EA040163B1 - LIQUIFICATION OF INDUSTRIAL AND HYDROCARBON GAS - Google Patents

Description

Technical field

This invention relates to systems and methods for liquefying industrial and hydrocarbon gases or gas mixtures.

State of the art

Industrial gases such as CO2, H2S, N2, O2, H2, He, Ar, air and other gases, and hydrocarbon gases such as methane, ethane, propane, ethylene and other hydrocarbon gases, or mixtures of gases, are traditionally liquefied using refrigeration cycles based on the well-known Carnot cycle, or turbo-expander cycles. The cryogenic temperatures reached during these commercial processes to allow liquefaction may require complex cascaded refrigeration cycles that are expensive in terms of cost, energy and operating costs.

Accordingly, there is a need in the art for alternative processes for liquefying industrial and hydrocarbon gases or gas mixtures that can be relatively energy efficient, economical, and practical to implement.

The essence of the invention

In one aspect, the invention includes a process for liquefying a gas using a gas liquefaction system, comprising the following non-sequential steps:

a) receiving an inlet gas having an inlet pressure at the stage of gas production;

b) cooling the gas obtained in step (a) in a refrigeration stage comprising an absorption refrigeration circuit with at least one absorption cooler;

c) performing an adiabatic pressure reduction of the cooled gas in the liquefaction stage using a Joule-Thomson valve to liquefy at least a portion of the gas;

d) heating the saturated aqueous ammonia fluid in a rectifier to separate ammonia gas, thereby obtaining an unsaturated aqueous ammonia fluid;

e) supercooling the water-ammonia fluid and circulating it to the top of the vapor absorption column (PAC);

f) condensing the ammonia gas produced in the rectifier and flashing the liquid ammonia to produce chilled ammonia gas for use in at least one absorption cooler;

g) absorbing the ammonia gas withdrawn from at least one absorption cooler with an unsaturated aqueous ammonia fluid in the PAK to obtain a saturated aqueous ammonia fluid supplied to step (d) while maintaining the working pressure in the upper part of the PAK below atmospheric by clamping an evaporative valve that controls the circulation of unsaturated aqueous ammonia fluid to the top of the PAK, thereby reducing the pressure of the suction pump in the pump to remove the saturated aqueous ammonia fluid from the PAK, while maintaining the suction pressure above the effective positive head required at the suction for operation pump, to guarantee the absorption of anhydrous ammonia vapor.

The gas may comprise industrial gas or hydrocarbon gas or any mixture of industrial or hydrocarbon gases. Said process may result in the liquefaction of at least one component of the gas, a portion of the gas, or substantially all of the gas.

In another aspect, the invention may include a gas liquefaction system comprising a production stage for receiving inlet gas at inlet pressure, a refrigeration stage comprising an absorption refrigeration circuit for cooling the produced gas, and a liquefaction stage comprising a Joule-Thompson throttle valve for at least partial liquefaction chilled gas. In one embodiment of the invention, said system may further comprise a compression stage for compressing the gas to the required pressure and a compression energy heat recovery stage for transferring heat from the compression stage to the absorption cooling circuit. In another embodiment of the invention, said system may comprise a gas recirculation stage for recirculating non-liquefied gas components in a low pressure steam recirculation circuit, which additionally cools the compressed and cooled gas, which is then sent to the compression stage.

In one embodiment of the invention, the absorption refrigeration circuit comprises a rectifier and a vapor absorption column.

Brief description of graphic materials

The invention will now be described by way of exemplary embodiments with reference to the accompanying simplified, schematic drawings.

Fig. 1 is a schematic representation of one embodiment of the present invention.

Fig. 2A and 2B together represent a process flow diagram (STP) of one embodiment of the invention in which gas is compressed to a subcritical pressure.

Fig. 3 is a Mollier diagram for carbon dioxide (CO2) used in one embodiment of the present invention. This and other Mollier diagrams depict individual enthalpy-pressure diagrams provided by ChemicaLogic Corporation, USA.

Fig. 4 is a process flow diagram (STP) using the method

- 1 040163 gas liquefaction 2 for low sulfur natural gas at an inlet pressure of 170 kPa, water saturated with 2% CO ₂ and 98% CH ₄ . The liquefaction cycle uses a one-time flash liquefaction application to maintain a pressure of 170 kPa.

Fig. 5A and 5B together are a process flow diagram (STP) using the recording main equipment and data processing of a modified absorption refrigeration cycle. The process diagram shows the key components for a 4-stage NH ₃ refrigeration system, a vapor absorption column (VAC), an unsaturated liquor chiller, waste heat exchangers, a generator, a distillation column, a reflux condenser (reflux condenser), an ammonia condenser, and other ancillary equipment.

Fig. 6A is a Molier diagram for methane (CH ₄ ) used in a liquefaction cycle illustrating one embodiment of the present invention. Fig. 6B is a Molier diagram for methane used in an alternative embodiment of the invention with an optional high pressure feed.

Fig. 7 is a Mollier diagram for anhydrous ammonia (NH ₃ ) which depicts the thermodynamic points for a 4-stage liquefaction-refrigeration system that correspond to the vapor pressure and temperature of the anhydrous ammonia as it flows back into the PAC. The system cooling temperature corresponding to the ambient temperature for this example assumes a condensing temperature of 22°C.

Fig. 8 is a PTX image for an aqueous ammonia solution showing operating points for a key process, in particular operating pressure, temperature and solution concentration in PAA and the remainder of the modified absorption cycle used in the invention. PTX-plot for aqueous ammonia was built using the processed data obtained on the PROMAX™ process simulator.

Fig. 9 is a process flow diagram (PTS) of one embodiment of the invention, in which the final cooling of the liquefied gas takes place in the vaporization heat exchanger of the liquefied gas.

Fig. 10 is a Molier diagram for the liquefaction of air used in one embodiment of the liquefaction cycle depicted in FIG. 9.

Detailed description of the preferred embodiments of the invention

In physics, the term gas refers to a state of matter in which the substance is characterized by excellent molecular mobility and the property of unrestricted expansion. In the context of this document, a gas contains substances that are gaseous at standard temperature and pressure, such as CO ₂ , H ₂ S, N2, O ₂ , H2, He, Ar, air, or hydrocarbon gases such as methane, ethane, propane, ethylene and other hydrocarbon gases, or any mixture of gases. In the context of this document, the term liquefied gas means any gas or mixture of gases that has been liquefied for sale, disposal, or commercial, research, or industrial use.

In the context of this document, the term Joule-Thompson valve or Joule-Thompson throttle valve means a gas valve capable of adiabatically expanding the gas in accordance with the Joule-Thompson effect. Joule-Thompson valves are well known in the art and commercially available.

In the context of this document, the term low pressure separator or LPS means a separation vessel that operates at a certain low pressure and outlet temperature after the Joule-Thompson throttle valve so that the liquefied gas can be removed from the flow path or further processed within the flow path.

In the context of this document, the term high pressure separator or HPS means a separation vessel that operates at the required pressure to cool the gas and is located before the Joule-Thompson throttle valve.

In the context of this document, the term dense phase, referring to any gas or mixture of gases, means the state of a gas resulting from its compression above its crikondenbara, which is the maximum pressure above which a gas cannot exist in the gas phase, regardless of temperature, at a temperature in a range determined approximately by the critical temperature, which is the temperature corresponding to the critical point, which is the combination of pressure and temperature at which the intense properties of the gaseous and liquid phases of a substance are the same, and approximately equal to its cricodentherm, which is the maximum temperature above which natural gas cannot exist in the liquid phase, regardless of pressure. In the dense phase, the gas has a viscosity similar to the gas phase, but may have a density closer to the liquid phase.

