope inhibiting agent added may, for example, be between about 0.4 and 5 mol% (e.g. about 1, 2, 3, 4 or 5 mol%) with respect to the total feed of the fifth distillation column.

[0069] The second azeotrope inhibiting agent may be added at any convenient time. For example, the second azeotrope inhibiting agent may be added to or mixed with the overhead from the third distillation column before it is transferred to (e.g. before it is introduced into or it enters) the fifth distillation column (e.g. to ensure good mixing occurs). However, the second azeotrope inhibiting agent could also be added by injecting it directly into the fifth distillation column, either at the same tray or a different tray.

[0070] In some embodiments, products of the fifth distillation column other than the carbon dioxide enriched product (e.g. ethane and a small proportion of the carbon dioxide which entered the column) may be recycled back to the third distillation column in order to reduce potential losses of valuable products (e.g. small amounts of ethane and carbon dioxide).

[0071] Typically, the carbon dioxide enriched product is separated together with the second azeotrope inhibitor agent as bottoms product of the fifth distillation column. In some embodiments, the carbon dioxide enriched product from the fifth distillation column is transferred to a sixth distillation column, which is operated at conditions whereby the carbon dioxide and the second azeotrope inhibitor agent are separated. The sixth distillation column can be used to produce a stream of carbon dioxide (typically as the overhead product of the sixth distillation column) that is substantially free of the second azeotrope inhibiting agent (typically formed as the bottoms product). The second azeotrope inhibitor agent separated from the carbon dioxide in the sixth distillation column may be recycled back to the fifth distillation column in order to even further improve the efficiency of the process of the present invention.

[0072] The operating conditions in the sixth distillation column may be influenced by a number of factors, the most important of which are the temperature and pressure conditions in the sixth distillation column. The temperature and pressure conditions in the sixth distillation column will depend on the composition of the carbon dioxide enriched product from the fifth distillation column.

[0073] The sixth distillation column may be operated at any temperature which (in combination with other factors which affect the conditions in the sixth distillation column) results in substantial separation of carbon dioxide and the second azeotrope inhibitor. Generally, the sixth distillation column may be operated with a top temperature in a range of about -40°C to about 5°C (e.g. about -40°C to about 0°C or about 0°C to about 5°C), for example, at about -40°C, - 35°C, -30°C, -25°C, -20°C, -15°C, -10°C, -5°C, 0°C, 1°C, 2°C, 3°C, 4°C or 5°C, and a bottom temperature in a range of about 137°C to about 180°C (e.g. about 170°C to about 180°C), for example, at about 137°C, 140°C, 145°C, 150°C, 155°C, 160°C, 165°C, 170°C, 172°C, 174°C, 176°C, 178°C or 180°C.

[0074] The sixth distillation column may be operated at any pressure which (in combination with other factors which affect the conditions in the sixth distillation column) results in substantial separation of carbon dioxide and the second azeotrope inhibitor. Generally, the sixth distillation column may be operated at a pressure of between about 5 and about 7 bar, for example, at about 6 bar.

[0075] The sixth distillation column may, in some embodiments, contain a plurality of vapour- liquid contact devices such as trays or packing. The number of such contact devices required in the distillation column may, for example, depend on the required operating conditions and may be readily determined by a person skilled in the art. In some embodiments, the location of the distillation column at which the carbon dioxide enriched product from the fifth distillation column is introduced may also influence the operating conditions of the sixth distillation column. The sixth distillation column may, for example, contain 25-35 trays with the feed tray being located intermediate a bottom and top tray of the sixth distillation column. [0076] As noted above, sour natural gasses may contain hydrogen sulphide in addition to (or instead of) carbon dioxide. In some embodiments, the proportion of hydrogen sulphide in the sour natural gas may necessitate that it be separated (as hydrogen sulphide is one of the primary sour gas contaminants, its removal from natural gas is typically vital). Alternatively, it may be advantageous for other reasons to separate any hydrogen sulphide from the sour natural gas.

