SPLIT FLOW INTEGRATED LNG PRODUCTION (SFI-LNG)

Abstract

Processes for purifying and liquefying natural gas in conjunction and integration with cryogenic processing natural gas to recover natural gas liquids (NGL) is disclosed. In the process, the natural gas stream to be purified and liquefied is taken from top outlet stream of demethanizer in the cryogenic NGL recovery plant, first purified and then cooled under moderate pressure to condense it as a liquefied natural gas (LNG) product stream. Some of the cooling required for the demethanizer reflux stream is provided by natural gas liquefaction section before supplied to top of the column to serve as reflux. The top outlet stream from the demethanizer preferentially contains up to 3 mole percent of CO.sub.2 and the majority of methane and small portion of any hydrocarbon heavier than methane, a split portion of this stream is taken and routed to cryogenic CO.sub.2 removal section, in which a molecular sieve that forms a physical adsorption column is used to extract pure CO.sub.2 as a product stream, then purified stream is routed to the liquefaction section where only two stages of coil-wound exchangers with a Semi-C3-MR cycle are used to liquefy natural gas. This present invention process is suited for LNG production in small-scale. This zeolite-based small-scale LNG process can be integrated with the design of any new natural gas facility and the technology can also be retrofitted to existing natural gas liquid (NGL) recovery plants, allowing for co-production of LNG and CO.sub.2 with high purity.

Claims

1. A process for removing CO.sub.2 from a cryogenic gas stream (213 K and cooler) comprising dry sweet methane having a CO.sub.2 concentration of up to 3 mole percent, the process comprising adsorptive carbon dioxide removal on a molecular sieve at a pressure of 1000-2000 kPag, wherein the molecular sieve is a zeolite, to obtain a carbon dioxide product stream and scrubbed natural gas stream, wherein the scrubbed natural gas stream has a carbon dioxide content of less than or equal to 50 part per million.

2. The process of claim 1, wherein the cryogenic gas stream is at 173-213 K.

3. The process of claim 1, wherein the cryogenic gas stream has a water content below 1 part per million.

4. The process of claim 1, wherein the cryogenic gas stream has a maximum H.sub.2S concentration of 4 part per million.

5. The process of claim 1, wherein the cryogenic gas stream has a CO.sub.2 concentration of up to 3 mole percent.

6. The process of claim 1, wherein adsorptive carbon dioxide removal on the molecular sieve is carried out at a pressure of 1000-2000 kPag.

7. The process of claim 1, wherein the molecular sieve is a basic zeolite.

8. The process of claim 1, wherein the molecular sieve is a LTA, FAU or GIS zeolite.

9. A process for integrated liquefaction of natural gas and recovery of natural gas liquids, said process comprising: introducing a dry sweet methane gas stream, having a CO.sub.2 concentration of up to 3 mole percent, into an upper region of a demethanizer column; removing a natural gas liquids (NGL) product stream from the bottom of said demethanizer column and flow rate and composition are same after SFI-LNG integration; removing an overhead gaseous stream from the top of said demethanizer column at a cryogenic temperature and flow rate and composition are same after SFI-LNG integration; subjecting a half portion of the overhead gaseous stream from said demethanizer column to cryogenic adsorptive carbon dioxide removal (cryogenic CO.sub.2 removal) with a molecular sieve, to obtain a carbon dioxide product stream and scrubbed natural gas stream, wherein the scrubbed natural gas stream has a carbon dioxide content of less than or equal to 50 part per million; liquefying said scrubbed natural gas stream by two stages of coil-wound heat exchangers to obtain a liquefied natural gas product stream in the absence of a propane pre-chilling process.

10. The process of claim 9, wherein liquefying said scrubbed natural gas stream is by two stages in coil-wound heat exchangers.

11. The process of claim 9, wherein liquefying said scrubbed natural gas stream comprises a propane precooled/mixed refrigerant (Semi-C3-MR) cycle, and no propane pre-chilling process is required.

12. A process for integrated liquefaction of natural gas and recovery of natural gas liquids, said process comprising: cooling a dry sweet methane gas feed stream containing light hydrocarbons in one or more primary heat exchangers, wherein said feed stream is cooled and partially condensed by heat exchange; introducing the partially condensed feed stream into a gas/liquid cold separator producing an overhead gaseous stream and bottoms liquid stream; segregating the overhead gaseous stream from the gas/liquid cold separator into first and second portions; expanding said first portion of the overhead gaseous stream from the gas/liquid cold separator and introducing said expanded overhead gaseous stream into an upper region of a demethanizer column; introducing at least a portion of the bottoms liquid stream from gas/liquid cold separator into said demethanizer column at an intermediate point thereof; removing a natural gas liquids (NGL) product stream from the bottom of said demethanizer column; removing an overhead gaseous stream from the top of said demethanizer column at a cryogenic temperature; subjecting a first portion of the overhead gaseous stream from said demethanizer column to indirect heat exchange in one of said primary heat exchangers, wherein heat is exchanged between said first portion of the overhead gaseous stream from said demethanizer column and said second portion of the overhead gaseous stream from the gas/liquid cold separator, to obtain a natural gas product stream; subjecting a second portion of the overhead gaseous stream from said demethanizer column to cryogenic adsorptive carbon dioxide removal in a molecular sieve, to obtain a carbon dioxide product stream and scrubbed natural gas stream, wherein the scrubbed natural gas stream has a carbon dioxide content of less than or equal to 50 part per million; liquefying said scrubbed natural gas stream by heat exchange to obtain a liquefied natural gas product stream.

