Techno-economic Assessment of Membrane Reactor Technologies for

Mar 5, 2013 - Techno-economic Assessment of Membrane Reactor Technologies for Pure Hydrogen Production for Fuel Cell Vehicle Fleets. Leonardo ...
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Techno-economic Assessment of Membrane Reactor Technologies for Pure Hydrogen Production for Fuel Cell Vehicle Fleets Leonardo Roses,*,† Giampaolo Manzolini,† Stefano Campanari,† Ellart De Wit,‡ and Michael Walter‡ †

Dipartimento di Energia, Politecnico di Milano, Via Lambruschini 4, 20156 Milano, Italy HyGear B.V., Post Office Box 5280, 6802 EG Arnhem, The Netherlands

‡

ABSTRACT: In the evolution toward a “carbon-neutral” energy economy, among the most promising solutions for replacing today’s greenhouse gas (GHG)-emitting vehicles is the use of hydrogen as an energy carrier. In the pathway toward a future infrastructure based on renewable energy sources, a medium-term step would rely on the use of fossil fuels for on-site production of hydrogen, feeding small fleets of fuel cell vehicles. Great interest is on natural gas as a primary source because of its high hydrogen/carbon ratio. State of the art technology for the production of hydrogen from natural gas includes a series of reacting steps typically involving steam reforming (at 800 °C or above), a water-gas shift reactor, and a final purification of hydrogen through pressure swing adsorption (PSA). An alternative that has been the subject of growing interest is the use of thin (2−50 μm thick) Pd-alloy materials as hydrogen perm-selective membranes for the embedded extraction of pure hydrogen from the chemical reactor; this system is usually known as the “membrane reactor”. This paper studies the adoption of palladium-based membrane reactor technologies for pure hydrogen production from natural gas. In particular, three system layouts are analyzed and compared to the traditional option: (i) autothermal reforming membrane reactor, (ii) steam reforming membrane reactor (externally heated), and (iii) water-gas shift membrane reactor downstream of a steam reformer. The comparison is made in terms of performances and techno-economic considerations for the design of compact systems for on-site production of hydrogen at filling stations. The systems are designed for 50 m3/h (1766 cfh) of hydrogen, which corresponds to refilling 25 vehicles a day with 4 kg of hydrogen (approximately 418 km driving range on fuel cell vehicles with a 70 MPa storage tank).

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INTRODUCTION In the challenge of reducing the carbon footprint in mobility, the use of hydrogen as an energy carrier has been envisioned for many years as a promising solution for avoiding the emission of carbon dioxide from vehicles. The transformation into a hydrogen-based economy raises some challenging obstacles, such as the required infrastructure in filling stations and the issue on the primary energy source for the production of hydrogen. Diverse regional and national programs are already active in the introduction of a hydrogen infrastructure for a future transport market with a significant share of hydrogen-based technology.1−3 In the long term, the prospective solution consists of using wind and solar or any other renewable energy sources to produce hydrogen by means of electrolysis of water. However, in a transition process moving toward this objective, hydrogen is likely to be produced from reforming of light hydrocarbons, such as natural gas (NG). The use of NG seems to be the most feasible option because of the fact that many countries already have a distribution network widely developed and also because its low carbon/hydrogen ratio leads to a lower specific CO2 emission compared to heavier hydrocarbons. This work will investigate different options for on-site hydrogen production. The assessment will be performed both from a thermodynamic and an economic point of view. The benchmark technology for the production of hydrogen from methane would consist of the production of a hydrogen-rich syngas by means of steam reforming at a temperature of around 800 °C, followed by a water-gas shift reactor, and later separation of hydrogen using pressure swing adsorption (PSA). © XXXX American Chemical Society

Besides the reference technology, three different configurations based on perm-selective membranes for the separation of pure hydrogen will be proposed. This technology has potential benefits in terms of compactness and cost reduction.