In the context of this document, the term non-condensable means any gas that does not liquefy at the operating pressure and temperature of a particular stage or stages for any SND within the flow path.

In the context of this document, the term absorption refrigeration process or PAO means a refrigeration system that utilizes a thermodynamic process known in the art.

- 2 040163 cooling process based on input heat for carrying out the cooling process.

In the context of this document, the term supplementary heat means the heat input to the system, the source of which is any waste heat recovery means, heat exchange medium, electric heaters, or other conventional means of providing heat input to the modified PAO saturated solution heating circuit according to this invention. Additional heat is usually supplied from low temperature heat sources. Low temperature heat means low and medium temperature heat which has a lower exergy density and cannot be efficiently converted in a conventional manner. Although there is no single standard for the temperature range of low temperature heat, it is understood that a heat source with a temperature below 37°C is considered a low temperature heat source, since it is believed that below this temperature, heat cannot be efficiently converted using the Rankine steam cycle. The main low-temperature heat sources are solar thermal energy, geothermal energy and industrial waste heat.

In the context of this document, the term mechanical refrigeration process means a refrigeration system that uses an inlet compression-based thermodynamic refrigeration process known in the art to carry out the refrigeration process.

In the context of this document, the term turbo-expander refrigeration process means a refrigeration system that uses a thermodynamic refrigeration process known in the art based on adiabatic expansion and work recovery for compression as the refrigeration process.

In one aspect, embodiments of the present invention include a system that includes a gas production stage, a refrigeration stage, a liquefaction stage or stages, and a modified PAO that controls the refrigeration stage. In a preferred embodiment, the invention may also comprise a compression stage, a compression energy heat recovery stage or stages, and a gas recirculation stage. One embodiment of the present invention focuses on the use of the potential energy (enthalpy) of the inlet gas stream, and on the heat recovery of the compression energy in the compression stage of the liquefaction process, to improve the overall thermodynamic efficiency of the gas liquefaction process.

In one embodiment of the invention, as schematically depicted in FIG. 1, the invention includes a gas liquefaction system that is combined with a modified aqueous ammonia absorption refrigeration system. The compression energy heat generated from the compression work on the gas or gas mixture to be liquefied can be recovered by using aqueous ammonia to absorb heat from the liquid gas working stream through a heat exchanger. In conventional gas treatment processes, this massively generated low temperature thermal energy is removed to the environment by either an air blower or a water-cooled system. Embodiments of the present invention utilize the heat recovered from compression energy in an absorption refrigeration cycle to provide cooling and the ability to liquefy gases.

The potential energy (enthalpy) available in the gas to be liquefied is directly related to the pressure and temperature of the gas entering the system and is used during pressure-reducing refrigeration processes such as the Joule-Thompson (JT) depressurization process for cooling a gas or gas mixture by auto-cooling due to adiabatic pressure reduction. The Joule-Thompson process is reliable and simple, and is suitable for refrigeration with no practical restrictions on operation within the gas-liquid zone of the phase diagram, and does not require the use of special cryogenic rotating equipment, which is complex, expensive, and has practical restrictions requiring operation. outside the gas-liquid zone of the phase diagram.

With a heat source, absorption cooling systems typically use less than 5% of grid electricity compared to the cooling energy generated by an absorption cooling system. The low temperature heat of compression energy that is released due to doing work of compression on the liquefied gas stream may provide some, all or excess cooling capacity depending on the specific gas liquefaction application and the liquefaction process. In applications where insufficient thermal energy is available to be recovered from the liquefaction cycle, additional thermal energy may be required in the form of other available low temperature waste heat streams and/or other conventional means of heat input to provide the necessary thermal energy to provide the necessary cooling capacity generated by the absorption system. cooling.

The absorption refrigeration system comprises a rectifier that uses thermal energy to separate ammonia from a saturated aqueous ammonia solution and a vapor absorption column (PAC) which, in one embodiment of the invention, allows the refrigeration device to operate at -71° C. at an operating pressure of 10 kPa. PAC design uses thermodynamic principles to eliminate the need for traditional mechanical

- 3 040163 vacuum pumps to achieve the required vacuum operating pressures. The design of the PAA also allows for at least some and, if possible, complete recovery of the heat of solution and heat of condensation energy when the anhydrous ammonia vapor is absorbed into the unsaturated aqueous ammonia at the top of the PAA, and optionally at additional PAA entry points. The concentration and temperature of the solution increase from the top to the bottom of the PAK, while the hydraulic head maintains the aqueous ammonia solution in a supercooled state until the final concentration of the saturated solution is reached. The heat of dissolution and condensation is stored as useful energy in the saturated solution, unlike traditional absorbers, which transfer this energy to a heat sink.

In the production step, the inlet gas stream is compressed or decompressed to the required pressure, which may be above or below the critical pressure of the gas, before starting the refrigeration/liquefaction process. If the inlet gas stream is overpressurized, it can be throttled with a Joule-Thompson valve to initiate the process at a lower pressure. In such cases, there is no compression heat recovery to transfer it to the modified PAO.

In one embodiment of the invention, said method is adapted to liquefy a gas that has an inlet pressure below the critical point for that gas. This method uses a compressor (one or more stages), a PAO-modified compression energy recovery system, one or more Joule-Thompson valves, one or more LPRs, and a recycle gas refrigeration compressor with one or more stages. This method of liquefying the gas reduces or eliminates the need for a second set of refrigeration compressors used in a conventional mechanical refrigeration system such as the Carnot cycle. The liquefied gas acts as a heat transfer fluid when the vapor phase component is recirculated by rapidly opening the Joule-Thompson valve and the liquid phase component is sent to storage. This example of said method can be used to liquefy CO2, H2S, propane or produce C3+ natural gas condensate liquid (NGL) from shallow depths where the temperature required for liquefaction is above -70°C.

For example, this method can be used to liquefy CO ₂ gas, as shown schematically in FIG. 2 in the form of STP and in Fig. 3 in the form of a Mollier diagram. The typical storage range for liquid CO2 is from about -15°C to -29°C. The process can produce liquid CO ₂ at a temperature of about -23° C., at a pressure of about 1600 kPa. CO ₂ is delivered to the inlet at atmospheric pressure and a temperature of about 30°C, which is well below the critical point of the gas. The gas is then compressed in stages through heat exchangers that recover the heat from the compression energy, with the heat exchangers in direct contact with saturated aqueous ammonia to provide all or part of the heat energy needed to power the modified absorption refrigeration system. The compressed CO2 is then cooled by at least one absorption chiller. The thermal energy to feed the absorption refrigeration system is provided by any combination of reclaimed compression energy heat and/or additional heat that can be generated through direct or indirect combustion heat exchange or other available reclaimed waste heat streams with the required temperature and mass transfer conditions.

The compressed and cooled CO ₂ then enters through a Joule-Thompson valve into a low pressure separator (LPS) at such a release pressure and such a release temperature that the CO2 is in a gas-liquid two-phase state, which, in some cases, may occur in a supercooled state. The liquid CO2 can be drained into a storage vessel while the gaseous portion containing any vent gases and/or non-condensable vapors is sent to a recirculation compressor, bleed for ventilation, fuel gas and/or post-treatment as appropriate. The recirculation compressor is part of a recirculation circuit in which the gas portion is introduced into the gas flow path during the compression stage, as shown in FIG. 2 on CO2-11 and CO2-11a.