13. The process of claim 9, wherein the dry sweet methane gas stream has a water content below 1 part per million.

14. The process of claim 9, wherein the dry sweet methane gas stream has a maximum H.sub.2S concentration of 4 part per million.

15. The process of claim 9, wherein the dry sweet methane gas stream has a CO.sub.2 concentration up to 3 mole percent.

16. The process of claim 11, wherein adsorptive carbon dioxide removal on the molecular sieve is carried out at a pressure of 1000-2000 kPag.

17. The process of claim 9, wherein the molecular sieve is a basic zeolite.

18. The process of claim 9, wherein the molecular sieve is a LTA, FAU or GIS zeolite.

Description

BRIEF DESCRIPTION OF THE DRAWINGS

[0030] The present invention and its advantages will be better understood by referring to the following detailed description and attached Figures.

[0031] FIG. 1 is a schematic block flow diagram of a conventional LNG production process, with a cryogenic LNG removal in three stages of coil-wound heat exchangers and C3-MR cycle with a propane refrigeration pre-chillers cycle, an alkanolamine based CO.sub.2 removal and a dehydration unit for purifying and drying the inlet natural gas into liquefaction section of LNG production facility.

[0032] FIG. 2 is a schematic process flow diagram of a conventional LNG production process, showing (red line equipment and process lines) with a cryogenic LNG removal in three stages of coil-wound heat exchangers and C3-MR cycle with a propane refrigeration pre-chillers cycle, an alkanolamine based CO.sub.2 removal and a dehydration unit for purifying and drying the inlet natural gas into liquefaction section of LNG production facility.

[0033] FIG. 3 is a schematic block flow diagram of a split flow integrated LNG production (SFI-LNG) in conjunction and integration of a cryogenic NGL recovery process.

[0034] FIG. 4 is a schematic process flow illustration of the split flow integrated LNG liquefaction unit (SFI-LNG), showing in conjunction and integration of an SFI-LNG unit (red lined equipment and process lines) to an existing cryogenic NGL recovery plant.

[0035] FIG. 5 is a schematic illustration of exemplified zeolitic models for the separation of CO.sub.2 from CO.sub.2CH.sub.4 mixtures: (a) Na-LTA, (b) Ca-LTA, (c) Na-FAU and (d) Ca-FAU. Green atoms represent sodium, red atoms represent oxygen, yellow atoms represent silicon, blue atoms represent calcium and pink atoms represent aluminum.

[0036] FIG. 6 is a schematic illustration of most probable adsorption sites, as modeled, for CO.sub.2 in the different zeolites (red or green dots): (a) Na-LTA, (b) Ca-LTA, (c) Na-FAU and (d) Ca-FAU. Green atoms represent sodium, red atoms represent oxygen, yellow atoms represent silicon, blue atoms represent calcium and pink atoms represent aluminum.

[0037] FIG. 7 is a graph illustrating sorption isotherms of the mixture of 3 mole percent CO.sub.297 mole percent CH.sub.4 in the Molecular Sieve Na-LTA.

[0038] FIG. 8 is a graph illustrating sorption isotherms of the mixture of 3 mole percent CO.sub.297 mole percent CH.sub.4 in the Molecular Sieve Ca-LTA.

[0039] FIG. 9 is a graph illustrating sorption isotherms of the mixture of 3 mole percent CO.sub.297 mole percent CH.sub.4 in the Molecular Sieve Na-FAU.

[0040] FIG. 10 is a graph illustrating sorption Isotherms of the mixture of 3 mole percent CO.sub.297 mole percent CH.sub.4 in the Molecular Sieve Ca-FAU.

DETAILED DESCRIPTION

[0041] Referring now to FIG. 1 and FIG. 2, illustrating conventional processes which can be straddled downstream of any gas facility as a standalone facility, for comparison purpose we begin with an example of a LNG production facility which is not integrated into a cryogenic NGL recovery plant and it was considered only fifty percent of produced sales gas by the cryogenic NGL recovery unit will be purified and liquefied. In this simulation of a prior art NGL recovery plant according to U.S. Pat. Nos. 4,278,457 and 5,983,664, inlet gas enters the LNG plant purification section at 316 K, 6,672 kPag and flowrate of 33448 kg/hr as stream 42 which contains about 1.3 mole percent CO.sub.2 and 4 part per million H.sub.2S. Typically, if the inlet gas contains impurities such as H.sub.2S, CO.sub.2 and organic sulfur compounds (COS, CS.sub.2, SO.sub.x and mercaptans) then feed gas is treated by appropriate treatment process (not illustrated and shown as the Conventional CO.sub.2 Removal box) in order to remove impurities and meet LNG liquefaction specification and in this case it is considered that CO.sub.2 content will be reduced to 50 part per million in order to avoid operating problems from CO.sub.2 freezing in LNG liquefaction section. In addition, purified natural gas stream is saturated with water in the Conventional CO.sub.2 Removal process and the stream is dehydrated to prevent hydrate (ice) formation under cryogenic condition in the liquefaction section. Typically, conventional dehydration process (not illustrated and shown as the Conventional Dehydration box) has used solid desiccant for this purpose.

