a preferred configuration or arrangement of coupling an oxygen transport membrane based reforming system to a methanol production process is shown in fig1 . as seen therein , the synthesis gas supplied to the methanol synthesis and purification system 400 is a combined synthesis gas stream 390 comprising the synthesis gas product stream 176 produced by the oxygen transport membrane based reforming system 100 and a synthesis gas stream 376 produced from a hydrocarbon feed 382 and a conventional synthesis gas generation system 300 such as a steam methane reformer ( smr ); combined reformer ; or autothermal reformer ( atr ). in this arrangement , the hydrocarbon containing stream 182 is received by the oxygen transport membrane based reforming system 100 , as generally described above with reference to fig2 - 4 . the two hydrocarbon feed streams 182 , 382 may be independent streams or , as illustrated , may originate from a common hydrocarbon feed 82 . the combined synthesis gas stream 390 is synthesized into crude methanol 562 in a methanol converter reactor 405 and subsequently purified in a methanol purification system 410 into the methanol product 420 . to enhance the overall efficiency of the methanol plant , a portion of the unreacted gas stream is usually recycled to the methanol conveyer reactor 405 via a circulator or compressor 570 . in addition , purge streams 430 a , 430 b comprising unreacted hydrogen and methane slip are recycled from the methanol synthesis and purification system 400 to the oxygen transport membrane based reforming system 100 or the conventional synthesis gas generation system 300 or both . this particular coupling arrangement , schematically shown in fig1 , is most suitable for the retrofit of existing methanol production plants having a conventional synthesis gas production system , and where the oxygen transport membrane based reforming system is constructed as a retrofit to the existing methanol production plant and integrated therein . turning now to fig2 , there is shown a schematic illustration of an embodiment of an oxygen transport membrane based reforming system 100 in accordance with the present invention . as seen therein , an oxygen containing stream 110 , such as air , is introduced to the system by means of a blower or fan 114 into a heat exchanger 113 for purposes of preheating the oxygen containing stream 110 . heat exchanger 113 is preferably a high efficiency , cyclic and continuously rotating regenerator disposed in operative association with the oxygen containing stream 110 and the heated retentate stream 124 . the heated and oxygen depleted retentate stream 124 can optionally be introduced into a duct burner region containing duct burner 126 and used to support combustion of a supplemental fuel stream 128 to produce supplemental heat introduced into the continuously rotating regenerator 113 to preheat the oxygen containing stream 110 . alternatively , the duct burner may also be disposed directly in the duct leaving heat exchanger 113 to pre - heat the oxygen containing stream 110 . exhaust stream 132 from heat exchanger 113 is discharged . the heated oxygen containing stream 115 is then directed via the intake duct to the oxygen transport membrane elements 120 incorporated into the oxygen transport membrane reactor 101 . each of the oxygen transport membrane elements 120 are preferably configured as a multilayered ceramic tube capable of conducting oxygen ions at an elevated operational temperature , wherein the retentate side of the oxygen transport membrane elements 120 is the exterior surface of the ceramic tubes exposed to the oxygen containing stream and the permeate side is the interior surface of the ceramic tubes . although only six oxygen transport membrane elements 120 are illustrated in close proximity to three catalytic reforming tubes 140 , as would occur to those skilled in the art , there could be many of such oxygen transport membrane elements and many catalytic reforming tubes in each oxygen transport membrane assembly . likewise , there would be multiple oxygen transport membrane assemblies used in an industrial application of the oxygen transport membrane reactor 101 . a hydrogen containing stream is also introduced into the permeate side of the oxygen transport membrane elements 120 and is oxidized though reaction with the permeated oxygen to produce a reaction product stream 198 and heat . as described in more detail below , the hydrogen containing stream is preferably a recycled portion of the produced synthesis gas 163 . as a result of the separation of the oxygen and the reaction occurring at the permeate side of oxygen transport membrane elements 120 , a heated and oxygen depleted retentate stream 124 is also formed . the reaction of the hydrogen containing stream or recycled synthesis gas stream 163 at the permeate side of the oxygen transport membrane element 120 produces heat . radiation of this heat together with the convective heat transfer provided by heated retentate stream 124 heats the catalytic reactor tubes 140 to supply the endothermic heating requirements of the steam methane reforming occurring in catalytic reactor tubes 140 . as the heated retentate stream 124 exits the oxygen transport membrane based reforming system 100 , it also heats a reformer feed stream 138 to a temperature between about 475 ° c . and 650 ° c . via indirect heat transfer using one or more coils 191 disposed