WO2017020969A1

WO2017020969A1 - Process for treating a natural gas stream

Info

Publication number: WO2017020969A1
Application number: PCT/EP2015/068218
Authority: WO
Inventors: Arne Olav Fredheim; Bengt Olav Neeraas; Eivind Johannessen; Even SOLBRAA; Andrea Carolina Machado MIGUENS; Eleni PANTELI
Original assignee: Statoil Petroleum As
Priority date: 2015-08-06
Filing date: 2015-08-06
Publication date: 2017-02-09

Abstract

The present invention relates to a process for treating a natural gas stream to remove impurities such as carbon dioxide and water. The process may be carried out onshore or, alternatively, offshore in a floating environment, such as on a floating platform or ship which carries natural gas purification equipment. The invention further relates to an apparatus for performing such a process. In another aspect the invention relates to the use as a hydrate inhibitor of acid gas absorption compounds, such as methyl ethanol amine (MEA). In a further aspect the present invention provides a purified natural gas or liquefied natural gas obtained or obtainable by any of the processes herein described.

Description

PROCESS FOR TREATING A NATURAL GAS STREAM

Field of the Invention The present invention relates to a process and apparatus for removing impurities such as carbon dioxide and water from natural gas. Such a process may be carried out onshore or, alternatively, offshore in a floating environment, such as on a floating platform or ship which carries natural gas purification equipment. The invention further relates to the use as a hydrate inhibitor of acid gas absorption compounds, such as methyl ethanol amine (MEA).

Background of the Invention Natural gas is a fossil fuel mainly comprising a complex mixture of hydrocarbon and non- hydrocarbon components. The primary component of raw natural gas (i.e. natural gas directly taken from a geological formation) is typically methane. This is normally present together with varying amounts of heavier hydrocarbons such as ethane, propane, n-butane, iso-butane, pentanes and other higher molecular weight hydrocarbons as well as so-called "BTX" (benzene, toluene and xylene) components. Other components such as carbon dioxide (C0₂), hydrogen sulfide (H₂S), mercaptans, and other sulfur-containing compounds, are typically also present. C0₂ and sulfur-containing compounds are referred to collectively as "acid gases". Additionally, water (either in vapour or liquid form) and mercury (primarily as elemental mercury but also in the form of chlorides and other mercury compounds) may be present.

For transportation purposes, natural gas is desirably liquefied to form liquefied natural gas (LNG). LNG is a cryogenic liquid having a temperature at atmospheric pressure of about -162°C and a density of about 450 kg/m³. 600 m³ of natural gas at 1 atm pressure and 15°C corresponds to about 1 m³ of LNG. Transformation of natural gas to LNG therefore allows significant amounts of natural gas to be transported economically, e.g. by ship.

Impurities such as acid gases (C0₂ and sulfur-containing gases such as H₂S), water, mercury, BTX and heavier hydrocarbons, can have undesirable effects when present in natural gas. Water and acid gases can act together to corrode storage vessels, pipelines and other containers, presenting a safety risk and increasing maintenance and operational costs. H₂S is also toxic. The presence of acid gases also reduces the heating value of the natural gas. Mercury is toxic to humans and can cause corrosion problems, e.g. in aluminium heat exchangers. Water, C0₂, BTX and heavier hydrocarbons can condense and/or freeze during the liquefaction process, forming solids which interfere with liquefaction and block pipelines and other equipment such as heat exchangers.

It is therefore necessary to reduce the impurity content of natural gas in order to minimise these problems. Industry standards relating to natural gas production and transportation typically require that the C0₂ content of LNG should be no more than 50 parts per million by volume (ppmv) and the H₂S content should be no more than 4 ppmv. In addition, the water content of the gas should be reduced to no more than 1 ppmv, ideally in the range of 0.1 to 1 ppmv, to avoid freezing during cryogenic liquefaction.

Conventional methods for the removal of impurities prior to and/or during liquefaction to form LNG typically involve the following steps:

physical separation of distinct gas/liquid hydrocarbon/liquid water and/or solid phases in the raw gas (i.e. well stream), e.g. using an inlet separator;

- treatment of the separated gas phase to remove acid gases, typically using aqueous solutions of amines such as aMDEA (activated methyl diethanolamine);

- dehydration to reduce water content, typically using a molecular sieve adsorption plant downstream of the acid gas treatment plant;

removal of mercury, typically by adsorption, which may take place either upstream of the acid gas treatment plant or downstream of the dehydration plant;

removal of heavy hydrocarbons and "BTX" using a scrub column integrated with the liquefaction plant;

where hydrate inhibition is required, a hydrate inhibitor such as monoethylene glycol (MEG) is used, for example during upstream transportation of the raw gas to the purification plant and/or during other steps of the purification procedure. In natural gas production the well-stream is separated into gas and liquid phases in an inlet separator. The gas and liquid phases are in equilibrium at operational temperature and pressure, and the temperature will exceed that at which hydrates are formed. The gas leaving the inlet separator is saturated with water and so before cooling hydrate inhibition and/or removal of water is required. When cooling, liquid water will form and the gas will be stripped of water, i.e. the water content in the gas is reduced. However, if the cooling temperature is lower than the hydrate formation temperature a hydrate inhibitor needs to be added. In the existing technology, treatment to remove acid gases such as C0₂ typically involves direct contact between the natural gas and an absorption solution (e.g. an aqueous solution of an amine solvent such as aMDEA) in a large counter-current absorption column. The column dimensions are typically of the order of 3-5 metres in diameter and 20-30 metres in height. Natural gas which is discharged from the absorption column is saturated with water at the outlet temperature and pressure (typically this may contain 500-800 ppmv water). This necessitates a dehydration step as outlined above, for example in a downstream adsorption plant. Processing plants currently used for purification of natural gas before its conversion to LNG are therefore large, heavy and, as a result, hugely expensive. Their size and weight also limits their use in an offshore or floating environment.

There is therefore a need to develop alternative natural gas treatment processes which reduce the spatial requirements and cost of the purification plant. The need to do so is particularly acute in view of increasing interest in so-called "floating liquefied natural gas" (FLNG) systems. FLNG systems are those in which the necessary equipment for gas recovery, treatment, liquefaction and off-loading for storage or transportation is located on an offshore, floating environment such as a floating platform, ship or barge. FLNG systems are of particular interest for accessing offshore gas deposits which are isolated, remote from land, and potentially too small to warrant the installation of a permanent processing platform.

