Patent Publication Number: US-8535415-B2

Title: Refinery gas upgrading via partial condensation and PSA

Description:
RELATED APPLICATIONS 
     This application is a continuation of prior U.S. application Ser. No. 12/515,181, filed Jan. 12, 2010 now U.S. Pat. No. 8,262,772, which claims priority from International Application No. PCT/US07/086,488, filed Dec. 5, 2007, which claims priority from U.S. Provisional Application Ser. No. 60/872,919, filed Dec. 5, 2006, which is incorporated herein by reference. 
    
    
     TECHNICAL FIELD 
     The present invention relates generally to the recovery of valuable by-products from refinery gas streams, such as hydrogen. 
     BACKGROUND ART 
     The petroleum industry has often sought new integration opportunities for its refinery products with other processes. One of the areas of interest concerns refinery gases that are currently used as a fuel. In addition, refineries are processing heavier crude oils and sulfur specifications for both diesel and gasoline products are becoming more stringent. Hydrogen can be used within hydrotreaters to remove sulfur, oxygen, and nitrogen and also within hydrocrackers to produce lighter and more parafinic oils. Consequently, refineries are looking for low cost sources of hydrogen. 
     While the refinery gases are a potential source of hydrogen, many refinery gas streams are not used either for their hydrogen content or to generate hydrogen through known reforming techniques due to a variety of economic and practical reasons. For instance, the economics for separating hydrogen from refinery gases that contain less than about 30% hydrogen are generally unfavorable. Generally, the hydrogen concentration within such refinery off-gases is too low for the hydrogen to be economically recovered using current available separation technologies. 
     Several processes are known in the art for the separation and recovery of hydrogen from hydrogen-hydrocarbon feed gas streams. Among these are the following: 
     U.S. Pat. No. 3,838,553 (Doherty) describes a combination pressure swing adsorption (PSA) and cryogenic process to recover a light gas, especially hydrogen or helium, at high purity and at high recovery from a multicomponent gaseous mixture. The process described by Doherty utilizes both a low temperature separator unit (LTU) and a pressure swing adsorber with recompression of the regeneration stream from the latter and recycle of the recompressed stream to the LTU. However, Doherty requires upgrading the hydrogen stream to at least 95% hydrogen prior to PSA purification. Also, Doherty teaches the recycling of the PSA tailgas, which is uneconomical for PSAs with reasonable recovery (e.g., higher than 80%) due to high compression costs. Furthermore, Doherty does not consider the recovery of heavy hydrocarbons (ethane and heavier). 
     U.S. Pat. No. 3,691,779 (Meisler et al.) describes a process consisting of a low temperature refrigeration unit and a PSA for producing a high purity, 97 to 99.9% hydrogen. The hydrogen-rich feed, containing methane, nitrogen, carbon monoxide and traces of argon, oxygen, carbon dioxide and light hydrocarbons, passes through a series of cooling and condensation stages having successively lower temperatures, the lowest being −340° F. (120° R). Hydrogen containing vapors and condensate are separated between cooling stages. The hydrogen-enriched gas is sent for further upgrading to an adsorption system operating between −260 to −320° F. (200 to 140° R). A portion of the upgraded stream is expanded, passed through at least one refrigeration stage to provide refrigeration, and then used for regeneration in the PSA system. Most of the refrigeration is provided by Joule-Thomson expansion of the condensates. Meisler et al. teach the use of a capital intensive system, due to the multiple steps required for refrigeration in order to upgrade the feed to approximately 97% hydrogen before sending to PSA. Meisler et al. also teach very low temperature levels in the cold box, driven mostly by the amount of non-condensable compounds present in the feed and the high purity required for the hydrogen stream. 
     U.S. Pat. No. 7,041,271 (Drnevich et al.) discloses an integrated method for olefins recovery and hydrogen production from a refinery off-gas. After conventional pretreatment, the refinery gas is separated to obtain a light ends stream containing hydrogen, nitrogen and methane, and a heavy ends stream containing olefins. The light ends stream is mixed with natural gas and subjected to reforming and water gas shift reactions for hydrogen production. The heavy ends can be further processed for olefin recovery, such as ethylene and propylene. Drnevich et al. teach the recovery one light end stream from the refinery off-gas. Also, the present invention does not consider the further processing of C 2 + hydrocarbon stream for olefin/liquefied petroleum gas (LPG) recovery. Furthermore, Drnevich et al. teach the use of low temperature distillation, membrane, PSA and adsorption-desorption processes as means for light end separation, but partial condensation is not discussed. 