[0077] In some embodiments, the hydrogen sulphide may be removed from the second overhead product before the second overhead product is introduced into the third distillation column. The second overhead product may, for example, be transferred to the third distillation column via a seventh distillation column, where a third azeotrope inhibiting agent is added and the seventh distillation column operated at conditions whereby a product that is substantially free of hydrogen sulphide and the third azeotrope inhibiting agent (i.e. comprising primarily ethane and carbon dioxide) is produced. Typically, the product that is substantially free of hydrogen sulphide and the third azeotrope inhibiting agent is produced as the overhead of the seventh distillation column, with hydrogen sulphide and the third azeotrope inhibiting agent being the bottom product. The product of the seventh distillation column that is substantially free of hydrogen sulphide and the third azeotrope inhibiting agent may be transferred onto the third distillation column, as discussed above.

[0078] The operating conditions in the seventh distillation column may be influenced by a number of factors, the most important of which are the temperature and pressure conditions in the seventh distillation column. The temperature and pressure conditions in the seventh distillation column will depend on the composition of the second overhead product (which will include ethane, hydrogen sulphide and most likely carbon dioxide).

[0079] The seventh distillation column may be operated at any temperature which (in combination with other factors which affect the conditions in the seventh distillation column) results in production of a product that is substantially free of hydrogen sulphide and the third azeotrope inhibiting agent. Generally, the seventh distillation column may be operated with a top temperature in a range of about -17°C to about -25°C, for example, at about -17°C, -18°C, - 19°C, -20°C, -21°C, -22°C, -23°C, -24°C, or -25°C, and a bottom temperature in a range of about 55°C to about 98°C, for example, at about 55°C, 60°C, 65°C, 70°C, 75°C, 80°C, 85°C, 90°C, 95°C or 98°C.

[0080] The seventh distillation column may be operated at any pressure which (in combination with other factors which affect the conditions in the seventh distillation column) results in production of a product that is substantially free of hydrogen sulphide and the third azeotrope inhibiting agent. Generally, the seventh distillation column may be operated at a pressure of between about 20 and about 25 bar, for example, at about 20, 21, 22, 23, 24 or 25 bar.

[0081] The seventh distillation column may, in some embodiments, contain a plurality of vapour-liquid contact devices such as trays or packing. The number of such contact devices required in the distillation column may, for example, depend on the operating conditions and may be readily determined by a person skilled in the art. In some embodiments, the location in the distillation column at which the second overhead product is introduced may also influence the operating conditions of the seventh distillation column. The seventh distillation column may, for example, contain 50-60 trays with the feed tray being located intermediate a bottom and top tray of the seventh distillation column.

[0082] The third azeotrope inhibiting agent acts to substantially prevent formation of an azeotropic mixture (e.g. an azeotrope between hydrogen sulphide and ethane) under the conditions experienced in the seventh distillation column. The formation of such an azeotrope would prevent effective separation of hydrogen sulphide and ethane in the distillation column. The presence of the third azeotrope inhibiting agent in the column enables a vast majority of the ethane to be separated from the mixture.

[0083] Typically, the third azeotrope inhibitor agent will be different to the azeotrope inhibitor agent (i.e. that used to break the carbon dioxide/ethane azeotrope in the third column) and second azeotrope inhibitor agent (i.e. that used to break the carbon dioxide/residual ethane azeotrope in the fifth column) because, as noted above, each agent changes the equilibrium diagram differently and is being used in gas mixtures containing different relative proportions of ethane, carbon dioxide and hydrogen sulphide. In some cases, however, it is possible that the third azeotrope inhibitor agent will be the same as the azeotrope inhibitor agent or second azeotrope inhibitor agent.

[0084] Any azeotrope inhibiting agent which is capable of inhibiting an azeotrope between components present in the seventh distillation column may be used in the process of the present invention. A particular azeotrope inhibiting agent the inventors have found to be suitable is ethanol.

