Patent Publication Number: US-6209350-B1

Title: Refrigeration process for liquefaction of natural gas

Description:
This application claims the benefit of U.S. Provisional Application No. 60/105,462, filed Oct. 23, 1998. 
    
    
     FIELD OF THE INVENTION 
     This invention relates generally to a process for conveying a natural gas stream, and more specifically to a process for conveying a natural gas stream through a pipeline to a liquefication plant which produces a pressurized liquefied natural gas (PLNG) for further conveyance. 
     BACKGROUND OF THE INVENTION 
     Because of its clean burning qualities and convenience, natural gas has become widely used in recent years. Many sources of natural gas are located in remote areas, great distances from any commercial markets for the gas. Sometimes a pipeline is available for transporting produced natural gas to a commercial market. Although the transportation of gas by pipeline normally takes place over fairly lengthy distances, this would be no problem where only transportation over land is encountered. However, in many instances the natural gas is separated from a suitable market by expansive bodies of water. When pipeline transportation is not feasible, produced natural gas is often processed into liquefied natural gas (which is called “LNG”) for transport to market. The liquefication plants are sometimes located at the source of the LNG, but the LNG plants are often located at ports from which the liquefied gas is shipped to foreign markets. 
     One of the distinguishing features of natural gas transportation systems is the large capital investment required. Pipelines, plants used to liquefy natural gas, and ships to carry the liquefied natural gas are all quite expensive. Pipeline materials and installation cost can be quite high and gas compressors and cooling systems arc required to move the gas through the pipeline. The liquefication plant is made up of several basic systems, including gas treatment to remove impurities, liquefication, refrigeration, power facilities, and storage and ship loading facilities. The design and operation of these systems can significantly increase the transportation cost of the natural gas. These systems can make transportation of the natural gas in some locations in the world economically prohibitive. 
     The development of natural gas fields in arctic regions, such as the North Slope gas and oil fields of the State of Alaska, present special challenges. The natural gas pipelines that are buried in frozen soil or permafrost must be taken into account. If such pipelines arc transmitting gas at temperatures above 0° C. (32° F.), the frozen ground in which the pipelines are buried will eventually thaw, and the resulting settlement or heaving action could possibly cause pipeline failure. Accordingly, preservation of the frozen soil or permafrost is a major concern to pipeline installers and operators, not only with a view to protecting the environment, but also with a view to minimizing damage and failure of the pipelines. 
     Various pipelines systems for conveying the natural gas in arctic environments have been suggested. U.S. Pat. No. 4,192,655 to von Linde discloses one example of a pipeline system for transporting natural gas over long distances in arctic regions by a pipeline to a liquefication plant at a port. The von Linde patent suggests using a pipeline having a number of sections in series with intermediate compressor stations. The pressure and temperature of the gas at the entry to each pipeline section is such that the drop in pressure of the gas in each section creates a drop in gas temperature and this low temperature gas is used to re-cool the gas heated by compression before it enters the next pipeline section. Von Linde suggests conveying the gas at an initial pressure of between 7,500 kPa (1,088 psia) and 15,000 kPa (2,175 psia) and at an initial temperature of below −10° C. (14° F.). The gas exiting the last pipeline section can be −45.2° C. (−50° F.) or lower. The liquefication plant, being located at the end of the last pipeline section, takes advantage of the low temperature in the liquefication process. From the liquefication plant the liquefied gas is pumped into tankers for transport to market. 
     Conventional gas liquefaction processes are required to produce a liquefied product that is below about −156.7° C. (−250° F.) for transportation via ships to the customer. As a result, more of the gas is consumed in the CO 2  removal, gas liquefaction, and liquid regasification processes, thereby making less of the gas available to the consumer as product. In addition, gas transportation to the liquefaction facilities in conventional steel pipelines limits the practical (economical) operating pressure of conventional pipelines to pressures in the range of 6,895 to 15,860 kPa (1,000 to 2,300 psia), thereby requiring the use of gas recompressor stations along the pipeline route. The pipeline recompressors consume additional fuel and add heat of compression to the gas in the pipeline, so that the gas reaches the liquefaction plant at a warmer temperature than it would if pipeline recompression were not required. 
     The industry has a continuing need for an improved process for conveying natural gas which minimizes the amount of treating equipment required and the overall power consumption. By reducing the overall cost of conveying natural gas over long distances will add to the amount of gas available for use by consumers. 
     SUMMARY 
     This invention relates to an improved process for conveying gas stream rich in methane, such as natural gas. In the first step of the process, gas is supplied to a pipeline at an entry pressure that is substantially higher than the output pressure of the pipeline. The drop in pressure in the pipeline causes a lowering of the gas temperature, preferably to a temperature below about −29° C. (−20° F.). The entry pressure of the gas to the pipeline is controlled to achieve a predetermined output pressure of the gas from the pipeline. Output gas from the pipeline is then liquefied to produce liquefied gas having a temperature above about −112° C. (−170° F.) and a pressure sufficient for the liquid to be at or below its bubble point temperature. The pressurized liquefied gas is then further transported in a suitable container. 
     The liquefaction plant receives the natural gas at a temperature below about −29° C. (−20° F.) and a pressure above about 3,450 kPa (500 psia). The natural gas is then introduced to a first phase separator to produce a first liquid stream and a first vapor stream. The pressure of the first liquid stream is adjusted to approximately the operating pressure of a third phase separator used in the process. This pressure adjusted liquid stream is passed to the third phase separator. The first vapor stream is passed through a first heat exchanger, thereby warming the first vapor stream. The first vapor stream is compressed and cooled. The compressed first vapor stream is passed through the first heat exchanger to further cool the compressed first vapor stream. The compressed vapor stream is passed through a second heat exchanger to still further cool the first vapor stream. This compressed vapor stream is expanded to thereby decreasing its temperature. This expanded stream is then passed to a second phase separator to produce a second vapor stream and a second liquid stream. The second vapor stream is recycled back to the first phase separator. The second liquid stream is expanded to further reduce the pressure and lower the temperature. The second liquid stream is passed to a third phase separator to produce a third vapor stream and a liquid product stream having a temperature above −112° C. (−170° F.) and having a pressure sufficient for the liquid to be at or below its bubble point. The third vapor stream is passed through the second heat exchanger to provide refrigeration to the second heat exchanger. The third vapor stream is passed through a third heat exchanger, the third vapor stream is compressed to approximately the operating pressure of the first phase separator, the compressed third vapor stream is cooled, and the cooled compressed third vapor stream is passed through the third heat exchanger and the compressed third vapor stream is passed to the first phase separator for recycling. 
     In the practice of this invention, natural gas can be transported at higher pressure (17,238 to 34,475 kPa) without the requirement of pipeline recompressor stations, thereby avoiding the addition of recompression heat along the pipeline. The natural gas arrives at the liquefaction plant at a colder temperature, which lessens the amount of refrigeration needed to liquefy the gas and it also lessens the amount of gas consumed as fuel in the liquefaction plant. 
    