in the retentate duct such that the oxygen depleted retentate stream 124 heats the feed streams passing through the coils 191 . the hydrocarbon containing feed stream 182 to be reformed is preferably natural gas . depending on the supply pressure , the natural gas feed is compressed or let down to the desired pressure via a compressor or valve arrangement ( not shown ) and then preheated in heat exchanger 150 that serves as a fuel preheater . also , since the natural gas typically contains unacceptably high level of sulfur species , the natural gas feed stream 182 undergoes a sulfur removal process such as hydro - treating , via device 190 , to reduce the sulfur species to h 2 s , which is subsequently removed in a guard bed using material like zno and / or cuo . to facilitate the desulfurization , a small amount of hydrogen or hydrogen - containing gas is added ( not shown ) to the natural gas feed stream 182 . the hydro - treating step also saturates any alkenes present in the hydrocarbon containing feed stream . further , since natural gas generally contains higher hydrocarbons that will break down at high temperatures to form unwanted carbon deposits that adversely impact the reforming process , the natural gas feed stream 182 is pre - reformed in an adiabatic pre - reformer 192 , which converts higher hydrocarbons to methane , hydrogen , carbon monoxide , and carbon dioxide . pre - reformers are typically catalyst - based systems . in the case of heated pre - reforming , it is contemplated that the catalyst based pre - reformer be thermally coupled with the oxygen transport membrane based reforming system . in the illustrated embodiment , the above - described heated reaction product stream 198 is combined with the pre - reformed reformer feed stream 195 to produce a combined feed stream 197 that contains steam and hydrocarbons . this combined feed stream is introduced into the catalytic reactor tubes 140 where the combined feed stream 197 is subjected to steam methane reforming to produce a synthesis gas stream 142 . the temperature of the combined feed stream 197 is between about 550 ° c . and 800 ° c ., and more preferably between about 600 ° c . and 800 ° c . steam 180 may also be added to the combined feed stream 197 , or the preheated pre - reformed reformer feed stream 195 , as required , to adjust the temperature of stream 197 as well as the steam to carbon ratio of stream 197 to between about 1 . 5 and 3 . 0 , and more preferably to steam to carbon ratio between about 2 . 0 and 2 . 8 . the steam is preferably superheated steam 180 between about 300 psia and 1200 psia and between about 300 ° c . and 600 ° c . and heated by means of indirect heat exchange with the heated retentate stream 124 using steam coils 179 disposed in the retentate duct . the superheated steam 180 is preferably added to the hydrocarbon containing feed stream 182 upstream of the pre - reformer 192 to adjust the steam to carbon ratio and final temperature of the combined feed stream 197 . also , as described in more detail with reference to fig7 , to reduce the methane slip and optimize the economic performance of the oxygen transport membrane based reforming system in a methanol production process , the oxygen transport membrane reactor 101 should preferably be maintained at an exit pressure of less than or equal to about 250 psia . the synthesis gas stream 142 produced by the oxygen transport membrane based reforming system 100 generally contains hydrogen , carbon monoxide , steam , carbon dioxide and other constituents such as unconverted methane . heat exchange section 104 is designed to cool the produced synthesis gas stream 142 and recycle a portion of the synthesis gas stream 162 a to form all or a part of the hydrogen containing stream 163 . in this illustrated embodiment , the synthesis gas stream 142 is preferably cooled before recycling such stream using a synthesis gas recycle compressor 177 or other blower means . the heat exchange section 104 is also designed such that in cooling the synthesis gas stream 142 , various feed streams are preheated and process steam is also generated . the initial cooling of synthesis gas stream 142 is accomplished with steam generation in a process gas boiler ( pg boiler ) 149 coupled to steam drum 157 and designed to reduce the temperature of the cooled synthesis gas 144 to about 400 ° c . or less . as illustrated in fig2 , the initially cooled synthesis gas stream 144 is successively further cooled in a heat exchange network that includes hydrocarbon feed preheater 150 , economizer 156 , feed water heaters 158 a and 158 b , synthesis gas cooler 161 and water cooled heat exchanger 164 . the initially cooled synthesis gas stream 144 is directed to the natural gas feed preheater 150 to heat the natural gas feed stream 182 and then is directed to the economizer 156 to heat boiler feed water 188 . the boiler feed water stream 188 is preferably pumped using a feed water pump ( not shown ), heated in economizer 156 and sent to steam drum 157 . the cooled synthesis gas stream 146 is then divided into a first portion 160 and a second or recycled portion 162 a . first portion 160 is further cooled in a series of steps including a feed water heater 158 a , used to heat feed water stream 159 , followed by a synthesis gas cooler 161 and a subsequent water cooled heat exchanger 164 cooled via a separate cooling water stream 166 . the heated feed water 159 is