Decarbonized hydrogen and electricity from natural gas

Transcription

1 International Journal of Hydrogen Energy 30 (2005) Decarbonized hydrogen and electricity from natural gas Stefano Consonni,1, Federico Viganò 1 1 Princeton Environmental Institute, Princeton University, Princeton, NJ , USA Available online 12 September 2004 Abstract This paper discusses configuration, attainable performances and thermodynamic features of stand-alone plants for the coproduction of de-carbonized hydrogen and electricity from natural gas (NG) based on commercially available technology. We focus on the two basic technologies currently used in large industrial applications: fired tubular reformer (FTR) and auto-thermal reformer (ATR). In both cases we assume that NG is pre-heated and humidified in a saturator providing water for the reforming reaction; this reduces the amount of steam to be bled from the power cycle and increases electricity production. Outputs flows are made available at conditions suitable for transport via pipeline: 60 bar for pure hydrogen, 150 bar for pure CO 2. To reduce hydrogen compression power requirements reforming is carried out at relatively high pressures: 25 bar for FTR, 70 bar for ATR. Reformed gas is cooled and then passed through two water gas shift reactors to optimize heat recovery and maximize the conversion to hydrogen. In plants with CO 2 capture, shifted gas goes through an amine-based chemical absorption system that removes most of the CO 2. Pure hydrogen is obtained by pressure swing absorption (PSA), leaving a purge gas utilized to fire the reformer (in FTR) and to boost electricity production. For the power cycle we consider conventional steam cycles (SC) and combined cycles (CC). The scale of plants based on a CC is determined by the gas turbine. To maintain NG input within the same range (around 1200 MW), we considered a General Electric 7FA for ATR, a 6FA for FTR. The scale of plants with SC is set by assuming the same NG input of the corresponding CC plant. Heat and mass balances are evaluated by a model accounting for the constraints posed by commercial technology, as well as the effects of scale. Results show that, from a performance standpoint, the technologies of choice for the production of de-carbonized hydrogen from NG are FTR with SC or ATR with CC. When operated at high steam-to-carbon ratios, the latter reach CO 2 emissions chargeable to hydrogen of kg of CO 2 per GJ LHV less than 20% of NG with an equivalent efficiency of hydrogen production in excess of 77% International Association for Hydrogen Energy. Published by Elsevier Ltd. All rights reserved. Keywords: Hydrogen production; Combined cycles; CO 2 capture; Efficiency; Natural gas 1. Background and scope Concern about rising concentrations of greenhouse gases in the atmosphere has spurred research into various ways Corresponding author. Fax: address: stefano.consonni@polimi.it (S. Consonni). 1 Current address: Dipartimento di Energetica, Politecnico di Milano, P.za L. Da Vinci 32, Milano 20133, Italy. of capturing and storing CO 2 before it is emitted. Due to strong scale effects on performance and cost, this technology is applied most naturally to large, point source emitters of CO 2 such as fossil fuel-based plants generating electricity or syntheticfuels. So far, most of the attention has been directed to the de-carbonization of electricity from fossil fuels, which can be realized according to three different capture strategies: pre-combustion [1 4], post-combustion [5 7], and oxy-fuel [8,9,40]. However, electricity production is responsible for only 30% of global CO 2 emissions, /$ International Association for Hydrogen Energy. Published by Elsevier Ltd. All rights reserved. doi: /j.ijhydene

2 702 S. Consonni, F. Viganò / International Journal of Hydrogen Energy 30 (2005) Nomenclature E electric power, MW F fuel, MW H hydrogen output, MW LHV m bl flow bled from steam turbine, kg/s p pressure, bar T temperature, CorK W LHV LHV power, MW γ E specific CO 2 emissions, kg/mwh γ H specific CO 2 emissions, kg/gj LHV η E net electric power/nat gas LHV power η H H 2 LHV power/nat gas LHV power η Ex Exergy (or 2nd-Law) efficiency η Ex Exergy efficiency, net of CO 2 capture Acronyms ASU air separation unit CC combined cycle ATR auto-thermal reformers FTR GT LHV HP, IP, LP HRSG HT, LT NG PSA RH, SH SC S/C ST TIT WGS WGSR Subscripts E H ref fried tubular reformers gas turbine lower heating value, MJ/kg high/intermediate/low pressure heat recovery steam generator high/low temperature natural gas pressure swing adsorption reheat, superheat steam cycle steam-to-carbon ratio (molar) steam turbine turbine inlet temperature water-gas shift water gas shift reactor Electricity Hydrogen Reforming and this share is expected to fall [10]; the remainder comes from fossil fuels used directly in distributed applications such as in transportation, residences, commercial and industrial facilities, where CO 2 capture, transport, and disposal would be problematicand extremely costly. In order to stabilize atmosphericco 2 at 450 ppm an apparently acceptable target the carbon emissions predicted by the business-as-usual IPCC IS92a scenario [11] for the period should be cut by about 50% and, by 2100, annual emissions would have to be reduced to about 3 Gigatons of carbon per year [12,13]. Even if the power sector could be completely de-carbonized, the achievement of this emission level would require a five-fold reduction of carbon emissions from fuels used directly [10] a level achievable only by a massive use of synthetic, de-carbonized fuels. Despite its many severe handicaps low density, large compression work, wide flammability limits, low ignition energy, containment problems, safety concerns, etc. hydrogen is attractive both because it is completely de-carbonized and because it is most suited to fuel cells. In a carbon-constrained world, the production of hydrogen from fossil fuels (other than via electrolysis, nuclear energy or other carbon-free processes) makes sense only if its production is associated to carbon capture. This has been investigated in a number of recent studies. Table 1 reports the basic technological characteristics and performance of the systems considered in these studies. In all cases, hydrogen is the only basic output; net electricity production is generally negative, i.e. it must be imported to satisfy auxiliary power consumption. This follows from three circumstances shared by many of the natural gas (NG)-based hydrogen plants operating today, where (i) electricity is less valuable than hydrogen; (ii) export of electricity is unattractive because the plant is remote from consumption centers and/or the grid has limited capacity; (iii) export of electricity is constrained by legislation. Should hydrogen be produced on a very large scale to feed a large fractions of the transportation, industrial and residential sectors, then these circumstances are likely to vanish because (i) hydrogen will have to compete for less profitable uses (ii) hydrogen production will take place in a relatively large number of plants all across industrial countries, with good interconnections to the electric grid; (iii) the liberalization of electricity markets being carried out in many countries will facilitate electricity exports. This is why in this paper we wish to clarify whether the co-production of significant amounts of electricity together with hydrogen can be beneficial in terms of primary energy consumption and/or reduced CO 2 emissions, and whether such practice is subject to technological and/or thermodynamicconstraints. Co-production can be accomplished by a number of technologies. In this paper, we limit our attention to established commercial technologies: fired tubular reformers (FTR) and auto-thermal reformers (ATR) for hydrogen synthesis; ironand copper-based water gas shift reactors (WGSR) for hydrogen enrichment; Rankine steam cycles (SC) and combined cycles (CC) for power production; pressure swing absorption (PSA) for hydrogen separation; amine absorption for CO 2 removal. Focusing on commercial technology allows us to utilize data and draw from the expertise developed from actual industrial plants. Results can be used as a reference to assess the merits of alternative feedstocks like coal or those of advanced technologies [14 16,42].