[0085] The amount of the third azeotrope inhibiting agent that needs to be added will depend on factors such as the volume of gas in the seventh distillation column, the relative proportions of the components in the gas and the operating conditions in the seventh distillation column. The amount of the third azeotrope inhibiting agent could readily be determined for any given process by a person skilled in the art. The amount of the third azeotrope inhibiting agent added may, for example, be between about 4 and 10 mol% (e.g. about 4, 5, 6, 7, 8, 9 or 10 mol%) with respect to the total feed of the seventh distillation column.

[0086] The third azeotrope inhibiting agent may be added at any convenient time. For example, the third azeotrope inhibiting agent may be added to or mixed with the second overhead product before it is transferred to (e.g. before it is introduced into or it enters) the seventh distillation column (e.g. to ensure good mixing occurs). However, the third azeotrope inhibiting agent could also be added by injecting it directly into the seventh distillation column, either at the same tray or a different tray.

[0087] Typically, hydrogen sulphide and the third azeotrope inhibiting agent are the bottom products of the seventh distillation column. In some embodiments, the hydrogen sulphide and third azeotrope inhibiting agent bottom products may be transferred to an eighth distillation column, which is operated at conditions whereby the hydrogen sulphide and third azeotrope inhibitor agent are separated. The eighth distillation column can be used to produce a stream of hydrogen sulphide (typically as the overhead product of the eighth distillation column) that is substantially free of the third azeotrope inhibiting agent (typically formed as the bottoms product). The third azeotrope inhibitor agent separated from the hydrogen sulphide in the eighth distillation column may be recycled back to the seventh distillation column in order to even further improve the efficiency of the process of the present invention.

[0088] The operating conditions in the eighth distillation column may be influenced by a number of factors, the most important of which are the temperature and pressure conditions in the eighth distillation column. The temperature and pressure conditions in the eighth distillation column will depend on the composition of the bottoms product from the seventh distillation column.

[0089] The eighth distillation column may be operated at any temperature which (in

combination with other factors which affect the conditions in the eighth distillation column) results in the substantial separation of hydrogen sulphide and the third azeotrope inhibitor. Generally, the eighth distillation column may be operated with a top temperature in a range of about -14°C to about -31°C, for example, at about -14°C, -19°C, -20°C, -25°C, -29°C or -31°C, and a bottom temperature in a range of about 125°C to about 142°C, for example, at about 125°C, 130°C, 135°C, 140°C or 142°C.

[0090] The eighth distillation column may be operated at any pressure which (in combination with other factors which affect the conditions in the eighth distillation column) results in substantial separation of hydrogen sulphide and the third azeotrope inhibitor. Generally, the eighth distillation column may be operated at a pressure of between about 4 and about 7 bar, for example, at about 4, 5,6 or 7 bar.

[0091] The eighth distillation column may, in some embodiments, contain a plurality of vapour- liquid contact devices such as trays or packing. The number of such contact devices required in the distillation column may, for example, depend on the required operating conditions and may be readily determined by a person skilled in the art. In some embodiments, the location in the distillation column at which the bottoms product from the seventh distillation column is introduced may also influence the operating conditions of the eighth distillation column. The eighth distillation column may, for example, contain 10-20 trays with the feed tray being located intermediate a bottom and top tray of the eighth distillation column.

[0092] Specific embodiments of the present invention will now be described with reference to the process flow diagrams shown in the accompanying drawings. These process flow diagrams were developed using the HYSYS ^® software package sold by Aspen Tech. The configurations of distillation columns used for separating natural gas components described below were verified by HYSYS ^® simulation in the steady state mode.

[0093] Figure 1 shows a process flow diagram for a process 10 in accordance with one embodiment of the present invention where components of a sour natural gas are separated. The sour gas feed 12 includes N ₂ , CI, C2, C0 ₂ and C3+ gases. The feed gas 12 enters tower one 14 where the N ₂ and CI gases are separated by distillation as top products, while the other gases (i.e. C2, C0 ₂ and C3+) leave tower one 14 at the bottom of the column.