[0042] It is considered that the purified and dried natural gas is liquified by a C3-MR process. In this example and, the simulation is based on co-production of a nominal 1887 m.sup.3/day of LNG, with the volume of LNG measured at flowing (not standard) conditions at approximately 116 K and 449 kPag as stream 45. The treated and dried feed gas enters the liquefaction section at 316 K and 6500 kPag as stream 44 and is cooled in PRECHILLER heat exchangers 1, 2 and 3 by heat exchange with a propane cycle and cooled down the gas to approximate 278 K, 258 K and 238 K in three stages respectively. The pre-cooled stream then enters the COIL-WOUND 1 heat exchanger to be fully condensed at 186 K and 6210 kPag as stream 49 then enter the COIL-WOUND 2 heat exchanger to subcooled down at 159 K and 6010 kPag as stream 50 then enter the last COIL-WOUND 3 heat exchanger to further subcooled down at 117 K and 5810 kPag as stream 51. Stream 51 enters a work expansion machine (LNG-EXPANDER) in which mechanical energy is extracted from this high-pressure stream. The LNG-EXPANDER expands the subcooled fluid substantially isentropically from a pressure about 5810 kPag to the 449 kPag. The work expansion cools the expanded stream 45 to a temperature of approximately 157 K, whereupon it is then directed to the LNG-TANK which holds the final LNG product. In the cooling side, all of the cooling for streams 49, 50 and 51 are provided by a closed cycle mixed refrigerant (MR) loop. The working fluid for this cycle is a mixture of nitrogen, methane, ethane and propane, with composition of the mixture tuned as needed to provide the required refrigerant temperature as desired to provide the required refrigerant temperature while condensing at an economical pressure using the available cooling medium. The mixed refrigerant (MR) loop undergoes a three-stage compression MR-COMP 1, 2 and 3 with inter-stage cooling by air after MR-COMP2 and 3. The composition of the mixed refrigerant (MR) coolant stream 46, in approximate mole percent, is 8 percent nitrogen, 50 percent methane, 35 percent ethane, and 7 percent propane. The compressed mixed refrigerant (MR) coolant, similarly to the natural gas, passes through three propane pre-cooling stages MR-CHILLER 1, 2 and 3, the outlet temperature is approximately 238 K. The two-phase pre-cooled mixed refrigerant (MR) coolant then enters a gas-liquid separator (MR SEP.-1) to be separated into vapour and liquid streams at 238 K and 3200 kPag. Both vapour (stream 53) and liquid (stream 52) streams enter the COIL-WOUND 1 heat exchanger for further cooling. The liquid mixed refrigerant (MR) coolant stream passes through JTV-1 at the exit of COIL-WOUND 1 heat exchanger, and it returns back through COIL-WOUND 1 heat exchanger as coolant at 164 K and 128.7 kPag. The mixed refrigerant (MR) coolant vapour stream entering COIL-WOUND 1 heat exchanger and turns into a two-phase flow coolant, MR SEP.-2 segregate vapor from liquid phase at 186 K and 3000 kPag where the liquid and vapour outlets streams enter COIL-WOUND 2 heat exchanger for further cooling. The liquid phase after exiting COIL-WOUND 2 heat exchanger is further cooled through JTV-2 and returns back to COIL-WOUND 2 heat exchanger as coolant at 132 K and 158.7 kPag. The vapour phase enters COIL-WOUND 3 heat exchanger and it is further cooled. The refrigerant then passes through JTV-3 and returns back to COIL-WOUND 3 heat exchanger to serve as coolant at 106 K and 200 kPag.

[0043] An energy consumption summary for the process illustrated in FIG. 2 is set forth in the following table:

TABLE-US-00001 TABLE 1 Energy Consumption Summary for the process illustrated in FIG. 2 Power Consumption (kW/hr) Mixed Refrigerant (MR) Cycle 32113.8 Natural gas Pre-Chillers 2097.3 Sales Gas Compressor 5040.1 Sales Gas Compressor Air Cooler 5923.7 Turbo-Expander Air Cooler 1255.3 LNG-Expander 87.6 Propane Air Cooler 18387.7 Propane Compressor 6404.0 Total 71309.4