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MEMBRANE REACTOR Membrane reactor technology generally refers to the integration of a reacting process with membranes having the characteristics of being selectively permeable to a specific species, allowing for the addition or removal of that species from the reacting mixture. In this work, the application of hydrogen selective membranes based on palladium alloys to steam reforming is evaluated: the H2 is separated from other components, such as CO2, CO and steam, and is sent to the permeate side of the membrane. Figure 1 shows the case of a steam reforming membrane reactor, where the extraction of hydrogen from the reforming reaction shifts the equilibrium conversion of methane toward higher values with respect to the conventional reforming reaction. In this way, higher values of conversion can be achieved at a lower reacting temperature. This is in fact one of the key effects that turn the membrane reactor fuel processing attractive when compared to the traditional approach. The process intensification obtained with the use of membrane reactors reduces the number of Special Issue: Accelerating Fossil Energy Technology Development through Integrated Computation and Experiment Received: November 30, 2012 Revised: February 22, 2013

A

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s−1 m−2 Pa−n set by the United States Department of Energy (U.S. DOE) as a target for 2015 in its Hydrogen and Fuel Cell Program.8 In practical applications, this value of permeance is still far from the performances achieved by many developers; however, some have already proven this permeance to be a feasible objective,5,9,10 and it is expected that evolution of the research in membrane technology will move the state of the art performance forward. In this work, the calculations are performed assuming a membrane having the permeance set by the above-mentioned U.S. DOE 2015 target. Nevertheless, a sensitivity analysis is also carried out showing the influence of this assumption on the economic results.

Figure 1. Schematic view of the steam reforming membrane reactor (SR-MR).

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components and eliminates the need for batch operating PSA units with potential equipment savings, besides allowing for a more compact system. Separation of pure hydrogen can be achieved through the use of dense palladium-based membranes of reduced thickness (2− 50 μm).4−7 The minimum thickness that can be used in a membrane is linked to the preparation techniques, kind and quality of supports used, and operating conditions required for the membrane reactor in terms of pressure, temperature, and selectivity target. The permeation of hydrogen across the membrane is usually a function of the membrane permeance and hydrogen partial pressure difference, as shown in eq 1 JH = PH2(pHn ,ret − pHn ,prm ) 2

2

2

SYSTEM LAYOUT

Three innovative configurations and one reference case are considered in this work. They are as follows: (1) SR-PSA, steam reforming with pressure swing adsorption; (2) SR-MR, steam reforming membrane reactor; (3) ATR-MR, autothermal reforming membrane reactor; and (4) WGS-MR, water-gas shift membrane reactor. Starting from the reference case, the SR-PSA layout is shown in Figure 2. NG is compressed to the operating pressure and preheated prior to the desulfurization unit. The steam reforming reactor is controlled at an outlet temperature, which is optimized for performance, falling typically within 800 and 900 °C. Then, the reformate is cooled prior to the water-gas shift reactor (WGSR). In the reforming reactor, the reactions R1 and R2 (see below) take place, whereas in the WGSR, the operation is selective only to R2 over a Febased catalyst.

(1)

where “PH2” is the permeance of the membrane measured in mol s−1 m−2 Pa−n), the term between parentheses is the driving force for permeation depending upon the respective partial pressure of hydrogen “pH2” on the retentate and permeate side, and “n” is the exponent for the driving force, which is experimentally determined to be in the range of 0.5−1.0. Many factors may affect the permeance of a membrane, with the most important being the type of alloy, the manufacturing procedure, type of support, surface quality, heat treatment, and thickness of the perm-selective layer. A technological challenge is addressed by the objective of a permeance of 8.52 × 10−3 mol

CH4 + H 2O ⇒ CO + 3H 2

CO + H 2O ⇒ CO2 + H 2

° K = 206 kJ/mol ΔH298

(R1)

° K = − 41 kJ/mol ΔH298

(R2)

The syngas is then cooled to the operating temperature of the PSA unit of 50 °C. The impurities from the PSA and possibly additional NG are sent to the burner required for sustaining the endothermic reforming reaction and producing the required steam. The configuration involving SR-MR is shown in Figure 3. In this architecture, the separation of hydrogen takes place in a single reactor operating at 600 °C. In the SR-MR reactor, the equilibrium of both reactions R1 and R2 is shifted toward the products thanks to the simultaneous extraction of hydrogen from the reacting section. The

Figure 2. Layout traditional system “SR-PSA”. B

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Figure 3. Layout steam reforming membrane reactor system “SR-MR”.