[0094] The mixture of gases C2, C0 ₂ and C3+ then enter tower two 16, where C3+ components are separated as bottom products from C2 and C0 ₂ as top products. If conditions require it, a portion (e.g. 10%) of the C3+ components can be recycled back into tower one 14 in order to prevent freezing of sour species (not shown in Figure 1). The mixture of gases of C2 and C0 ₂ then enter tower three 18, where an azeotropic inhibitor (e.g. ethanol) is added at 5-10 mol % to the feed gas rate to this tower 18 to remove the azeotropic behaviour between the C2 and C0 ₂ gases. In tower three 18, the enriched C2 stream is separated together with the azeotropic inhibitor as the bottom product, while the C0 ₂ and residual C2 leave as top products.

[0095] Should it be desired to recover the azeotrope inhibitor for re-use in tower three 18, the C2 gas may be separated from the azeotrope inhibitor in tower four 20. The (now relatively pure) C2 gas is the top products of tower four 20, with the azeotrope inhibitor being recovered as the bottom product and subsequently recycled back to the feed of tower three 18. [0096] The mixture of residual C2 and CO2 top products from tower three 18 then enter tower five 22, where a second azeotropic inhibitor (e.g. propargylOl) is added at 0.4-5 mol % to the feed gas rate to this tower to remove the azeotropic behaviour between the CO2 and residual C2 gases. In tower five 22, the enriched CO2 stream is separated together with the second azeotrope inhibitor as the bottom product, while the residual C2 and a small proportion of the CO2 gases leave as top products.

[0097] Should it be desired to recover the second azeotrope inhibitor from the bottom product of tower five 22 for re-use, the CO2 gas may be separated from the second azeotrope inhibitor in tower six 24. The CC gas is the top product of tower six 24, with the azeotrope inhibitor being recovered as the bottom product and subsequently recycled back to the feed of tower five 22.

[0098] Advantageously, the volume of the feed gas to towers two 16 and three 18 is relatively small due to separation of CI in tower one 14. As such, the amount of azeotrope inhibiting agent required is correspondingly relatively small. The relatively smaller amount of gas to be treated post-tower one 14 also enables much simpler (and smaller) subsequent towers to be used (e.g. towers capable of being operated at much lower design pressures, requiring less thicknesses). Further, the amount of energy required for processing, handling and recycling of components may be significantly less compared with that required for existing processes.

[0099] Furthermore, this process results in the C2 component of the sour gas feed 12 being separated from the C 1 component, thereby providing separate and relatively pure streams of these components (which is not possible in many prior art separation processes).

[0100] Figure 2 shows a process flow diagram for a process 50 in accordance with another embodiment of the present invention where components of a sour natural gas are separated. Process 50 is similar to process 10, with the exception that a distillation system for removing H2S is also present. The sour gas feed 52 includes N2, CI, C2, CO2, H2S and C3+ gases. The feed gas 52 enters tower one 54, where the N2 and CI gases are separated by distillation as top products, while the other gases (i.e. C2, CO2, H2S and C3+) leave tower one 54 at the bottom of the column.

[0101] The mixture of gases C2, CO2, H2S and C3+ then enter tower two 56, where C3+ components are separated as bottom products from C2, H2S and CO2 as top products. If conditions require it, a portion (e.g. 10%) of the C3+ components can be recycled back into tower one 54 in order to prevent freezing of sour species. Typically, the C3+ components would be introduced back into tower one 54 at a level above that of the feed 52. [0102] The mixture of gases of C2, H ₂ S and C0 ₂ then pass through hydrogen sulphide distillation system 57 where, as will be described below, H ₂ S is removed before the C2 and C0 ₂ carry onto enter tower three 58. An azeotropic inhibitor (e.g. ethanol) is added at 5-10 mol % to the feed gas rate to tower three 58 to remove the azeotropic behaviour between the C2 and C0 ₂ gases. In tower three 58, an enriched C2 stream is separated together with the azeotropic inhibitor as the bottom product, while the C0 ₂ and residual C2 leave as top products.