3 S. Consonni, F. Viganò / International Journal of Hydrogen Energy 30 (2005) Table 1 Technologies and performances quoted in recent studies on the production of hydrogen from NG with FTR Source CO 2 capture technology NG input H 2 output (LHV) Electric output % of LHV input CO 2 capture MW LHV % of input (%) gross net Foster Wheeler [14] Foster Wheeler [14] Amines Jacobs [15] Amines n.a Jacobs [15] Selexol n.a Klett et al. [16] n.a Klett et al. [16] Amines n.a The values in italics in the last two rows are estimates based on the information given in [16]. 2. Technologies, plant configurations and basic assumptions The variety of options analyzed in the paper comes from the combination of the technologies considered for: (a) hydrogen synthesis: FTR vs. ATR; (b) power production: SC vs. CC; (c) CO 2 capture: venting vs. amine absorption. This originates a total of eight technological options. For all of them we have assumed the following basicfeatures, meant to be representative of large-scale plants feeding a network for H 2 distribution and, when applicable, CO 2 storage. With respect to industrial standards, the plant is very large: NG input is in the range MW LHV, giving an hydrogen output in the range MW LHV. This compares with the hydrogen output of MW LHV of the largest FTRs built to date and the range up to 2000 MW LHV of hydrogen apparently interesting for ATRs [17]. NG input is determined by imposing that the gas turbine (GT) of the plants with CC, CO 2 capture and reference steam-tocarbon 1.0 for ATR, or 2.0 for FTR runs at the design turbine inlet temperature (TIT) while fed only by purge gas; the NG input of the other cases follows as illustrated in Table 2. The selection of two different GTs GE 7FA for ATR, GE 6FA for FTR allows maintaining the NG input of both reforming technologies within comparable ranges. The plant is stand alone, i.e. there is no steam or chemical integration with an adjoining process. Feedstock is commercial NG with enough sulfur and paraffins to require de-sulfurization and pre-reforming. De-sulfurization is carried out by catalytic hydrogenation and sulfur absorption over zinc-oxide plates at 380 C; the hydrogen required for this process (10% by volume of the gas fed to the desulfurization unit) 1 is supplied by recycling a small fraction of the pure hydrogen generated by the PSA unit. To maximize electricity production, NG is humidified (and preheated) by a saturator. Plants based on a CC also include another saturator to humidify and preheat the purge gas fed to the GT. Ahead of the reformer, the mixture of NG and water vapor is heated to 620 C and passed through an adiabaticcatalyticpre-reformer. By getting rid of higherorder hydrocarbons ethane, propane, butane, etc. the pre-reformer strongly reduces the likelihood of soot formation downstream of the reformer. The gas exiting the pre-reformer is pre-heated to 670 C with combustion gases (FTR) or with reformed syngas (ATR). Water gas shift (WGS) is carried out in two steps: the first is promoted by an iron-based catalyst in an adiabatic reactor operating between 320 and 470 C; the second is promoted by a copper-based catalyst in a cooled (ATR) or adiabatic(ftr) reactor operating between 200 and 230 C. Pure hydrogen separated by PSA is delivered at the plant gate at 60 bar. For the plants with CO 2 capture, nearly pure liquid CO 2 is delivered at the plant gate at 150 bar. For each technological option, the complex heat exchanger network needed to recover heat from the reformed syngas and from combustion gases has been arranged to limit heat transfer irreversibilities (keep low ΔT ) while meeting the assumptions on temperatures, pressures, ΔT and Δp adopted to represent the state-of-the-art of reformers and power plants (see Chapter 3). 1 This value, which is appreciably higher than that typically adopted for natural gas (2 5%), is presumably necessary to insure no carbon formation in the pre-heaters that follow desulfurization. High hydrogen concentration in the hydrogenator saturates the unsaturated hydrocarbons more prone to cracking [18], allowing higher pre-heating temperatures.