In another embodiment, the present invention provides a process for liquefying a gas that is produced above the critical point or compressed above the critical point using a compressor (one or more stages), a PAO-modified compression energy heat recovery system, one or more Joule- Thompson, one or more LPCs and a refrigeration compressor for recycle gas with one or more stages. The method utilizes a flow path including, but not limited to, a refrigeration cycle utilizing compression of the gas to substantially dense phase pressure to provide liquefaction by cooling the dense phase gas with any combination of compression energy heat recovery system, absorption refrigeration system and a heat exchanger with a low pressure recycle gas vapor flow of one or more LPCs, one or more Joule-Thompson valves, and a recycle gas refrigeration compressor with one or more compression stages. Depending on the pressure and temperature of the inlet gas entering the liquefaction process, it may be advisable to additionally compress the gas to a dense phase well above the critical pressure and temperature of the gas in order to optimize heat removal (enthalpy change) during the cooling process for a specific gas or gas mixture to be liquefied.

The pressure chosen for the cooling process for a particular gas or gas mixture is directly related to the change in the slope of the isotherm of the gas or gas mixture above the critical point, as shown in the Mollier diagram for pressure versus specific enthalpy. The point at which the slope of the isotherm is vertical (infinite slope) provides the maximum potential for significant heat transfer for a given gas or gas mixture at a given temperature. The actual pressure selected need not be at this point, as a combination of factors must be considered, such as practical pressure and temperature limits for compression and heat exchange equipment, and the minimum temperature available or provided by the absorption cooling system. The slope of the isotherm for a particular gas can be determined from a Mollier diagram (x-axis is specific enthalpy, y-axis is absolute pressure) to help select the optimum pressure to cool a given gas or gas mixture prior to the liquefaction step. This selection process will be further described below.

In the examples of this embodiment, the gas is produced at an inlet pressure corresponding to the desired dense phase pressure, or if the inlet pressure does not match the desired dense phase pressure, the gas is compressed or decompressed to the desired dense phase pressure required for liquefaction. In the case of compression, the heat from the compression energy can be recovered via a heat exchanger and transferred to a saturated aqueous ammonia solution to provide all or part of the heat energy needed to power the modified absorption cooling system. If the thermal energy recovered from the compression heat is not sufficient, additional heat can be provided through any direct or indirect combustion heat exchange or other available recovered waste heat streams with the required temperature and mass transfer conditions.

The inlet gas can be compressed in a single or, if necessary, multi-stage compressor to achieve the required final pressure equal to the inlet pressure of the Joule Thompson valve. In general, in one embodiment of the invention, the outlet temperature of any particular compression stage is limited to about 150 to 160°C, depending on the specifications of the particular compression equipment.

The compressed gas is cooled by at least one and preferably 2, 3 or 4 stages of an absorption chiller to a minimum temperature of -70°C. In one embodiment of the invention, the compressed gas may first be cooled by a low pressure recycle steam stream from the LPC, as discussed below.

The cooled industrial gas or gas mixture is fed through a Joule-Thompson valve to a low pressure separator (LPS) at such a release pressure and such a release temperature that the gas is in the X quality vapor phase within the two-phase region of the gas. In FIG. 6A and 6B X = 0.53 at M17 and M12, respectively.

The liquid flows out of the SND into the storage tank, and the vapor is sent to the recirculating steam stream at low pressure. This recycle stream contains a heat exchanger that initially cools the industrial gas or dense phase gas mixture to the desired temperature before cooling the dense phase gas in an absorption chiller. Therefore, the recycle vapor stream at low pressure is heated to a temperature suitable for inlet to the compression equipment, then compressed in one or more stages until the required dense phase liquefaction pressure is reached, and then combined with the inlet gas stream.

Non-condensable vapors from LIS can be directed to the bleed stream for: ventilation, post-treatment, or as fuel gas, depending on the properties of the particular industrial gas or gas mixture and process application.

It is possible to carry out additional stages of rapid evaporation of the liquid removed from the SND to further reduce the temperature and pressure of the liquefied industrial gas or gas mixture, if necessary, using additional Joule-Thompson valves, SND and compression stages.

For some gases, absorption refrigeration units do not operate at a temperature low enough to allow simple flash evaporation through a Joule-Thompson valve of a dense-phase liquid cooled to a supercooled state at the desired liquefaction end temperature, but provide flash evaporation to the desired end temperature and pressure to a certain quality. vapor X within the phase state for a gas in a gas-liquid diagram. The liquid portion is removed from the LPC and sent to a storage tank, while the gas phase is removed from the separator, while the cold gas phase at low pressure can be used to further cool the warmer dense phase gas stream that was cooled in the last stage of the absorption refrigeration heat exchanger. installation.

In one embodiment of the invention, by cooling the dense phase gas stream, the low pressure recycle vapor stream from the LPC is heated to a temperature approaching the final operating temperature of the absorption chiller. It can then be directed to another heat exchanger which further heats the recycle gas in the compressor circuit to a temperature acceptable for the selected recirculation compressor equipment (-29°C or more in order to use standard ductile iron or carbon steel materials and avoid the need for stainless steel required for cryogenic operations). Once there has been sufficient energy exchange and suitable heating of the low pressure recycle gas, it can be mixed with the inlet gas stream and compressed as described herein.

Depending on the particular application, excess cooling available for other processes may be generated, or additional heat energy may be generated to ensure sufficient cooling capacity is generated by absorption cooling equipment. After the heat recovery of compression energy from the inlet stream and the recirculation stream (mixed stream equals the inlet stream + recirculation stream (Y)) this stream is further cooled using one or more stages of the absorption cooling system until the desired final temperature is reached using the absorption cooling system. The low pressure steam recycle stream from the LPS has a mass flow equal to Y or X/(1-X) times the inlet stream to be liquefied. The mass flow of liquid leaving the LPC is equal to the inlet mass flow of the gas or gas mixture entering the system, minus any fuel gas or bleed to prevent the formation of non-condensable gases under the desired system liquefaction conditions.

The gas liquefaction methods described herein can minimize the need for additional equipment required for traditional refrigeration processes with cascaded multi-stage external cooling processes or mixed refrigeration systems, which are currently used in large-scale LNG liquefaction plants and which require significant costs of electricity and money. funds for construction, and working capital in the process of its operation and operational support. In addition, there is no need for a brazed aluminum heat exchanger (BAHE) and cryogenic rotating equipment.

If more cooled and lower pressure liquefied gas or mixed gas products are required, additional flash stages can be added via a Joule-Thompson valve, which result in additional gas recirculation or bleed streams. Depending on the properties of the gas or gas mixture to be liquefied, it may be advantageous to use one or a combination of steam streams for fuel gas or as an inlet stream to recover non-condensable gases in another liquefaction process at different operating pressures and temperatures that liquefy the non-condensable gas or gas mixtures. One example of this method is suitable for applications at liquefaction temperatures up to 170° C., and is particularly suitable for LNG production or C2+ production from great depths.