[0103] Should it be desired to recover the azeotrope inhibitor for re-use in tower three 58, the C2 gas may be separated from the azeotrope inhibitor in tower four 60. The C2 gas is the top product of tower four 60, with the azeotrope inhibitor being recovered as the bottom product and subsequently recycled back to the feed of tower three 58.

[0104] The mixture of C0 ₂ and residual C2 from tower three 58 flow onto tower five 62, where a second azeotropic inhibitor (e.g. propargylOl) is added at 0.4-5 mol % to the feed gas rate to this tower to remove the azeotropic behaviour between the C0 ₂ and residual C2 gases. In tower five 62, an enriched C0 ₂ stream is separated together with the second azeotrope inhibitor as the bottom product, while the residual C2 and a small proportion of the C0 ₂ gases leave as top products.

[0105] Should it be desired to recover the second azeotrope inhibitor for re-use in tower five 62 (e.g. to improve the overall efficiency of the process or improve the purity of the C0 ₂ ), the C0 ₂ gas may be separated from the second azeotrope inhibitor in tower six 64. The C0 ₂ gas is the top product of tower six 64, with the second azeotrope inhibitor being recovered as the bottom product and subsequently recycled back to the feed of tower five 62.

[0106] As noted above, the mixture of gases of C2, H ₂ S and C0 ₂ from tower two 56 pass through hydrogen sulphide distillation system 57, where the H ₂ S is removed. Hydrogen sulphide distillation system 57 includes tower seven 66 where, with the aid of a third azeotrope inhibitor agent, the H ₂ S is separated from the ethane (and C0 ₂ ). The overhead from tower seven includes C2, and C0 ₂ , which flow onto tower three 58. The bottoms product includes H ₂ S and the third azeotrope inhibitor, which can be separated (should it be desired to recover the third azeotrope inhibitor for re-use in tower seven 66) in tower eight 68. The H ₂ S gas is the top product of tower eight 68, with the third azeotrope inhibitor being recovered as the bottom product and subsequently recycled back to the feed of tower seven 66.

[0107] Referring now to Figure 3, shown is a process flow diagram for a process 100 in accordance with another embodiment of the present invention, where components of a sour natural gas are separated. The sour gas feed 102 includes N ₂ , CI, C2, C0 ₂ and C3+ gases, and has the composition set out below in Table 2.

Table 2 - Material balance of feed gas 102 (mph = kgmoles per hour)

[0108] The gas feed 102 is passed through a chiller 104 (as well as heat exchangers 116 and 160, described below) where the mixture is cooled to about -40 °C, and then introduced under a pressure of about 84 bar to a first distillation tower 106. The feed 102 enters the first distillation tower 106 through line 108 at a feed tray at about two thirds of the height of the tower 106.

[0109] The first distillation tower 106 contains a plurality of vapour-liquid contact devices such as trays or packing. The number of such contact devices required in the distillation column depends on the required operating conditions and may be readily determined by a person skilled in the art. In this particular embodiment, the first distillation tower 106 contains 40-50 trays with line 108 feeding into a tray located intermediate a bottom and top tray of the first distillation tower 106.

[0110] The first distillation tower 106 into which the sour natural gas 102 has been introduced is operated at conditions whereby a methane enriched overheard product stream substantially free of other feed gas components (except for nitrogen) and a bottoms product which is substantially free of methane are produced (see Table 3). In this embodiment, the temperature at the bottom of the first distillation tower 106 is about 44°C, and the temperature at the top about -85 °C.

Upon entry to the first distillation tower 106, methane contained in the feed 102 is distilled off as a vapour stream (along with any nitrogen present). The methane vapour stream exits the first distillation tower 106 through line 110 and enters an overhead vapour-liquid separator 112, where it is partially condensed. Resulting overhead vapour product, which is rich in methane, is separated in separator 112 and leaves the process under pressure through line 114, via a heat exchanger 116 to cool down the feed 102. Any resulting condensate from the separator 112 is pumped back to the top tray of the first distillation tower 106.

[0111] First distillation tower 106 also has a reboiler 113, which heats the liquid at the bottom of the column to generate vapours to drive the distillation separation.