directed to a de - aerator ( not shown ) that provides boiler feed water 188 . the resulting fully cooled synthesis gas stream 148 is then introduced into a knock - out drum 168 from which a condensate stream 170 is drained to produce a fully cooled synthesis gas stream 172 . the fully cooled synthesis gas stream 172 is optionally compressed in a synthesis gas compressor 174 to produce a synthesis gas product 176 . the second or recycled portion 162 a of the initially cooled synthesis gas stream 144 is directed to a second feed water heater 158 b , used to heat feed water stream 159 , and this cooled recycle synthesis gas stream 163 is recirculated back to the permeate side of the oxygen transport membrane element 120 by means of a recycle compressor 177 . also note that any superheated steam not added or used in natural gas feed 182 or recycle synthesis gas stream 163 is exported steam 181 that may be used for power generation . when customized for a methanol production process , the oxygen transport membrane produced synthesis gas should have a module of between about 1 . 5 and 2 . 0 . in addition , such produced synthesis gas stream ideally has a methane slip of less than about 4 . 5 percent by volume where the exit pressure of the oxygen transport membrane based reforming system is 250 psia or less . turning now to fig3 , there is shown a schematic illustration of an alternate embodiment of an oxygen transport membrane based reforming system . in many regards , this embodiment is similar to the embodiment of fig2 and , for sake of brevity , the description of the common aspects of the two embodiments will not be repeated here , rather , the following discussion shall focus on differences between fig2 and fig3 . the primary difference between the embodiments in fig2 and fig3 is the use of a hot synthesis gas recycle 162 b in fig3 embodiment in lieu of the cold gas recycle 162 a in the embodiment of fig2 . as a result , the heat exchange section 104 in fig3 is designed to only cool the produced synthesis gas stream 142 and need not recycle a portion of the cold synthesis gas stream 162 a . by using the hot synthesis gas recycle 162 a , there is no need to use the synthesis gas recycle compressor 177 or the second feed water heater 158 b , used to heat feed water stream 159 potentially further reducing the capital cost of the oxygen transport membrane based reformer reactor and system . the hot synthesis gas recycle involves recycling a portion of the heated synthesis gas stream 142 exiting the catalyst reforming tubes 140 or reactor and directing the hot recycled synthesis gas 162 b to the permeate side of the oxygen transport membrane elements 120 to react the portion of heated synthesis gas stream 162 b with the oxygen permeate to generate the heated reaction product stream and radiant heat . the temperature of the hot synthesis recycled gas is preferably above 800 ° c . so as to avoid problems associated with metal dusting corrosion . the hot synthesis gas stream 162 b is driven or pulled to the permeate side of the oxygen transport membrane elements 120 by means of an ejector , eductor or venturi based device 199 operatively coupled to the permeate side of the oxygen transport membrane elements 120 . by suctioning the streams at the permeate side of the oxygen transport membrane elements 120 into the ejector , eductor or venturi based device 199 with a motive fluid comprising the pre - reformed reformer feed stream 195 , the reaction product stream 198 mixes with the pre - reformed reformer feed stream 195 to produce the combined feed stream 197 , preferably having a steam to carbon ratio between about 1 . 5 and 3 . 0 and a temperature between about 550 ° c . and 800 ° c . essentially , device 199 moves the lower pressure hot synthesis gas recycle stream 162 b to the higher pressure combined feed stream 199 . turning now to fig4 , there is shown a schematic illustration of yet another alternate embodiment of an oxygen transport membrane based reforming system . in many regards , this illustrated embodiment is similar to the embodiments shown in fig2 and fig3 . thus , for sake of brevity ; the description of the common aspects of the embodiments will not be repeated here . rather , the following discussion shall focus on the differences present in the embodiment of fig4 . the primary difference between the embodiments in fig3 and fig4 is the addition of downstream processing of the cooled synthesis gas in the embodiment of fig4 . as seen therein , a portion of the initially cooled synthesis gas 350 is diverted to a synthesis gas conditioning system . this diverted portion of the cooled synthesis gas stream 350 is roughly between about 5 % and 25 % of the synthesis gas stream 142 . the diverted portion of the cooled synthesis gas stream 350 is subjected to a water gas shift reaction in reactor 352 and subsequently cooled in heat exchanger 354 using cooling water or boiler feed water and then compressed in a synthesis gas compressor 360 . alternatively , heat exchanger 354 could instead be used to preheat a portion of the natural gas feed 182 . the compressed gas feed stream 362 is directed to a hydrogen pressure swing adsorption ( psa ) unit 370 which takes the compressed feed stream 362 and produces a higher purity hydrogen stream 372 at or near the feed pressure while the carbon oxides , methane and other impurities are rejected at lower pressure tail gas stream 374 . stream 374 may be recycled to use with the duct burners 126 while the higher pressure and higher purity hydrogen stream 372 is recombined with the non - diverted portion of the synthesis gas stream 172 , preferably at some point mid - stage point within the synthesis gas compressor 174 . by re - combining the higher purity hydrogen stream 372 with the non - diverted portion of the synthesis gas stream 172 , one can adjust the module of the final synthesis gas product to about 2 . 