     U.S. Pat. No. 4,749,393 (Rowles et al.) describes a hybrid gas separation process which recovers both heavy hydrocarbon (C2+, C3+ or C4+) and high purity hydrogen products from a gas containing a relatively low concentration of hydrogen (&lt;40%). The process comprises a warm heat exchanger where the feed (together with recycle from the hydrogen purifier) is cooled to an intermediate temperature to allow the recovery of heavy hydrocarbons, a separator coupled with a dephlegmator (reflux condenser) for enriching the heavy hydrocarbons condensate, and a cold heat exchanger followed by a separator, from which an enriched hydrogen gas is obtained and, after warming to recover the refrigeration, is sent to the hydrogen purifier (e.g., PSA, membrane). An optional turboexpander or compressor (depending on the hydrogen purifier requirements) can be added on the hydrogen-enriched stream. The condensate from the cold end separator, rich in methane (80-85%) is used to generate the refrigeration through Joule-Thomson expansion and then is sent to fuel. Additional refrigeration is generated by Joule-Thomson expansion of the condensate from the warm end separator/dephlegmator. The tailgas from the hydrogen purifier (PSA) is recycled to the feed, which can be uneconomical due to recompression requirements. In addition, the use of the dephlegmator for enriching the hydrocarbon condensate can be expensive. 
     As will be discussed, the present invention provides a process for recovering valuable products, especially hydrogen, from refinery fuel gases in order to economically and practically produce such products and increase efficiency and lower costs. 
     SUMMARY OF THE INVENTION 
     The present invention relates to a process and system for recovering high purity hydrogen and optionally one or more hydrocarbon-rich products from a feed gas containing less than 50 mole % hydrogen via partial condensation in an auto-refrigeration cold box and PSA. The process can be used stand-alone, only for hydrogen recovery, or can be modified to allow the recovery of other valuable by-products such as methane-rich gas and raw LPG (methane depleted gas containing ethane and heavier hydrocarbons, usually used as feedstock for LPG recovery). In the stand-alone process, the feed is compressed, treated for water and impurities removal, cooled in a cold box and then sent to a separator. The separator liquid is used to provide refrigeration through Joule-Thomson expansion and, after warming in the heat exchanger against the incoming feed, is sent to fuel. The high pressure vapor from the separator, containing at least 60% hydrogen, is warmed against the incoming feed and sent to PSA for purification. The PSA tailgas is sent to fuel after recompression. 
     In another embodiment of the invention, the feed, after compression and treatment, is cooled in two stages. After the first cooling stage (warm end heat exchanger) the feed is separated at high pressure and a temperature chosen such as to maximize the recovery of the C2+ hydrocarbons in the liquid phase. The gas phase is sent to a second cooling stage (cold end heat exchanger), and then separated into a hydrogen-rich stream containing at least 60% hydrogen and a liquid containing at least 70% methane. The condensate, after Joule-Thomson expansion to provide refrigeration and warming in the cold end heat exchanger, is used as stripping gas in a low pressure wash column, to enhance hydrocarbon recovery in the condensate. The hydrogen-rich vapor is warmed against incoming feed in both cold end and warm end heat exchangers and then sent to PSA for purification. The PSA tailgas is sent, after recompression, to fuel. Wash column vapor product is, after warming against incoming feed, sent to fuel or, after further compression and cooling, sent as SMR feed. The liquid product from the wash column is, after warming against incoming feed, sent to further processing for LPG recovery. 
     Compared to the prior art, the present invention is more efficient. For example, Doherty teaches that the hydrogen stream must be at 95% prior to PSA purification, whereas the present invention requires only around 60% hydrogen. The present invention also uses a wash column for recovering C2+ hydrocarbons from the gaseous phase. Using the wash column arrangement, a simpler configuration is used to obtain basically the same quality products, leading to lower capital costs. As compared to the invention of Drnevich et al., the present invention separates two light end streams, a low pressure methane-rich stream, which is blended with natural gas and sent as feed to a steam methane reformer (SMR) and a hydrogen-rich stream at high pressure, which can be fed directly to the PSA. This allows for more flexibility in blending in order to meet certain requirements for SMR feed. For example, the hydrogen needs for the hydrotreating step can be satisfied by mixing a part of the hydrogen-rich stream with the SMR feed. 
    