Figure 4. Layout autothermal reforming membrane reactor system “ATR-MR”. The layout of the system with WGS-MR is shown in Figure 5. In this configuration, a traditional steam reforming reactor is used for methane conversion, with the syngas then cooled to 300 °C as the inlet temperature to a WGS-MR. In the shift reactor only, R2 takes place over a Fe-based catalyst, together with the extraction of hydrogen through the perm-selective membrane. The hot gas produced by combustion of the retentate and additional NG is use for sustaining the reforming reaction. A summary of the conditions obtained in specific points of the layouts as described in Figures 2−5 is shown in Table 1. As seen in the layouts, points “a−e” refer to the inlets/outlets of the reactors, pressurized hydrogen, and off-gas sent to the burner. From the results in the table, it is possible to discriminate the effect of the extraction of hydrogen in the membrane reactor; for instance, methane contents in points “b” and “e” for the SR-PSA and SR-MR systems are 2.5 and 1.5%, respectively, showing higher reforming conversion in the SR-MR system, despite the lower temperature. Similarly, the effect of extraction of hydrogen in a water-gas shift reactor is evidenced in

extracted hydrogen is then cooled prior to the inlet of the compressor. The retentate together with additional NG is sent to the burner required for sustaining the endothermic reforming reaction and for steam production. The layout of the system with the autothermal membrane reactor is shown in Figure 4. In this system, the reactions in the ATR are thermally balanced: the exothermal reaction of methane partial oxidation R3 and partly R2 supply the heat to support the reforming R1, which is endothermic. This configuration therefore requires an air compressor. The fraction of stoichiometric air provided to the reactor is determined to keep the outlet temperature of the reactor at 600 °C. The retentate, which has unconverted species and the remaining hydrogen, is combusted with additional air and then cooled while producing steam. CH4 + 2O2 ⇒ CO2 + 2H 2O

° K = − 802 kJ/mol ΔH298 (R3) C

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Figure 5. Layout water-gas shift membrane reactor system “WGS-MR”.

Table 1. Results of Compositions and Properties in Specified Points of the Layout a point

SR-PSA

composition (%mol) H2 0 H2O 75 CH4 25 CO 0 CO2 0 O2 0 N2 0 T (°C) 550 P (barg) 25 F (mol/s) 1.00

c

d

SR- MR

ATR-MR

WGS-MR

SR-PSA

b WGS-MR

SR-PSA

all

SR-PSA

SR- MR

ATR-MR

WGS-MR

0 75 25 0 0 0 0 350 25 0.672

0 44.9 20.9 0.0 0.0 7.2 27.0 650 25 1.05

0 75 25 0 0 0 0 540 25 1.00

50.2 32.3 2.5 9.8 5.2 0 0 884 25 1.43

43.9 37.4 6.1 6.6 6.0 0 0 800 25 1.34

57.9 24.6 2.5 2.2 12.8 0 0 386 25 1.43

100 0 0 0 0 0 0 50 24 0.620

45.0 0.5 7.8 6.7 40.0 0 0 50 0.5 0.459

5.7 49.7 1.5 2.1 41.0 0 0 600 25 0.377

5.2 29.3 1.8 1.9 24.9 0 36.9 600 25 0.767

7.8 57.5 11.3 0.2 23.2 0 0 372 25 0.723

“c” and “e” for the SR-PSA and WGS-MR systems, respectively, showing a much lower CO content for the membrane reactor solution.