    
     BRIEF DESCRIPTION OF THE DRAWINGS 
     The present invention and its advantages will be better understood by referring to the following detailed description and the attached Figures. 
     FIG. 1 is a schematic diagram of one embodiment of the liquefaction process of the present invention. 
     FIG. 2 is a schematic diagram of a second embodiment of the liquefaction process of the present invention. 
    
    
     The Figures present two embodiments of practicing the process of this invention. The Figures are not intended to exclude from the scope of the invention other embodiments that are the result of normal and expected modifications of these specific embodiments. Various required subsystems such as valves, control systems, sensors, clamps, and riser support structures have been deleted from the Figures for the purposes of simplicity and clarity of presentation. 
     DESCRIPTION OF THE INVENTION 
     The present invention is an improved process for conveying natural gas over long distance by first passing the natural gas through a pipeline and then liquefying the gas in a liquefication plant to produce a methane-rich liquid product having a temperature above about −112° C. (−170° F.) and a pressure sufficient for the liquid product to be at or below its bubble point temperature. This methane-rich product is sometimes referred to in this description as pressurized liquid natural gas (“PLNG”). The term “bubble point” is the temperature and pressure at which a liquid begins to convert to gas. For example, if a certain volume of PLNG is held at constant pressure, but its temperature is increased, the temperature at which bubbles of gas begin to form in the PLNG is the bubble point. Similarly, if a certain volume of PLNG is held at constant temperature but the pressure is reduced, the pressure at which gas begins to form defines the bubble point. At the bubble point, the mixture is saturated liquid. 
     The gas liquefication process of the present invention requires less total power for transporting through a pipeline and then liquefying the natural gas in a liquefication plant than processes used in the past and the equipment used in the process of this invention can be made of less expensive materials. By contrast, prior art processes that produce conventional LNG at atmospheric pressures having temperatures as low as −160° C. (−256° F.) require process equipment made of expensive materials for safe operation. The invention is particularly useful in arctic applications, but the invention can also be used in warm climates. 
     The energy needed for liquefying the natural gas in the practice of this invention is greatly reduced over energy requirements of a conventional LNG plant which produces LNG at atmospheric pressure and a temperature of about −160° C. (−256 ° F.). The reduction in necessary refrigeration energy required for the process of the present invention results in a large reduction in capital costs, proportionately lower operating expenses, and increased efficiency and reliability, thus greatly enhancing the economics of producing liquefied natural gas. 
     Referring to FIG. 1, a feed gas produced from a natural gas reservoir, from associated gas from oil production or from any other suitable source is fed as stream  5  to a compression zone  45  comprising one or more compressors. Although not shown in the FIG. 1, before the feed gas is passed to the compressors, the feed gas will normally have passed through treatment stage to remove contaminants. 
     The first consideration in cryogenic processing of natural gas is contamination. The raw natural gas feed stock suitable for the process of this invention may comprise natural gas obtained from a crude oil well (associated gas) or from a gas well (non-associated gas). The composition of natural gas can vary significantly. As used herein, a natural gas stream contains methane (C 1 ) as a major component. The natural gas will typically also contain ethane (C 2 ), higher hydrocarbons (C 3+ ), and minor amounts of contaminants such as water, carbon dioxide, hydrogen sulfide, nitrogen, butane, hydrocarbons of six or more carbon atoms, dirt, iron sulfide, wax, mercury, helium, and crude oil. The solubilities of these contaminants vary with temperature, pressure, and composition. At cryogenic temperatures, CO 2 , water, or other contaminants can form solids, which can plug flow passages in cryogenic heat exchangers. These potential difficulties can be avoided by removing such contaminants if conditions within their pure component, solid phase temperature-pressure phase boundaries are anticipated. In the following description of the invention, it is assumed that the natural gas stream being fed to the compressor zone  45  has been suitably treated to remove unacceptably high levels of sulfides and carbon dioxide and dried to remove water using conventional and well-known processes to produce a “sweet, dry” natural gas stream. If the natural gas stream contains heavy hydrocarbons that could freeze out during liquefication or if the heavy hydrocarbons are not desired in PLNG, the heavy hydrocarbon may be removed by a fractionation process prior to liquefaction of the natural gas. At the operating pressures and temperatures of PLNG, moderate amounts of nitrogen in the natural gas can be tolerated since the nitrogen will remain in the liquid phase with the PLNG. 