0 to 2 . 2 , the preferable range for methanol production or other synthesis gas characteristics such as hydrogen to carbon monoxide ratio , etc . it may also be possible to combine a portion or all of the purge 430 a ( see fig1 ) from the methanol loop with stream 362 prior to feeding the psa unit 370 . the embodiment shown in fig4 is particularly applicable to use in a methanol production facility , where the oxygen transport membrane based reforming system is the only or primary source of synthesis gas supply to the methanol synthesis process . fig5 provides a schematic illustration of an alternate embodiment of an oxygen transport membrane based reforming system 200 in accordance with the present invention . as seen therein , an oxygen containing stream 210 , such as air , is introduced to the system by means of a forced draft ( fd ) fan 214 into a heat exchanger 213 for purposes of preheating the oxygen containing feed stream 210 . heat exchanger 213 is preferably a high efficiency , cyclic and continuously rotating ceramic regenerator disposed in operative association with the oxygen containing feed stream 210 and the heated retentate stream 224 . the ceramic regenerator 213 which heats the incoming air feed stream 210 to a temperature in the range of about 850 ° c . to 1000 ° c . the oxygen depleted air leaves the oxygen transport membrane reforming tubes as a heated retentate stream 224 at the same or slightly higher temperature than the heated air feed stream 215 . any temperature increase , typically & lt ; 30 ° c ., is attributable to the portion of energy generated by the oxidizing reaction of hydrogen and carbon monoxide in the oxygen transport membrane tubes and transferred by convection to the air stream . the heated , oxygen depleted retentate stream 224 is first used to heat the mixed feed stream to a temperature between 450 ° c . and 650 ° c ., and more preferably to a temperature between 525 ° c . and 600 ° c ., and used to further heat the superheated steam . the temperature of this oxygen depleted retentate stream 224 preferably needs to be then increased back to a temperature between about 1050 ° c . and 1200 ° c . prior to being directed to the ceramic heat exchanger or regenerator 213 . this increase in temperature of the retentate stream 224 is preferably accomplished by use of a duct burner 226 , which facilitates combustion of a supplemental fuel stream 228 using some of the residual oxygen in the retentate stream 224 . it is conceivable that the mixed feed heater and steam superheater could alternatively be located in a separate fired heater ( not shown ). in that case , the fuel requirements of the duct burner 226 will be substantially less . in the ceramic heat exchanger or regenerator 213 , the heated , oxygen depleted retentate stream provides the energy to raise the temperature of the incoming feed air stream from ambient temperature to a temperature between about 900 ° c . and 1000 ° c . the resulting cold retentate stream exiting the ceramic heat exchanger , typically containing less than 5 % oxygen , leaves the oxygen transport membrane based reforming system 200 as exhaust gas 232 at a temperature of around 150 ° c . the oxygen transport membrane based reforming system 200 comprises two sets of reforming tubes , including primary reforming tubes 240 where the primary reforming occurs and oxygen transport membrane tubes 220 where the secondary reforming occurs . although only six secondary reforming oxygen transport membrane tubes 220 are illustrated in close proximity to three primary reforming tubes 240 , as would occur to those skilled in the art , there could be many of such secondary reforming oxygen transport membrane tubes and many primary reforming tubes in each oxygen transport membrane assembly . likewise , there would be multiple oxygen transport membrane assemblies used in an industrial application of the oxygen transport membrane based reforming system 200 . the heated oxygen containing stream 215 is directed via the intake duct 216 to a plurality of secondary reforming oxygen transport membrane tubes 220 incorporated into the oxygen transport membrane reactor 201 . the secondary reforming oxygen transport membrane tubes 220 are preferably configured as multilayered ceramic tubes capable of conducting oxygen ions at an elevated operational temperature , wherein the retentate side of the secondary reforming oxygen transport membrane tubes 220 is the exterior surface of the ceramic tubes exposed to the heated oxygen containing stream 215 and the permeate side is the interior surface of the ceramic tubes . within each of the secondary reforming oxygen transport membrane tubes 220 are one or more catalysts that facilitate secondary reforming . the hydrocarbon containing feed stream 282 , preferably natural gas , to be reformed is typically preheated to around 370 ° c ., as described in more detail below . as natural gas typically contains unacceptably