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e

reforming reactor was set to 2.5 MPag for all configurations studied, and a 0.1 MPa pressure drop was assumed (being the predominant) across the PSA. The resulting pure hydrogen pressure is at 2.4 MPag. For automotive applications, highpurity hydrogen is desired (grade 5.0); hence, the operation of the PSA would operate at a lower recovery factor than for the production of technical-grade hydrogen (e.g., 2.5).16 Of the overall hydrogen flow rate at the PSA inlet, a recovery of 75% was assumed. The compressors used for NG, air, and hydrogen are reciprocating multi-stage-type with intercooling stages to keep the discharge temperature below approximately 150 °C (300 °F). Air and NG are compressed to 2.5 MPag in the ATR-MR system, whereas in the other cases, air is only required at a low pressure at the burner. With regard to the compression of hydrogen, in the models of all four system layouts the outlet pressure of hydrogen was set to 2.4 MPag to match the conditions set by SR-PSA. The consumption of a booster, which compresses the hydrogen from 2.4 to 70 MPag required by the refilling station, was calculated independently being that it is the same (five stage intercooled) in any of the four cases studied.

DESIGN AND MODELING CONSIDERATIONS

The size of the hydrogen production system is set to supply a fleet of 25−30 refills a day of hydrogen-fueled vehicles. This definition of scale of production corresponds to an early market stage, where the hydrogen-fueled vehicle is not yet fully immersed in the market. A fuel cell powered family car performing at approximately 105 km/kgH2,11,12 a value taken from real operation experiences and more conservative than typical results from standard driving cycles,13 assuming a range of autonomy of 483 km, requires storage for 4.6 kgH2. Therefore, a plant with a production rate of 4.5 kgH2/h would supply hydrogen for about 30 vehicles a day topping up their tanks by 75−80% of their capacity. This indication is consistent also with recent reports from the car industry for the preliminary distribution of hydrogen into small retail stations.14 The most traditional technology for the separation of pure hydrogen is based on the PSA unit. The equilibrium-based process of PSA typically operates at a pressure ranging from 2 to 3 MPa.15 In this work, the operating pressure in the D

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Table 2 summarizes the operating parameters of the compressors.

Table 3. Design Parameters parameter steam reforming control temperature in SR-PSA steam reforming control temperature in WGS-MR ATR-MR and SR-MR control temperature WGSR and WGS-MR inlet temperature desulfurization temperature minimum temperature difference for all heat exchangers PSA hydrogen recovery excess feedwater for complete hydrogen production excess air for combustor reactors operating pressure hydrogen output pressure hydrogen production capacity membrane permeance

Table 2. Compression Parameters

inlet pressure outlet pressure number of stages compression ratio isentropic efficiency mechanical−electrical efficiency maximum discharge temperature maximum temperature after intercooling

units

NG

hydrogen

air (ATR-MR case)