Compared to a land-based LNG facility, an FLNG system will have a greater need for reduced equipment size as well as reduced weight. FLNG systems must also be suitable for coping with motion of the platform, ship or barge caused by movement of the water (e.g. due to tidal and/or wave motion) and the weather conditions (e.g. wind) out at sea. Counter- current absorption columns typically used for acid gas removal are particularly susceptible to motion such as rocking, tilting and oscillation, which can disrupt motion of the gases and liquids inside the column and adversely affect the efficiency of the absorption process. In order to minimise any impact on performance of the process, current offshore facilities tend to be scaled-up in size to provide for an increase in solvent circulation. This results in an increase in size and weight of the process units required for effective offshore removal of C0₂ and water from natural gas streams (the weight of these units, including bulk and structural steel, is several thousand tons). It is therefore desirable not only to reduce the size and weight of the gas purification plant but also to make this less susceptible to reduced efficiency when used in an FLNG system. Summary of the Invention

The present inventors now propose that a solvent which has the ability not only to absorb an acid gas (especially C0₂), but also to function as a hydrate inhibitor may be used to simplify treatment of a natural gas stream in a pre-processing plant for natural gas or LNG. Lowering of the hydrate formation temperature using the solvent allows for effective removal of water from the gas stream by condensation and separation. Use of a solvent having a multipurpose application in this way enables the production of a natural gas product having regulation amounts (50 ppmv or below) of C0₂ and low amounts (e.g. 10 to 20 ppmv) of water vapour.

The process herein described represents a significant improvement in the reduction of C0₂ and water content of natural gas compared to methods known in the art and thereby simplifies the processing steps which need to be undertaken to remove acid gas and water. By simplifying the processing steps, the size, weight and cost of natural gas processing plants can be significantly reduced. This is advantageous in the context of both onshore and offshore processing. The process according to the invention is also less susceptible to reductions in efficiency caused by tilting, rocking or oscillating movements when carried out in a floating

environment. These features make the process of the invention particularly advantageous for employment in FLNG systems. In one aspect the present invention provides a process for treating a natural gas stream comprising the following steps:

(a) supplying a natural gas feed to a first gas/liquid separator whereby to produce a natural gas stream having a reduced water content;

(b) contacting the resulting natural gas stream with an aqueous solution which

comprises an acid gas absorption compound whereby to absorb C0₂, wherein said acid gas absorption compound also acts as a hydrate inhibitor;

(c) cooling the resulting gas/liquid mixture whereby to produce a gaseous phase comprising purified natural gas and a liquid phase comprising said acid gas absorption compound, C0₂ and water; and

(d) separating said gaseous phase and said liquid phase in a second gas/liquid

separator. The resulting gaseous phase will be on specification for subsequent liquefaction in terms of its C0₂ content (< 50 ppmv). The water content of the gaseous phase will be low, typically in the range of 10-50 ppmv, or preferably 10-20 ppmv. Prior to liquefaction, however, it is generally desirable to remove the remaining water to a level of 0.1-1 ppm. Optionally, the process may thus further comprise an additional dehydration step. This may be carried out using conventional dehydration methods, such as molecular sieve adsorption. Due to the low amount of additional water which needs to be removed, however, the adsorption beds or columns required can be significantly smaller and lighter than those used for traditional preprocessing concepts (where, for example, as much as 500-800 ppm water needs to be removed).

The process of the invention may comprise additional processing steps such as mercury removal, and removal of heavy hydrocarbons (HHC) and BTX. Such steps may be carried out according to conventional methods known to those skilled in the art. In principle the mercury removal step, where present, may be carried out either upstream of the C0₂ absorption step, or downstream of the dehydration step. HHC and BTX may, for example, be removed downstream of the dehydration step, for example using adsorption methods or a conventional turbo expander process. Alternatively the removal of HHC and BTX may be integrated with a refrigeration process such as those described herein. In a preferred aspect, HHC and/or BTX removal is achieved by means of the cooling step (c) as described herein, as HHC/BTX will be liquid at the reduced temperatures attained in this step and therefore any HHC/BTX present may be separated as a liquid phase, e.g. as part of the liquid phase comprising the acid gas absorption compound, C0₂ and water, or as an additional liquid phase which is immiscible with, but can be separated simultaneously with, the liquid phase formed in step (c).

Following step (d) and, optionally, any other downstream processing steps as may be desired, the purified natural gas stream may be subjected to liquefaction using conventional technology.

In a further aspect the present invention provides a purified natural gas or liquefied natural gas obtained or obtainable by any of the processes herein described.

The invention also provides apparatus specifically adapted for performing a process as herein described. In a further aspect the invention thus provides an apparatus for treating a natural gas stream comprising:

(a) a first gas/liquid separator arranged to receive a natural gas feed; (b) a mixer arranged to receive a natural gas stream having a reduced water content from said first gas/liquid separator and to contact said natural gas stream with a solution which comprises an acid gas absorption compound which is capable of absorbing C0₂ and which also acts as a hydrate inhibitor;

(c) cooling means arranged to receive a gas/liquid mixture from said mixer and to cool said mixture to form a gaseous phase comprising purified natural gas and a liquid phase comprising said acid gas absorption compound, C0₂ and water; and

(d) a second gas/liquid separator arranged to separate said gaseous phase and said liquid phase.

Additional components of the apparatus may include a dehydrator arranged to extract water from the gaseous phase produced in step (d) and/or liquefying means capable of liquefying the resulting purified natural gas to produce LNG. As described above, the inventors have also found that the acid gas absorption compounds herein described are capable of acting as hydrate inhibitors, thereby effectively replacing the need for any separate hydrate inhibitor (such as MEG) in the processing of natural gas. This finding represents an additional aspect of the invention. In a yet further aspect the invention thus provides the use of an acid gas absorption compound as herein described as a hydrate inhibitor, for example in a process for the production of purified natural gas or the production of LNG. A method of inhibiting hydrate formation in a natural gas stream which comprises the step of contacting said gas stream with an acid gas absorption compound as herein described forms a further aspect of the invention.

As used herein "hydrate" refers to a solid form of water formed at high pressure in the presence of gas molecules normally found in hydrocarbon gases and liquids. Hydrates form a crystalline phase, similar to ice, and can potentially plug gas and liquid pipelines and production equipment used in the processing of hydrocarbon gases and liquids. A typical example is transport of hydrocarbons (gas and/or liquid) in a pipeline where the temperature drops due to cold surroundings (e.g. sea water or cold air). A water phase may be present at all times or water may condense as the temperature drops. At high pressure and low temperature, the water may form hydrates stabilised by the gas molecules in the

hydrocarbon phase. Detailed description of the invention

The first gas/liquid separator (also referred to herein as an "inlet separator") may be one commonly employed for physical separation of liquid hydrocarbon, liquid water and/or solid phases from a gas feed, for example a "raw" gas feed from a wellhead. Suitable inlet separators will be known to those skilled in the art.