     After being compressed in compression zone  45 , the natural gas is preferably passed through an aftercooler  46  to cool the gas stream by indirect heat exchange before the gas enters pipeline  47 . Aftercooler  46  may be any conventional cooling system that cools the natural gas to a temperature below about −1.1° C. (30° F.) for applications in which the pipeline will be buried in frozen soil or permafrost. Aftercooler  46  preferably comprises a combination of air or water-cooled heat exchangers and a conventional closed-cycle propane refrigeration system. 
     The natural gas is compressed by compression zone  45  to a pressure sufficient to produce a predetermined pressure and temperature at the output of the pipeline (stream  7 ). The pressure of the natural gas at the entry to the pipeline (stream  6 ) is controlled so that lowering of natural gas temperatures results from the Joule-Thomson effect created by the drop in pressure in the pipeline. The gas pressure at the entry to the pipeline can be determined by those skilled in the art taking into account the length of the pipeline, gas flow rate, and frictional losses incurred in conveyance of the gas through the pipeline. The pressure of the entry gas (stream  6 ) will preferably range between about 17,238 kPa (2,500 psia) and about 48,265 kPa (7,000 psia), and more preferably between 20,685 kPa (3,000 psia) and 24,133 kPa (3,500 psia). 
     The pipeline, which may be composed of alloy steel, is preferably provided with thermal insulation which is designed to ensure that temperature of the output gas is lower than the temperature of the input gas. Suitable insulating materials are well known to those skilled in the art. The pipeline metal is preferably a high-strength, low-alloy steel containing less than about three weight percent nickel and having strength and toughness for containing the natural gas at the operating conditions of this invention. Example steels for use in constructing the pipeline of this invention are described in U.S. Pat. Nos. 5,531,842; 5,545,269; and 5,545,270. 
     The pipeline  47  may be buried in the ground or in the sea floor, or laid on the ground or sea floor, or elevated above the ground or sea floor, or any combination of the foregoing, depending on where the gas is being transported. 
     The pressure of the pipeline output gas (stream  7 ) preferably ranges between about 3,450 kPa (500 psia) and 10,340 kPa (1,500 psia), and more preferably between about 3,790 kPa (550 psia) and 8,620 kPa (1,250 psia). If the output gas pressure is below about 500 psia, the gas pressure can be pressurized by a suitable compression means (not shown), which may comprise one or more compressors that compress the gas to at least 500 psia before the gas enters the liquefaction plant. The temperature of the natural gas output from pipeline  47  preferably ranges between about −29° C. (−20° F.) and −73° C. (−100° F.), and more preferably between about 29° C. (−20° F.) and −62° C. (−80° F.). Although the output gas from the pipeline may be introduced directly to phase separator  54 , the pipeline output gas is preferably further cooled by an external refrigeration system and it is preferably still further cooled by pressure expansion. As shown the FIG. 1, the pipeline output gas is preferably cooled by a cooling system  48  which may comprise any conventional closed-circuit refrigeration system, preferably a closed-cycle propane refrigeration system, and more preferably a closed-cycle refrigeration system containing a mixture of C 1 , C 2 , C 3 , C 4 , and C 5  as a refrigerant. The output from the cooling system  48  is further cooled by an expander zone  49  which comprises a mechanical expander or a throttling valve, or both, to achieve a predetermined final output pressure and temperature. Expander zone  49 , preferably comprising one or more turboexpanders, which at least partially liquefies the gas stream. 