    
     
       BRIEF DESCRIPTION OF THE DRAWINGS 
       While the specification concludes with claims distinctly pointing at the subject matter that applicants regards as their invention, it is believed that the invention will be better understood when taken in connection with the accompanying drawings: 
         FIG. 1  is a schematic for an embodiment of the invention for the recovery of hydrogen. 
         FIG. 2  is a schematic for an embodiment of the invention for the recovery of hydrogen, raw LPG and methane-rich gas. 
         FIG. 3  is a schematic for an embodiment of the invention for the recovery of hydrogen, raw LPG and SMR feed. 
     
    
    
     DETAILED DESCRIPTION 
     The present invention relates to a process and system for recovering high purity hydrogen and optionally one or more hydrocarbon-rich product from a feed gas containing less than 50 mole % hydrogen via partial condensation in an auto-refrigeration cold box and PSA. The process can be used stand-alone, only for hydrogen recovery, or can be modified to allow the recovery of other valuable by-products such as methane-rich gas and raw LPG (methane depleted gases containing ethane and heavier hydrocarbons, usually used as feedstock for LPG recovery). As used herein, the percentage of hydrogen or other components may be expressed as “%” or “mole %”. In the stand-alone process ( FIG. 1 ), the feed is compressed, treated for water and impurities removal, cooled in a cold box and then sent to a separator. The separator liquid is used to provide refrigeration through Joule-Thomson expansion and, after warming in the heat exchanger against the incoming feed, is sent to fuel. The high pressure vapor from the separator, containing at least 60 mole % hydrogen, is warmed against the incoming feed and sent to PSA for purification. The PSA tailgas is sent to fuel after recompression. 
     In another embodiment of the invention ( FIGS. 2 and 3 ) the feed, after compression and treatment, is cooled in two stages. After the first cooling stage (warm end heat exchanger) the feed is separated at high pressure and a temperature chosen such as to maximize the recovery of the C2+ hydrocarbons in the liquid phase. The gas phase is sent to a second cooling stage (cold end heat exchanger), and then separated into a hydrogen-rich stream containing at least 60% hydrogen and a liquid containing at least 70% methane. The condensate, after Joule-Thomson expansion to provide refrigeration and warming in the cold end heat exchanger, is used as stripping gas in a low pressure wash column, to enhance hydrocarbon recovery in the condensate. The hydrogen-rich vapor is warmed against incoming feed in both cold end and warm end heat exchangers and then set to PSA for purification. The PSA tailgas is sent, after recompression, to fuel. Wash column vapor product is, after warming against incoming feed, sent to fuel or, after further compression and cooling, sent as SMR feed. The liquid product from the wash column is, after warming against incoming feed, sent to further processing for LPG recovery. 
     One embodiment of the present invention is a process for recovering high purity hydrogen and optionally one or more hydrocarbon-rich products from a feed gas containing less than 50 mole % hydrogen, as represented by  FIG. 1 , comprising the following steps:
         a) cooling a feed gas  1  in one or more heat exchangers and producing a partially condensed feed gas  4 ;   b) separating the partially condensed feed gas  4  to form a hydrogen-rich stream  5  and a hydrocarbon-rich stream  11 ;   c) generating refrigeration by Joule-Thomson expansion of the hydrocarbon-rich stream  11  and forming a colder hydrocarbon-rich stream  12 ;   d) recovering refrigeration from the colder hydrocarbon-rich stream  12  to cool the feed gas  1  in step (a);   e) recovering refrigeration from the hydrogen-rich stream  5  to cool the feed gas  1  in step (a);   f) recovering a gas stream  6  containing at least 60 mole % hydrogen from the hydrogen-rich stream  5 ; and   g) feeding the gas stream  6  to a PSA system  107  and recovering high purity hydrogen product  8  with or without feeding an external hydrogen containing stream  7  to the PSA system.       

     Another embodiment of the present invention is a process for recovering high purity hydrogen and optionally one or more hydrocarbon-rich products from a feed gas containing less than 50 mole % hydrogen, as represented by  FIGS. 2 and 3 , comprising the following steps:
         a) cooling a feed gas  21 ,  51  in one or more heat exchangers and producing a partially condensed feed gas  24 ,  54 ;   b) separating the partially condensed feed gas  24 ,  54  to form a hydrogen-enriched stream  25 ,  55  and a hydrocarbon-enriched stream  26 ,  56 ;   c) generating refrigeration by Joule-Thomson expansion of the hydrocarbon-enriched stream  26 ,  56  and forming a colder hydrocarbon-rich stream  34 ,  61 ;   d) further cooling the hydrogen-enriched stream  25 ,  55  to produce a partially condensed hydrogen-enriched stream  27 ,  57 ;   e) separating the partially condensed hydrogen-enriched stream  27 ,  57  into a hydrogen-rich stream  27 ,  66  and a hydrocarbon-rich stream  35 ,  58 ;   f) generating refrigeration by Joule-Thomson expansion of hydrogen-rich stream  28 ,  66  and hydrocarbon-rich stream  35 ,  58 ;   g) recovering refrigeration from hydrogen-rich stream  28 ,  66  to cool the hydrogen-enriched stream  25 ,  55  in step (d) and the feed gas  21 ,  51  in step (a);   h) recovering a gas stream  29 ,  69  containing at least 60 mole % hydrogen from the hydrogen-rich stream  28 ,  66 ;   i) feeding gas stream  29 ,  69  to a PSA system  219 ,  317  and recovering high purity hydrogen;   j) recovering refrigeration from cold hydrocarbon-rich stream  36 ,  59  to cool the hydrogen-enriched stream  25 ,  55  in step (d);   k) producing a recontacting stream  38 ,  60  from the partially condensed hydrogen-enriched stream  27 ,  57 ;   l) feeding the cold hydrocarbon-enriched stream  34 ,  61  and the recontacting stream  38 ,  60  to a wash column and producing a methane-rich vapor stream  39 ,  62  and a hydrocarbon-rich liquid stream  41 ,  75 ;   m) recovering refrigeration from the methane-rich vapor stream  39 ,  62  and compressing it to form methane-rich vapor product  40 ,  63 ; and   n) recovering refrigeration from the hydrocarbon-rich liquid stream  41 ,  75  and compressing it to form raw LPG product  42 ,  76 .       