kPag MPag

% %

4 2.5 3 2.92 80 98

9 2.4 5 1.87 80 98

0 2.5 4 2.25 80 98

°C

155

130

155

°C

50

50

50

With regards to the design of the membrane reactor, all membrane cases were designed with equal permeate pressure set slightly above ambient pressure (9 kPag). In linear membrane reactors, such as the one shown schematically in Figure 1, the fraction of hydrogen on the pressurized side of the membrane decreases toward the exit of the reactor, reducing the driving force for permeation. Therefore, the flux of hydrogen is the lowest toward the end of the reactor, as shown in Figure 8. The criteria used for membrane reactor design is that of keeping as minimum flux between the 25 and 30% of average flux. A sensitivity analysis on this assumption will be performed in the last part of this work. The reforming reactors are designed for operating at temperatures between 800 and 900 °C in the SR-PSA and WGS-MR systems and at 600 °C in the ATR-MR and SR-MR. In the ATR-MR reactor, a maximum of 50 °C above the operating temperature is admitted at the reactor inlet (see Figure 4). The inlet temperature for the WGS reactors is of 300 °C in the corresponding systems either with or without membranes. It is to be noted that 300 °C is the temperature to the inlet to the WGS reaction section. A WGS-MR is designed so that the membrane begins after an initial section of conventional shifting without permeation. The gas arrives to the membrane reactor section at 360 °C. In all cases, the reactors are considered adiabatic. For all four investigated cases, a steam excess in the reforming reaction of 50% is considered, leading for example to a steam/carbon ratio of 3.0 at the inlet of the steam reforming reactor. A desulfurization step is required prior to any reactor because sulfur poisons the reforming catalyst. For effective sulfur removal, this process typically takes place between 300 and 400 °C over a ZnO bed. Further design considerations are summarized in Table 3. The model is coded using the Aspen Plus simulation software, a well-known modeling tool for process conceptual design for chemical industries, which was found to be quite suitable for simulating the process described here. The thermodynamic properties were evaluated using a standard Peng−Robinson cubic equation of state,17 except for liquid molar volume evaluation, where the Rackett model is used,18,19 and the NBS/NRC steam tables are used for steam properties.20 For simplicity, the fuel inlet has been considered as pure methane. Figure 6 shows a snapshot of the implementation of the membrane reactor in Aspen Plus. The model consists of a series of reactors and hydrogen separators

units

value

°C

884

°C

800

°C

600

°C

300

°C °C

330−350 40

% %

75 50

% MPag MPag m3/h mol s−1 m−2 Pa−0.5

40 2.5 2.4 50 8.52 × 10−3

in series, coupled with heat exchangers for thermal transfer with the hot gas from an external stream where applicable (as in the SR-MR layout). In the reactors, the conversions through reforming and shift reactions are assumed to achieve chemical equilibrium and determined by Gibbs free energy minimization. In the hydrogen separators, the model calculates the average driving force for permeation and calculates the required membrane for a given hydrogen flow. The number of separators are set for the total amount of permeated hydrogen, and the total membrane area is determined as a sum of those on the single separators. The example shown in Figure 6 applies for the case of the SR-MR system. The inlet on the left side of the figure corresponds to the feed inlet to the reactor as in point “a” of the layout in Figure 3; the outlet on the lower right represents point “e” of the layout as the retentate leaving the reactor, and on the upper right is the outlet of permeated hydrogen.

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PERFORMANCE COMPARISON An initial comparison regarding operating performance in terms of efficiency and cost was performed. Table 4 shows the results for energy cost performance. Results in terms of hydrogen output and thermal and total efficiency in this table are determined as follows: H 2 output (kW) = FH2 HHVH2

ηth = ηtot =

(2)

FH2 HHVH2 FNG HHVNG

(3)

FH2 HHVH2 FNG HHVNG + Pel

(4)

where FH2 and FNG are the flows of hydrogen and NG, respectively. The table shows the results in terms of hydrogen recovery factor (HRF). Such a term is defined as the ratio between the permeated hydrogen and the maximum amount of hydrogen that could be theoretically produced given the feed to the reactor. This result describes the performance of the membrane reactor for the purpose of hydrogen separation. Table 4 shows the total input of natural gas (NG input) in terms of thermal E

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Figure 6. Snapshot of the membrane reactor implemented in Aspen Plus for the SR-MR system.