4 704 S. Consonni, F. Viganò / International Journal of Hydrogen Energy 30 (2005) Table 2 Criteria adopted to size plants with ATR at 950 C and FTR at 850 C Steam cycle Combined cycle Criterion adopted to NG W LHV, MW Criterion adopted to NG W LHV,MW determine NG input determine NG input CO 2 CO 2 capture CO 2 CO 2 capture venting venting ATR S/C = 1.0 Same NG input of plant Purge gas generated by PSA= with CC and CO 2 capture flow needed to fully load GE 7FA S/C#1.0 a FTR S/C = 2.0 Purge gas generated by PSA= flow needed by FTR burners+flow needed to fully load GE 6FA S/C > 2.0 Same NG input of plant Same NG input of case with with CC and CO 2 capture S/C = 2.0 b When the same criteria are applied to plants with CC and higher reforming temperature, NG input increases to MW (ATR at 1050 C, S/C=0.80) and MW (FTR at 880 C, S/C=2.0). The situations corresponding to the shaded cells have not been considered. a NG input varies from MW, for S/C = 1.0, to MW, for S/C = b Since the purge gas needed for the FTR burners increases with S/C, when S/C > 2.0 the purge gas to the GT is not enough, and a fraction of the NG input must be used directly in the GT combustor. When S/C > 2.75 the purge gas does not even meet the demand of the FTR burners; in this case the GT runs 100% on NG and some NG also goes to the FTR burners. Since thermal power and temperature profile of heat exchangers depend on mass flow rates, the plant configuration does not only depend on the technologies considered for reforming, power production and CO 2 capture, but also on the steam-to-carbon ratio (S/C). The plant schemes in Figs. 1 4 give the configurations considered as base cases for each technology, i.e. FTR and SC with S/C = 3.07, FTR and CC with S/C = 2.0, ATR with S/C = 1.0. Some of the results for different S/C have been obtained for a different arrangement of the heat exchanger sequence Fired tubular reformers FTR are the dominant technology for NG steamreforming [17 19,37]. They consist of a furnace filled with an array of super-alloy tubes containing a nickel-based catalyst which promotes the reforming and WGS reaction of the mixture of methane and water vapor flowing inside the tubes: CH 4 + H 2 O kj/mol CH4 CO + 3H 2 CO + H 2 O CO 2 + H kj/mol C The heat required by the highly endothermicreforming reaction is provided by burners fed with purge gas left after hydrogen extraction and some NG. Due to the large heat of reaction of steam reforming, 2 hydrogen production is 2 The catalyst brings the mixture exiting the tubes close to equilibrium, and the equilibrium constant varies with temperature according to vant Hoff equation. The large heat of reaction makes the equilibrium constant a very steep function of temperature. very sensitive to temperature, so that it is imperative that the reacting mixture inside the tubes reach high temperatures. High pressure is detrimental, because the reaction increases volume flow, yet desirable because it reduces the size (and the cost) of the equipment. In state-of-the-art FTRs the reformed syngas exits the tubes at bar and C. High temperatures, high stresses and high heat fluxes create most challenging operating conditions for the tube material, consisting of a nickel-based super-alloy at the leading edge of metallurgical science. Reforming conditions are ultimately determined by the requirements on tube creep and life, which relate to temperature and stress through correlations like Larson Miller s. Our choice of 25 bar, 850 C conforms to the state-of-the-art of large plants for the refining industry. Figs. 1 and 2 report the configuration of the plants based on a SC and a CC, respectively. NG (at 50 bar and ambient temperature) is heated to the desulfurization temperature first in a regenerative heat exchanger fed with de-sulfurized gas, then in a heater fed by reformed syngas. After being desulfurized and cooled in the regenerator, 3 NG is saturated, pre-heated and mixed with steam bled from the power cycle to achieve the specified S/C; then, the mixture to be reformed is heated to 620 C with syngas exiting the reformer, prereformed and finally heated to 670 C with the exhausts of the FTR burners. The gas gas heat exchanger heating the mixture to be reformed to 620 C ahead of the pre-reformer substitutes the quench boiler typically placed at the exit of the reformer. This appears viable because the pre-reformer 3 The temperature at the outlet of the regenerator is set by imposing a minimum ΔT of 30 C.

5 S. Consonni, F. Viganò / International Journal of Hydrogen Energy 30 (2005) H14 PSA Steam for amine stripping Steam Turbine CO 2 Absorption H2 IC Compr. H18 HP Recycled H 2 to Hydrogenation 60bar Stack NG H15 H2 Rec. Compr. Air H1 Hydrogenation and Sulfur Absorption H13 H2 H12 LT WGSR H11 H17 Deareator H16 H6 Prereformer NG Saturator H9 H10 H4 BFWH FTR Burners H5 H7 H3 H8 HT WGSR Process steam Fig. 1. Configuration of plant with FTR and SC for S/C = makes soot formation unlikely. The S/C ratio is controlled by adjusting the amount of superheated vapor bled from the steam turbine (ST) and injected ahead of the gas gas heater mentioned above. The combustion gases of the FTR burners exit the furnace at 1000 C and enter a series of heat exchangers including, besides the heater of the mixture exiting the pre-reformer, a steam superheater, an HP evaporator, an economizer and an air heater bringing combustion air to 300 C. 4 In the plants with SC, some NG (6 9% of LHV input to FTR burners) must be fed to the furnace together with PSA purge gas to control the flame. Instead, in the plants with CC the burners are fed solely by purge gas because the flame is controlled by purge gas bled (at 3.5 bar) from the intercooled compressor feeding the GT Auto-thermal reformers ATR consist of an adiabatic vessel with a section filled with nickel catalyst where the heat for the reforming reaction 4 Except in the scheme with SC and CO 2 capture, where the combustion air temperature is limited to 130 C to leave enough low-temperature heat for the economizer. Lowering the combustion air temperature increases the amount of purge gas needed for the FTR furnace and thus the attainable S/C; in turn, lower S/C give lower η H and higher η E. is provided by the partial oxidation of methane with oxygen CH O 2 CO + 2H kj/molch 4 CO + H 2 O O 2 + H kj/molc CH 4 + H 2 O kj/molch 4 CO + 3H 2 These very same reactions could be carried out also using air as the oxidant. However, unless nitrogen is needed to generate useful products (like in ammonia production) the cost increase due to the much larger flow rate across the whole plant makes air unattractive. Without a heat transfer surface, the operating temperature is not constrained by material properties and it is typically chosen to warrant adequate activity and life of the catalyst or the desired syngas composition, a situation typically encountered in plants for urea production. Our base case at 950 C is within the range of current industrial applications and allows operating at S/C favorable for CO 2 capture. 5 Despite its detrimental effects on the reforming 5 Higher ATR temperatures greatly increase the amount of steam generated in the evaporator placed downstream of the reformer (heat exchanger H7 in Figs. 3 and 4). If this steam flow is too large, superhear/reheat (SH/RH) temperature must decrease because the heat recoverable for superheating and reheating is limited. To prevent this reduction of SH/RH temperature and the consequent de-rating of the power cycle one must decrease S/C.