     A third embodiment of the present invention is represented by  FIG. 3  and comprises the steps set forth above and represented by  FIG. 2 , plus the additional step of processing methane-rich product  63  in a steam methane reforming process to produce hydrogen containing feed stream  69  to the PSA system  219 ,  317 . 
     Process 1: Hydrogen Only ( FIG. 1 ) 
     Referring to  FIG. 1 , a refinery gas  1  containing between 30-50 mole % hydrogen, at about 100° F. and 50-120 psia, is compressed in a multistage compressor  101 , cooled to about 100° F. in compressor cooler  102 , and sent through line  2  to a pretreatment system  103 , comprising a drier and one or multiple adsorbing beds for impurities removal (e.g., CO 2 , H 2 S, NH 3 , benzene, n-hexane, mercury). The treated gas is passed by line  3  through heat exchanger  104 , where it is cooled to a temperature such that the stream is partially condensed and the vapor contains at least 60 mole % hydrogen. The partially condensed stream is fed by line  4  to separator  105 , where it separates into a hydrogen-rich gas and a practically hydrogen-free condensate. The hydrogen-rich gas, containing at least 60 mole % hydrogen, leaves the separator  105  by line  5 , is expanded if needed to PSA pressure in valve  106 , then warmed to about 90° F. in heat exchanger  104 , and fed through line  6  to pressure swing adsorption (PSA) system  107 , operating at about 350 psia. Optionally, a high purity hydrogen stream containing 70-90 mole % hydrogen can be fed, together with hydrogen containing stream  6 , to PSA system  107  through feed line  7  after compression, if required, in compressor  108  and cooling to about 100° F. in compressor cooler  109 . A high pressure, high purity hydrogen gas, containing at least 99 mole % hydrogen, leaves the PSA through product line  8 , and a low pressure tailgas (at about 20 psia) is rejected through line  9 . The PSA tailgas, comprising methane, residual hydrogen, and other non-readily condensable compounds in lesser amounts (e.g., N 2 , CO, CO 2 ) is sent to fuel through line  10 , after recompression to the fuel header pressure (i.e., 50-120 psia) in compressor  110  and cooling to about 100° F. in compressor cooler  111 . The hydrogen-free condensate containing methane and heavier hydrocarbons, leaves the separator  105  through line  11  and then passes through expansion valve  112 , where it is pressure reduced to about 50-120 psia to generate refrigeration. The condensate is then returned to heat exchanger  104  through line  12 , where it provides cooling for the incoming feed, and sent to fuel through line  13 , after optional compression and cooling in compressor  113  and heat exchanger  114 , respectively. A slip stream is taken through line  14  from the high pressure hydrogen-rich vapor  5  and passed through an expansion valve  115 , allowing for temperature control in heat exchanger  104 . 
     Example 1 
     Process scheme shown in  FIG. 1  was modeled using a commercial process simulator for recovering high purity hydrogen from a refinery gas stream (for example, from a fuel gas header) containing 29.66 mole % hydrogen. The results are summarized in Table 1. The fuel gas header pressure is 80 psia and the PSA system operates at 350 psia. The feed  1  is compressed to a pressure slightly higher than the PSA pressure (about 360 psia), cooled to −194° F., and then separated in a hydrocarbon-rich liquid and a hydrogen-rich gas. The liquid is expanded to a pressure close to the fuel gas header pressure (82 psia), generating enough refrigeration to obtain a hydrogen-rich gas  6  containing 66.8 mole % hydrogen. The hydrogen-rich gas is then sent to the PSA system for purification. No recompression is required for the hydrocarbon-rich gas before sending it to fuel. 
                     TABLE 1                  Example 1 Simulation Results                                             Name   1   4   6   8   9   12   13                                                     Temperature   100   −194.1   89.2   110   110   −199.3   90.5       (F.)       Pressure   80   357   355   345   20   82   80       (psia)       Flow   65   65   28.5   16   12.5   36.5   36.5       (MMSCFD)                 Composition - Mole Fractions (Dry Basis)                                             Hydrogen   0.2966   0.2966   0.6678   0.9999   0.2434   0.0078   0.0078       Methane   0.3143   0.3143   0.2092   0.0000   0.4766   0.3961   0.3961       Ethane   0.1652   0.1652   0.0018   0   0.004   0.2924   0.2924       Ethylene   0.077   0.0770   0.0027   0   0.0061   0.1349   0.1349       Propane   0.0291   0.0291   0   0   0   0.0517   0.0517       Propene   0.031   0.0310   0   0   0.0001   0.0551   0.0551       i-Butane   0.0022   0.0022   0   0   0   0.0039   0.0039       n-Butane   0.0084   0.0084   0   0   0   0.0149   0.0149       l-Butene   0.0089   0.0089   0   0   0   0.0158   0.0158       i-Pentane   0.0025   0.0025   0   0   0   0.0044   0.0044       n-Pentane   0.0007   0.0007   0   0   0   0.0012   0.0012       n-Hexane   0.0003   0.0003   0   0   0   0.0005   0.0005       Nitrogen   0.0591   0.0591   0.1139   0.0001   0.2594   0.0165   0.0165       Oxygen   0.0034   0.0034   0.0043   0   0.0099   0.0027   0.0027       CO   0.0001   0.0001   0.0002   0.0000   0.0004   0.   0       CO2   0.0011   0.0011   0.0001   0.0000   0.0001   0.0019   0.0019                    
Process 2: Hydrogen, Raw LPG and Methane-Rich Gas ( FIG. 2 )
 