With the methods described above, which use dense phase gas, it is possible to cool the gas down to -71° C. before adiabatically expanding, which is sufficient to liquefy the methane. In another embodiment of the invention where the required liquefaction temperature is lower, the invention may include an additional cooling step in which evaporation of the separately liquefied industrial gas further cools the gas to be liquefied. This gas liquefaction process uses a compressor (with one or more stages), a PAO-modified compression energy recovery system, one or more Joule-Thompson valves, one or more LCVs, a refrigeration recirculation compressor with one or more stages, and one or more gas evaporative heat exchangers.

In this embodiment, a liquefied gas is produced using the steps described above and additionally adding the step of using a heat exchanger with an liquefied gas vaporizer to cool another gas in the dense phase from the temperature of the last stage of the absorption refrigeration plant to a temperature low enough so that the cooled gas in the dense phase can be additionally cool with a recirculating steam flow from the LPS to ensure the liquefaction of industrial gas or gas mixture through adiabatic expansion through a Joule-Thompson valve to quality X steam at the required temperature and pressure. If LNG is used in an LPG vaporizer, then for every 0.35 kg of LNG vaporized, approximately one kg of air can be liquefied using the process embodiment shown in FIG. 9 and 10.

Accordingly, stages for alternative embodiments of the invention are similar, but may differ in required operating temperatures, pressures, and heat and material balances for gas liquefaction applications. The concentration of the unsaturated and saturated aqueous ammonia solution and the flow rates depend mainly on the external (heat sink temperature) and the required operating temperature of the final stage of the refrigeration plant. The circulation rate of a given mixture of solutions depends on the total required heat load and the available heat input to the system. The calculation and determination of these parameters is within the competence of a person skilled in the art.

- 6 040163 technicians who need the information of this application.

One of the features of the present invention involves the recovery of a significant amount and, in some cases, all of the heat of solution and heat of condensation energy in PAA, which in traditional configurations of PAA is discharged to the environment or to a heat sink. Another feature of one embodiment of the PAO segment according to the invention is the ability to achieve very strong cooling down to -71°C at the final stage of the refrigeration plant without the need for a rotary vacuum pump, thus providing a simpler and less costly solution to achieve LNG liquefaction with a minimum rotating equipment and, in particular, the absence of rotating cryogenic equipment.

Examples

The following examples are described to depict specific embodiments of the claimed invention and are not intended to limit the claimed invention.

In FIG. 2 and 3, CO ₂ gas is produced at atmospheric pressure at about 30°C and then compressed to a pressure of about 4400 kPa through three compression stages (STG-1, STG-2 and STG-3) with cooling by means of heat exchangers (WHX-1 , WHX-2 and WHX 3). The gas is then cooled first by a recycle steam stream from the final separator (Sep. MP) and then by an absorption chiller (NH3-CH1 (10)). WHX-1, WHX-2 and WHX 3 transfer heat to an ammonia water system to power an absorption refrigeration system.

The cooled CO2 then passes through a Joule-Thompson valve into a separator (Sep. MP) at a release pressure and a release temperature such that the _CO2 is in a gas-liquid two-phase state, which in some cases may occur in a supercooled state. The liquid portion is poured into a storage vessel while the gaseous portion, containing any weathering gases and/or non-condensable vapors, is directed to a recirculating compressor, by-pass stream for: ventilation, fuel gas and/or post-treatment as appropriate;

In FIG. 4 and 5 are STPs depicting a liquefaction process and system for liquefying low sulfur natural gas, while FIGS. 6 is a Mollier diagram for the flow path of natural gas. In the attached table. 1 summarizes the heat and material balance for these examples.

In FIG. 8 shows the flow path of modified PAO and PAA on a PTX plot at down scale pressures up to 10 kPa to depict the flow path in this example. Conventional HRT curves for aqueous ammonia generally do not extend below 100 kPa and do not take into account the operation of PTO systems operating below atmospheric pressure. In FIG. 8 shows the flow path of a 4-stage anhydrous ammonia gas refrigeration system to operate at the top of the PAH at pressures as low as 10 kPa and -71°C. In table. 1 summarizes the properties of gas, methanol, aqueous ammonia and anhydrous ammonia as they pass through the flow path, obtained from the available Mollier diagrams for methane and anhydrous ammonia, publicly available tables, graphs and diagrams with thermodynamic properties of aqueous solutions of ammonia, steam and steam tables. Calculations of expected performance and operating parameters for the modified PAO and PAK were made by the inventor as part of the invention. Manual calculations are subject to rounding, simplification, estimation and approximation, which is necessary to obtain key parameters and key operating parameters of the system. For example, non-condensable gases were excluded, and anhydrous ammonia was considered pure to simplify the necessary manual calculations (compared to a target purity of 99.5%). Mathematical modeling with process simulation software can lead to refinement of the calculations to provide a detailed process design of the necessary bypass flows for the gas liquefaction loop and the modified PAO and PAK system.

The process for liquefying water-saturated sweet natural gas (98% CH ₄ and 2% CO2) is shown in FIG. 4, 5 and 6A. Natural gas is fed into the flow path through the inlet separator at M1 at a pressure of 170 kPa and 17°C. The gas is compressed at the inlet in the 1st stage (COMP-IN) and compressed to 650 kPa (M2), the same pressure as the recycle gas in the first stage (STG-1, M3). The compression heat from the 1st stage inlet is recovered in the WHX-IN (M2 to M2a), the compression heat from the 1st stage inlet (M3 to M3a) is used to heat the recycle gas (M20 to M21) to at least -29° C, which is the minimum acceptable temperature for operation in a compressor constructed from standard materials (non-cryogenic). The combined inlet temperature of the 2nd stage recirculation compressor is 47°C (M4). The temperature and pressure on the M4 should be checked to ensure that hydrates or freezing are not a problem; this is not a problem in this example, but the intake gas recirculation ratio and water content may vary depending on the application.

The combined intake and recirculation gas is compressed in the 2nd recirculation stage to 2200 kPa (M5), the gas is cooled and the compression heat is recovered in the WHX-1 (M5 to M6). The gas is further compressed in the 3rd recirculation stage (STG-3) and the compression heat is recovered starting at 160°C in WHX-2 (M7 to M8) up to 47°C.

The gas now enters the flow path for pre-treatment in preparation for the liquefaction process. HSX-5 is used to provide temperature control before the gas travels through

- 7 040163 current enters the amine contactor (M9), in which the CO2 content decreases from 20,000 ppm. to less than 50 ppm to prevent solidification of CO ₂ during the liquefaction process. The gas leaving the amine contactor at point M10 is saturated with water when it enters the TEG-glycol dehydrator, in which the water vapor content was reduced to 0.065 kg/10 ³ m ³ when leaving point M11. At point M12, 11.7 kg of methanol is injected to ensure that there is approximately a 75/25 methanol/water mixture when condensation begins along the flow path to the HPS at 8200 kPa and -88° C. (M14). The condensed mixture of methanol and water is removed from the HPS (M16), leaving a dehydrated steam stream with trace amounts of MeOH/H ₂ O, which will solidify as the gas evaporates through the J-T-1 valve to 170 kPa and -152°C, and can be removed by filtering the liquid product. This dehydration method is not used or recognized in the prior art as a method for dehydrating a gas stream during the production of cryogenic gases such as LNG, and is a method that can be used in the invention to eliminate the need for traditional molecular sieve dehydration methods.