Prior to passage into the inlet separator the raw natural gas feed may be treated with a conventional hydrate inhibitor. This is desirable to minimise the risk of hydrates (e.g.

clathrates) forming during transport of the gas feed from the wellhead to the inlet separator (e.g. via a pipeline), as hydrate formation can lead to blockages and/or interruptions in the gas feed. A commonly-used hydrate inhibitor is monoethylene glycol (MEG) although other hydrate inhibitors will be known to those skilled in the art. The natural gas feed to the inlet separator may have a pressure of 50 bar or greater, e.g. 80 bar or greater. In one embodiment, the natural gas feed is a "high-pressure" gas feed, e.g. a gas feed having a pressure of about 100 to about 300 bar. A high-pressure gas feed enables subsequent cooling of the gas by expansion to operational pressures (typically 50 to 80 bar). In the case where cooling is achieved by expansion (i.e. using an expander), the work generated by expansion can be used to drive other components of the process, e.g. to drive a compressor and/or to produce electrical power. However, the invention is not limited to treatment of high pressure natural gas feeds (i.e. those having a pressure above about 100 bar). If sufficient gas pressure is not available, cooling can be achieved by other means as herein described, e.g. by refrigeration which may be integrated with the LNG liquefaction process.

As used herein, the term "acid gas absorption compound" is used to describe a compound which is capable of absorbing and/or reacting with an acid gas such as C0₂ and optionally H₂S. The term "acid gas absorption solution" refers to a solution comprising at least one such acid gas absorption compound.

Suitable acid gas absorption compounds for use in the invention are those which, when present as a component of an acid gas absorption solution, allow the acid gas absorption solution to remain liquid at temperatures between 0°C and -60°C, preferably at temperatures between 0°C and -35°C, e.g. between 0°C and -30°C, e.g. at about -25°C. The acid gas absorption compound is one which is also capable of lowering the hydrate formation point when mixed with natural gas and water (e.g. any water vapour which may be present in the natural gas feed). This allows the mixture of natural gas, water and acid gas absorption compound to be cooled to a temperature below the temperature at which hydrates would normally form. This process may be referred to as "hydrate inhibition". Water which would normally be "locked up" in hydrate form instead remains as free water in the mixture, consequently becoming easier to separate from the mixture.

Addition of an acid gas absorption compound causes a lowering of the freezing point of the aqueous solvent (the term "aqueous solvent" may encompass not only the water present in the added acid gas absorption solution but also any water present in the natural gas stream). The extent to which the freezing point is lowered will depend on the composition of the mixture, i.e. the ratio of acid gas absorption compound to water. The ratio of acid gas absorption compound to water will also influence the extent of hydrate inhibition. The ratio of acid gas absorption compound to water will therefore influence both the hydrate inhibition properties and the freezing point depression properties of the resulting mixture. As will be understood, the desired ratio will depend on various parameters, such as the water content of the gas feed, the pressure of the gas feed (before and/or after expansion, if cooling by expansion is employed), the nature of the acid gas absorption compound, and the desired temperature after cooling, etc. In determining the amount of acid gas absorption compound to be employed, it is important not only that the acid gas absorption compound achieves the desired degree of acid gas absorption and hydrate inhibition, but also that sufficient acid gas absorption compound should be present in the liquid phase after cooling to ensure that the liquid phase remains liquid rather than freezing.

In an embodiment, the amount of the acid gas absorption compound in the liquid phase obtained following cooling should be sufficient to ensure that the freezing point of the liquid phase is below 0°C. In other embodiments, this should be sufficient to ensure the freezing point of the liquid phase is below -10°C, below -25°C, or below -50°C.

By way of example, the concentration of acid gas absorption compound in the liquid phase may be about 65 to about 80 wt% (e.g. about 70 to about 75 wt%) if cooling to -25°C is required, about 50 to about 90 wt% (e.g. about 55 to about 85 wt%, such as about 60 to about 80 wt%, about 65 to 75 wt%, or about 65 to 70 wt%) if cooling to about -10°C is required, or about 40 to about 95 wt% (e.g. about 45 to about 90 wt%, about 50 to about 85 wt%, about 55 to about 85 wt%, such as about 60 to about 80 wt%, about 65 to 75 wt%, or about 65 to 70 wt%) if cooling to about 0°C is required. Where cooling is achieved by expansion such that the inlet pressure in the separator is in the range of from about 60 to about 150 bar, e.g. about 60 to about 1 10 bar (e.g. about 104.5 bar), and where the acid gas absorption compound employed is

N-methylethanolamine (MEA), the concentration of the acid gas absorption compound in the liquid phase may, for example, be about 65 to about 80 wt% (e.g. about 70 to about 75 wt%) if cooling to -25°C is required, about 50 to about 90 wt% (e.g. about 55 to about 85 wt%, such as about 60 to about 80 wt%, about 65 to 75 wt%, or about 65 to 70 wt%) if cooling to about -10°C is required, or about 40 to about 95 wt% (e.g. about 45 to about 90 wt%, about 50 to about 85 wt%, about 55 to about 85 wt%, such as about 60 to about 80 wt%, about 65 to 75 wt%, or about 65 to 70 wt%) if cooling to about 0°C is required. In one example, where the inlet pressure in the separator may be about 60 bar, and the cooling temperature is about -25°C, the concentration of MEA in the liquid phase will be about 75 wt% (e.g. 70 to 80 wt.%).

The required amount of acid gas absorption compound will also be dependent on other factors such as its molecular weight. As will be understood, the required amount will be dependent approximately linearly on the molecular weight of the compound. For example, a compound having about half the molecular weight of MEA would need to be present at about half of any of the concentrations specified above for any given level of freezing point depression.

Once the desired concentration of acid gas absorption compound in the resulting (i.e.

separated) liquid phase has been determined (i.e. that required to avoid hydrates and freeze-out), it is possible to determine the amount of acid gas absorption compound which needs to be added to the natural gas stream prior to the cooling step. For example, 1 MSm³ (where Sm³ = standard cubic metre and M = 1 ^χ 10⁶, in line with normal SI nomenclature, so that 1 MSm³ is 1 x 10⁶ Sm³ ) of natural gas may contain about 450 kg of water at 70 bar and 25°C if it is saturated with water vapour. If this is cooled to -25°C while maintaining a pressure of 70 bar, the gas will contain only about 14 kg of water and the remaining 436 kg of water will condense into the liquid phase, diluting the acid gas absorption compound which was injected upstream. It is therefore important to ensure that the amount of acid gas absorption compound injected upstream is sufficient that even when it is diluted by the condensation of water from the gas into the liquid phase, its concentration remains sufficient to achieve the desired freezing point suppression as discussed above. Knowing the identity of the acid gas absorption compound, the degree of freezing point depression required, the water content and flow rate of the gas feed, and other parameters which will vary from process to process and from one gas well to another, the skilled person will be able to determine the desired amount of acid gas absorption compound to employ in the processes described herein.