     The partial condensation arrangement in  FIG. 1  can be modified to account for hydrocarbon recovery. This arrangement is especially useful when both BTU removal and hydrogen recovery are of interest. Referring to  FIG. 2 , a refinery gas  21  containing between 30-50 mole % hydrogen, at about 100° F. and 50-120 psia, is compressed in a multistage compressor  201 , cooled to about 100° F. in compressor cooler  202 , and sent through line  22  to pretreatment system  203 . The treated gas is passed by line  23  through heat exchanger  204 , where it is cooled to an intermediate temperature between −40° F. to −120° F. and partially condensed, such that most of the C2+ hydrocarbons will be in the liquid phase. The partially condensed stream is fed by line  24  to separator  205 , where it separates in a vapor containing mostly hydrogen and methane, and a condensate rich in C2+ hydrocarbons. The gas leaves the separator  205  by line  25 , and is further cooled in heat exchanger  208  to a temperature between −170° F. and −230° F., such as the stream is partially condensed and the vapor contains at least 60 mole % hydrogen. The partially condensed stream is sent by line  27  to separator  209 , where it separates in a hydrogen-rich gas and a methane-rich liquid. The hydrogen-rich gas leaves separator  209  by line  28 , is expanded, if needed, to the pressure required by the PSA system in expansion valve  211 , is warmed in heat exchangers  208  and  204  to about 90° F., and fed through line  29  to PSA system  219 , operating at about 350 psia. An optional high purity hydrogen stream containing 70-90 mole % hydrogen can be fed, together with hydrogen-containing stream  29 , to PSA system  219  through feed line  30  after compression, if required, in compressor  217  and cooling to about 100° F. in compressor cooler  218 . A high pressure, high purity hydrogen gas, containing at least 99 mole % hydrogen, leaves the PSA through product line  31 , and a low pressure tailgas (at about 20 psia) is rejected through line  32 . The tailgas is sent to fuel through line  33 , after recompression to the fuel header pressure (i.e., 50-120 psia) in compressor  220  and cooling to about 100° F. in compressor cooler  221 . The C2+ rich liquid stream leaving separator  205  by line  26  and is expanded through valve  206 , generating refrigeration, and then fed by line  34  to the top of wash column  207 . The methane-rich condensate leaving separator  209  through line  35  is passing through an expansion valve  210 , where it is pressure reduced to about 50-120 psia, generating refrigeration, and then is returned to heat exchanger  208  through line  36 , where it vaporizes providing cooling for the incoming stream. A slip stream is taken through line  37  from the high pressure hydrogen-rich vapor  28  and passed through an expansion valve  212  to the methane-rich stream  36 , allowing for temperature control in heat exchanger  208 . The vaporized stream is then fed by line  38  to the bottom of wash column  207 , where it comes into contact with the C2+ rich liquid, thus maximizing the recovery of C2+ hydrocarbons in the liquid, as well as increasing the recovery of methane from the C2+ rich liquid. A methane-rich product containing between 70-95% methane, leaves column  207  by line  39 , is warmed in heat exchanger  204  against incoming feed and, after optional compression in  215  and cooling in  216 , is sent to fuel through line  40 . The C2+ rich liquid, containing less than 10% methane, leaves column  207  by line  41 , is warmed in heat exchanger  204  against incoming feed and, after optional compression and cooling in  213  and  214 , respectively, is sent by line  42  to further processing for LPG recovery. 