Returning to M12, the methanol injection point along the flow path, the gas is sequentially cooled from 22°C (M12) to -88°C (M14) by means of ammonia refrigeration units NH3-CH1, NH3-CH2, NH ₃ CH3, NH3-CH4 and GGX -2. In this example, due to the higher ambient temperature, the HSX-3 does not provide any preferential heat transfer. In colder climates, HSX-3 can provide significant cooling, which can reduce the cooling function of NH3-CH1. As a result, the thermal efficiency of the gas liquefaction process will increase as the outside temperature drops during the winter months.

After the cooled dense phase gas reaches the HPS, the condensed MeOH/H ₂ O liquid is removed at M16 as previously stated, the high pressure dehydrated cooled gas leaves the HPS at M15 (-88°C, 8200 kPa) and evaporates through the Joule-Thompson J-T1 valve, passing further through point M17 (170 kPa, -152°C and vapor fraction X = 0.53) in the SND. Liquid LNG is removed from the LNG by an M-18 by gravity into the LNG storage system (whereby trace MeOH/H ₂ O solids are filtered and removed from the LNG) and cold recycle gas vapor is recycled back to act as a heat transfer fluid, by cooling the gas stream in GGX2 (from M13 to M-15) and heating from M-19 to M20 (from -152°C to -71°C), similar temperature drops are obtained using a cryogenic high pressure heat exchanger. The recycle gas is then further heated in the GGX-1, a low pressure cryogenic heat exchanger, to -29°C to allow the use of non-cryogenic compression equipment, which can be either reciprocating or centrifugal as the gas liquefaction plant grows.

A saturated solution is obtained at the inlet of the pump for a saturated aqueous ammonia solution at point Aq1 at the bottom of the PAK in a supercooled state at 50°C or less and 10.4 wt.% in this application. Warmer external conditions compared to cold winter external conditions lead to a reduction in the total amount of saturated and unsaturated solutions used for modified absorption PAO. In this example, the concentration for a saturated solution is 5 wt.%, and for unsaturated - 10.4 wt.%.

PAH fed with anhydrous ammonia vapor from four gas refrigeration units (NH3-CH1, NH3-CH2, NH ₃ -CH ₃ , NH3-CH4), in this example operating at 10 kPa at the top and an unsaturated solution subcooling temperature of 22 °C.

In general, the lower the outside temperature or the more heat sinks available, the smaller the NH3-CH1 cooling function becomes, in this example, since the HSX-3 cannot lower the flow path temperature in M12a below 22°C, its function is minimal. The lower outside temperature also results in a lower operating temperature for the HSX-2 condenser (anhydrous ammonia condenser). The effect of lowering the outside temperature on the modified PAO and PAA is shown graphically in FIG. 8 (PTX diagram for absorber operating pressure up to 10 kPa). The lower ammonia condensing pressure (HSX-2) and reduced functional load on the unsaturated aqueous ammonia solution (NH3-CH1) provide the opportunity for further optimization of the concentration and circulation rate of the saturated and unsaturated solutions.

Returning to the saturated aqueous ammonia pump, the pump outlet pressure is a direct function of the condensing temperature (and pressure) of the ammonia condenser (HSX-2). In this example, a pressure of 950 kPa is required for Aq-2, at this point in the flow path, a 10.4 wt% saturated solution is subcooled. First, the saturated solution enters the compression heat recovery stage, splitting in parallel with the separation of flow rates proportional to the waste heat recovery function of each exchanger (WHX-IN, WHX-1, WHX-2), while the temperature rises from 50°C (Aq2) to 72.5°C (Aq10) at 10.4 wt.% and 940 kPa, while the saturated solution is still supercooled.

The next point on the flow path is the saturated/unsaturated solution exchanger where the saturated solution is further heated to 143°C and at point Aq12 the saturated solution enters the PAO distillation column.

As a result of the operating pressure at a condensing temperature of 22°C (HSX-2), it was calculated that

- 8 040163 The modified PAO system in this example requires 924 kW of additional heat capacity, which can come from the available low temperature reclaimed waste heat streams, but a final temperature of 159°C is required to achieve an unsaturated solution concentration of 5 wt.%. Additional waste heat can be fed directly to the generator/surge tank or through the saturated solution heat exchanger heating circuit.

After providing sufficient additional heat in the distillation column, the necessary reflux and steam movement is achieved. The DPX reflux condenser requires 436 kW of power to cool and achieve a temperature of 50°C at the outlet, resulting in an ammonia stream that is assumed to be 99.5 wt.% ammonia based on the assumed reflux ratio of 2 and the concentration of a weakly saturated solution of 5 wt.% (Aq14 ).

The unsaturated solution Aq14 is supercooled in the exchanger for unsaturated and saturated solution, and the temperature is reduced from 159 to 85°C (Aq15). In this example, the unsaturated solution is further cooled in HSX-1 to 22° C. at point Aq16 of the flow path. A subcooled 5 wt. % unsaturated solution is injected into the top of the PAA column, approximately 10.6 m above the suction side of the pump saturated with Aq. The unsaturated aqueous ammonia solution at 5 wt.% is supercooled at 22°C to ensure that ammonia from NH3-19 enters the flow path (at -71°C and 10 kPa) to completely dissolve in the supercooled unsaturated solution and to remain in the supercooled state after taking into account the increase in temperature due to the heat of the solution and the heat of the condensation energy, as well as the mixing of the enthalpy of ammonia vapors and the unsaturated solution.

An operating pressure of 10 kPa is obtained by wedging the valve to quickly admit unsaturated aqueous ammonia, thereby lowering the pump suction pressure in the saturated aqueous ammonia pump, but maintaining the suction pressure above NPSHR and subcooled unsaturated solution to ensure absorption of anhydrous ammonia vapours.

Returning to the point in the flow path where the ammonia vapor (NH3-1) leaves the top of the DPX, the vapor continues to the ammonia condenser (HSX-2). It is the condensing temperature of said heat exchanger that sets the operating pressure for the PAO saturated solution side. In this example, the HSX-2 consumes 230 kW to condense the required ammonia vapor stream.

After NH3-2, the ammonia is completely condensed, and in FIG. 7, the Mollier diagram for anhydrous ammonia details the thermodynamic aspects of this part of the invention. The ammonia is at a condensing pressure of 900 kPa and a temperature of 22° C. and is quickly converted to the appropriate pressures for 4 refrigeration units, as depicted in the STP for the modified PAO and PAA aspects of the invention. As shown in FIG. 5, NH3 bleed valves 1, 2, 3, and 4 are associated with the ammonia chillers. In this example, bleed flows are not used in this example to simplify manual calculations. However, in practice, approximately 5% bypass flow for each ammonia chiller may be required to prevent build-up of H ₂ O in the ammonia chillers, which can lead to system failure. The actual bypass stream depends on the purity of the ammonia produced by the distillation column, which in this example was set to 99.5% purity.

At the top of the PAA, the ammonia entering the PAA is at a height, temperature, and mass flow rate that results in an aqueous ammonia solution with increasing concentration and temperature as the solution flows down the PAA. As depicted in FIG. 8, in this example, the PTX diagram of the solution remains supercooled for streams of NH3-19 (Aq1 7 24.6°C, 10 kPa, 5.5 wt%), NH3-15 (Aq 18 28.1 °C, 13 kPa, 6.2 wt.%), NH3-11 (Aq19 34.1°C, 30 kPa, 7.3 wt.%) and NH3-6 (Aq20, 49.8°C, 72 kPa , 10.4 wt.%) for four refrigeration units of the gas flow path. If an overheated solution occurs at the lowest mixing point (Aq20), an HSX-4 heat exchanger can be used to remove excess heat to supercool the saturated aqueous ammonia before being sucked into the pump (Aq1) to maintain the required working pressure at the top of the PAH.