     The metallurgy, diameter, and operating pressure of pipeline  47  and the gas feed conditions (stream  6 ) to the pipeline  47  can be optimized by those skilled in the art in view of the teachings of this description to eliminate costly pipeline recompression systems and thereby minimize the overall cost of the pipeline system. The temperature and pressure conditions for the cooling system  48  and the expander zone  49  can also be optimized by those skilled in the art taking in account the teaching of this description to fully use the Joule-Thomson cooling in the pipeline  47  and thereby maximize the gas volume available to consumers. 
     Natural gas introduced to phase separator  54  is separated into a liquid stream  13  and a vapor stream  12 . The liquid stream  13  will typically need to be pressure regulated in pressure adjustment zone  70  to a pressure approximately the same as the operating pressure of the phase separator  65 . In most applications of this invention, the pressure of stream  13  will not be the same as the operating pressure of phase separator  65 . If the pressure of stream  13  is less than the operating pressure of separator  65 , pressure adjustment zone  70  preferably comprises a pump to increase the pressure of stream  13  to approximately the same pressure of fluid in separator  65 . If the pressure of stream  13  is greater than the operating pressure of separator  65 , pressure adjustment zone  70  preferably comprises an expander, such as a hydraulic turbine, to lower the pressure to the pressure of fluid in separator  65 . 
     The vapor stream  12  from the phase separator  54  is passed to a compression zone  55  to pressurize stream  12 . The compression zone preferably comprises a heat exchanger  56  through which stream  12  is warmed before passing as stream  15  to at least two compressors  57  and  59 , with at least one heat exchanger  58  between compressors  57  and  59  and one at least one heat exchanger  60  after the last compressor  69 . The vapor stream  19  exiting heat exchanger  60  is passed through heat exchanger  56  to be further cooled by indirect heat exchange with the incoming vapor stream  12 . 
     This invention is not limited to any type of heat exchanger, but because of economics, plate-fin, spiral wound, and cold box heat exchangers are preferred, which all cool by indirect heat exchange. The term “indirect heat exchange,” as used in this description and claims, means the bringing of two fluid streams into heat exchange relation without any physical contact or intermixing of the fluids with each other. 
     From the compression zone  55 , the compressed gas stream  20  passes through heat exchanger  61  which is cooled with overhead vapor stream  26  from the phase separator  65 . From the heat exchanger  61 , stream  21  then passes through an expander zone  62 , preferably one or more hydraulic turbines to reduce the pressure and temperature of the gas stream and thereby at least partially liquefying the gas stream. The at least partially liquefied gas (stream  22 ) then passes to phase separator  63  which separates the liquid and vapor, producing vapor stream  24  and liquid stream  23 . A fraction of vapor stream  24  is returned to the phase separator  54  for recycling. A second fraction of stream  24  is withdrawn as stream  36  and passed through heat exchanger  61  to heat stream  36 . From the heat exchanger  61 , the heated stream (stream  37 ) is further heated by heat exchanger  67  to produce a heated stream  31  suitable for use as fuel. This fuel may provide energy for powering turbines that partially power the compressors in compression zone  55 . 
     The liquid stream  23  produced by separator  63  is passed to another expander zone  64 , preferably one hydraulic turbine, to further reduce the pressure and temperature of the liquid stream. Stream  25  from the expander zone  64  then passes to phase separator  65 . The expanders of expander zones  62  and  64  are preferably used to provide at least part of the power for the compressors  57  and  59 . 
     Phase separator  65  produces a vapor stream  26  and a liquid stream  27 . The liquid stream  27  passes to a suitable container such as a stationary storage vessel or a suitable carrier such as a ship, barge, submarine vessel, railroad tank car, or truck. In accordance with the practice of this invention, liquid stream  27  will have a temperature above about −112° C. (−170° F.) and a pressure sufficient for the liquid to be at or below its bubble point. 
     The vapor stream  26  passes through heat exchanger  61  to provide cooling to vapor stream  20  by indirect heat exchange. From heat exchanger  61 , stream  29  passes through another heat exchanger  67  and is then compressed by compressor  68  to a pressure approximately the same as the pressure of phase separator  54 . The compressed gas (stream  32 ) is then cooled in a conventional aftercooler  69  by air or water, and then further cooled by heat exchanger  34  before being combined with stream  24  and returned to phase separator  54  for recycling. 