     Example 2 
     The process scheme shown in  FIG. 2  was modeled, using a commercial process simulator, for recovering high purity hydrogen, methane-rich gas and C2+ hydrocarbons from a refinery gas stream and the results are summarized in Table 2. The product quality targets for this case were a maximum 10% olefins in the methane-rich gas  40 , less than 10% methane in the C2+ stream  42  and a minimum of 60 mole % hydrogen in the hydrogen-rich gas  28 . To achieve the first two targets, the feed stream has to be cooled in the first step to about −100° F. Applying a minimum pressure differential over the Joule-Thomson valve, i.e., from the pressure required by PSA to the fuel gas header pressure (as in Example 1), will not generate enough refrigeration in the system, causing the warm end heat exchanger  204  to pinch. To avoid this, either the feed needs to be compressed to a higher pressure, or the condensates  26  and  35  to be expanded across the Joule-Thomson valves to a lower pressure. However, the latter will require recompression of the methane-rich gas before sending it to fuel, thus increasing capital costs. It is generally more economical to increase the pressure on the feed side, as this will have less impact on capital cost. In this case, to eliminate the pinch on heat exchanger  204 , the feed needs to be compressed to 450 psia. The positive effect of the higher pressure differential is that a higher hydrogen purity is obtained for stream  28  (73%), which will have a positive effect on PSA recovery. 
                     TABLE 2                  Example 2 Simulation Results                                                         Name   21   26   25   35   28   29   31   33   40   42                                                                 Temperature   100   −100   −100   −220.9   −220.9   90   110   110   78.82   90       (° F.)       Pressure   80   446.5   446.5   445   445   354   345   20   80   80       (psia)       Molar Flow   65   22.45   42.55   16.92   25.63   25.61   15.77   9.834   20.66   18.74       (MMSCFD)                 Composition - Mole Fractions (Dry Basis)                                                         Hydrogen   0.2966   0.01   0.4479   0.0156   0.7333   0.7333   0.9999   0.3055   0.0244   0.0001       Methane   0.3143   0.1956   0.377   0.7247   0.1475   0.1475   0.0000   0.384   0.7173   0.0982       Ethane   0.1652   0.3942   0.0444   0.1114   0.0002   0.0002   0   0.0006   0.1082   0.4535       Ethylene   0.0770   0.1543   0.0362   0.0901   0.0006   0.0006   0   0.0017   0.094   0.1626       Propane   0.0291   0.082   0.0012   0.003   0   0   0   0   0.0021   0.0987       Propene   0.0310   0.0866   0.0017   0.0043   0   0   0   0   0.0028   0.1044       i-Butane   0.0022   0.0063   0   0.0001   0   0   0   0   0   0.0076       n-Butane   0.0084   0.0242   0.0001   0.0002   0   0   0   0   0.0001   0.0291       l-Butene   0.0089   0.0256   0.0001   0.0002   0   0   0   0   0.0001   0.0308       i-Pentane   0.0025   0.0072   0   0   0   0   0   0   0   0.0087       n-Pentane   0.0007   0.002   0   0   0   0   0   0   0   0.0024       n-Hexane   0.0003   0.0009   0   0   0   0   0   0   0   0.001       Nitrogen   0.0591   0.008   0.086   0.0434   0.1142   0.1142   0.0001   0.2974   0.0437   0.0008       Oxygen   0.0034   0.0011   0.0046   0.0055   0.004   0.004   0   0.0104   0.0055   0.0003       CO   0.0001   0   0.0001   0.0001   0.0002   0.0002   0.0000   0.0004   0.0001   0       CO2   0.0011   0.0019   0.0007   0.0016   0   0   0.0000   0   0.0017   0.0019                    
Process 3: Hydrogen, Raw LPG and SMR Feed ( FIG. 3 )
 