[0044] The liquefaction of natural gas typically involves distinct steps of purification, principally to remove carbon dioxide and water (typically, the water/moisture content in the natural gas pipeline should not exceed 4 lb water per million standard cubic feet per day, MMSCFD, of gas), therefore, a dehydration unit is required downstream of acid gas treatment unit in the conventional LNG processes as stated earlier. For instance, see propane pre-cooled mixed refrigerant processes (Remeljej et al., 2006; He & Ju, 2014). A variety of processes patent have been suggested for integrating NGL recovery processes with LNG liquefaction processes and LNG production in cryogenic natural gas processing plants (see U.S. Pat. Nos. 4,525,185, 4,548,629, 4,545,795, 5,537,827, 5,600,969, 5,615,561, 6,016,665, 6,119,479, 6,526,777 B1, 6,662,589, 6,742,358B2, 6,889,523B2, 7,237,407B2, 7,204,100B2; and U.S. Patent Application Publication Nos. US 2003/0089125A1, US 2007/0012072A1, US 2007/0157663, US 2008/0271480A1, US 2009/0025422A1, US 2010/0024477 and US 2013/0061633).

[0045] A summary of stream flow rates for the process illustrated in FIG. 2 is set forth in the following table:

TABLE-US-00002 TABLE 2 Stream Flow Summary - kg moles/h for the process illustrated in FIG. 2 Stream N.sub.2 CO.sub.2 Methane Ethane Propane Butane+ Total 42 10.3 26.7 1972.6 48.8 0.9 0.2 2059.5 44 10.3 0.1 1972.6 48.8 0.9 0.2 2032.9 45 10.3 0.1 1972.6 48.8 0.9 0.2 2032.9 46 408.9 0.0 2555.6 1788.9 357.8 0.0 5111.2 47 10.3 26.7 1972.6 48.8 0.9 0.2 2059.5 52 21.2 0.0 377.6 939.9 290.4 0.0 1629.1 53 387.7 0.0 2178.0 849.0 67.4 0.0 3482.1

[0046] The cryogenic NGL recovery plants operates exactly the same in the FIG. 2 and FIG. 4 and the recovery levels for ethane, propane and butanes plus are just the same. In this example based on FIG. 4 and, the simulation is based on purified and dried natural gas which is acceptable by a cryogenic NGL recovery operation condition and the feed gas enters at approximately 230 K, 6115 kPag and 94867 kg/h as stream 1 (stream 31 in FIG. 2). A typical analysis of a natural gas stream to be processed in accordance with this invention would be, in approximate mole percent, 82 percent methane, 10.6 percent ethane, 4.3 percent propane, 1 percent butanes, 1.5 percent carbon dioxide, with the balance made up of pentanes plus, nitrogen, helium. Sulfur containing gases such as hydrogen sulfide (maximum 4 part per million) are sometimes present. In this simulation of a prior art NGL recovery plant according to U.S. Pat. Nos. 4,278,457 and 5,983,664, the cryogenic extraction of NGL from natural gas typically involves the use of a cryogenic column (referred to as a demethanizer, typically operating at a top pressure of about 1300-2100 kPag depends on inlet gas quality and applied process technology), in processes that often involve the following steps. Inlet feed gas stream 1 (stream 31 in FIG. 2) split into two streams and 60 percent of gas flow went to stream 2 (stream 32 in FIG. 2) where it cooled in Pass A of the COLD BOX heat exchanger to 254 K before gas-liquid mixture enters the COLD-SEP approximately at 252 K and 6078 kPag. A significant fraction of the heavier components, such as butanes, are condensed and segregated in COLD-SEP. The liquid flow through a throttling expansion valve that results in lowering the pressure then enters the demethanizer approximately at 222 K, 1419 kPag and 4284 kg/h as stream 7 (stream 37 in FIG. 2).

[0047] The vapour stream from the COLD-SEP split into two streams; 27 percent of the flow went to stream 8 (stream 38 in FIG. 2) before enters the pass F of COLD-BOX heat exchanger where it cooled down before sent to the top of the DEMETHANIZER and enters rectification section as a column cold reflux at 164 K, 1417 kPag and 24440 kg/h as stream 18 (stream 48 in FIG. 2). The 73 percent of the vapour stream from the COLD-SEP enters the EXPANDER-COMP., a turbine in which the energy released by gas expansion is used to drive a compressor, lowering the vapour stream temperature and pressure as it enters the DEMETHANIZER approximately at 200 K, 1419 kPag and 66145 kg/hr as stream 6 (stream 36 in FIG. 2).

[0048] Liquid side streams from a lower section of the DEMETHANIZER may also be routed through the COLD BOX as Pass B and C in order to heat exchange and precooled the inlet gas, then heated up side streams are returned to the DEMETHANIZER bottom section. The DEMETHANIZER side streams are similar in concept but with different temperature and flow rate profiles. The bottom NGL product from the DEMETHANIZER warmed in the COLD BOX Pass D, as it cools the inlet gas, before goes to the NGL product tank approximately at 252 K, 3684 kPag and 25618 kg/h as stream 5 (stream 35 in FIG. 2).