     In the storage, transportation, and handling of liquefied natural gas, there can be a considerable amount of boil-off vapor resulting from evaporation. The process of this invention can optionally liquefy the boil-off gas. Referring to FIG. 1, the boil-off vapor  28  is preferably introduced to the liquefication process by being combined with vapor stream  26 . Although not shown in FIG. 1, the boil-off vapor preferably is introduced to the process at the same pressure as stream  26 . Although not shown in FIG. 1, the boil-off gas will typically need to be pressurized by a compressor or de-pressurized by an expander before being introduced to stream  26 . 
     FIG. 2 illustrates another embodiment of this invention, and in this embodiment the parts having like numerals to those in FIG. 1 have the same process functions. Those skilled in the art will recognize, however, that the process equipment from one embodiment to another may vary in size and capacity to handle different fluid flow rates, temperatures, and compositions. The embodiment of FIG. 2 is similar to the embodiment of FIG. 1 except that the cooling zone  48  and expansion zone  49  of FIG. 1 are not used in the embodiment of FIG.  2  and in FIG. 2 the fuel gas (stream  31 ) is withdrawn from vapor overhead of separator  65  whereas in FIG. 1 fuel gas (stream  38 ) is withdrawn from vapor overhead of separator  63 . 
     To minimize compression power required for liquefaction when appreciable nitrogen exists in natural gas feed stream  5  and/or in the boil-off vapor stream  28 , the nitrogen concentration is preferably concentrated and removed at some location in the process. The process of this invention concentrates nitrogen as vapor streams  24  and  26 , with vaporous stream  24  having a higher concentration of nitrogen than vaporous stream  26 . In FIG. 1, a portion of vapor stream  24  is removed as a fuel gas (stream  31 ) and in FIG. 2 a portion of vapor stream  26  is removed as fuel gas. 
     EXAMPLE 
     A simulated mass and energy balance was carried out to illustrate the embodiment illustrated in the Figures, and the results are set forth in Tables 1 and 2 below. Table 1 corresponds to the embodiment shown in FIG.  1  and Table 2 corresponds to the embodiment shown in FIG.  2 . The temperatures, pressures, and flow rates presented in the Tables are not to be considered as limitations upon the invention which can have many variations in temperatures and flow rates in view of the teachings herein. 
     In both simulations, it was assumed that natural gas was fed to a 284 mile, 21 inch pipeline that was buried in permafrost in the North Slope of Alaska. In the first simulation (Table 1), it was assumed that the gas composition comprised 85.9 mole percent methane, 13.5 mole percent ethane and heavier hydrocarbons, 100 parts per million CO 2 , and 0.6 mole percent N 2 . In the second simulation (Table 2), it was assumed that the gas composition comprised 94.5 mole percent methane, 5 mole percent ethane and heavier hydrocarbons, 100 parts per million CO 2  and 0.5 mole percent N 2 . 
     In the first simulation, the pipeline inlet pressure (stream  6  of FIG. 1) was assumed to be 22,754 kPa (3,300 psia) In the second simulation, the pipeline inlet pressure (stream  6  of FIG. 2) was assumed to be 48,266 kPa (7,000 psia). FIG. 2 is optimum when the overall cost of the pipeline system is minimized for 3,450 kPa (500 psia) delivery with a starting pressure of 48,266 kPa (7,000 psia). 
     The data were obtained using a commercially available process simulation program called HYSYS™, marketed by Hyprotech Ltd. of Calgary, Canada; however, other commercially available process simulation programs can be used to develop the data, including for example HYSIM™, PROII™, and ASPEN PLUS™, all of which are familiar to those of ordinary skill in the art. 
     A person skilled in the art, particularly one having the benefit of the teachings of this patent, will recognize many modifications and variations to the specific processes disclosed above. For example, a variety of temperatures and pressures may be used in accordance with the invention, depending on the overall design of the system and the composition of the feed gas. Also, the feed gas cooling train may be supplemented or reconfigured depending on the overall design requirements to achieve optimum and efficient heat exchange requirements. As discussed above, the specifically disclosed embodiments and examples should not be used to limit or restrict the scope of the invention, which is to be determined by the claims below and their equivalents. 
     