     The methane-rich gas obtained as vapor product from the wash column can be used either as fuel (as in embodiment in  FIG. 2 ), or as feed to a SMR.  FIG. 3  shows such an integration where a methane-rich gas is blended with natural gas and fed to a SMR. Referring to  FIG. 3 , a refinery gas  51  containing between 30-50 mole % hydrogen, at about 100° F. and 50-120 psia, is compressed in a multistage compressor  301 , cooled to about 100° F. in compressor cooler  302 , and sent through line  52  to pretreatment system  303 . The treated gas is passed by line  53  through heat exchanger  304 , where it is cooled to an intermediate temperature between −40° F. to −120° F. and partially condensed, such as most of the ethylene will be in the liquid phase. The partially condensed stream is fed by line  54  to separator  305 , where it separates in a vapor containing mostly hydrogen and methane, and a condensate rich in C2+ hydrocarbons. The gas leaves the separator  305  by line  55 , and is further cooled in heat exchanger  308  to a temperature between −170° F. and −230° F., such as the stream is partially condensed and the vapor will contain at least 60 mole % hydrogen. The partially condensed stream is sent by line  57  to separator  309 , where it separates in a methane-rich liquid and a hydrogen-rich gas. The methane-rich condensate leaving separator  309  through line  58  is passing through expansion valve  310 , where it is pressure reduced to about 50-120 psia, generating refrigeration, and then is returned to heat exchanger  308  through line  59 , where it vaporizes providing cooling for the incoming stream. The vaporized stream is then fed by line  60  to the bottom of wash column  307 . The C2+ rich liquid stream leaving separator  305  by line  56  is expanded through valve  306 , generating refrigeration, and then fed by line  61  to the top of wash column  307 . A methane-rich product containing between 70-95% methane, leaves column  307  by line  62 , is warmed in heat exchanger  304  against incoming feed and, after compression in  313  and cooling in  314 , is sent through line  63  as feed for the steam methane reformer (SMR)  315 , operating at about 450 psia. Natural gas is also added to the SMR feed through line  64 , the ratio between the natural gas and the upgraded methane-rich gas depending on the olefin and nitrogen content of the refinery gas. The gaseous mixture leaving the reformer by line  65  is then sent to the water gas shift reactor  316  and then for purification to PSA system  317 , operating at about 350 psia. The hydrogen-rich gas leaving separator  309  by line  66  is, if needed, expanded to the SMR pressure in expansion valve  311 . A slip stream is taken through line  67  from the high pressure hydrogen-rich vapor  66  and passed through an expansion valve  312  to the low pressure condensate, allowing for temperature control in heat exchanger  308 . After warming in heat exchangers  308  and  304  to about 90° F., a part of the high pressure hydrogen-rich stream is taken by line  68  to the SMR feed  64 ′, to provide the hydrogen needs for the hydrotreating step of the SMR, and the rest is expanded to the PSA pressure in valve  318  and sent by line  69  to PSA system  317 . Optionally, a high purity hydrogen stream containing 70-90 mole % hydrogen can be fed to PSA system  317  through feed line  70  after compression, if required, in compressor  319  and cooling to about 100° F. in compressor cooler  320 . A high pressure, high purity hydrogen gas, containing at least 99 mole % hydrogen, leaves the PSA through product line  71 , and a low pressure tailgas (at about 20 psia) is rejected through line  72 . Part of the PSA tailgas is sent by line  73  back to the SMR where it is used as fuel, and the rest is, after recompression in compressor  321  and cooling to about 100° F. in compressor cooler  322 , sent by line  74  to the main fuel system. The C2+ rich liquid, containing less than 10% methane, leaving column  307  by line  75 , is warmed in heat exchanger  304  against incoming feed and, after optional compression in  323  and cooling in  324  is sent by line  76  to further processing for LPG recovery. 
     Example 3 
     Process scheme shown in  FIG. 3  was modeled using a process simulator and the results are summarized in Table 3. For SMR integration, tighter constraints are applied on the ethylene and nitrogen contained in the methane-rich stream  63 . It is desired that the SMR feed should not have more than 2% olefins and 2-3% nitrogen. Generally, there is a tradeoff between the amount of methane recovered in the methane-rich stream and its olefin content. In this case, the proposed process recovered 90% of the methane in the feed, with an olefin content low enough such as it could replace as much as 20% of a SMR feed producing 100 MMSCFD hydrogen. The hydrogen produced does not take into account the additional hydrogen recovered from the hydrogen-rich stream  69 . 
     