[0112] Acid gas solids are prevented from freezing in first distillation tower 106 at low temperate by recycling (via line 120) about 10% of the C3+ stream from the bottom of the second distillation tower 118 to tray number 4 from the top of first distillation tower 106.

[0113] As can be seen from Table 3, below, the operating conditions of distillation tower 106 are such that a recovery of 99.97% methane from feed gas 102 could be achieved at optimum conditions, giving a very high purity methane product, basically free of the other components of the feed gas 102 (e.g. C2, C0 ₂ and C3+).

Table 3 - Material balance of the overhead product from the first distillation tower 106

[0114] A liquid bottoms product from first distillation tower 106 is passed via line 122 to second distillation tower 118. The second distillation tower 118 into which these liquid bottoms are introduced is operated at conditions whereby the overhead produced 124 is substantially free of C3+ components, and the bottoms product 126 comprises C3+ components, substantially free of the lighter products in the sour gas feed 102. In this embodiment, second distillation tower 118 operates at about 21 bar, with the temperature at the bottom of tower 118 being about 90 °C and at the top about -22.6 °C. Typically, the second distillation tower 118 contains about 50 trays and is arranged to receive liquid feed from line 122 at about the seventh tray. [0115] The bottoms product 126 is removed via line 128, with 10% being recycled via line 120 to the top section of the first distillation tower 106 to prevent freezing of C0 ₂ in that column (as discussed above). The remaining 90% of the bottoms product 126 is collected as propane and heavier product for beneficial reuse. The operating conditions of distillation tower 118 are such that over 99 mol% recovery of the C3+ product is achievable (see Table 4).

Table 4 - Material balance of the bottoms product from the second distillation tower 118

[0116] The overhead vapours 124 contain carbon dioxide and ethane, and exit the second distillation tower 118 through line 130, where they are cooled and partially condensed. The resulting liquid-vapour mixture is directed to separator 132. Condensate separated in separator 132 is pumped back to the top tray of the second distillation tower 118.

[0117] The overhead vapour stream separated in separator 132 is passed via line 134 to a compressor 136, where the pressure is increased from about 21 bar to about 56 bar. The stream is then passed through a mixer 138, into which an azeotrope inhibitor agent (e.g. ethanol, at a concentration of less than 10 mol% of the feed rate to the third tower 142) is also added via recycle line 140. The mixture is then transferred into a third distillation tower 142, which may be operated under conditions where over 98.5 mol% of the total C2 product obtained from the top of the second distillation tower 118 is separated with the azeotrope inhibitor agent as bottom product 144 of the third distillation tower 142. In this embodiment, the third distillation tower 142 operates at about 47 bar, with the temperature at the bottom of tower 142 being about 52.5°C and at the top about -7°C. Typically, the third distillation tower 142 contains about 50 trays. [0118] The bottom product 144 of the third distillation tower 142 contains primarily ethane and the azeotrope inhibiting agent, and may be transferred via line 146 to a fourth distillation tower 148. The fourth distillation tower 148 may be operated at conditions where all of the C2 fraction of bottom product 144 is stripped as top product 150, while the azeotrope inhibiting agent is separated as bottom product 152 and recycled back to mixer 138 via line 140.

Operating distillation tower 148 under such conditions may achieve over 98.5 mole% recovery for the C2 product (based on its feed quantity to the first distillation tower 106), with the C2 product having a purity over 99mol% (see Table 5). In this embodiment, the fourth distillation tower 148 operates at about 5 bar, with the temperature at the bottom of tower 148 being about 137°C and at the top about -52°C. Typically, the fourth distillation tower 148 contains about 30 trays.

Table 5 - Material balance of the overhead product from the fourth distillation tower 148

[0119] The top product 150 (i.e. ethane) of the fourth distillation tower 148 exits the fourth distillation tower 148 through line 154, where it is cooled and partially condensed. The resulting liquid-vapour mixture is directed to separator 156. Condensate separated in separator 156 is pumped back to the top tray of the fourth distillation tower 148. The pure ethane product is then recovered from the process through line 158, via heat exchanger 160 to cool down the feed 102.