The required flow rate of acid gas absorption compound will depend on the exact process conditions, including the temperature of the gas feed leaving the inlet separator, the flow rate of the gas feed from the well, the desired temperature reduction to be attained in the cooling step, and the extent of hydrate inhibition required, etc. For example, the flow rate may be about 2,000 kg/hr or more, such as about 2,100 kg/hr or more, about 2,200 kg/hr or more (e.g. 2,250 kg/hr or more), about 2,300 kg/hr or more, about 2,400 kg/hr or more, or about 2,500 kg/hr or more. Preferably a flow rate greater than the minimum flow rate will be employed (e.g. at least 1.5 times the minimum flow rate, preferably at least 2 times the minimum flow rate). Thus, for example, the flow rate may be about 4,000 kg/hr or more, about 4,200 kg/hr or more, about 4,400 kg/hr or more, about 4,500 kg/hr or more (e.g. about 4,574 kg/hr), or about 4,600 kg/hr or more.

The acid gas absorption compound may, for example, be any primary, secondary or tertiary amine, amino acid salt, ammonia, sodium carbonate, potassium carbonate, or lithium carbonate. In particular, it may be at least one primary, secondary or tertiary alkanolamine. Preferably it is at least one secondary alkanolamine such as MEA.

According to one embodiment the acid gas absorption solution may comprise at least one alkanolamine of formula (I):

NR¹R²R³ (I) wherein R¹, R² and R³ are independently selected from hydrogen, Ci_-6 alkyl and Ci_-6 alkanol, with the proviso that at least one of R¹, R² and R³ is a Ci_-6 alkanol.

As used herein the term "C_1-6 alkyl" refers to any straight-chain or branched alkyl group having one, two, three, four, five or six carbon atoms, such as methyl, ethyl, n-propyl, isopropyl, n-butyl, iso-butyl, sec-butyl, t-butyl, n-pentyl, tert-pentyl, neopentyl, isopentyl, sec-pentyl, 3-pentyl, 1-hexyl, 2-hexyl or 3-hexyl groups.

As used herein the term "C_1-6 alkanol" refers to any straight-chain or branched alkyl group having one, two, three, four, five or six carbon atoms such as those described above, wherein at least one hydrogen atom is substituted by an -OH functional group. Examples include methanol, ethanol, propanol (e.g. 1 -hydroxypropyl, 2-hydroxypropyl and

3-hydroxypropyl), butanol, pentanol and hexanol groups. In an embodiment the acid gas absorption solution may comprise at least one amine selected from among 2-amino-2-methyl-1 -propanol (AMP); 2-amino-2-hydroxymethyl-1 ,3- propanediol (TRIS); diethyl monoethanolamine (DEMEA); dimethyl monoethanolamine (DMMEA); N-methyl diethanolamine (MDEA); hydroxethyl piperazine (HEPZ);

2-aminoethanol (MEA); 2,2'-dihydroxydiethylamine (DEA); diethylene glycol monoamine (DGA); di-isopropyl amine (Dl PA); triisopropanolamine (Tl PA); and mixtures thereof. MEA, DEA, DGA, DIPA and/or MDEA are preferred for use in the invention. A particularly preferred acid gas absorption compound is MEA (methyl ethanol amine or 2-aminoethanol).

The acid gas absorption solution which contacts the natural gas stream may comprise a single acid gas absorption compound such as those described above, or a mixture of two or more different acid gas absorption compounds. The acid gas absorption solution is an aqueous solution, i.e. one which contains at least a proportion of water.

The pKa of the acid gas absorption solution, when measured at 25°C, is typically in the range of about 8 to about 1 1.

The acid gas absorption compounds for use in the invention are also capable of inhibiting hydrate formation in a gas feed, e.g. inhibiting the formation of gas hydrates caused by the presence of water vapour. Of particular interest are compounds which are capable of inhibiting the formation of methane hydrates, e.g. methane clathrate hydrates.

Contact between the gas feed and the acid gas absorption compound may be carried out in a conventional mixer, for example in either a co-current or counter-current mixer. In one embodiment, contact is effected in a co-current mixer or column in which the components (i.e. the gas feed and acid gas absorption solution) are both added at the same inlet point (or inlet points which are in close proximity to one another) and removed at the same outlet point. It is preferred that this step is carried out in a single mixing step (i.e. with a single mixer), however, multiple mixing steps may be employed, for example employing a plurality of co-current mixers arranged in series. Where a plurality of co-current mixers are used, the resulting mixture will be separated into separate gas and liquid phases by cooling prior to the addition of further acid gas absorption compound and further co-current mixing (and separation). Use of a co-current mixer (or co-current mixers) effectively eliminates, or at least reduces, any impact on the efficiency of the process due to tilting, vibration or rocking of the processing equipment when this is sited on a floating structure, e.g. platform, ship or barge. Following the step of contacting the gas feed with the acid gas absorption compound, the gas/liquid mixture is cooled in order to produce a liquid phase comprising the acid gas absorption compound, absorbed C0₂ and other acid gases (where these are present, e.g. H₂S), and water. This liquid phase may also comprise other impurities such as heavier hydrocarbons, where these are present in the original gas feed. In the case where the step of acid gas absorption is carried out by means of a series of separate mixing steps, the cooling of the gas is also performed after the final such mixing step, i.e. after the mixture of the gas and the acid gas absorption solution has exited the final mixer.

The step of cooling is carried out directly following treatment with the acid gas absorption compound. By "directly" it is intended that the gas/liquid mixture produced following mixing with the acid gas absorption compound will be subjected to cooling without any additional intermediate processing steps. The mixture will, however, typically pass through a length of piping (i.e. a pipeline) during transit to the cooling means. During passage through the piping, the acid gas absorption solution and gas remain in contact thus increasing their contact time and thereby enhancing the efficiency of removal of acidic components from the gas. Passage of the liquid/gas mixture through one or more pipes is not considered to constitute an additional processing step.

The cooling should be performed without resulting in freezing of any phase or component of any phase. Typically, the gas/liquid mixture will be cooled to a temperature in the range 20 to -60°C, preferably 0 to -40°C, e.g. to about -25°C.

In one embodiment cooling may be effected by expansion of the gas. Cooling by expansion is more efficient where a significant reduction in pressure is possible, and therefore this is the preferred cooling method where the gas feed to be cooled is at high pressure. For example, the gas feed to be cooled may be provided at a pressure of about 100 to about 300 bar and then cooled by expansion to reduce the pressure. The exact degree of reduction in pressure required will depend on the starting pressure of the gas before expansion, but in any case the reduction in pressure should be selected in order to achieve the desired degree of cooling and thus separation. Typically the desired degree of cooling will be attainable by reducing the pressure by at least about 20 to 40 bar compared to the starting pressure. In general, the pressure after expansion will not be lower than about 30 bar, and for example may be in the range of about 50 to about 80 bar. Thus in the case where the gas feed is supplied at an initial pressure of about 100 to about 300 bar, cooling may be achieved by rapidly reducing the pressure to a pressure of about 30 to 150 bar, about 30 to 100 bar, about 40 to 100 bar, about 50 to 100 bar, or about 40 to 80 bar, e.g. to about 50 to 80 bar.

The step of expansion may involve reducing the pressure to an extent sufficient to achieve cooling to a temperature between about 0°C and about -60°C, e.g below 0°C, below -10°C, below -25°C, or below -50°C.