6 706 S. Consonni, F. Viganò / International Journal of Hydrogen Energy 30 (2005) Steam for amine stripping Purge Saturator PSA Gas Turbine H15 Purge IC Compr. CO 2 Absorption Recycled H 2 to Hydrogenation H2 IC Compr. 60bar H19 HP H18 Stack NG H16 H2 Rec. Compr. H1 Air Hydrogenation and Sulfur Absorption H14 H13 H2 H12 LT WGSR H11 Air Steam Turbine Stack HRSC HP Deareator H17 H6 Prereformer NG Saturator H9 H10 H4 FTR Burners H5 H3 HT WGSR H7 H8 Process steam Fig. 2. Configuration of plants with FTR and CC for S/C = 2.0. reaction, operating pressure is typically high to reduce the size and the cost of the reformer. In our case, reforming pressure is set to warrant the extraction of pure hydrogen from the PSA at 60 bar. Figs. 3 and 4 report the configurations considered for SCs and CCs. The path followed by NG through the desulfurization unit, the saturator and the pre-reformer is similar to that of FTRs. The basic differences consist in (i) ASU and oxygen compressor; (ii) extraction of steam for the reformer at high pressure ( 70 bar); (iii) no fuel gas from the reformer, so that heat recovery takes place solely from reformed syngas; (iv) cooled low-temperature (LT) WGS reactor, generating low pressure (LP) steam at bar 6 (v) two-pressure-level steam cycle. Extracting heat from the LT WGS reactor increases hydrogen production and the fraction of CO 2 which can be captured; these advantages will have to be weighed against the higher complexity and cost of the LT WGS reactor and of the two-pressure-level steam cycle. 6 For the cooled WGS reactor it is assumed that the evaporation temperature is 10 C lower than that of the reacting mixture. At high S/C, the pressure of LP steam is increased to warrant a temperature of the reacting mixture at least 10 C above its dew point, to avoid condensate formation in the reactor Pressure swing absorption High purity (99.99%) H 2 is assumed to be removed from the syngas at 35 C using PSA, a proprietary process commonly used in syngas processing [20 22,41]. Industrial installations based on FTRs operate at about 20 bar and reach H 2 separation efficiencies in the range 85 90%. In this paper, we have assumed 88% for all FTR plants. Maintaining the same H 2 separation efficiency at the pressure of 60 bar considered for ATRs requires a more complex (and more expensive) system arrangement. The actual design and thus the separation efficiency will be determined by economic considerations. To be consistent with our focus on best available technology and given the information provided by Jacobs [15] and Allam [23], for ATRs we have assumed a separation efficiency of 85%, a value significantly higher than that assumed by Foster Wheeler [16] and Doctor et al. [24]. On the hydrogen side, the flow incurs a 2% pressure drop, while the purge gas (remaining H 2 along with CO, H 2 O and CH 4 ) is discharged at 1.3 bar independently of the inlet pressure. In ATR plants, the pressure of the reformer is set to give a hydrogen pressure at the PSA exit of 60 bar, a value suited for long-range transport; in FTR plants, hydrogen is discharged from the PSA at

7 S. Consonni, F. Viganò / International Journal of Hydrogen Energy 30 (2005) H3 H9-10 Stack HT WGSR H4 H5 H6 Hydrogenation and Sulfur Absorption NG Saturator H2 H8 LT WGSR H12 LP H13 SH Steam Turbine HP HRSC Deareator Prereformer H7 ATR H14 H18 O2 IC Compr. H1 Burners Air Process steam ASU Vent H18 H16 Air NG NG Compressor H2 Rec. Compr. Recycled H 2 to Hydrogenation CO 2 Absorption H17 PSA Steam for amine stripping 60bar Fig. 3. Configuration of plants with ATR and conventional SC for S/C = bar and is compressed to 60 bar by an intercooled compressor Power plant Given our interest in exploring the potential for power production, we have considered a steam cycle with operating conditions typical of power plants: evaporation at bar, SH temperature C, RH at 30 bar, C. These conditions are more advanced than those typically adopted in hydrogen plants, but certainly feasible in large stand-alone plants like those considered here. For the plants based on a CC we have considered two GTs conforming to the latest generation of heavy-duty, 60 Hz engines: General Electric 6FA for FTR, 7FA for ATR. The smaller power output of the turbine considered for FTRs matches one of the intrinsic features of this technology, where the fraction of NG input converted to electricity (η E ) is smaller than that converted by ATRs. Considering the same GT for both technologies would lead to comparisons where the NG input of FTRs is much larger than that of ATRs, a situation opposite to the experience dictated by economies of scale. Table 3 compares the performances quoted by GE with those predicted by our model for operation on NG, as well as with those for syngas-fired operation. When running on syngas it is assumed that TIT, compression ratio and turbine reduced mass flow are the same of the NG version, where the last condition is met by adjusting the air flow (close Inlet Guide Vanes) CO 2 separation and compression The relatively low partial pressure of CO 2 in the syngas to be decarbonized points to chemical absorption as the technology which is presumably most cost effective. Following 7 Gas turbines typically run in choked conditions, so that an increase in fuel flow requires either a higher pressure ratio or a lower air flow. Higher pressure ratio is preferable because it gives the largest increase in power output; however, compressor stall limits its maximum allowable increase to few percent. When air flow is decreased by closing the Inlet Guide Vanes, pressure ratio might have to be decreased to avoid stall. Our assumption of constant pressure ratio and lower air flow is somewhat intermediate between what is most desirable constant air flow, higher pressure ratio and most unfavorable lower air flow and lower pressure ratio. The pressure ratio of the actual gas turbine will be determined by its compressor map, which is kept strictly confidential by manufacturers.