       
         
           
               
             
               
                 TABLE 3 
               
             
            
               
                   
               
               
                 Example 3 Simulation Results 
               
            
           
           
               
               
               
               
               
               
               
               
               
               
               
            
               
                 Name 
                 51 
                 59 
                 61 
                 63 
                 64 
                 64′ 
                 69 
                 71 
                 74 
                 76 
               
               
                   
               
            
           
           
               
               
               
               
               
               
               
               
               
               
               
            
               
                 Temperature 
                 100 
                 −229.9 
                 −121.3 
                 100.0 
                 100.0 
                 99.2 
                 89.8 
                 110 
                 100 
                 90 
               
               
                 (° F.) 
               
               
                 Pressure 
                 80 
                 83 
                 81.5 
                 450 
                 80 
                 450 
                 355 
                 345 
                 80 
                 80 
               
               
                 (psia) 
               
               
                 Molar Flow 
                 22 
                 7.1 
                 8.8 
                 8.8 
                 32.0 
                 42.0 
                 8.6 
                 5.4 
                 3.2 
                 7.3 
               
               
                 (MMSCFD) 
               
            
           
           
               
            
               
                 Composition - Mole Fractions (Dry Basis) 
               
            
           
           
               
               
               
               
               
               
               
               
               
               
               
            
               
                 Hydrogen 
                 0.2966 
                 0.0156 
                 0.0102 
                 0.0315 
                 0.0000 
                 0.0297 
                 0.7487 
                 0.9999 
                 0.3228 
                 0.0001 
               
               
                 Methane 
                 0.3143 
                 0.7205 
                 0.1926 
                 0.7019 
                 0.9552 
                 0.8770 
                 0.1327 
                 0.0000 
                 0.3577 
                 0.0958 
               
               
                 Ethane 
                 0.1652 
                 0.1130 
                 0.3947 
                 0.1125 
                 0.0130 
                 0.0333 
                 0.0002 
                 0 
                 0.0005 
                 0.4531 
               
               
                 Ethylene 
                 0.0770 
                 0.0900 
                 0.1534 
                 0.0958 
                 0.0000 
                 0.0200 
                 0.0006 
                 0 
                 0.0015 
                 0.1585 
               
               
                 Propane 
                 0.0291 
                 0.0032 
                 0.0830 
                 0.0022 
                 0.0038 
                 0.0034 
                 0.0000 
                 0 
                 0.0000 
                 0.1010 
               
               
                 Propene 
                 0.0310 
                 0.0045 
                 0.0876 
                 0.0030 
                 0.0000 
                 0.0006 
                 0.0000 
                 0 
                 0.0000 
                 0.1068 
               
               
                 i-Butane 
                 0.0022 
                 0.0001 
                 0.0064 
                 0.0000 
                 0.0004 
                 0.0003 
                 0.0000 
                 0 
                 0.0000 
                 0.0078 
               
               
                 n-Butane 
                 0.0084 
                 0.0002 
                 0.0246 
                 0.0001 
                 0.0005 
                 0.0004 
                 0.0000 
                 0 
                 0.0000 
                 0.0298 
               
               
                 l-Butene 
                 0.0089 
                 0.0002 
                 0.0260 
                 0.0001 
                 0.0000 
                 0.0000 
                 0.0000 
                 0 
                 0.0000 
                 0.0316 
               
               
                 i-Pentane 
                 0.0025 
                 0.0000 
                 0.0073 
                 0.0000 
                 0.0002 
                 0.0002 
                 0.0000 
                 0 
                 0.0000 
                 0.0089 
               
               
                 n-Pentane 
                 0.0007 
                 0.0000 
                 0.0021 
                 0.0000 
                 0.0001 
                 0.0001 
                 0.0000 
                 0 
                 0.0000 
                 0.0025 
               
               
                 n-Hexane 
                 0.0003 
                 0.0000 
                 0.0009 
                 0.0000 
                 0.0001 
                 0.0001 
                 0.0000 
                 0 
                 0.0000 
                 0.0011 
               
               
                 Nitrogen 
                 0.0591 
                 0.0454 
                 0.0080 
                 0.0456 
                 0.0090 
                 0.0199 
                 0.1139 
                 0.0001 
                 0.3068 
                 0.0008 
               
               
                 Oxygen 
                 0.0034 
                 0.0056 
                 0.0011 
                 0.0055 
                 0.0000 
                 0.0013 
                 0.0038 
                 0 
                 0.0103 
                 0.0003 
               
               
                 CO 
                 0.0001 
                 0.0001 
                 0.0000 
                 0.0001 
                 0.0000 
                 0.0000 
                 0.0002 
                 0.0000 
                 0.0005 
                 0.0000 
               
               
                 CO2 
                 0.0011 
                 0.0016 
                 0.0019 
                 0.0017 
                 0.0177 
                 0.0138 
                 0.0000 
                 0.0000 
                 0.0000 
                 0.0018 
               
               
                   
               
            
           
         
       
     
     Although the invention has been described in detail with reference to certain preferred embodiments, those skilled in the art will recognize that there are other embodiments within the spirit and the scope of the claims.