high level of sulfur species , some hydrogen gas 725 is added prior to desulfurization . the mixture 283 of the hydrogen gas 725 and hydrocarbon containing feed stream 282 is heated in heat exchanger 250 that serves as a pre - heater and then undergoes a sulfur removal process via device 290 such as hydro - treating to reduce the sulfur species to h 2 s , which is subsequently removed in a guard bed using material like zno and / or cuo . the hydro - treating step also saturates any alkenes present in the hydrocarbon containing feed stream . although not shown , the heated feed stream 282 may also undergo pre - reforming step an adiabatic pre - reformer , which converts higher hydrocarbons to methane , hydrogen , carbon monoxide , and carbon dioxide or a heated pre - reforming step . in the case of heated pre - reforming , it is contemplated that the catalyst based pre - reformer be thermally coupled with the oxygen transport membrane based reforming system . superheated steam 280 is added to the pre - treated natural gas and hydrogen feed stream , as required , to produce a mixed feed stream 238 with a steam to carbon ratio between about 1 . 0 and 2 . 5 , and more preferably between about 1 . 2 and 2 . 2 . the superheated steam 280 is preferably between about 300 psia and 1200 psia and between about 300 ° c . and 600 ° c . and heated by means of indirect heat exchange with the heated retentate stream 224 using steam coils 279 disposed in the retentate duct 225 . any superheated steam 280 not added or used in the natural gas and hydrogen feed 282 is exported steam 281 used for power generation . the mixed feed stream 238 is heated , by means of indirect heat exchange with the heated retentate stream using coils 289 disposed in the retentate duct 225 , to preferably between about 475 ° c . and 650 ° c ., and more preferably between about 550 ° c . and 600 ° c . the heated mixed feed stream 238 is then sent to the reforming tubes 240 , which contain conventional reforming catalyst . the temperature of the partially reformed hydrogen - rich synthesis gas 298 leaving the reforming tubes 240 is typically designed to be between 650 ° c . and 900 ° c . this synthesis gas is then fed to the oxygen transport membrane tubes 220 filled with one or more catalysts that would facilitate partial oxidation and reforming . oxygen from the heated intake air permeates through the oxygen transport membrane tubes 220 and facilitates reaction of a portion of the hydrogen and carbon monoxide , and possibly some methane . a portion of the energy or heat generated by this reaction is used for in - situ secondary reforming of the residual methane in the partially reformed synthesis gas 298 . the rest of the energy or heat is transferred by radiation to the reforming tubes 240 to drive the primary reforming reactions and by convection to the oxygen - depleted air stream . the synthesis gas 242 leaving the oxygen transport membrane tubes 220 , which essentially function as a secondary reformer , is at a temperature between about 900 ° c . and 1050 ° c . the endothermic heating requirements of the reforming process occurring in the primary reforming tubes 240 is supplied through radiation of some of the heat from the secondary reforming oxygen transport membrane tubes 220 together with the convective heat transfer provided by heated retentate stream 224 . in addition , as the heated , oxygen depleted retentate stream 224 exits the oxygen transport membrane based reforming system 200 , it also heats the mixed feed stream 238 to a temperature between about 475 ° c . and 650 ° c . via indirect heat transfer using one or more coils 289 disposed in the retentate stream duct 225 . the synthesis gas stream 242 produced by the oxygen transport membrane based reforming system 200 generally contains hydrogen , carbon monoxide , unconverted methane , steam and carbon dioxide other constituents . a significant portion of the sensible heat from the synthesis gas stream 242 can be recovered using a heat exchange section or recovery train 204 . heat exchange section 204 is designed to cool the produced synthesis gas stream 242 exiting the oxygen transport membrane based reforming system 200 . in this illustrated embodiment , the heat exchange section 204 is also designed such that in cooling the synthesis gas stream 242 , process steam is generated , hydrocarbon feed stream is preheated , and boiler feed water and feedwater are heated . the hot synthesis gas 242 is directly cooled to about 400 ° c . or less in a process gas ( pg ) boiler 249 . the initially cooled synthesis gas stream 244 is then used to preheat the mixture of natural gas and hydrogen feed stream 282 in a pre - heater 250 and subsequently to pre - heat boiler feed water 288 in the economizer 256 and to heat the feed water stream 259 . in the illustrated embodiment , the boiler feed water stream 288 is preferably pumped using a feed water pump ( not shown ), heated in economizer 256 and sent to steam drum 257 while the heated feed water 259 is sent to a de - aerator ( not shown ) that provides boiler feed water 288 . synthesis gas leaving the feedwater heater 258 is preferably around 160 ° c . it is cooled down to 38 ° c . using a fin - fan cooler 261 and a synthesis gas cooler 264 fed by cooling water 266 . the cooled synthesis gas 248 then enters a knock - out drum 268 where water is removed from the bottoms as process condensate stream 270 which , although not shown , is recycled for use as feedwater , and the cooled synthesis gas 272 is recovered overhead . the cooled synthesis gas stream 272 is optionally compressed in a synthesis gas compressor 274 to produce a synthesis gas product 276 . depending on the operating pressure of the oxygen transport membrane based reforming system , pressure of the recovered synthesis gas is preferably in the range of about 150 and 550 psia and more preferably in the range of 175 and 400 psia . one of the likely disadvantages of the prior art oxygen transport membrane reforming reactors and systems is the potential for corrosion , and in particular metal dusting corrosion . metal dusting is a severe form of corrosion that occurs when surfaces of certain metal and metal alloy components and piping are exposed to severe gas environments with a high carbon activity or content . the metal dusting corrosion is manifested by a disintegration of bulk metals , such as iron , nickel and cobalt to metal powders . the typical metal dusting process results from a series of sequential steps , including ( i ) rapid uptake of carbon into the metallic phase leading to saturation of the alloy matrix with carbon ; ( ii ) formation of metastable carbides ; and ( iii ) decomposition of these carbides into a loose film of carbon and metallic particles , which acts as catalyst for further carbon deposition . the exact mechanism may vary depending on the type of metal being used . the temperatures normally associated with metal dusting are about 400 ° c . to 800 ° c . at temperatures generally below 400 ° c . the rate of reaction to form the metastable carbide species is too low to be significant , while at temperatures above 800 ° c . the carbon formation is minimal . to avoid the metal dusting corrosion in the oxygen based transport membrane based reforming system and associated components and piping , the system should be designed to avoid contact of any synthesis gas with metal surfaces having temperatures between about 400 ° c . to 800 ° c . while it is possible to provide corrosion resistant coatings on all high temperature metal surfaces exposed to synthesis gas , such solution would be cost prohibitive . the alternative corrosion prevention technique employed in the present embodiments is to manage the temperatures of the synthesis gas so as to avoid contacting bare metal surfaces with synthesis gas in the deleterious temperature range where metal dusting corrosion occurs . the embodiments illustrated in fig3 and 4 achieve this temperature control of the metal surfaces exposed to the synthesis gas by preferably recycling the synthesis gas in a high temperature state , generally above 800 ° c . where the non - recycled portion of the synthesis gas is cooled in the pg boiler until the synthesis gas is cooled below 400 ° c . the metal surfaces of the synthesis gas piping exiting the oxygen transport membrane based reforming system are either maintained at a temperature above 800 ° c . or are refractory surfaces . the metal surfaces in the pg boiler and other elements in the heat recovery system that are exposed to the synthesis gas are generally maintained at temperatures below about 400 ° c . in the embodiment of fig2 , the cooled synthesis gas is recycled back to the oxygen based transport membrane based reactor and the cooled recycled synthesis gas stream is preferably maintained at temperatures below about 400 ° c . the addition of superheated steam to the recycled synthesis gas stream should be avoided unless the addition of superheated steam maintains the temperatures below about 400 ° c . turning now to fig6 , region 310 on the feed temperature versus feed steam to carbon ratio graph generally corresponds to a combined feed stream at a temperature near or below about 550 ° c . and having a steam to carbon ratio of between 1 . 5 and 2 . 4 . because the present combined feed stream preferably contains carbon monoxide produced in the pre - reformer , a combined feed stream having characteristics of region 310 , may undergo unwanted boudouard reactions thereby depositing excessive amounts of soot on the oxygen transport membrane based reforming system as well as associated components and piping . the boudouard reaction is a redox reaction of a mixture of carbon monoxide and carbon dioxide at a given temperature and involves the disproportionation of carbon monoxide into carbon dioxide and carbon ( i . e . soot ). on the other hand , region 330 generally corresponds to a region of high temperature and low steam to carbon ratios where the combined feed stream is subject to carbon lay down in the reformer tubes , most likely at the inlet . region 330 is generally depicted as the region to the right of the curve on the feed temperature versus steam to carbon ratio defined by two points , namely a feed temperature near 550 ° c . with a steam to carbon ratio of 1 . 5 and a feed temperature of 700 ° c . with a steam to carbon ratio of 2 . 