8 708 S. Consonni, F. Viganò / International Journal of Hydrogen Energy 30 (2005) H3 H9-10 Stack HT WGSR H4 H5 H6 Hydrogenation and Sulfur Absorption NG Saturator H2 H8 LT WGSR H12 LP H13 SH Steam Turbine HP HRSC Deareator Prereformer ASU H7 ATR O2 IC Compr. Vent H1 H16 H14 H18 Air Purge Saturator Gas Turbine Process steam Air NG H2 Rec. Compr. NG Compressor Recycled H 2 to Hydrogenation CO 2 Absorption H17 PSA Steam for amine stripping 60bar Purge IC Compressor Fig. 4. Configuration of plants with ATR and CC for S/C = 1.0. the performances quoted in [15] and the absorption properties of amines reported in [25], we have assumed that the amine absorber removes 100% of the CO 2 in the syngas (apparently, CO 2 content in treated gas is lower than 100 ppm) and requires 1 MJ of low-temperature heat per kg of CO 2 captured. Total electricity consumption with a compressor with four intercoolers is 440 kj el per kg of CO 2. The low-temperature heat needed by the amine stripper is provided by 2.1 bar steam bled from the ST. Alternatively, with FTR one could supply this heat by running the reformer at high S/C and condense the steam left in the reformed syngas, as it is done in the studies carried out by Foster Wheeler [14] and Jacobs [15]. This configuration increases η H and the fraction of CO 2 captured due to higher S/C but gives much smaller, possibly negative η E. Which configuration is preferable will obviously depend on capital costs and the values of H 2, electricity and CO Calculation model and assumptions Heat and material balances have been estimated by a computer code originally developed to assess the performances of gas/steam cycles for power production with natural gas fuel [26 29] and later extended to handle gasification of coal and biomass [30], unconventional fuels [31], chemical reactors [2], fuel cells [32], steam-cooled GT expansion [33] and essentially all processes encountered in advanced power plants. The system of interest is defined as an ensemble of components, each belonging to one of 14 basic types: pump, compressor, turbine, heat exchanger, combustor, steam cycle, etc. Basic characteristics and mass/energy balances of each component are calculated sequentially and iteratively until the conditions at all interconnections converge toward stable values. After converging, the code can carry out a Second-Law analysis to trace irreversibilities and exergy flows. The performances estimated by the model are for design conditions, i.e. it is assumed that all plant components have been specifically designed to operate at the conditions reached upon convergence of the heat/mass balances. In other words, our results apply to greenfield construction. Tables 4 and 5 summarize the assumptions maintained throughout all calculations. They are representative of stateof-the-art reforming and power plant technology, but are not necessarily optimal. Optimization of specific features like pressure drops, heat losses, temperature differences in heat exchangers or enthalpy differences in saturators require the detailed design of each component, which is beyond the scope of this work.

9 S. Consonni, F. Viganò / International Journal of Hydrogen Energy 30 (2005) Table 3 Comparison between ISO GT performances quoted by General Electric ( May 2003) and those predicted by our model for operation on NG and purge gas Turbine GE 6FA GE 7FA Fuel Natural gas Purge gas Natural gas Purge gas Source GE Our estimate GE Our estimate Input assumed to generate our estimates Air flow (kg/s) Pressure ratio TIT ( C) n.a. 1,316 1,316 n.a. 1,316 1,316 Δp compressor inlet (kpa) n.a n.a Δp turbine outlet (kpa) n.a n.a Fuel flow (kg/s) n.a n.a Ar n.a CH C 2 H C 3 H Fuel C 4 H 10 (n) Composition CO n.a n.a % volume CO H H 2 O N Fuel LHV (MJ/kg) n.a n.a Output Compressor outlet T ( C) n.a n.a Exhaust flow (kg/s) n.a n.a TOT ( C) Power output (MW) LHV efficiency (%) NG composition is the same assumed for all hydrogen plants. The purge gas composition reported for the 6FA is the one calculated for FTR at 850 C, S/C = 2.0; the composition reported for the 7FA is the one calculated for ATR at 950 C, S/C = 1.0. n.a. = notavailable. The detailed input required to calculate each GT has been fine-tuned to best reproduce manufacturer s data for operation on NG. The composition of the syngas exiting the reactors is calculated as follows: ATR and LT WGSR generate mixtures at equilibrium; the gas exiting the pre-reformer contains no hydrocarbon other than methane and meets the conditions for the equilibrium of the WGS reaction. The methane content is set to the value giving a temperature 10 C higher than that reached at equilibrium under adiabatic conditions. 8 FTR converts 88.5% of the methane that would be converted if the outlet gas were at equilibrium; 8 Due to economic considerations, the pre-reformer is typically designed for a methane conversion lower than that reached at equilibrium; on the other hand the WGS reaction, being somewhat faster, does get very close to equilibrium. The 10 C difference between the actual temperature and the one that could be reached at equilibrium is often called chemical approach. HT WGSR converts 97% of the CO that would be converted if the outlet gas were at equilibrium. Given the characteristics of the catalysts [39] and the operating practices adopted in industrial plants, 9 these assumptions reflect the average of well-maintained and welloperated plants Rationale of calculation scheme Heat and mass balances are calculated through a complex iterative algorithm. One basicindependent variable is always the amount of steam m bl bled from the ST and added to the gas to be reformed; such flow is always non-zero because the 9 For example, in a FTR a loss of catalyst activity due to catalyst poisoning affects mainly the first portion of the tubes, near to the inlet. Some FTR configurations allow to compensate these activity loss by re-positioning the flames, while maintaining the same plant capacity [18].