Expansion of the gas may be effected by way of any technique or combination of techniques known to those skilled in the art. Suitable apparatus for achieving the desired degree of expansion will be readily identified by those skilled in the art and include, for example, a Joule-Thomson valve (which provides isenthalpic expansion) or an expander, such as a rotating expander (which provides almost isentropic expansion). In an embodiment the work generated by the expander may be used to produce electrical power, e.g. to drive a compressor, thus improving the energy efficiency of the overall process.

Alternatively, cooling may be performed by externally applied cooling means, e.g. by refrigeration. This may be employed, for example, in the case where a high-pressure gas feed is not available, for example where the gas feed has a pressure of below about 100 bar, e.g. about 80 bar or below. Refrigeration may be achieved by any conventional mechanical refrigeration means. Optionally, refrigeration may be "integrated" with the step of liquefying the gas to produce LNG. In such an integrated process the refrigeration step and the liquefaction step are discrete steps in a single continuous process. In such a process, refrigeration and subsequent cooling to effect liquefaction may be carried out using the same or different refrigerants. In one embodiment the same refrigerant may be used. Refrigeration and subsequent liquefaction may be carried out in different vessels, although in one aspect the same vessel may be employed in order to simplify the overall process. Suitably the vessel or vessels employed is/are heat exchangers. A non-limiting example of an "integrated" process is illustrated in Figure 2. In an alternative embodiment, the step of liquefaction may be performed in a separate, independent, process (i.e. a "non-integrated" process), e.g. in a separate LNG liquefaction plant. For example, where cooling to effect removal of C0₂ and water is carried out offshore (e.g. on a floating platform), subsequent liquefaction may be carried out onshore. Onshore liquefaction may be preferred where, for example, for safety reasons this is difficult to perform offshore. However, if cooling is carried out offshore and liquefaction is carried out onshore, the temperature of the cooled gas phase may undesirably increase during transport from the cooling step to the liquefaction step, for example if the cooled gas is transported to the liquefaction step by means of an undersea pipeline which is at the same temperature as the surrounding sea water. In such an embodiment the gas may require further cooling onshore prior to liquefaction. Preferably therefore the complete treatment process is carried out onshore, or alternatively the complete treatment process is carried out offshore; in either case this means that the process is not separated into onshore and offshore sections which could give rise to undesirable temperature changes in between those sections. For similar reasons, it is preferred that the complete treatment process is also carried out onshore, or the complete treatment process is carried out offshore, in the case of "integrated" processes as discussed above.

Cooling may also be achieved using a combination of expansion and externally applied cooling means (e.g. mechanical cooling). The degree of cooling which can be achieved by expansion is determined by the available pressure difference (which depends principally upon the feed pressure) and on the type of expansion employed (for example, either when using a Joule-Thomson valve or an expander). In instances where insufficient cooling can be achieved by expansion alone, a combined use of expansion and other cooling means (e.g. refrigeration) may be appropriate. This could be an option, for example, where the feed gas pressure is in the range of about 90 to 150 bar, e.g. 100 to 140 bar, and thus the degree of cooling which can be achieved by expansion alone may not be sufficient.

Following cooling, the liquid phase is separated from the gas phase. In this way the process of the invention provides a gas phase having a C0₂ content of 50 ppmv or below, and a water content in the range of 1 to 50 ppmv, preferably below 30 ppmv, e.g. about 10-20 ppmv. This represents a significant reduction in water content compared to gas phases obtained after acid gas removal steps known in the art (typically these contain 500-800 ppm water).

To reduce the water content further in order to meet industry standards, the cooling step may be followed by an additional dehydration step in order to reach a final water level of 0 to 2 ppmv, preferably 0.1 to 1.5 ppmv, e.g. about 0.1 to 1 ppmv. Dehydration may be performed according to methods known to those skilled in the art, e.g. by molecular sieve absorption. Following dehydration the gas phase may be liquefied according to conventional methods in order to produce LNG having low concentrations of impurities in line with industry

specifications. Alternatively the gas feed may be retained in gaseous form, e.g. for supply to a pipeline (so-called "pipeline gas").

For pipeline gas the industry standard specifies a C0₂ content of at most 2 to 3 mole %. In one embodiment, the process of the invention provides a gas phase having a C0₂ content of 3 mole % or below and a water content in the range of 1 to 50 ppmv. In other embodiments, the C0₂ content of the resulting gas phase may be 2.5 mole % or below, e.g. 2 mole % or below, and the water content may be below 30 ppmv, e.g. about 10-20 ppmv.

Although it is preferred that the step of contacting the natural gas with the acid gas absorption compound takes place in one or more co-current mixers, as described above, a conventional counter-current absorber may be employed in addition to the co-current mixer or mixers. Such processes are particularly appropriate in applications where co-current mixing may not be sufficient to remove enough C0₂ to reach the industry standard of 50 ppmv or below, and/or if the kinetics of the co-current absorption are too slow. A preferred sequence of mixing is for the feed gas to contact the acid gas absorption solution first in a counter-current absorber before exiting the counter-current absorber and passing to a co-current mixer. In the context of FLNG applications, a combination of a counter-current absorber and a co-current mixer may serve as a safeguard in the event that the

counter-current absorber functions sub-optimally due to motion and tilt of the FLNG platform as described above. In cases where a combination of co-current and counter-current mixing is employed, the feed gas temperature to the counter-current absorber may be in the range from about 20°C to about 50°C. Gas leaving the counter-current contactor should preferably then be cooled by ambient cooling whereby to remove water by condensation and separation. The processes herein described are especially suitable where the raw natural gas feed has a low C0₂ content, preferably about 1 to 2 mol% C0₂ or less, e.g. about 2 mol% C0₂ or less, or 1 mol% C0₂ or less.

As discussed, the processes herein described are particularly suited to meet FLNG requirements. In one embodiment, therefore, the process of treating a natural gas stream may further comprise the step of sending a natural gas stream to a floating platform, ship or barge that is carrying equipment for purification of natural gas. A further aspect of the invention relates to the use of an acid gas absorption compound as herein described in inhibiting hydrate formation in a natural gas or hydrocarbon feed. In one embodiment the use is in a process for treating natural gas as herein described.

In one embodiment the acid gas absorption compound may be employed as a hydrate inhibitor in a natural gas pipeline. In such applications, the acid gas absorption compound acts both to inhibit hydrate formation in the pipeline flow as well as absorbing acid gases such as C0₂ in the pipeline. When used as a hydrate inhibitor, the compound may suitably be added to the raw natural gas feed at the wellhead and prior to feeding to an inlet separator.