[0049] Overhead residue gas produced in the DEMETHANIZER at 176 K, 1415 kPag and 69249 kg/h as stream 41 (stream 11 in FIG. 1 explained later part of the innovation) pass through COLD BOX Pass E and heated up to approximate at 239 K and 1400 kPag as stream 39. Then flow to compressor part of EXPANDER-COMP. which pressurized the residue gas to approximate at 344 K and 2146 kPag as stream 33 before enter air cooler and SALES GAS COMP. and after cooler to boost the residue gas pressure to approximate at 316 K, 6672 kPag and 69249 kg/h as stream 34. For shipment by pipeline, the natural gas product may be further compressed in a second-stage gas compressor. At this stage natural gas is sufficiently free of NGL for distribution to end-use consumers by pipeline and may be further processed to produce LNG as stated earlier and illustrated in FIGS. 1 and 2. A variety of processes patent have been suggested for cryogenic NGL recovery process which are adoptable and applicable with presented innovation (see U.S. Pat. Nos. 4,278,457, 4,887,545, 5,275,005, 5,325,673, 5,881,569, 5,555,748, 5,983,664, 6,453,698B2, 7,051,553B2, and U.S. Patent Application Publication Nos. US 2006/0283207, US 2010/0011810A1, US 2009/0107175A1, US 2010/0251764A1, US 2008/0190136A1).

[0050] The present invention relates to a process for processing natural gas to produce liquified natural gas (LNG) in conjunction and integration with a cryogenic natural gas recovery (NGL) process. In particular, this invention is well suited to LNG production in small-scale by conjunction and integration into natural gas processing plant that recover natural gas liquids (NGL) using a cryogenic process and produce a rich in methane natural gas.

[0051] In the following description, various examples are set out of embodiments, together with experimental procedures that may be used to implement a wide variety of modifications and variations in the practice of the present invention. For clarity, a variety of technical terms are used herein in accordance with what is understood to be the commonly understood meaning, as reflected in definitions set out below.

[0052] C3/MR: Propane/Mixed Refrigerant

[0053] J-T: Joule-Thompson

[0054] MMSCFD: Million Standard Cubic Feet per Day

[0055] MR: Mixed Refrigerant

[0056] NG: Natural Gas

[0057] LTA: Linde Type A (zeolite structure)

[0058] LNG: Liquefied Natural Gas

[0059] NGL: Natural Gas Liquids

[0060] FAU: Faujasite (zeolite structure)

EXAMPLES

Example 1: Integrated NGL and LNG Production Processes

[0061] FIGS. 3 and 4 illustrate a flow diagram of Split Flow Integrated LNG Production (SFI-LNG) in accordance with the present invention which is well suited for a small-scale LNG production plant. As stated earlier, the simulation is based on purified and dried natural gas and the inlet feed gas enters at approximately 316 K, 6115 kPag and 94867 kg/h as stream 1 then split into two streams and 60 percent of gas flow went to stream 2 where it cooled in Pass A of the COLD BOX heat exchanger to 254 K before gas-liquid mixture enters the COLD-SEP approximately at 254 K and 6078 kPag. The liquid flow through a throttling expansion valve that results in lowering the pressure then enters the demethanizer approximately at 222 K, 1419 kPag and 4284 kg/h as stream 7. The gas stream from the COLD-SEP split into two streams; 27 percent of the flow went to stream 8 then enters the pass F of COLD-BOX heat exchanger where it cooled down to approximate at 250 K and 6036 kPag as stream 10. Then the stream left the NGL plant and enters into the liquefaction section where enters the COIL-WOUND 1 and then COIL-WOUND 2 heat exchanger and cooled down before return to NGL recovery plant approximate at 177 K and 5636 kPag as stream 17, then gas flows undergo a throttling expansion valve (in an isenthalpic process that results in lowering the temperature of the reflux stream) before enters as reflux stream into the rectification section of the DEMETHANIZER at 164 K, 1417 kPag and 24440 kg/h as stream 18. The 73 percent of the vapour stream from the COLD-SEP enters the EXPANDER-COMP., lowering the vapour stream pressure as it enters the DEMETHANIZER approximately at 200 K, 1419 kPag and 66145 kg/h as stream 6. The bottom NGL product from the DEMETHANIZER warmed in the COLD BOX Pass D, as it cools the inlet gas, before goes to the NGL product tank approximately at 294 K, 3684 kPag and 25618 kg/h as stream 5. There will be no change in stream 5 condition as part of this invention.

[0062] Overhead residue gas produced in the DEMETHANIZER split between stream 11 and 12 with 50/50 percent, by a split flow controller, each approximate at 175.7 K, 1415 kPag and 34624.5 kg/h. Stream 11 pass through COLD BOX Pass E and heated up to approximate at 307 K and 1400 kPag as stream 9. Then flow to compressor part of EXPANDER-COMP. which pressurized the residue gas to approximate at 379 K and 3152 kPag as stream 3 before enter air cooler and SALES GAS COMP. and after cooler to boost the residue gas pressure to approximate at 316 K, 6672 kPag and 34624.5 kg/h as stream 4. For shipment by pipeline, the natural gas product may be further compressed in a second-stage gas compressor.