10 710 S. Consonni, F. Viganò / International Journal of Hydrogen Energy 30 (2005) Table 4 Assumptions adopted for the reformer, the pre-reformer, WGSR, and the hydrogenation-desulfurization section Hydrogenation & desulfurization H 2 in feed gas to hydrogenator (% vol) 10 T of gas to sulfur absorber ( C) 380 Pressure drop (%) 1 Reformer ATR FTR Syngas outlet pressure (bar) Syngas outlet temperature ( C) Inlet T of gas to be reformed ( C) Inlet T of O 2 /air ( C) a Reacting gas Δp (%) 4 4 Overpressure of O 2 /air (%/bar) Heat loss: ΔT, C/% of heat transferred 4 1 CH 4 conversion, % to equilibrium T gas at furnace exit ( C) 1000 %O 2 at furnace exit 1 Pre-reformer Inlet T of gas to be pre-reformed ( C) 620 ΔT due to heat loss ( C) 2 ΔT chemical approach ( C) 10 Water gas shift reactors Pressure drops (%) 2 ΔT due to heat losses ( C) AdiabaticLT 1 AdiabaticHT 2 Heat loss: % of heat transf. cooled LT 0.7 CO conversion to equilibrium LT/HT, % 100/97 These same assumptions have been maintained also for reforming temperature 1000 C (ATR) and 880 C (FTR). a Except FTR/SC plant with S/C = 3.07, where air pre-heat temperature is 130 C. saturator never provides enough water to reach the steamto-carbon ratios considered here. In FTRs with SC, all the purge gas is used in the furnace and m bl is adjusted until the temperature of the combustion gases exiting the furnace is 1000 C (see Table 4); this also gives S/C, which cannot be chosen at will. In ATRs and FTRs with CC, m bl is varied to meet the specified S/C. Then, in ATRs the specified reformer outlet temperature is met by adjusting the oxygen flow. Instead, in FTRs with CC the fraction of purge gas used in the furnace burners and the fraction of NG used in the reformer (the rest are used in the gas turbine) are adjusted to achieve: (i) temperature of combustion gases exiting the furnace 1000 C; (ii) full GT power. Table 5 Assumptions maintained for all calculations Air separation unit ASU power consumption (kj el /kg PURE O2 ) O 2 purity (% vol) 95 Pressure of O 2 delivered by ASU (bar) 1.01 Saturators Pressure drop of gas stream (%) 2 Pressure drop of water at nozzles (%) 10 Max ΔT of water at inlet, C below boiling point 10 Relative humidity of gas at exit (%) 100 Minimum Δh for mass transfer (kj/kg (dry gas) ) 25 Heat exchangers Gas-side pressure drop LP/IP/HP (%) 2/1/0.5 Liquid-side pressure drop (%) 2 4 Min ΔT, C Gas gas at LP/HP 40/30 Gas liquid 15 Pinch point ΔT for evaporators ( C) 10 Heat losses, % of heat transferred 0.7 PSA Separation efficiency FTR/ATR (%) 88/85 Pressure of purge gas (bar) 1.30 Pressure drop of permeated H 2 (%) 2 Compressors Polytropicefficiency of O 2 compressor (%) 82 # of intercoolers set to maintain O 2 below 120 C 6 Polytropic efficiency of fuel compressors (%) 77 # of intercoolers set to maintain fuels below 150 C Organic Electric efficiency of motor drives (%) 92 Pressure drop in intercoolers (%) 1 Pressure of fuel to GT combustor pressure 1.5 CO 2 capture and compression CO 2 concentration in treated gas (ppmv) 100 Gas pressure drop across absorber (%) 1.5 Heat duty of amine reboiler(kj steam /kg CO2 ) 1000 Total electricity consumption (kj el /kg CO2 ) 440 # of intercoolers in the compression train 4 Heat rejected to compression work (%) 150 Final delivery pressure (bar) 150 Boiler and steam cycle SC/CC evaporation pressure HP level (bar) 110/130 SC/CC SH and RH temperature ( C) 540/565 RH pressure (bar) 30 Max gas temperature at SH/RH inlet ( C) 1000 a Evaporating pressure LP level (bar) Deaerator pressure 1.40 Condensing pressure (bar) 0.05 a 1100 C in ATR with SC and CO 2 capture, to make avail- able enough heat to SH steam up to 540 C.