The amount of acid gas absorption compound added for use as an inhibitor depends on the necessary degree of protection required, which for example will be dependent on the exact process conditions such as the amount of C0₂ and/or H₂S to be removed, the required degree of freezing point depression, the level of water in the feed gas, and so on. Typically this may be added in an amount of about 2 vol% of the combined gas and liquid flow or less, e.g. from about 1 to about 2 vol%, preferably about 1 vol% or less, e.g. from about 0.1 to 1 vol%, about 0.8 vol% or less, e.g. about 0.5 to about 0.8 vol%, of the combined gas and liquid flow. The skilled person will readily be able to determine the appropriate flow rates and amounts having regard to the exact process conditions associated with any given gas feed. The hydrate inhibitor may be added to the gas or hydrocarbon stream either batchwise or continuously. Embodiments of the invention will now be described in further detail with reference to the accompanying non-limiting figures in which:

Fig. 1 illustrates a process and apparatus according to the invention for use with a high feed gas pressure, for example in an LNG or FLNG application;

Fig. 2 illustrates a process and apparatus according to the invention for use with a medium or low feed gas pressure comprising an integrated refrigeration step, for example in an LNG or FLNG application; Fig. 3 illustrates a process and apparatus according to the invention for use in producing natural gas suitable for transportation without liquefaction, for example transportation in a gas pipeline; Fig. 4 illustrates a process and apparatus according to the invention employing a

combination of counter-current and co-current mixing to remove the desired amount of C0₂; and

Fig. 5 illustrates a process and apparatus according to the invention in which the acid gas absorption compound is employed as a hydrate inhibitor in the gas feed to the inlet separator. Figs. 1 to 5 illustrate the steps which may be used to treat a natural gas stream in accordance with the invention. The processes are primarily intended to be carried out on a floating structure, e.g. on a floating platform such as a raft, or on a ship, but could also be carried out at a land based facility. In the various embodiments shown, acid gases (e.g. C0₂) are removed from an impure gas feed, such as natural gas which predominantly comprises methane, using MEA as the acid gas absorption compound in a co-current mixer. As will be understood from the description of the invention provided herein, other acid gas absorption compounds may be used in place of MEA. Subsequent cooling of the resulting gas/liquid mixture to a temperature of about -25°C (e.g. by expansion or refrigeration) provides a treated gas feed which is on specification for C0₂ content (< 50 ppmv C0₂) and nearly on specification for water content (~ 10-50 ppmv H₂0) for liquefaction. If required, the resulting gas feed can be passed to a "finishing" or "polishing" zone to more completely remove any residual water before liquefaction.

Referring to Fig. 1 , a feed gas 101 from an offshore well at high pressure is inhibited with a chemical, typically MEG (Mono Ethylene Glycol) 102 to avoid hydrate formation in the pipeline 103 from the well-head 100 to the FLNG installation or the LNG plant which may for example be an FLNG topside or an LNG onshore plant. At the inlet facility 104 of the FLNG/LNG plant, hydrate inhibitor, water and hydrocarbon condensate 105 is separated from the natural gas in the inlet separator or slug catcher 106. The amount of liquid formed will depend on the actual pressure and temperature under the receiving conditions. The water saturated natural gas feed 107 exiting at high pressure from the separator (top) is mixed with a solvent containing an acid gas absorption compound 108 (e.g. MEA, although alternative compounds as described herein are suitable). The MEA (in this case) and the natural gas are mixed in a co-current mixer 109. Mixing is preferably performed in a single- step process (as illustrated in Fig. 1 ), although a multi-step process may also be employed for this stage. After the mixing of MEA/water and natural gas, the mixture is expanded to a lower pressure, e.g. using a Joule-Thomson valve or a rotating expander 1 10. Expansion will lower the temperature of the liquid/gas mixture, typically down to -25°C. At this temperature, liquid phases of MEA, water and hydrocarbon condensate will form and can be separated from the natural gas in a second separator 1 1 1 (the liquid phase 1 12 containing MEA, water and hydrocarbon condensate leaves through conduit 1 12a). In the example shown a pressure of about 60-100 bar and a temperature of about -25 °C are employed in the separator although the skilled person will appreciate that other conditions are also applicable. The natural gas 1 13 leaving the second separator will be on specification for C0₂ content (< 50 ppm) and nearly on specification for water content (about 10-50 ppm). The remaining water will typically be removed by an adsorption process (not shown).

Referring to the process shown in Fig. 1 , a typical example of a procedure which may be used to calculate appropriate amounts of acid gas absorption compound (hydrate inhibitor) is illustrated below:

A three-step process is employed: 1. Specify temperatures, pressures and gas flow rates.

2. Determine the wt% MEG (as a hydrate inhibitor in the feed gas flow from well head to the FLNG or LNG facility) and wt% MEA (as acid gas absorption compound/hydrate inhibitor 108) which are needed in order to avoid hydrates and freezing out of MEA during the cooling step. 3. Calculate the rates and concentrations of MEG and MEA which need to be injected into the gas flow in order to obtain the concentrations which were determined in point 2 and result in < 50 ppm C0₂ in the treated gas.

Steps 2 and 3 do not affect the temperatures, pressures or gas flow rates significantly because C0₂ and water are trace components in this particular example and only small quantities of MEG and MEA are therefore needed.

The following parameters are assumed: feed gas 101 flow rate of 30 MSm³ per day from the well-head 100; pressure of 130 bar and temperature of 5°C for gas 107 leaving the inlet separator 106; expansion through a Joule-Thomson valve 1 10 to a reduced pressure of 60 bar which produces a temperature of -25°C in the separator 1 1 1 . The gas flow 101 from the well-head 100 is assumed to contain about 200 ppmv C0₂, a flow rate of about 464 kg/hour (10.6 kmol/hour) of C0₂. The treated gas 1 13 leaving the separator 1 1 1 should have a C0₂ content of <50 ppmv, which equates to a flow rate of < 1 16 kg/hour or <2.64 kmol/hour of C0₂. To maintain these temperatures without freezing-out of MEG, MEA or other liquid

components it is determined that there must be about 50 wt% MEG in the liquid phase 105 in the inlet separator 106 and about 75 wt% MEA in the liquid phase 1 12 in the separator 1 1 1. Under these conditions the gas 107 leaving the inlet separator 106 has a water content of about 95 ppmv (a flow of about 90 kg/hour) and the treated gas 1 13 leaving separator 1 1 1 has a water content of about 21 ppmv (a flow of about 20 kg/hour).

The MEG injection rate will depend on the temperature and pressure in the well, and whether water will be produced or not. It is common to inject 90 wt% MEG. A significant part of the water dissolved in the gas at well conditions, plus any water which is produced, will dilute the MEG. The calculation of the amount/rate of MEG needed to give a resulting concentration of about 50 wt% MEG in the inlet separator 106 therefore depends heavily on the conditions at the production well.