Table 4. Performance Comparison NG input reformer burner total electric consumption NG compression air compression H2 compression at 2.4 MPag H2 compression at 70 MPag BOP and controls membrane area H2 output HRF thermal efficiency (ηth) total efficiency production cost

units

SR-PSA

SR-MR

ATR-MR

WGS-MR

kW(HHV) % % kW kW kW kW kW kW m2 kW(HHV) % %(HHV) % $/kgH2

222.6 100.0

203.0 73.7 26.3 18.0 2.0

195.9 100.0

234.6 95.2 4.8 19.1 3.0

11.7 3.0

79.6 75.6 1.27

7.3 7.9 0.8 0.308 177.1 92.2 87.3 80.2 1.33

22.8 2.6 4.4 7.3 7.9 0.6 0.437 177.1 84.9 90.4 81.0 1.41

7.3 7.9 0.9 0.150 177.1 61.7 75.5 69.8 1.50

7.9 0.8 177.1

NG cost

$/kgH2

1.00

0.92

0.88

1.06

electricity cost

$/kgH2

0.27

0.41

0.52

0.44

power (as FNGHHVNG) divided between the fractions sent to the reformer and the burner. In terms of thermal efficiency, both SR-MR and ATR-MR show higher efficiency than the traditional option. This benefit comes from the lower exergy losses associated with a lower operating temperature of the reformer and simultaneous hydrogen separation, which reduces temperature swing losses. The electric power consumption, as expected, is higher in any of the membrane reactor systems with respect to SR-PSA, because of the higher demand for hydrogen compression. The ATR-MR system is also penalized for the air compression. After evaluation of the thermodynamic performance of the three innovative cases with the reference case, the comparison was then performed in terms of H2 production costs. In this assessment, only the two main operating costs, e.g., those related to NG and electricity consumption, were considered. The prices of NG and electricity were taken as follows: $5.87/ GJ of gas ($6.34/Mcf) from the U.S. industrial average price in the period from January 2001 to June 201121 and $0.103/kWh from the U.S. commercial sector 2011 average.22 Table 4 also shows that the production costs in the traditional system remain lower than in any of the three membrane-reactor-based configurations. The latter are penalized by the higher electricity consumption and, consequently, affected by the cost of

electricity. However, it is seen that with the SR-MR, the production cost is relatively close to the SR-PSA option. The production costs are strongly related to the prices of NG and electricity. The NG price, for instance, shows strong variability linked to the location or political reasons; for example, the yearly averaged industrial price in 2010 was 4.3 and $11.2/GJ in Texas and Rhode Island, respectively. Moreover, NG prices in other countries are substantially higher than the reference adopted here, as it happens for instance in several European Union (EU) countries, as well as Japan. Illustrating this effect, Figure 7 shows the results of the production costs recalculated for 4.3 and $11.2/GJ compared to the reference case of $5.87/GJ. The figure shows how a relatively high price of NG would put the SR-MR solution as convenient over the traditional option in terms of variable production costs.

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COST COMPARISON

Until the technology for hydrogen distribution at filling stations is not a settled market, the investment and operation costs of the system remain uncertain. Cost estimations were taken from a detailed study23 conducted in 2002 for a system similar in configuration and scale to the SR-PSA described in this work F

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based systems balances the higher production costs as a consequence of the higher cost for hydrogen compression.

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CRITERIA FOR MEMBRANE UNIT DESIGN As previously mentioned, the membrane reactor was designed to set the flux at the exit of the reactor between 0.25 and 0.30 of the average one. Figure 8 shows the profiles of the hydrogen

Figure 7. Change in hydrogen production costs with NG price.

and updated to 2012 values using the Chemical Engineering Plant Cost Index24 (CEPCI). Capital investment for the SRPSA system results in $392 745, from which $28 057 corresponds to the PSA unit and $32 256 corresponds to the hydrogen compressor. Considering these equipment prices for the traditional system, the deviations from this cost for the innovative cases were considered, as shown in Table 5, distinguishing the costs associated with key components. In particular, the cost of the hydrogen compressor varies from the low inlet pressure used in the MR systems to the 2.4 MPag inlet pressure of the SR-PSA system. The membrane cost was taken as $5382/m2 ($500/ft2) as in the U.S. DOE target for 2015,8 which will be discussed below in a sensitivity analysis. The economic and lifecycle conditions adopted in the analysis consider a 10 year lifetime of the system, and operation and maintenance (O&M) costs include 5% of the capital cost, which covers yearly replacement of the desulfurization bed absorbent, replacement of reformer and shift catalyst after 5 years, and general maintenance for pumps, compressors, burner, valves, and instrumentation. The membranes are assumed to be replaced after 5 years of operation,8,25 and such a cost is amortized through the 10 years of operation and included in the O&M. The table shows how the total cost of hydrogen in $/kg for the SR-MR solution can be competitive with the traditional system, whereas both in ATR-MR and WGS-MR, MR features higher costs. The potential equipment savings of membrane-