In another example, shown in FIG. 8 and 9, which depicts the liquefaction of air, the intake gas enters at a pressure below the critical point. Liquefied air is produced using an liquefied gas evaporator to provide additional cooling in the flow path after the last stage absorption chiller (which operates at -70°C) to ensure that temperature and pressure conditions are reached that cause the vaporized gas or the gas mixture has the required temperature and pressure corresponding to the gas-liquid phase diagram at a certain quality X.

For example, natural gas can be liquefied using the methods described above, and then the LNG can be vaporized to provide additional cooling to the air stream in addition to the cooling provided by the last stage of the refrigeration plant of the modified absorption refrigeration system. The vaporized natural gas can then feed into an LNG liquefaction loop using an alternative embodiment as described above, or a gaseous fuel source if the air liquefaction plant is located at the location of the LNG fuel source. This method may be suitable for liquefying a gas requiring very low liquefaction temperatures (below -170° C.) and minimizing the additional equipment required for traditional refrigeration processes with external refrigeration cascaded multi-stage processes.

Definitions and interpretation

All references to temperature and pressure in this description should be taken as modified by the term about, which means a deviation of ±5%, ±10%, ±20% or ±25% from the specified value. For example, about 50% in some embodiments of the invention may include a deviation from 45 to 55%. In the case of numeric ranges, the term about may include one or two numbers greater than and/or less than the specified numeric value on each side of the range. Unless otherwise indicated herein, the term about is considered to include meanings and ranges close to the specified range, which are equivalent in the context of the functionality of the described composition or embodiment of the invention. The term near can also represent any inaccuracy of the instruments, devices, or methods used to measure the specified value.

As will be understood by those skilled in the art, various modifications, adaptations, and variations of the above description may be made without departing from the scope of the invention claimed herein. The various elements of the invention described herein may be combined in a manner other than the specific examples described or claimed herein without departing from the scope of the invention. In other words, any element or feature can be combined with any other element or feature in different embodiments of the invention, unless there is an obvious or inherent incompatibility between them, or unless they are explicitly excluded.

The singular form includes plural references unless otherwise clearly indicated by the context. Thus, for example, a reference to a setting includes a plurality of such settings. Additionally, it should be noted that the claims may be drafted so as to exclude any optional element. Therefore, this statement should serve as the basis for the use of exclusive terminology, such as exclusively, only, etc., in connection with the enumeration of the claimed elements or the application of a negative limitation. The terms preferred, preferred, prefer, optional, may and similar terms are used to indicate that the subject, condition or step referred to is an optional feature of the invention.

The term and/or means any of the items, any combination of items, or all of the items to which the term is associated. The expression one or more is clear to a person skilled in the art, in particular when used in the context of its application.

As one skilled in the art will appreciate, all numerical values, including those expressing amounts of reactants or ingredients, properties such as molecular weight, reaction conditions, and so on, are approximations and are considered to be optionally modified in all cases by the term about. These values may vary depending on the desired properties expected to be obtained by those skilled in the art using the teachings of the description given herein. It will also be understood that such values naturally contain variability necessarily arising from standard deviations at the time of the respective measurements.

As one skilled in the art would understand, for any and all purposes, and particularly in the context of providing a written description, all ranges listed herein also include all possible subranges and combinations of subranges, as well as the individual values that make up the range, in particular, integer values. Listed ranges (eg, mass percentages or carbon groups) include each single value, integer, decimal, or identity within the range. Any given range may be considered to be adequately descriptive of and inclusive of the same range, divided at least half into equal halves, thirds, quarters, fifths, or tenths. As a non-limiting example, each range discussed herein can be easily divided into lower third, middle third, and upper third, and so on.

As will also be appreciated by one of ordinary skill in the art, all expressions such as up to at least more than, less than, more, or more, etc., include the number indicated, and such terms refer to ranges that may subsequently be be broken down into subranges as discussed above. Likewise, all relationships listed herein also include all intermediate relationships included in the broader relationship. Accordingly, the specific meanings given for radicals, substituents, and ranges are for illustrative purposes only; they do not exclude other defined meanings or other meanings within the defined ranges for radicals and substituents.

One skilled in the art will also appreciate that when elements are grouped together

- 10 040163 in a general way, as in the case of the Markush group, the invention includes not only the entire group, given as a whole, but also each representative of the group separately and possible subgroups of the main group. Furthermore, for all purposes, the invention includes not only a main group, but also a main group with one or more members of the group missing. Therefore, the invention expressly excludes any one or more members of the above group. Accordingly, a condition may be applied to any of the disclosed categories or embodiments, implying that any one or more of the listed elements, types or embodiments may be excluded from such categories or embodiments, for example, as in the case of applying an unambiguous negative restriction .

Table 1

Named □opzha-Szhnipel-Kovtur mtssha M1 M2 M2a M3 MZA M4 M5 MB M7 M8 me M10 M11 Pressure (nPa) 170 650 600 650 600 600 2200 2100 8600 8520 8500 8450 8400 Temperature (°C) 17 123 65 63 32 47 157 47 160 47 22 22 22 Vapor density (pUm ³ ) Liquid density (ng/s ³ ) 102 33 3 25 33 4 37 10 5 12 5 38 57 61 61 61 Shares of steam MeOH (% mass from liquid; 1 0 1 0 1 0 1 0 1 0 1 0 1 0 1 0 1 0 1 0 1 0 1 0 1 0 Enthalpy (kJ/kg) 892 1140 EE5 1005 320 355 1220 940 1200 885 820 820 820 Approximate mass in kt with 28D e3sh3; d-drivers C1 (0.98 dry vol.) (zhg / d) 18500 18500 18500 20860 20860 33360 39360 39360 33360 33360 39360 39360 39360 Сot (0.02 mol gute vol.) (nTsp) 575 575 575 0 0 575 575 575 575 575 575 05 05 Ht O (pyr) (kg/d) 250 250 250 0 0 250 250 250 250 18 18 18 33 H ₂ O (liquid) (nt/d) 0 0 0 0 0 0 0 0 0 232 0 0 0 MUN(kg,'d) 0 0 0 0 0 0 0 0 0 0 0 0 0 Estimated total mass flow (kg/d) 13325 13325 13325 20860 20860 40185 40185 40185 40185 40185 33353 33373 33364 mass notes (nUs) 0 224 0 224 0 224 0 241 0 241 0 465 0 465 0 465 0 465 0 465 0 462 0 456 0 456 Name pvizha-Szhihitel-Koitur melka M12 M12a M12b M12s M12d M13 M14 1415 1416 M17 M18 M1E M20 1421 Pressure (kPa) 8400 8300 8275 8200 8225 8200 8200 8200 8200 200 200 200 185 170 Temperature (O.Q. 22 22 -33 -54 -64 -70 -88 -88 -88 -152 -152 -152 -70 -23 Vapor density (i / m ³ ) 61 61 105 155 205 245 300 300 3 3 33 16 13 Liquid density (noise ³ ) 245 300 800 410 410 Steam share 1 1 1 1 1 0 333 0 333 1 0 053 0 1 1 1 MeOH (% May of liquid) 0 0 0 0 0 0 75 0 75 0 75 0 75 0 0 0 0 0 'Enthalpy (kJ/kg) 820 820 610 500 435 385 230 285 285 thirty 525 705 730 Approximate weight in kg with 28J e3w3 ^-BHTSCHSH C1 <0.98 mol d.v.) (kg/d) 33360 33360 33360 33360 33360 33360 33360 39360 0 33360 18500 20860 20860 20860 CO , (0.02 min dry vol.) (nTd) 05 05 05 05 05 05 05 05 0 05 05 0 0 0 H _T O (steam) (kg/d) 33 33 33 33 33 0 04 0 04 0 04 0 0 04 0 04 0 0 0 P?O (liquid) (iud) 0 0 0 0 0 3 86 3 86 0 3 86 0 0 0 0 0 MeOH(zhg/d) 117 117 117 117 117 117 117 012 11 58 012 0 12 0 0 0 Estimated total mass flow (kg/d) 33376 1 33376 33376 33376 33376 33376 33376 33361 15 44 39361 18500 66 20860 20860 20860 mass flow (kTus) 0 456 0 456 0 456 0 456 0 456 0 456 0 456 0 456 0 000 0 456 0 214 0 241 0 241 0 241