In order to achieve the desired level of <50 ppmv C0₂ in the treated gas 1 13 leaving separator 1 1 1 , the minimum flow rate of aqueous MEA needed is 2250 kg/hour. The composition of the lean solvent is 161 1 kg/hr MEA plus 467 kg/hr H₂0 plus 174 kg/hr C0₂. It is assumed that the lean solvent contains 0.15 mol C0₂ per mol MEA (written coventionally as 0.15 mol/mol) and that the rich solvent contains 0.45 mol/mol (this is approximately the theoretical maximum capacity for C0₂ under these conditions). The injected solvent 108 will also absorb 70 kg/hour H₂0 from the gas 107 on its way to the low temperature separator 1 1 1. The injected solvent 108 therefore contains 77.5 wt% MEA and 22.5 wt% H₂0 on a C0₂-free basis whereas the composition of the solvent 1 12 leaving the low temperature separator 1 1 1 will be 75 wt% MEA and 25 wt% H₂0 on a C0₂-free basis.

In practice more than the minimum quantity of MEA would normally be injected. For example, the quantity of MEA may be double the minimum quantity. The rate of injected solvent 108 will then be 4574 kg/hr, containing 3222 kg/hr MEA, 1004 kg/hr H₂0 and 348 kg/hr C0₂. The quantity of C0₂ left in the treated gas 1 13 in this case will be <50 ppm. The exact value will depend on the rate of C0₂ absorption by the solvent. The absorption rate depends on the exact design and parameters of the system. The solvent will also absorb 70 kg/hr of H₂0 from the gas. The injected solvent 108 (where double the minimum quantity is used) therefore contains 76.6 wt% MEA and 23.4 wt% H₂0 on a C0₂-free basis whereas the composition of the solvent 1 12 leaving the low temperature separator 1 1 1 will be 75 wt% MEA and 25 wt% H₂0 on a C0₂-free basis. With reference to Fig. 2, a feed gas 201 entering the FLNG installation or the LNG plant has a medium to low pressure. The feed gas 201 is inhibited with a chemical, typically MEG (Mono Ethylene Glycol) 202 to avoid hydrate formation in the pipeline 203 from the wellhead 200 to the FLNG installation or the LNG plant. At the inlet facility 204 of the

FLNG/LNG plant, hydrate inhibitor, water and hydrocarbon condensate 205 is separated from the natural gas in the inlet separator or slug catcher 206. The water saturated natural gas feed 207 exiting at high pressure from the separator (top) is mixed with an acid gas absorption compound 208 (e.g. MEA, although alternative compounds as described herein are suitable). The MEA (in this case) and the natural gas are mixed in a co-current mixer 209. Mixing is preferably performed in a single-step process (as illustrated in Fig. 2), although a multi-step process may also be employed for this stage.

Due to the medium to low pressure of feed gas 201 , it is less desirable to effect cooling by expansion (the reason for this is that the feed pressure needs to be maintained to be able to liquefy the natural gas effectively). In this situation, it is preferred that cooling (typically down to about -25°C) is carried out by external cooling means 210 (e.g. by means of a refrigerant), although a combination (not shown) of cooling by expansion and cooling by external cooling means may also be employed. Following cooling, liquid phases of MEA, water and hydrocarbon condensate will form and can be separated from the natural gas in a second separator 21 1 (the liquid phase 212 containing MEA, water and hydrocarbon condensate leaves through conduit 212a). In the example shown a pressure of about 60-100 bar and a temperature of about -25 °C are employed in the separator although the skilled person will appreciate that other conditions are also applicable. The natural gas 213 leaving the second separator will be on specification for C0₂ content (<50 ppm) and nearly on specification for water content (about 10-50 ppm). The remaining water will typically be removed by an adsorption process (not shown).

Fig. 3 shows an embodiment of the invention in which the process herein described is used in the production of natural gas which is not liquefied (i.e. pipeline gas). The process which is shown is suitable for producing purified natural gas which having a specification suitable for transportation (e.g. in a pipeline) or a final product specification. With reference to Fig. 3, a feed gas 301 is inhibited with MEG 302, or other chemical if preferred, and transported in pipeline 303 from well-head 300 to the platform or FPSO inlet separator 304 where the liquid phase(s) 305 is separated. An acid gas absorption compound (e.g. MEA or another suitable chemical) 306 is mixed with the gas flow in one or more co-current mixers 307 for removal of C0₂ and H₂S. The resulting gas/liquid mixture is expanded (e.g. via expander 308) to a lower pressure/temperature to separate a liquid phase 310 water, MEA with C0₂, and condensate/heavy hydrocarbons from separator 309. In the example shown the separator employs a pressure of about 60 to 150 bar and a temperature of about -25°C although the skilled person will appreciate that other conditions are also applicable. The resulting purified gas stream is 31 1 suitable for transportation (e.g. to shore via a pipeline 312). If the final pressure of the gas stream exiting the second separator 309 is not high enough, this may be compressed (not shown) for further transportation.

Referring to Fig. 4, this embodiment of the invention comprises both co-current and counter- current mixing of the acid gas absorption compound (e.g. MEA or another suitable compound). Such a process may be used in cases where co-current mixing of the solutions is not sufficient to remove the desired amount of C0₂ and/or if the kinetics of the co-current absorption process is too slow.

In the process shown in Fig. 4, a feed gas 401 is inhibited with MEG 402 or other chemical (if required) and transported via pipeline 403 from well-head 400 to the platform or FPSO inlet separator 404 where the liquid phase(s) 405 is separated. The separated gaseous phase 406 is heated in heater 407, for example to a temperature of ~20-50°C), prior to being fed to a conventional counter-current mixer 408 for gas sweetening. An MEA stream 409 is introduced at the top of the counter-current mixer 408 and brought into contact with the gaseous phase 406. MEA with absorbed C0₂ is removed from the base of the mixer through conduit 410. A gaseous/liquid phase 41 1 is discharged from the top of the mixer, cooled in cooler 412 (e.g. to ambient temperature) to remove water by condensation and fed into a separator 413. The gaseous phase 414 from the top of the separator is mixed with a second portion 415 of acid gas absorption compound (e.g. MEA as illustrated in Fig. 4, although alternative compounds as described herein are suitable) and mixed in a co-current mixer 416. The resulting mixture is expanded to a lower pressure, e.g. using a Joule- Thomson valve or a rotating expander 417. Expansion will lower the temperature of the liquid/gas mixture, typically down to -25°C. At this temperature, liquid phases of MEA, water and hydrocarbon condensate will form and can be separated via conduit 418 from the natural gas in a third separator 419, for example using a pressure of 60-150 bar and a temperature of -25°C. The natural gas 420 leaving the third separator will be on

specification for C0₂ content and nearly on specification for water content. Where the purified gas is intended for liquefaction prior to transportation, the remaining water may be removed by an adsorption process (not shown but represented by route 420b).