[0063] Stream 12 enters the purification section of LNG production plant, a typical analysis of a natural gas stream to be processed in accordance with this invention would be, in approximate mole percent, 95.8 percent methane, 2.4 percent ethane, 1.3 percent carbon dioxide, with the balance made up of propane plus, nitrogen, helium. Sulfur containing gases such as hydrogen sulfide (maximum 4 part per million) are sometimes present. The CRYOGENIC CO.sub.2 REMOVAL serves as an LNG purification unit. The feed gas into the CRYOGENIC CO.sub.2 REMOVAL unit is treated by appropriate treatment process (not illustrated and shown as the CRYOGENIC CO.sub.2 REMOVAL box) in order to remove impurities and meet LNG liquefaction specification. In this case, it is considered that CO.sub.2 content will be reduced to 50 part per million in order to avoid operating problems from CO.sub.2 freezing in LNG liquefaction section. The CRYOGENIC CO.sub.2 REMOVAL unit is a physical zeolite-based adsorption bed in order to extract CO.sub.2 (impurities) as a co-product with high purities and in this simulation 1176.7 kg/h pure carbon dioxide produced as stream 13. A variety of molecular sieves, such as Faujasite (FAU), 5-Angstrom and 4-Angstrom beds, can be used to remove carbon dioxide. The solid adsorbent used in a molecular sieve bed may for example be formed from zeolites (crystalline materials composed of silicon and aluminum). After adsorption bed saturated with CO.sub.2, the molecular sieve material may be regenerated or cleaned, for example by lowering pressure on the bed and releasing acid gases, or by passing a hot gas through it. The hot gas may be, for example, a hot methane gas or heated nitrogen.

[0064] The scrubbed natural gas stream left the CRYOGENIC CO.sub.2 REMOVAL unit approximate at 176 K, 1315 kPag and 33448 kg/h as stream 14. It is considered that the purified natural gas is liquified by a Semi-C3-MR process. In this example and, the simulation is based on co-production of a nominal 1892 m.sup.3/day of LNG, with the volume of LNG measured at flowing (not standard) conditions at approximately 117.3 K and 449 kPag as stream 15. The treated feed gas enters the liquefaction section at 176 K and 1315 kPag as stream 14 and then enters the COIL-WOUND 1 heat exchanger to be almost fully condensed at 153 K and 1115 kPag as stream 19 then enter the COIL-WOUND 2 heat exchanger to subcooled down at 117.4 K and 915 kPag as stream 20. Stream 20 enters a work expansion machine (LNG-EXPANDER) in which mechanical energy is extracted from this stream. The LNG-EXPANDER expands the subcooled fluid substantially isentropically from a pressure about 915 kPag to the 449 kPag. The work expansion cools the expanded stream 20 to a temperature of approximately 117.3 K, whereupon it is then directed to the LNG-TANK which holds the final LNG product. In the cooling side, all of the cooling for streams 19 and 20 are provided by a closed cycle mixed refrigerant (MR) loop. The working fluid for this cycle is a mixture of nitrogen, methane, ethane and propane, with composition of the mixture tuned as needed to provide the required refrigerant temperature as desired to provide the required refrigerant temperature while condensing at an economical pressure using the available cooling medium. The mixed refrigerant (MR) loop undergoes a three-stage compression MR-COMP 1, 2 and 3 with inter-stage cooling by air after MR-COMP1, 2 and 3. The composition of the mixed refrigerant (MR) coolant stream 16, in approximate mole percent, is 8 percent nitrogen, 50 percent methane, 35 percent ethane, and 7 percent propane. The compressed mixed refrigerant (MR) coolant, similarly to the natural gas, passes through three propane pre-cooling stages MR-CHILLER 1, 2 and 3, the outlet temperature is approximately 238 K. The two-phase pre-cooled mixed refrigerant (MR) coolant then enters a gas-liquid separator (MR SEP.-1) to be separated into vapour and liquid streams at 238 K and 6000 kPag. Both vapour (stream 22) and liquid (stream 21) streams enter the COIL-WOUND 1 heat exchanger for further cooling. The liquid mixed refrigerant (MR) coolant stream passes through JTV-1 at the exit of COIL-WOUND 1 heat exchanger, and it returns back through COIL-WOUND 1 heat exchanger as coolant at 134 K and 130 kPag. The mixed refrigerant (MR) coolant vapour stream entering COIL-WOUND 1 heat exchanger and turns into a fully condensed coolant at 153 K and 5800 kPag where the fully liquid stream enters COIL-WOUND 2 heat exchanger for further cooling. The stream after exiting COIL-WOUND 2 heat exchanger is further cooled through JTV-2 and returns back to COIL-WOUND 2 heat exchanger as coolant at 108.9 K and 160 kPag.