4 . it has been found that conditioning the combined feed stream to a particular temperature range and steam to carbon ratio which avoids regions 310 and 330 translates into an optimum operating regime with noticeably less reliability problems in the oxygen transport membrane based reformer system ( otm tubes as well as reformer tubes ) due to carbon formation . this window of preferred operating characteristics for the combined feed stream is depicted generally as region 320 in fig6 . as indicated above , the presently disclosed oxygen transport membrane based reforming systems are preferably tailored for use in methanol production plants , either as a retrofit or modification of existing methanol plants or in the design and construction of new methanol plants . as illustrated in the graph of fig7 , when the embodiments of the oxygen transport membrane based reforming system described with reference to fig2 - 4 are tailored for a methanol production process , the exit pressure of the oxygen transport membrane based reforming system is preferably maintained at a pressure of about 250 psia or less . it has been found , through modeling and estimation , that such pressure range optimizes the overall performance and economics of the oxygen transport membrane based reforming system for methanol production applications . specifically , when the exit pressure of the oxygen transport membrane based reforming system is maintained at a pressure of about 250 psia or less , the estimated utility cost , expressed in dollars per ton of methanol produced , is optimized at or below about $ 117 , in part due to the low methane slip of less than about 4 . 5 % realized at such exit pressures . a similar curve is observed when one plots the estimated utility cost expressed in dollars per ton of methanol produced versus methane slip . specifically , when the methane slip is reduced to levels of 4 . 5 % or less , the utility costs are about $ 117 per ton of methanol produced . on the other hand , when the embodiments of the oxygen transport membrane based primary and secondary reforming system and reactor described with reference to fig5 are tailored for a methanol production process , the exit pressure of the secondary reformers in the oxygen transport membrane based reforming system are preferably maintained at a pressure of between about 100 psia to about 500 psia . maintaining the exit pressure of the secondary reformers at the higher pressures within this range ( e . g . 400 to 500 psia ) optimizes the overall performance and economics of the oxygen transport membrane based primary and secondary reforming reactor and system for methanol production applications . specifically , when the exit pressure of the of the secondary reformer in oxygen transport membrane based reforming system of fig5 is maintained at about 200 psia , the estimated utility cost , expressed in dollars per ton of methanol produced , is about $ 113 and when the exit pressure is maintained at about 400 psia , the estimated utility cost is reduced to about $ 110 per ton of methanol produced . using the above - described process that employs both an oxygen transport membrane ( otm ) based reforming system similar to the embodiments shown in fig2 , fig3 and fig5 together with a base smr synthesis gas production system , three example cases were modeled . all three modeled cases are compared to the base modeled case involving a traditional smr reactor based methanol plant that is designed to produce 2000 tons / day of methanol . in the base case scenario , the methanol reactor operates at 100 % capacity to produce the 2000 tons / day of methanol product and the smr reactor also is modeled to operate at 100 % capacity to supply the synthesis gas needed for the methanol reactor . the required natural gas flow rate to the smr reactor necessary to operate as suggested above is roughly 61 million scfd . the resulting measure of standard cubic feet ( scf ) of natural gas ( ng ) required per ton of methanol produced is roughly 30530 ( scf ng / ton meoh ). the comparative results of the modeled cases as shown in table 1 . in the base modeled smr only case , the 61 million standard cubic feet per day ( scfd ) of natural gas is used for both the reformer as well as the burner . at an energy conversion for natural gas of about 1000 btus per scf and an estimated cost of about $ 4 per million btus , the total cost of the natural gas is roughly $ 244 , 240 per day . in addition , the power requirements for the base smr case include supplying power to the air blower , boiler feed water pump , synthesis gas / recycle compressor , and methanol product pump which is offset by the power generation associated with the steam expander , liquid expander , and gas expander . as seen in table 1 , the net power consumption for the modeled base case is roughly 7102 kw which , at $ 45 . 8 per mwh translates to a power cost of approximately $ 7 , 807 per day . considering only the utility costs ( i . e . natural gas , electricity and make - up water ), the traditional smr reactor based methanol plant requires $ 258 , 648 per day to operate or roughly $ 129 per ton of methanol product produced . the steam to carbon ration in the natural gas feed is roughly 3 resulting in a module of the synthesis gas produced by the base smr modeled case of about 2 . 9 and a hydrogen to carbon monoxide ratio of about 5 . in case 1 , the base case smr system is arranged in parallel with an oxygen transport membrane ( otm ) based reformer system , similar to the embodiment described above with reference to fig2 , to produce the synthesis gas required to increase the methanol production by about 20 % or to a total production of about 2400 tons per day of methanol . in the case 1 modeled scenario , the 71 . 52 million cubic feet per hour of natural gas is generally split between the smr system ( 61 . 06 mmscfd ), and the otm based reforming system ( 10 . 