       
         
           
               
               
             
               
                   
                 TABLE 1 
               
             
            
               
                   
                   
               
               
                   
                 Composition 
               
            
           
           
               
               
               
               
               
               
               
               
               
               
               
               
            
               
                   
                   
                 Pressure 
                 Pressure 
                 Temp. 
                 Temp. 
                 Flowrate 
                 Flowrate 
                 C 1   
                 C 2+   
                 CO 2   
                 N 2   
               
               
                 Stream 
                 Phase 
                 kPa 
                 psia 
                 Deg C. 
                 Deg F. 
                 KgMol/hr 
                 #mol/hr 
                 mol % 
                 mol % 
                 ppmv 
                 mol % 
               
               
                   
               
            
           
           
               
               
               
               
               
               
               
               
               
               
               
               
            
               
                  6 
                 vapor 
                 22,754  
                 3,300   
                 −0.8 
                 30.0 
                 37,534 
                 82,747 
                 85.9 
                 13.5 
                 100  
                 0.6 
               
               
                  7 
                 vapor 
                 8,619 
                 1,250   
                 −29.2 
                 −21.1 
                 37,534 
                 82,747 
                 85.9 
                 13.5 
                 100  
                 0.6 
               
               
                  9 
                 vapor/liquid 
                 3,517 
                 510 
                 −65.2 
                 −85.9 
                 37,534 
                 82,747 
                 85.9 
                 13.5 
                 100  
                 0.6 
               
               
                 12 
                 vapor 
                 3,517 
                 510 
                 −68.6 
                 −92.0 
                 54,523 
                 120,200  
                 94.3 
                 4.1 
                 64 
                 1.6 
               
               
                 13 
                 liquid 
                 3,517 
                 510 
                 −68.6 
                 −92.0 
                  6,904 
                 15,220 
                 55.7 
                 44.1 
                 133  
                 0.2 
               
               
                 14 
                 vapor/liquid 
                 2,675 
                 388 
                 −76.3 
                 −106.0 
                  6,904 
                 15,220 
                 55.7 
                 44.1 
                 133  
                 0.2 
               
               
                 15 
                 vapor 
                 3,496 
                 507 
                 13.7 
                 56.0 
                 54,523 
                 120,200  
                 94.3 
                 4.1 
                 64 
                 1.6 
               
               
                 16 
                 vapor 
                 7,240 
                 1,050   
                 79.8 
                 175.1 
                 54,523 
                 120,200  
                 94.3 
                 4.1 
                 64 
                 1.6 
               
               
                 17 
                 vapor 
                 7,205 
                 1,045   
                 15.9 
                 60.0 
                 54,523 
                 120,200  
                 94.3 
                 4.1 
                 64 
                 1.6 
               
               
                 18 
                 vapor 
                 24,133  
                 3,500   
                 127.7 
                 261.2 
                 54,523 
                 120,200  
                 94.3 
                 4.1 
                 64 
                 1.6 
               
               
                 19 
                 vapor 
                 24,064  
                 3,490   
                 15.9 
                 60.0 
                 54,523 
                 120,200  
                 94.3 
                 4.1 
                 64 
                 1.6 
               
               
                 20 
                 vapor 
                 24,043  
                 3,487   
                 −42.7 
                 −45.4 
                 54,523 
                 120,200  
                 94.3 
                 4.1 
                 64 
                 1.6 
               
               
                 21 
                 vapor 
                 24,009  
                 3,482   
                 −51.2 
                 −60.7 
                 54,523 
                 120,200  
                 94.3 
                 4.1 
                 64 
                 1.6 
               
               
                 22 
                 vapor/liquid 
                 3,517 
                 510 
                 −89.5 
                 −129.7 
                 54,523 
                 120,200  
                 94.3 
                 4.1 
                 64 
                 1.6 
               
               
                 23 
                 liquid 
                 3,517 
                 510 
                 −89.5 
                 −129.7 
                 40,860 
                 90,080 
                 93.7 
                 5.2 
                 76 
                 1.1 
               