11 S. Consonni, F. Viganò / International Journal of Hydrogen Energy 30 (2005) Table 6 Comparison between the performances of the case study carried by Foster Wheeler [14] and those predicted by our computer code for a plant with the same configuration and the same design parameters Test case Our data Assumptions NG input (MW LHV ) T reforming ( C) p reforming (bar) S/C NG input to burners (%) H 2 in hydrogenation feed (%Vol) T hydrogenation ( C) T inlet pre-reformer ( C) T inlet reformer ( C) T inlet WGSR ( C) p inlet H 2 compressor (bar) PSA separation efficiency (%) p steam at turbine inlet (bar) T steam at turbine inlet ( C) Condensation p (bar) Results T outlet pre-reformer ( C) CH 4 at reformer outlet, (%Vol) T outlet WGSR ( C) CO at WGSR outlet (%Vol) η H (%) Gross η E (%) Net η E (%) Case study To verify the agreement between the projections generated by our model and our assumptions with those recently appeared in the literature, we have modeled one of the plants considered by Foster Wheeler in a report prepared for the IEA [14]. The plant is significantly different from those described above: there is a single WGSR operating at medium temperature ( C) and no saturator; the SC features no reheat and modest design parameters (48 bar, 380 C). The comparison shown in Table 6 indicates excellent agreement between the heat/mass balances developed by Foster Wheeler and those generated by our code. with CC because: ATRs with SC are always inferior to ATRs with CC (see par. 5.2); in order to feed the reformer with just the purge gas, FTRs with SC must operate at only one value of S/C; higher S/C require feeding the reformer with more NG than the minimum required to control the flame, while lower S/C make available extra purge gas for export a situation beyond the scope of this analysis. For FTR/SC plants, the S/C ratio giving the required match between purge gas production and reformer heat input can be (slightly) varied by varying the air pre-heat temperature, the pre-reformer inlet/outlet temperature, the reformer inlet/outlet temperature. For the plant with CO 2 capture (third leftmost column in Table 7) the air preheat temperature has been adjusted to give the same zero net power output of the plant proposed by Foster Wheeler [14], resulting in a air pre-heat temperature of 130 C (rather than 300 C, see Table 4) and S/C = As reported in Table 2, when ATR is coupled with CC all the purge gas goes to the gas turbine; consequently, the size of the reformer varies with S/C to match the requirements of the GE 7FA assumed for our calculations. For FTR with CC and S/C = 2.0, purge gas production is enough to meet the requirements of both the reformer and the GE 6FA gas turbine; when S/C > 2.0, the reformer needs more purge gas (and the purge gas LHV is lower), so that some NG must be fed to the GT to run it at full power. This is shown in Figure 5. When S/C is larger than 2.75, purge gas production cannot even meet the requirements of the reformer; then, the GT is fed solely with NG and the reformer and the CC are coupled only through the steam section. To maintain the tightest integration between the FTR and the CC, when S/C increases one should let NG input rise until purge gas production can match the gas turbine heat input or, for the same NG input, adopt a smaller GT. This has not been considered because very large NG inputs (above 2000 MW LHV ) appear impractical and are likely to make FTRs more expensive than ATRs; on the other hand, small gas turbines would increase CC specific costs, reducing even further the attractiveness of FTR/CC plants. These conjectures need to be verified by an adequate economic analysis. 4. Performance estimates Heat and mass balances of the plants depicted in Figs. 1 4 have been calculated for the conditions in Tables 2 5, allowing for some variations of reforming temperature and S/C ratio. Table 7 gives a summary of overall performances. Fuel inputs and outputs are always evaluated on a LHV basis. Variations of S/C have been considered only for plants 4.1. Overall performance indicators From the point of view of the First Law, the quality of the thermodynamicsystem is expressed by: η H = fraction of NG input exportable as pure hydrogen at 60 bar; η E = fraction of NG input exportable as electricity.

12 Table 7 Overall performances of the configurations investigated in this paper Power plant fate of CO 2 Steam cycle Combined cycle Vent Capture Vent Capture ATR T ref C S/C NG input MW LHV E from GT MW E from ST MW η H2 % LHV η E % LHV η Ex % η Ex % CO 2 capture % H/F H % γ H kg CO 2 per GJ LHV FTR T ref C S/C NG input MW LHV E from GT MW E from ST MW η H2 % LHV η E % LHV η Ex % η Ex % CO 2 capture % H/F H % γ H kg CO 2 per GJ GJ LHV S. Consonni, F. Viganò / International Journal of Hydrogen Energy 30 (2005) For comparison, shaded columns report the performances of the case study considered by Foster Wheeler [14]. H/F H is the equivalent efficiency of hydrogen production; γ H is the specific emission chargeable to H 2 (see par. 5.6). η Ex accounts for the exergy of the pure, pressurized CO 2 made available at the plant gate; instead, η Ex disregards the exergy of CO 2. The difference between η Ex and η Ex is the CO 2 exergy divided by the exergy of the NG input.