46 mmscfd ), with a portion of the natural gas being used a fuel to fire the burners . the total methanol produced is estimated to be about 2401 tons per day . the steam to carbon ratio in the natural gas feed to the otm based reforming system is roughly 1 . 5 resulting in a module of the synthesis gas produced by the combined smr and otm arrangement of about 2 . 7 and a hydrogen to carbon monoxide ratio of about 3 . 4 . at an energy conversion of about 1000 btus per scf of natural gas and an estimated cost of about $ 4 per million btus , the estimated cost of the natural gas for the arrangement in case 1 is roughly $ 286 , 078 per day to produce the 2401 tons of methanol . in addition , the electricity or power requirements for case 1 is roughly 10694 kw or about $ 11 , 755 per day at $ 45 . 8 per mwh and includes supplying power to the air blower , boiler feed water pump , synthesis gas / recycle compressor , and methanol product pump and offset by any power generation associated with the steam expander , liquid expander , and gas expander . the total utility costs ( i . e . natural gas , electricity and make - up water ) for the case 1 arrangement is roughly $ 304 , 602 per day to operate or roughly $ 115 per ton of methanol product produced which represents a 10 . 9 % lower cost than the base smr case . in case 2 , the base case smr system is arranged in parallel with an oxygen transport membrane ( otm ) based reformer system , similar to the embodiment described above with reference to fig3 , to produce the synthesis gas required to increase methanol production by about 20 % or to a total production of about 2400 tons per day of methanol . in the case 2 modeled scenario , the 71 . 58 million cubic feet per hour of natural gas is generally split between the smr system ( 61 . 06 mmscfd ), and the otm based reforming system ( 10 . 52 mmscfd ), with a portion of the natural gas directed to each system being used a fuel to fire the associated system burners . the total methanol produced is estimated to be about 2400 tons per day . the steam to carbon ratio in the natural gas feed to the otm based reforming system is roughly 1 . 5 resulting in a synthesis gas module of about 2 . 7 with a hydrogen to carbon monoxide ratio of about 3 . 5 . at an energy conversion of about 1000 btus per scf of natural gas and an estimated cost of about $ 4 per million btus , the estimated cost of the natural gas for the arrangement in case 2 is roughly $ 286 , 327 per day to produce the 2400 tons of methanol . in addition , the electricity or power requirements for case 2 is roughly 9560 kw or about $ 10 , 508 per day at $ 45 . 8 per mwh and includes supplying power to the air blower , boiler feed water pump , synthesis gas compressor , and methanol product pump and offset by any power generation associated with the steam expander , liquid expander , and gas expander . the total utility costs ( i . e . natural gas , electricity and make - up water ) for the case 2 arrangement is roughly $ 303 , 609 per day to operate or roughly $ 112 . 40 per ton of methanol product produced which represents a 12 . 9 % lower cost than the base smr case . in case 3 , the base case smr system is arranged in parallel with an oxygen transport membrane ( otm ) based reformer system , similar to the system described above with reference to fig5 , to produce the synthesis gas required to increase the methanol production by 20 % or to a total production of about 2400 tons per day of methanol . in the case 2 modeled scenario , the 72 . 02 million cubic feet per hour of natural gas is generally split between the smr system ( 61 . 06 mmscfd ), and the otm based reforming system ( 10 . 96 mmscfd ), with a portion of the natural gas directed to each system being used a fuel to fire the associated system burners . the total methanol produced is estimated to be about 2400 tons per day . the steam to carbon ratio in the natural gas feed to the otm based reforming system is roughly 2 . 0 resulting in a module of the synthesis gas produced by the combined smr and otm arrangement of about 2 . 7 with a hydrogen to carbon monoxide ratio of about 3 . 2 . at an energy conversion of about 1000 btus per scf of natural gas and an estimated cost of about $ 4 per million btus , the estimated cost of the natural gas for the arrangement in case 3 is roughly $ 288 , 076 per day to produce the 2399 tons of methanol . in addition , the total net electricity or power requirements for case 3 is less than the base case at roughly 7042 kw or about $ 7 , 759 per day at $ 45 . 8 per mwh and includes supplying power to the air blower , boiler feed water pump , synthesis gas compressor , and methanol product pump and offset by any power generation associated with the steam expander , liquid expander , and gas expander . the total utility costs ( i . e . natural gas , electricity and make - up water ) for the case 3 arrangement is roughly $ 302 , 703 per day to operate or roughly $ 110 . 40 per ton of methanol product produced which represents a 14 . 4 % lower cost than the base smr case . while the inventions herein disclosed have been described by means of specific embodiments and processes associated therewith , numerous modifications and variations can be made thereto by those skilled in the art without departing from the scope of the invention as set forth in the appended claims or sacrificing all of its features and advantages .