               
                 24 
                 vapor 
                 3,517 
                 510 
                 −89.5 
                 −129.7 
                 13,313 
                 29,350 
                 96.2 
                 0.8 
                 29 
                 3.0 
               
               
                 25 
                 vapor/liquid 
                 2,620 
                 380 
                 −98.3 
                 −145.5 
                 41,187 
                 90,800 
                 93.7 
                 5.2 
                 76 
                 1.1 
               
               
                 26 
                 vapor 
                 2,620 
                 380 
                 −95.7 
                 −140.9 
                  8,777 
                 19,350 
                 96.5 
                 0.7 
                 25 
                 2.8 
               
               
                 27 
                 liquid 
                 2,620 
                 380 
                 −95.7 
                 −140.9 
                 39,314 
                 86,670 
                 86.4 
                 13.0 
                 97 
                 0.6 
               
               
                 28 
                 vapor 
                 2,658 
                 386 
                 −94.1 
                 −138.0 
                  2,780 
                  6,129 
                 97.2 
                 1.0 
                 33 
                 1.8 
               
               
                 29 
                 vapor 
                 2,586 
                 375 
                 −44.6 
                 −48.9 
                 11,558 
                 25,480 
                 96.6 
                 0.8 
                 27 
                 2.6 
               
               
                 30 
                 vapor 
                 2,565 
                 372 
                 11.4 
                 52.0 
                 11,558 
                 25,480 
                 96.6 
                 0.8 
                 27 
                 2.6 
               
               
                 32 
                 vapor 
                 3,585 
                 520 
                 41.3 
                 105.8 
                 11,558 
                 25,480 
                 96.6 
                 0.8 
                 27 
                 2.6 
               
               
                 33 
                 vapor 
                 3,565 
                 517 
                 15.9 
                 60.0 
                 11,558 
                 25,480 
                 96.6 
                 0.8 
                 27 
                 2.6 
               
               
                 34 
                 vapor 
                 3,544 
                 514 
                 −42.4 
                 −44.9 
                 11,558 
                 25,480 
                 96.6 
                 0.8 
                 27 
                 2.6 
               
               
                 35 
                 vapor 
                 3,517 
                 510 
                 −70.3 
                 −95.1 
                 23,873 
                 52,630 
                 96.4 
                 0.8 
                 28 
                 2.8 
               
               
                 37 
                 vapor 
                 3,517 
                 505 
                 −44.6 
                 −48.9 
                   998 
                  2,200 
                 96.2 
                 2.8 
                 29 
                 3 
               
               
                 38 
                 vapor 
                 3,461 
                 502 
                 11.4 
                 52.0 
                 16,488 
                 36,350 
                 96.2 
                 0.8 
                 29 
                 3.0 
               
               
                   
               
            
           
         
       
     
     
       
         
           
               
               
             
               
                   
                 TABLE 2 
               
             
            
               
                   
                   
               
               
                   
                 Composition 
               
            
           
           
               
               
               
               
               
               
               
               
               
               
               
               
            
               
                   
                   
                 Pressure 
                 Pressure 
                 Temp. 
                 Temp. 
                 Flowrate 
                 Flowrate 
                 C 1   
                 C 2+   
                 CO 2   
                 N 2   
               
               
                 Stream 
                 Phase 
                 kPa 
                 psia 
                 Deg C. 
                 Deg F. 
                 KgMol/hr 
                 Lb mol/hr 
                 mol % 
                 mol % 
                 ppmv 
                 mol % 
               
               
                   
               
            
           
           
               
               
               
               
               
               
               
               
               
               
               
               
            
               
                  6 
                 vapor 
                 48,266  
                 7,000   
                 −0.8 
                 30.0 
                 34,750 
                  76,610 
                 94.5 
                 5.0 
                 100  
                 0.6 
               
               
                  9 
                 vapor/liquid 
                 3,448 
                 500 
                 −76.2 
                 −105.8 
                 34,750 
                  76,610 
                 94.5 
                 5.0 
                 100  
                 0.6 
               
               
                 12 
                 vapor 
                 3,448 
                 500 
                 −76.2 
                 −105.8 
                 49,715 
                 109,600 
                 96.3 
                 2.7 
                 75 
                 1.0 
               
               
                 13 
                 liquid 
                 3,448 
                 500 
                 −76.2 
                 −105.8 
                  1,383 
                  3,048 
                 65.0 
                 34.8 
                 189  
                 0.2 
               