13 S. Consonni, F. Viganò / International Journal of Hydrogen Energy 30 (2005) % of Natural Gas Input S/C Vent Capt SC To FTR Burners Vent Capture CC To Reformer To Gas Turbine Fig. 5. Breakdown of NG input of FTR plants. The fraction to burners shown for plants with SC ( 7.5 of LHV input to burners) is needed to control the flame. In the plants with CC, the flame is controlled via a slipstream bled from the purge gas compressor. The fraction to burners shown for the plant with CC and S/C=3.5 is needed to complement purge gas production, which cannot meet the requirements of the reformer. while the Second-Law perspective is given by: η Ex = ratio between exergy out (hydrogen+electricity+ CO 2 ) and exergy in (NG); η Ex =ratio between the exergy of (hydrogen+ electricity) and the exergy of the NG feedstock. Note that exergy [34] includes the work generated by reversibly mixing the output flows into the atmosphere, i.e. by expanding each species to its atmospheric partial pressure [38]. Such work ΔEx mix, which accounts for most of theexergyofco 2, cannot be recovered with the technologies currently available. Its inclusion into the exergy flows is necessary to close the Second-Law balance but overemphasizes the relevance of output flows particularly for the CO 2 captured. Table 8 gives LHV, exergy and ΔEx mix of H 2, NG and CO 2, calculated from JANAF data [35] and, for ipercritical CO 2, from [36]. Table 7 points out the following: CO 2 capture reduces η E by 2 3% points but has negligible impact on hydrogen production. Yet, in plants with FTR and SC the loss of electric output is partially compensated by higher hydrogen output: comparing the 1st and 3rd column from the left in the bottom (FTR) section of the table, one can see that CO 2 capture reduces η E from 3.2% to zero, but η H2 increases from 75.7% to 78%. Compared to FTRs, ATRs give lower η H but higher η E. When power is provided by a CC, the GT power output of an ATR plant is a much larger fraction of NG input (and H 2 output), thus justifying the different GT sizes adopted for the two reforming technologies. ATRs attain higher fractions of CO 2 capture, particularly when coupled with CC and operated at high S/C; when S/C 2, CO 2 capture is larger than 85%. Instead, FTRs can achieve relatively low CO 2 capture. Only by totally giving up electricity export does one go over 70% with a plant based on a steam cycle (see FTR with S/C = 3.07). As explained in par. 4.2, this follows from imposing that the burners are fed solely with purge gas, which increases η H but limits CO 2 capture. The Second-Law efficiency η Ex slightly increases when S/C and thus the fraction of CO 2 captured increase. Even if the contribution to η Ex of compressed CO 2 may be questionable, because it does not correspond to the work recoverable with current technology, this trend points out that CO 2 capture does not entail specific thermodynamic drawbacks. This is highlighted also by the trend of η Ex, the Second-Law efficiency net of CO 2 capture Electricity production and CO 2 capture vs. H 2 output Fig. 6 reports η E vs η H, as well as the fraction of CO 2 removed. Points in the left part of the diagram (50% < ηh < 65%) refer to ATR; the ones in the right part of the diagram (65% < ηh < 80%) refer to FTR. The figure shows that for ATR/CC plants the variation of S/C introduces a trade-off among electric output which decreases when S/C increases hydrogen production and CO 2 capture which both increase when S/C increases. Instead, FTR/CC plants operate in a very narrow range of η E, η H and CO 2 capture independently of the S/C ratio. ATRs with SC are definitely inferior to ATRs with CC because: for the same S/C, the plant with SC gives the same η H and the same CO 2 capture of the plant with CC, but η E is much lower; if one lets S/C of the plant with CC vary, it is possible to reach a condition where η E of the SC and the CC are the same, but then the CC features higher η H and higher CO 2 capture. For the same S/C, the reforming section of the plants with SC and CC are identical: heat for the saturator and the reformer is taken solely from the reformed gas, which does not vary with the power plant; then, for the same amount (and composition) of purge gas and the same heat made available to the power plant, the CC is obviously more efficient. With FTR the CC does not show a clear advantage because, despite its higher η E, it exhibits lower η H and lower carbon capture. In fact, FTR/SC plants where the purge

14 714 S. Consonni, F. Viganò / International Journal of Hydrogen Energy 30 (2005) Table 8 Conditions, LHV and exergy of input and output flows p (bar) T ( C) LHV (MJ/kg) Ex (MJ/kg) ΔEx mix (MJ/kg) H NG CO Reference ambient conditions are 15 C, 1 atm, 60% relative humidity. ΔEx mix is the work recoverable by reversibly mixing the combustion products into the atmosphere. η E =electric output as % of NG input SC CC S/C=1.0 S/C=1.0 SC S/C=1.0 ATR E/NG input (left y-axis) S/C=1.5 CC S/C=1.5 S/C=2.1 S/C=2.1 CC, S/C=1.0 Tref=1000 C CO 2 capture yes no S/C=2.0 CC S/C=2.0 S/C=2.0 S/C=3.5 S/C=3.5 FTR CC, S/C=2.0 Tref=880 C SC S/C=3.0 % CO 2 captured (right y-axis) η H =hydrogen output as % of NG LHV input CC SC S/C= CO2 Captured [%] Fig. 6. η E and fraction of CO 2 removed as function of η H. The figure reports the same data in Table 7. Solid dots to be read on the right y-axis are given only for plants with CO 2 capture (with venting, CO 2 capture is always zero). gas matches the requirements of the reformer (as we have assumed here) reach η Ex ranging from 76% to 78% (see Table 7) an unsurpassed thermodynamicperformance. Plant performances are very sensitive to the reforming temperature. For ATR with CC, increasing the reforming temperature from 950 to 1000 C generates about the same effect of doubling S/C (from 1 to 2) at constant T ref =950 C Second-lawanalysis Fig. 7 reports the breakdown of η Ex, as well as the exergy losses of four configurations representative of each of the reforming/power generation technologies considered here. FTRs with SC stand out as the most efficient, whereas the opposite holds for ATRs with SC. The excellent performance of FTR/SC follows from very low losses in the reformer, while the poor outcome of ATR/SC is mainly due to large heat transfer losses. In a FTR, the reforming reaction is thermodynamically more efficient because it is fed by a lowgrade fuel like purge gas; instead, in a ATR the heat for steam reforming is provided by burning (although partially) NG. As for heat transfer, the larger losses shown for ATRs with SC are due to the quench boiler placed downstream of the reformer and the boiler fed by the purge gas burners, which operate under large temperature differences. When coupled to a CC, ATR combustion losses decrease because the purge gas burners are replaced by the GT combustor. On the other hand, at the high S/C needed to achieve high CO 2 capture, reforming losses are higher because more NG must be burnt to feed the reaction; at the same time, combustion losses are lower because of lower purge gas flow. For FTRs the CC ends up with lower η Ex because some extra- NG must be burnt into the gas turbine combustor, i.e. there is a fundamental imbalance between the reforming section and the power section. The sum of the losses due to combustion and chemical reaction is always smaller for FTR because the heat generated by the combustion of purge gas is recycled to