[0065] An energy consumption summary for the process illustrated in FIG. 4 is set forth in the following table:

TABLE-US-00003 TABLE 3 Energy Consumption Summary for the process illustrated in FIG. 4 Power Consumption (kW/hr) Mixed Refrigerant (MR) Cycle 36818.1 Natural gas Pre-Chillers 0 Sales Gas Compressor 1559.6 Sales Gas Compressor Air Cooler 1899.8 Turbo-Expander Air Cooler 1477.9 LNG-Expander 6.8 Propane Air Cooler 14944.8 Propane Compressor 5213.9 Total 61920.9

[0066] In select embodiments, Split Flow Integrated LNG Production (SFI-LNG) comprises two fundamental sections: natural gas purification (CRYOGENIC CO.sub.2 REMOVAL) and natural gas liquefaction (Semi-C3-MR). The exemplified process is illustrated in FIGS. 3 and 4. This example illustrates a savings in energy consumption in LNG production available through the use of in integrated NGL and LNG production process. Compared to a conventional natural gas liquefaction processes, the present innovations may provide opportunities to economize on equipment, for example by using two stages of coil-wound heat exchangers instead of three stages. In this simulation the mass flow of mixed refrigerant cycle (MR) is reduced compared to the conventional processes, for example by up to 16.1%. In particular, in conventional C3-MR processes (the most widely used liquefaction process to date and has about 75% of the liquefaction market), pre-chilling by propane is typically required. However, in exemplary SFI-LNG systems, no pre-chilling is required. In conjunction with this, the propane refrigerant loop is smaller than conventional installations, for example up to 18.8% smaller. This provides significant reductions in liquefaction power consumption, for example providing reductions of up to 13.2% compared to conventional technologies

[0067] The split flow integrated LNG production process (SFI-LNG) is a nanotechnology-based small-scale LNG process which can be integrated with the design of any new natural gas facility and the technology can also be retrofitted to existing natural gas liquid (NGL) recovery plants, allowing for co-production of LNG from natural gas facilities and also maximize the economic value of the cryogenic NGL recovery plant. SFI-LNG can make small-scale LNG production facility more environmentally and economically attractive.

[0068] One skilled in the art will recognize that the present invention can be adapted for use with all types of cryogenic NGL recovery plants to allow co-production of LNG and CO.sub.2 with high purity. The example presented earlier have all depicted the use of the present invention with a cryogenic NGL recovery plant employing the process disclosed in U.S. Pat. Nos. 4,278,457 and 5,983,664 in order to facilitate comparisons of the present invention with the prior art. Nevertheless, the present invention is generally applicable for use with any NGL recovery process that produces a residue gas stream that is at temperature of 203 K or cooler. Examples of such cryogenic NGL recovery processes are described and illustrated in U.S. Pat. Nos. 4,278,457, 4,887,545, 5,275,005, 5,325,673, 5,881,569, 5,555,748, 5,983,664, 6,453,698B2, 7051553B2, and U.S. Patent Application Publication Nos. US 2006/0283207, US 2010/0011810A1, US 2009/0107175A1, US 2010/0251764A1, US 2008/0190136A1).

[0069] The exemplified Cryogenic CO.sub.2 removal reduces CO.sub.2 content in the gaseous product of the demethanizer at cryogenic temperatures from up to 3 mole percent to less than 50 part per million. The purification process does not require additional water nor chemicals and consume substantially less energy and can produced CO.sub.2 as co-product with very high purity. For physical adsorption, the CO.sub.2 adsorbent will have high selectivity and adsorption capacity, provide good thermal and mechanical stability, remain stable in life cycle of adsorption/desorption and also provide adequate adsorption/desorption kinetics. A variety of molecular sieves, such as Faujasite (FAU), 5-Angstrom and 4-Angstrom beds, can be used to remove CO.sub.2.

[0070] A summary of stream flow rates for the process illustrated in FIG. 4 is set forth in the following table:

TABLE-US-00004 TABLE 4 Stream Flow Summary - kg moles/h for the process illustrated in FIG. 4 Stream N.sub.2 CO.sub.2 Methane Ethane Propane Butane+ Total 12 10.3 26.7 1972.6 48.8 0.9 0.2 2059.5 13 0 26.6 0 0 0 0 26.6 14 10.3 0.1 1972.6 48.8 0.9 0.2 2032.9 15 10.3 0.1 1972.6 48.8 0.9 0.2 2032.9 16 343.2 0.0 2144.9 1501.4 300.3 0.0 4289.8 18 5.5 19.4 1048.8 129.3 47.5 10.5 1261.0 21 127.2 0.0 1213.4 1237.2 277.1 0.0 2854.9 22 215.9 0.0 931.4 264.2 23.2 0.0 1434.7

Example 2: Separation of CO.SUB.2 .from CH.SUB.4.CO.SUB.2 .Mixtures

[0071] This Example illustrates with numerical modeling that basic zeolites, such as Na-LTA, Ca-LTA, Na-FAU and Ca-FAU, may be used to separate CO.sub.2, in the example 3 mole percent CO.sub.2, from a mixture of CO.sub.2CH.sub.4, modeled at 173 K and up to 1400 kPa. The results set out below illustrate the use of 4 types of zeolite under selected conditions. As modeled, the Ca-LTA and Na-LTA zeolites demonstrated the ability to extract CO.sub.2 without CH.sub.4 contamination. In some embodiments, Na-LTA may have an advantage as being relatively easily prepared, in particular not requiring an ionic exchange procedure that is required to produce pure Ca-LTA. Na-FAU and Ca-FAU are also demonstrably able to remove CO.sub.2 from the mixture, in modeled processes in which the CO.sub.2 retained traces of CH.sub.4.