               
                 14 
                 vapor/liquid 
                 2,675 
                 388 
                 −83.8 
                 −119.4 
                  1,383 
                  3,048 
                 65.0 
                 34.8 
                 189  
                 0.2 
               
               
                 15 
                 vapor 
                 3,427 
                 497 
                 9.8 
                 49.0 
                 49,715 
                 109,600 
                 96.3 
                 2.7 
                 75 
                 1.0 
               
               
                 16 
                 vapor 
                 7,240 
                 1,050   
                 77.8 
                 171.4 
                 49,715 
                 109,600 
                 96.3 
                 2.7 
                 75 
                 1.0 
               
               
                 17 
                 vapor 
                 7,205 
                 1,045   
                 11.4 
                 52.0 
                 49,715 
                 109,600 
                 96.3 
                 2.7 
                 75 
                 1.0 
               
               
                 18 
                 vapor 
                 24,133  
                 3,500   
                 122.8 
                 252.5 
                 49,715 
                 109,600 
                 96.3 
                 2.7 
                 75 
                 1.0 
               
               
                 19 
                 vapor 
                 24,064  
                 3,490   
                 11.4 
                 52.0 
                 49,715 
                 109,600 
                 96.3 
                 2.7 
                 75 
                 1.0 
               
               
                 20 
                 vapor 
                 24,043  
                 3,487   
                 −50.4 
                 −59.4 
                 49,715 
                 109,600 
                 96.3 
                 2.7 
                 75 
                 1.0 
               
               
                 21 
                 vapor 
                 24,009  
                 3,482   
                 −57.4 
                 −71.9 
                 49,715 
                 109,600 
                 96.3 
                 2.7 
                 75 
                 1.0 
               
               
                 22 
                 vapor/liquid 
                 3,517 
                 510 
                 −90.2 
                 −131.0 
                 49,715 
                 109,600 
                 96.3 
                 2.7 
                 75 
                 1.0 
               
               
                 23 
                 liquid 
                 3,517 
                 510 
                 −90.2 
                 −131.0 
                 42,865 
                  94,500 
                 96.1 
                 3.1 
                 82 
                 0.8 
               
               
                 24 
                 vapor 
                 3,517 
                 510 
                 −90.2 
                 −131.0 
                  6,863 
                  15,130 
                 97.4 
                 0.5 
                 32 
                 2.1 
               
               
                 25 
                 vapor/liquid 
                 2,620 
                 380 
                 −99.0 
                 −146.8 
                 42,865 
                  94,500 
                 96.1 
                 3.1 
                 82 
                 0.8 
               
               
                 26 
                 vapor 
                 2,620 
                 380 
                 −98.5 
                 −145.9 
                  7,689 
                  16,950 
                 97.6 
                 0.4 
                 26 
                 2.0 
               
               
                 27 
                 liquid 
                 2,620 
                 380 
                 −98.5 
                 −145.9 
                 36,560 
                  80,600 
                 94.6 
                 4.8 
                 97 
                 0.6 
               
               
                 28 
                 vapor 
                 2,658 
                 386 
                 −94.1 
                 −138.0 
                  2,573 
                  5,672 
                 97.2 
                 1.0 
                 33 
                 1.8 
               
               
                 29 
                 vapor 
                 2,599 
                 377 
                 −52.6 
                 −63.2 
                 10,260 
                  22,620 
                 97.5 
                 0.5 
                 28 
                 2.0 
               
               
                 30 
                 vapor 
                 2,579 
                 374 
                 8.1 
                 46.0 
                 10,260 
                  22,620 
                 97.5 
                 0.5 
                 28 
                 2.0 
               
               
                 31 
                 vapor 
                 2,579 
                 374 
                 8.1 
                 46.0 
                   768 
                  1,693 
                 97.5 
                 0.5 
                 28 
                 2.0 
               
               
                 32 
                 vapor 
                 3,585 
                 520 
                 37.3 
                 98.6 
                  9,494 
                  20,930 
                 97.5 
                 0.5 
                 28 
                 2.0 
               
               
                 33 
                 vapor 
                 3,565 
                 517 
                 11.4 
                 52.0 
                  9,494 
                  20,930 
                 97.5 
                 0.5 
                 28 
                 2.0 
               
               
                 34 
                 vapor 
                 3,544 
                 514 
                 −50.8 
                 −60.1 
                  9,494 
                  20,930 
                 97.5 
                 0.5 
                 28 
                 2.0 
               
               
                 35 
                 vapor 
                 3,517 
                 510 
                 −70.4 
                 −95.3 
                 16,357 
                  36,060 
                 97.4 
                 0.6 
                 30 
                 2.0