I The name of the stream is Abe o rb. o1ladzh.Ats-N A _H 1 Aq2 Aq3 Aq4 Aq5 Aq6 Aq7 Aq8 A _H E Aq1O Aq11 Aq12 Aq13 Aq14 Pressure (kPa) 110 350 350 350 350 350 340 340 940 340 330 300 300 310 ^Temperature (°Q 50 0 50 0 50 0 50 0 50 0 50 0 72 5 72 5 72 5 72 5 143 0 50 0 50 0 160 0 GGeypm; sh ura 122 0 122 0 122 0 122 0 122 0 122 0 162 5 162 5 162 5 162 5 283 4 122 0 122 0 320 0 (Supercooled Yes Yes Yes Yes Yes Yes Yes Yes Yes Yes Yes Yes Yes Yes p4wt. solution to water. ΝΗ3 0104 0104 0 104 0 104 0104 0 104 0104 0104 0104 0104 0 104 0 560 0 050 (Density of solution to water ΝΗ3 ’Obshchag yatalysh ved mixes (kJDg) (note 1) 337 128 126 126 126 126 126 225 225 225 225 527 7 638 The general evtalytnya of the Vedas. blends (Eta/lb) (straight G 54 54 54 54 54 54 38 38 98 38 227 3 275 Approximate! mass in kg s 28 D eSwL Mass of aqueous ammonia solution (kg/s) 'y-vkshshi I 31715 31715 15030 16685 1 3350 0 3330 0 3330 13350 1 5030 31715 31715 2 9900 Mass of steam things, ammonia (kg/s) 0 0000 0 0000 0 0000 0 0000 0 0000 0 0000 0 0000 0 0000 0 0000 0 0000 0 0000 0 0000 Total mass of water, ammonia (kTus) | 31715 31715 15030 16685 1 3350 0 3330 0 3330 13350 1 5030 31715 31715 2 3300

The stream name is Absorb. <plada.-Kvyatur Aq-NH3 Aql5 M16 AC7 Aql8 Aql9 Aq2O Dailtie (kPa) zed 880 10 13 thirty 72 Temperature (°C) 850 22.0 24 5 281 341 49.8 Temperature (°F) 1850 71 6 7BZ 82 6 93 4 1216 Changed the burnt Saturated Yes Yes Yes Yes Yes Yes % mass, water solution. NH3 Leading solution density. NH3 0050 0.050 0055 0.062 0073 0104 General endadtia schsch. mix (vJ/kg) (note 1} 319 35 58 39 74 125 Obprya enthalpy of water. blends (Btu/lb) (remark 1} Approximate mass in kg with 28E eZtpZ/d-ntod 137 )M 15 25 17 32 54 Mass of dispersal water, ammonia (kg/s) 29900 29900 3 0058 3 0274 3 0B63 31715 Ammonia mass (kg/s) 00000 0 0000 00000 00000 0.0000 0 0000 Total amount of ammonia padl (kg/s) 29900 2.9900 3 0058 30274 3.0663 31715

The name of the pyagoka is Absorb. <ilaj.-Kshur ΝΗ3 ΝΗ3-1 ΝΗ3-2 ΝΗ3-3 MH 3-4 MH 3-5 ΝΗ3-6 ΝΗ3-7 ΝΗ3-8 MNZ-9 MNZ-10 GShZ-P MNZ-12 NH3-13 NH3-14 Daileyave (kPa) 900 900 72 72 72 72 72 thirty thirty thirty thirty thirty 13 13 Maintenance temperature 500 22.0 -400 -40 0 -400 -400 -40.0 -55.0 -55.0 -550 -550 -55.0 -65.0 -65.0 Temperature (°F) supercooled 1220 71 6 -400 -40.0 -400 -400 -400 -670 -B70 -670 -670 -67 0 -85 0 -85 0 Saturated superheated Yes Yes Yes Yes Yes Yes Yes Yes Yes Yes Yes Yes Yes Yes Steam share of pure ammonia (95.5% wt.) 1 0 02 0 35 1 1 0 0.05 0.098 1 1 0 0.06 1 Vapor density (kt / m ³ ) Liquid charge (kg / m ³ ) 57 610 0.61 B85 0.61 685 061 061 685 0 27 710 0.27 710 027 0 27 710 0.13 715 0.13 Entalytnya ΝΗ3-Ι (kJ/kg, from Malier chart) 440 160 160 160 95 95 95 50 Eptalpy NH3-V (kJ/kg, from Mbpier chart) 1705 1550 1550 1550 1550 1525 1525 1525 1525 1510 1510 Siorr. enthalpy ΝΗ3-Ι (kJ/kg) -193 S 92 -188 -188 -188 -253 -253 253 -298 Siorr. enthalpy Nffl-ν (kJ/kg) -1938 1357 1202 1202 1202 1202 1177 1177 1177 1177 1162 1162 Approximate mass in kg with 283 u3t3/d-iodine! Weight ΝΗ3-Ι (kg/s) )M 0.1815 01452 00B89 0.0000 0.0000 00763 00725 010351 OOMYU o.oyu 0.0374 0 0203 000W Mass of NH3-V (kg/s) 01815 ΟΌ363 00363 01052 01052 0.0000 OSO38 00038 0.0389 00389 0 0000 0.0013 0Ό216 Total mass ΝΗ3 (kg/s) 0.1815 01815 01815 01052 0.1052 01052 00763 00763 00389 OOZV9 0.0389 00374 0.0216 0.0216

- 11 040163

Nkzvanne then-Absorb. oido-Koshur NB3 [Davzhyaye (k11e)

Temperature (OC)

[Saturated

superheated

Share of pure ammonia

Steam density (pUm ¹ ) ³

Liquid density (eg/m ³ )

Enthalpy NH3-I (EJ/kt; from Mollier chart) Eptatry ΝΗΤ-ν (EJ/p; from Mollier chart) Swarr. eeplpshTPV! (EJ/Eu) -1938 (Note 2)

Croppi enthalpy Nffi-ν (vJ/kg) -1938 (approx: 2)

Printed weight in kg with 2 8.3

Mass of NHJ-I (kg/s)

Mass of NH3-V (kg/s)

Total mass of NH3 (kg/s)