[0120] The top product 162 of the third distillation tower 142 may be transferred via line 164 to a mixer 168, into which a second azeotrope inhibitor agent (e.g. PropargylOl, at a concentration of less than 5 mol% of the feed rate to this tower) is also added via recycle line 170. The mixture is then transferred into a fifth distillation tower 172, which is operated under conditions whereby over 93% of the total amount of C0 ₂ of the top product 124 of the second distillation tower 118 is separated with the PropargylOl as bottom product 174 of the fifth distillation tower 172. In this embodiment, the fifth distillation tower 172 operates at about 29 bar, with the temperature at the bottom of tower 172 being about 1.5°C and at the top about -23°C.

Typically, the fifth distillation tower 172 contains about 40 trays.

[0121] The top product 176 of the fifth distillation tower 172 primarily comprises residual amounts of the ethane and carbon dioxide. About 99% of top product 176 is recycled via line 178 and compressor 180 back into mixer 138 (and hence third tower 142), with the other 1% being vented into the atmosphere via vent 181.

[0122] The bottom product 174 (mainly carbon dioxide and propargylOl) of the fifth distillation tower 172 may be transferred to a sixth distillation tower 182 via line 184. Sixth distillation tower 182 may be operated at conditions whereby substantially all of the captured CO2 fraction is stripped as top product 186, while the second azeotrope inhibiting agent is separated as bottom product 188, and recycled back to mixer 168 via line 170. As can be seen in Table 6, operation of the sixth distillation column 182 at such conditions can achieve over 93 mol% recovery of CO2 (based on the overhead product 124 of the second distillation tower 118), having a purity over 99mol%. In this embodiment, the sixth distillation tower 182 operates at about 5 bar, with the temperature at the bottom of tower 182 being about 176°C and at the top about 1°C. Typically, the sixth distillation tower 182 contains about 30 trays.

Table 6 - Material balance of the overhead product from the sixth distillation tower 182 [0123] As will be appreciated, the process described above with respect to Figure 3 enables substantially pure streams of methane, ethane, C3+ components and C0 ₂ to be recovered from a sour natural gas. The azeotrope inhibiting agents used in the process are recyclable, thereby reducing the quantity of these agents required to perform the process. The process can be optimised to take advantage of temperature and pressure differentials throughout in order to recover heat for beneficial use in the process.

[0124] The configurations and material balances described above in respect of the process depicted in Figure 3 are the optimum case based on HYSYS ^® simulation using ethanol and PropargylOl and a typical natural gas composition as shown in Table 2. Ethanol could be replaced by acetone and PropargylOl by ammonia, however, recoveries and purities of the respective products would then need to be re-optimized. Other agents like sulfolane could also be used to enhance separation, and other natural gas feed compositions could also be considered using the same agents and configuration. However, products purities and recoveries will be adjusted for each case.

[0125] Calculations using systems similar to that described above with reference to Figure 3 were applied to natural gasses having different compositions to that described above. Some of the natural gasses include hydrogen sulphide, in which case the process of Figure 3 was adapted to include a distillation system for separating hydrogen sulphide, as described above. The compositions of these natural gasses are shown in the following Table 7.

Table 7 - Material balance of natural gas feeds [0126] The compositions of the overhead product from the first distillation tower (Table 8), bottoms product from the second distillation tower (Table 9), overhead product from the fourth distillation tower (Table 10), overhead product from the sixth distillation tower (Table 11) and overhead product from the eighth distillation tower (Table 12) during separation of the components of each of the natural gasses set out in table 7 are set out below. As can be seen, the process of the present invention is applicable to a wide range of natural gas compositions.