[0072] The exemplary modeling was of processes that did not include H.sub.2O or H.sub.2S in the feed. In alternative embodiments, for example where levels of these compounds may be detrimental for CO.sub.2 uptake, adsorbents may for example be pre-treated at high temperature to remove any hydration water within the zeolitic channels.

Zeolitic Models

[0073] The program SORPTION within the BIOVIA 2017 Materials Studio Software was employed for the calculations presented in this Example. The COMPASSII forcefield was used for all calculations. The first step was to produce the zeolitic models required for the adsorption calculations. For this task, the aluminosilicate frameworks LTA and FAU, with an atomic ratio Si/Al=1, within the MS2017 database were selected. The addition and location of the sodium and calcium ions within the channels and cavities of each framework were carried out with the LOCATE module within the SORPTION software. After the calculations were carried out, the most stable structure for each case was taken as the model for the adsorption calculations. FIG. 5 shows the four zeolitic materials generated: Na-LTA, Ca-LTA, Na-FAU and Ca-FAU.

Sorption Models

[0074] The second step was to generate the CO.sub.2 and CH.sub.4 molecules and optimize them with the FORCITE module using the COMPASSII forcefield. Then, the optimized molecules were fed to SORPTION to produce the sorption isotherms of the mixture CO.sub.2CH.sub.4 with 3 mole percent CO.sub.2 at a temperature of 173 K with a maximum pressure of 1400 kPa. For this task, the ADSORPTION ISOTHERM module within the SORPTION software was used. FIG. 6 shows the most probable sorption sites for CO.sub.2 (in red or green dots) within the cavities of the four tested zeolites.

[0075] FIGS. 7, 8, 9 and 10 show the sorption isotherms for the separation of the CO.sub.2 from the CO.sub.2CH.sub.4 mixture at 173 K and up to 1400 kPa for Na-LTA, Ca-LTA, Na-FAU and Ca-FAU, respectively. As can be seen in these figures, the four zeolites are able to selectively uptake CO.sub.2 from the mixture in the whole range of pressures. In modeled embodiments, the Na-LTA and Ca-LTA are in some respects advantageous materials for the preferential separation of CO.sub.2 by adsorption, as the uptake of CH.sub.4 molecules is zero in these two zeolites.

[0076] The following table summarizes the predicted values for the uptake of CO.sub.2 from the CO.sub.2CH.sub.4 mixture at 173 K and 1400 kPa (assuming there is not H.sub.2O or H.sub.2S on the feed and that the zeolites were activated at high temperature to remove the hydration water inside the zeolite channels and cavities). This will be the maximum expected uptake.

TABLE-US-00005 Zeolite mg CO.sub.2/g zeolite mg CH.sub.4/g zeolite Na-LTA 445 Ca-LTA 471 Na-FAU 448 1.2 Ca-FAU 521 2.4

[0077] As the results in the table above illustrate, the 100% ion exchange of Na by Ca in zeolite LTA induces an increase of CO.sub.2 uptake of 26 mg/g zeolite which is almost 6% more than that for Na-LTA. In practice, 100% ion exchange of Na by Ca may for example be achieved through several ion exchange, drying and calcination processes. In the case of zeolite FAU, the perfect exchange of Na ions by Ca ions produces an increase of 73 mg CO.sub.2/g zeolite, which is an increase of the CO.sub.2 uptake of about 16% compared with that for Na-FAU.

[0078] The present example illustrates that CO.sub.2 separation from a mixture of CO.sub.2CH.sub.4 at 173 K and at 1400 kPa using basic zeolites is feasible.

[0079] This example illustrates the feasibility of separating 3 mole percent CO.sub.2 from a mixture of CO.sub.2CH.sub.4 at 173 K and 1400 kPa. LTA zeolites, as modeled, uptake only CO.sub.2 and the exchange of Na by Ca increases the CO.sub.2 uptake up to 6%. FAU zeolites have a slightly higher uptake than LTA zeolites; however, FAU zeolites as modeled take traces of CH.sub.4. Changing the Na by Ca increases the uptake of both CO.sub.2 and CH.sub.4 for the FAU zeolites.

[0080] SFI-LNG overcomes two key limitations of the conventional LNG production design-high energy consumption and the need for additional gas treatment for CO.sub.2 removal and drying the gas, i.e., the SFI-LNG process uses less energy compared with the conventional process and obviates the need for amine treatment and gas dehydration steps.