NH3-E NHM6 NHJ-17 NH3-18 NH349

Yes

0.13

1510

1162 yes o

715

-296 yes

0016

0.1

725

1500 -316 1152

1010

-71-71 dada

0.101

1500 1500

1152 1152

00000 0 0156 0 0155 0 0000 0 0000

0.0216 0 0000 0 0000 0 0158 0 0158

00216 00158 00155 00158 00158

Note. 1-Enthalpy of di vodi, ammonia w 1938-Jenningsand Shannon Tables - Lehigh University Bethlehem, Pennsylvania

Note. 2 - (down to zero (eJ/kg) Mollier diagram for anhydrous ammonia and tab. cpn. ammonia 1938 - 348

Equipment name RAqPump DPX AqRLX Trim-Heat WHX-1 WHX-2 Total mass flow (kg/s) Enthalpy change (EJ/p Heat in/vppc (kW) Work in/out (kW) 5.172 1 10 -456 2990 -519 -954 924 0465 -280 -150 0465 -515 -147

Name of equipment Total mass flow (kg/s) Engagement change (EJ/kg Ta™in/yield(kW) Work nvd/vnod (kW) 1 WHX-M 0.224 -145 52 COMP-IN 0.224 248 55 STG-1 0.241 215 52 STG-2 0465 265 125 STG-3 0.465 260 121

Equipment name Total mass flow (kg/s) Enthalpy change (kJ/kg) Heat input/output (kW) Work output (kW) HSX-1 2.990 -284 -849 H8X-2 0.182 -1265 -250 HSX-3 0.456 0 0 HSX-4 5172 0 0 HSX-5 D.462 -65 -thirty

Equipment name Total mass flow (kg/s) Change enthalty (kJ/kg Heat inf/exp (kW) Work inwd (kW) NHS CHl 0456 -210 -96 NH5-CH2 0456 110 50 NH3-CH5 0 455 -65 -thirty NH3-CH4 0456 -50 -23 GGX-1 0 241 85 21 GGX-2 0 241 180 45

CLAIM

Claims (17)

1. A method of gas liquefaction using a gas liquefaction system, including the following non-sequential steps: (a) receiving an inlet gas having an inlet pressure in the stage of gas production; (b) cooling the gas obtained in step (a) in a refrigeration stage comprising an absorption refrigeration circuit with at least one absorption cooler; (c) adiabatically reducing the pressure of the cooled gas in the liquefaction stage using a Joule-Thomson valve to liquefy at least a portion of the gas; (d) heating the saturated aqueous ammonia fluid in a rectifier to recover ammonia gas, thereby obtaining an unsaturated aqueous ammonia fluid; (e) subcooling the unsaturated aqueous ammonia fluid and circulating it to the top of the vapor absorption column (PAC); (f) condensing the ammonia gas produced in the rectifier and flashing the liquid ammonia to produce chilled ammonia gas for use in at least one absorption cooler; (g) absorbing the ammonia gas removed from at least one absorption cooler with an unsaturated aqueous ammonia fluid in the PAK to obtain a saturated aqueous ammonia fluid supplied to step (d) while maintaining the operating pressure in the upper part of the PAK below atmospheric using pinching the evaporator valve that controls the circulation of the unsaturated aqueous ammonia fluid to the top of the PAK, thereby reducing the pressure of the suction pump in the pump to remove the saturated aqueous ammonia fluid from the PAK, while maintaining the suction pressure above the effective positive head required at the suction for pump operation to ensure the absorption of anhydrous ammonia vapor. 2. The method according to claim 1, while the gas is partially liquefied in step (c), further comprises the steps of removing the liquefied gas product and recirculating the remaining gas in a steam recirculation loop, which cools the gas stream before adiabatically reducing the pressure, after which the gas is compressed and mixed with the inlet gas stream. 3. The method according to claim 1, characterized in that the inlet gas contains industrial gas or hydrocarbon gas, or any mixture of industrial or hydrocarbon gases. 4. Method according to claim 1, characterized in that the gas in step (a) is further compressed below the critical point of the gas, whereby a gas liquefaction temperature of greater than -71 °C is reached before adiabatic expansion. 5. Method according to claim 1, characterized in that the gas in step (a) is further compressed above the critical point of the gas, whereby a gas liquefaction temperature of about -71 °C is reached before adiabatic expansion. 6. The method of claim 1, wherein the at least one absorption cooler comprises a liquid stream outlet to prevent an increase in the concentration of water in the ammonia refrigerant. 7. The method according to claim 1, characterized in that the operating pressure, which is below atmospheric, in the upper part of the PAK is maintained by using a sufficient mass of supercooled unsaturated aqueous ammonia fluid at the point of mixing of anhydrous ammonia vapor and unsaturated aqueous ammonia fluid to allow gaseous ammonia to completely dissolve in the supercooled water-ammonia fluid in PAA. 8. The method according to claim 1, characterized in that some or all of the heat of dissolution and heat - 12040163 the condensation energy of mixing anhydrous ammonia vapor and unsaturated aqueous ammonia fluid is recovered in the PAK. 9. The method according to claim 8, characterized in that heat recovery in PAK is provided by hydrostatic head and pre-cooling of the raw stream of unsaturated aqueous ammonia PACK. 10. The method of claim 1 further comprising the steps of cooling the cooled gas stream from step (b) using gas/gas heat exchangers to further reduce the temperature of the gas stream prior to step (c). 11. The method of claim 1 further comprising dehydrating the gas after compression and before entering the absorption refrigerator. 12. The method according to claim 11, characterized in that the gas is dehydrated by adding a sufficient amount of alcohol to the flow path and subsequent condensation of alcohol and water is carried out before the adiabatic depressurization step. 13. The method according to claim 1, characterized in that the liquefied gas is filtered after an adiabatic pressure reduction to remove any solid particles from the liquefied gas product. 14. The method according to claim 1, further comprising the following non-consecutive step: cooling the second gas using a heat exchanger with a liquefied gas vaporizer using liquefied gas. 15. A gas liquefaction system for carrying out the method according to claim 1, comprising a production stage for producing inlet gas, a refrigeration stage containing an absorption refrigeration circuit for cooling the gas, and a liquefaction stage containing a Joule-Thomson valve for at least partial liquefaction gas, while the absorption cooling circuit contains: (i) at least one absorption refrigerator for cooling the gas; (ii) a rectifier for releasing gaseous ammonia from a saturated aqueous ammonia fluid to obtain an unsaturated aqueous ammonia fluid; saturated water-ammonia solution; (iv) an evaporator valve to control the circulation of unsaturated aqueous ammonia fluid from the rectifier to the top of the PAK; and (v) a pump to remove the saturated aqueous ammonia fluid from the PAK; wherein the evaporator valve is configured to quickly inlet unsaturated aqueous ammonia fluid to reduce the pressure of the suction part of the pump to remove the saturated aqueous ammonia fluid, and at the same time maintain the pressure of the suction part above the effective positive head required at the suction to operate the pump, which removes the saturated aqueous ammonia fluid to maintain subatmospheric pressure at the top of the PAK. 16. The system of claim 15 further comprising a compression stage for compressing the gas and a compression energy recovery stage for transferring heat from the compression stage to the absorption refrigeration circuit. 17. The system of claim 16, further comprising a gas recirculation stage for recirculating non-liquefied gas components in a low pressure vapor recirculation loop, said loop further cooling the compressed and cooled gas, which is then sent to the compression stage.