Table 8 - Material balance of the overhead product from the first distillation towers

Table 9 - Materia balance of the bottoms product from the second distillation towers Component Overhead o: ^" fourth distillation tower

3.11% Without H2S 10% Without H2S 10% With ffiS 50% With ffiS mph mol% mph mol% mph mol% mph mol%

N2 0 0 0 0 0 0 0 0

C02 0.49 0.12 0.0018 0 0.0011 0 0.00465 0

H2S 0 0 0 0 0.6170 0.16 0.0001 0

CI 0 0 0 0 0 0

C2 406.61 99.88 385.6919 98.74 383.2172 99.84 211.3731 100

C3 0 0 4.9320 1.26 0 0 0 0 iC4 0 0 0 0 0 0 0 0 nC4 0 0 0 0 0 0 0 0 iC5 0 0 0 0 0 0 0 0 nC5 0 0 0 0 0 0 0 0

C6+ 0 0 0 0 0 0 0 0

Ethanol 0 0 0 0 0 0 0 0

Total 407.1 100 390.6 100 383.83 100 211.377 100

Table 10 - Material balance of the overhead product from the fourth distillation towers

Table 12 - Materia balance of the overhead product from the eighth distillation towers [0127] As will be evident from the foregoing description, as a relatively pure stream of carbon dioxide can be obtained using the process of the present invention, practice of the invention may facilitate a reduction of greenhouse gas emissions in comparison with conventional technologies for sweetening sour gas streams.

[0128] A financial instrument tradable under a greenhouse gas Emissions Trading Scheme (ETS) may be created by exploitation of the processes of the present invention. The instrument may be, for example, one of either a carbon credit, carbon offset or renewable energy certificate. Generally, such instruments are tradable on a market that is arranged to discourage greenhouse gas emission through a cap and trade approach, in which total emissions are 'capped', permits are allocated up to the cap, and trading is allowed to let the market find the cheapest way to meet any necessary emission reductions. The Kyoto Protocol and the European Union ETS are both based on this approach. One example of how credits may be generated by using the process of the present invention follows. A person in an industrialised country wishes to get credits from a Clean Development Mechanism (CDM) project, under the European ETS. The person contributes to the establishment of a gas sweetening plant comprising a gas sweetening plant employing the processes of the present invention. Credits (or Certified Emission

Reduction Units where each unit is equivalent to the reduction of one metric tonne of C0 ₂ or its equivalent) may then be issued to the person. The number of CERs issued is based on the monitored difference between the baseline and the actual emissions. It is expected by the applicant that offsets or credits of a similar nature to CERs will be soon available to persons investing in low carbon emission energy generation in industrialised nations, and these could be similarly generated.

[0129] It will be appreciated that the present invention provides a number of new and useful results and advantages over existing processes. For example, specific embodiments of the present invention may provide one or more of the following advantages:

• as methane, which often forms a vast majority of natural gas, is removed in a first step of the process, subsequent steps involve a much smaller volume of gas, and therefore require less power for heating/cooling etc. and can operate using smaller and less complicated processing equipment;

• as the methane and C3+ components of the natural gas have been substantially

removed in preliminary steps, the amount of azeotrope inhibiting agent(s) required to prevent (or hinder) azeotrope formation can be significantly reduced compared to conventional separation processes; • as the azeotrope is removed between ethane and carbon dioxide, the ethane recovery can reach above 98mol%, and with a purity greater than 99mol%; and

• methane, ethane and the C3+ components of a sour natural gas can be separated, producing separate streams of relatively pure methane, ethane and C3+ components.

[0130] It will be understood to persons skilled in the art of the invention that many

modifications may be made without departing from the spirit and scope of the invention. All such modifications are intended to fall within the scope of the following claims.

[0131] It will be also understood that while the following description refers to specific sequences of process steps, pieces of apparatus and equipment and their configuration to perform such processes in relation to particular gas compositions, operating pressures and temperatures, and so forth, such detail is provided for illustrative purposes only and is not intended to limit the scope of the present invention in any way.

[0132] In the claims which follow and in the preceding description of the invention, except where the context requires otherwise due to express language or necessary implication, the word "comprise" or variations such as "comprises" or "comprising" is used in an inclusive sense, i.e. to specify the presence of the stated features but not to preclude the presence or addition of further features in various embodiments of the invention.