Alternatively, where the purified natural gas is intended for direct transportation to shore (i.e. without liquefaction), this may be transported without further processing (not shown but represented by route 420a). Fig. 5 illustrates an embodiment of the invention in which the acid gas absorption compound is employed as a hydrate inhibitor and in the removal of C0₂. Here the acid gas absorption compound 502, such as MEA, is added to the feed gas 501 in place of the traditional hydrate inhibitor, such as MEG. In this process the acid gas absorption compound absorbs C0₂ during the mixing and pipeline transportation in pipeline 503 and further C0₂ removal should not be required. By replacing the traditional hydrate inhibitor (e.g. MEG) in the pipeline with the same acid gas absorption compound which is employed in later processing steps, only one chemical needs to be handled. This provides advantages with respect to storage, regeneration, transportation, etc. of the various chemicals used during the production of LNG. At the inlet facility 504 of the FLNG/LNG plant, hydrate inhibitor, water and

hydrocarbon condensate 505 is separated from the natural gas in the inlet separator or slug catcher 506. The amount of liquid formed will depend on the actual pressure and temperature under the receiving conditions. The water saturated natural gas feed 507 exiting at high pressure from the separator (top) is mixed with further amounts of an acid gas absorption compound 508 which is identical to acid gas absorption compound 502 (e.g. MEA, although alternative compounds as described herein are suitable) to compensate for the earlier removal of acid gas absorption compound 502 in condensate 505 and therefore ensure that the concentration of acid gas absorption compound is replenished to a level capable of hydrate inhibition and freezing point depression. The MEA (in this case) and the natural gas are mixed in a co-current mixer 509. Mixing is preferably performed in a single- step process (as illustrated in Fig. 5), although a multi-step process may also be employed for this stage. After the mixing of MEA/water and natural gas, the mixture is expanded to a lower pressure, e.g. using a Joule-Thomson valve or a rotating expander 510. Expansion will lower the temperature of the liquid/gas mixture, typically down to -25°C. At this temperature, liquid phases of MEA, water and hydrocarbon condensate will form and can be separated from the natural gas in a second separator 51 1 (the liquid phase 512 containing MEA, water and hydrocarbon condensate leaves through conduit 512a). In the example shown a pressure of about 60-150 bar and a temperature of about -25°C are employed in the separator although the skilled person will appreciate that other conditions are also applicable. The natural gas 513 leaving the second separator will be on specification for C0₂ content (< 50 ppm) and nearly on specification for water content (about 10-50 ppm). The remaining water will typically be removed by an adsorption process (not shown).

Claims

Claims:

A process for treating a natural gas stream comprising the following steps:

(b) contacting the resulting natural gas stream with an aqueous solution which

comprises an acid gas absorption compound whereby to absorb C0₂ and optionally H₂S, wherein said compound also acts as a hydrate inhibitor;

(d) separating said gaseous phase and said liquid phase in a second gas/liquid

separator.

A process as claimed in claim 1 further comprising:

(e) dehydrating said gaseous phase whereby to reduce its water content to a level of 0.1 -1 ppmv; and

(f) optionally liquefying the resulting gaseous phase to produce liquefied natural gas.

A process as claimed in claim 1 or claim 2 wherein the separated gaseous phase obtained in step (d) has a C0₂ content of 50 ppmv or less, and optionally a water content of 20 ppmv or less.

A process as claimed in claim 1 or claim 2 wherein the separated gaseous phase obtained in step (d) has a C0₂ content of 3 mole% or less and a water content of 1 to 50 ppmv.

A process as claimed in any one of claims 1 to 4 wherein cooling in step (c) is carried out such that the resulting liquid phase has a temperature between about 0°C and about -60°C.

A process as claimed in any one of claims 1 to 5 wherein the acid gas absoprtion compound is present in the liquid phase at a concentration between about 50 to about 90 wt%.

7. A process as claimed in any one of claims 1 to 6 wherein the aqueous solution which comprises said acid gas absorption compound is added at a flow rate of about 2,000 kg/hr or more.

8. A process as claimed in any one of claims 1 to 7 wherein the acid gas absorption compound is at least one primary amine, secondary amine, tertiary amine, amino acid salt, ammonia, sodium carbonate, potassium carbonate, lithium carbonate, primary alkanolamine, secondary alkanolamine or tertiary alkanolamine.

9. A process as claimed in any one of claims 1 to 8 wherein the acid gas absorption compound is at least one alkanolamine of formula (I):

NR¹R²R³ (I) wherein

R¹, R² and R³ are independently selected from hydrogen, C_1-6 alkyl and C_1-6 alkanol, with the proviso that at least one of R¹, R² and R³ is a C_1-6 alkanol.

10. A process as claimed in any one of claims 1 to 9 wherein the acid gas absorption compound is methyl ethanol amine (MEA).

1 1 . A process as claimed in any one of claims 1 to 10 wherein the cooling in step (c) is effected by expansion.

12. A process as claimed in claim 1 1 wherein the expansion involves reducing the

pressure by at least 20 to 40 bar.

13. A process as claimed in claim 1 1 or claim 12 wherein the expansion involves

reducing the pressure to a pressure of about 30 to 150 bar.

14. A process as claimed in any of claims 1 1 to 13 wherein the expansion involves

reducing the pressure to an extent sufficient to achieve cooling to a temperature between about 0°C and about -60°C. 15. A process as claimed in any one of claims 1 to 10 wherein the cooling in step (c) is effected by refrigeration.

16. A process as claimed in any one of claims 1 to 15 wherein the cooling in step (c) is effected by a combination of refrigeration and expansion.

17. A purified natural gas obtained or obtainable by a process as claimed in any of

claims 1 to 16.

18. A liquefied natural gas obtained or obtainable by a process as claimed in any one of claims 2 to 16.

19. Apparatus adapted for performing a process as claimed in any one of claims 1 to 16 comprising:

(a) a first gas/liquid separator arranged to receive a natural gas feed;

(b) a mixer arranged to receive a natural gas stream having a reduced water content from said first gas/liquid separator and to contact said natural gas stream with a solvent which comprises an acid gas absorption compound whereby to absorb C0₂;

20. Apparatus as claimed in claim 19 wherein said mixer comprises at least one co- current mixer.

21 . Apparatus as claimed in claim 19 or claim 20 wherein said mixer comprises at least one counter-current mixer.

22. Apparatus as claimed in any one of claims 19 to 21 wherein said cooling means comprises mechanical refrigeration means.

23. Apparatus as claimed in any one of claims 19 to 21 wherein said cooling means comprises expansion means.

24. Apparatus as claimed in any one of claims 19 to 23 which further comprises a

dehydrator arranged to extract water from the separated gaseous phase obtained in step (d) and/or liquefying means capable of liquefying the resulting purified natural gas to produce LNG.

25. Use of an acid gas absorption compound as a hydrate inhibitor, for example in a process for the production of purified natural gas or the production of LNG.

26. Use as claimed in claim 25 wherein the acid gas absorption compound is as defined in any one of claims 8 to 10.

27. Use as claimed in claim 25 or claim 26 wherein said use is in a process according to any one of claims 1 to 16.

28. Use as claimed in claim 25 or claim 26 wherein said use is in a natural gas pipeline.