Figure 8. Permeation flux profile for all three membrane reactors.

flux along the membrane resulting from the calculations in all three MR configurations. The figure shows the flux profile from the inlet up to the total membrane area (as also indicated in Table 4). As the figure shows, the WGS-MR gives the most efficient use of the membrane surface, followed by the SR-MR. This effect is a consequence of the higher hydrogen content because of the shift reaction. The adopted design criteria is not general, and it should be modified accordingly depending upon the application or the configuration, such as the use or not of a sweep gas on the permeate side (not considered in this work). To validate the use of the design criteria adopted, a sensitivity analysis on the membrane area design was performed. The results of this analysis are summarized in Table 6. The table shows the performance results of the SR-MR system for ±10% of the membrane area together with the designed configuration shown above in Table 4. An additional case, which minimizes the membrane surface area, is also summarized. The results show that the direct effect of the membrane area on production costs is not too significant, when a ±10% change in membrane area leads to a ±0.8% in production costs.

Table 5. Comparison of the Cost of Hydrogen units

SR-PSA

SR-MR

ATR-MR

WGS-MR

capital cost main components air compressor hydrogen compressor PSA membranesa O&M (annual)b insurance (annual)c annual uniform capital costd total annual total cost

$ $ $ $ $ $ $ $ $ $ $/kgH2

392745 332432

373552 332432

372704 332432

32256 28057

39465

382383 332432 8134 39465

19637 11782 63917 95337 4.73

1655 18845 11207 60794 90845 4.62

2352 19357 11471 62231 93059 4.78

807 18717 11181 60656 90554 4.78

capital recovery

$/kgH2

3.46

3.29

3.38

3.28

production cost

$/kgH2

1.27

1.33

1.41

1.50

a

39465

Considering membrane cost at $5382/m2 ($500/ft2). bA total of 5% of the capital cost + membrane replacement. cA total of 3% of the capital cost. Interest rate of 10%, lifetime of 10 years, and capacity factor of 70%.

d

G

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Table 6. Sensitivity to Membrane Design Criteria units

SR-MR optimized

SR-MR + 10% membrane

SR-MR − 10% membrane

SR-MR min membrane

kW(HHV) % % kW m2 % %(HHV) %(HHV) $/kgH2

203.0 73.7 26.3 18.0 0.31 92.2 87.3 80.2 1.33

202.4 71.9 28.1 17.9 0.34 94.8 87.5 80.4 1.32

204.6 78.3 21.7 18.1 0.28 86.1 86.5 79.5 1.34

211.9 100.0 0.0 18.8 0.23 65.1 83.6 76.8 1.39

NG cost

$/kgH2

0.92

0.91

0.92

0.96

electricity cost

$/kgH2

0.41

0.41

0.41

0.43

NG input (total) reformer burner total electric consumption (Pel) membrane area HRF thermal efficiency (ηth) total efficiency (ηtot) production cost

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With regard to the cost of the membrane, as already addressed, this work considered $5382/m2 as in the U.S. DOE target for 2015.8 The assumption on the cost of palladiumbased perm-selective membranes is a delicate matter, because it is associated with not only the requirement of noble metals but also elaborate preparation techniques. Moreover, as addressed above, the permeance of the membrane as defined by the U.S. DOE target results in an ambitious performance compared to today’s achievements. If a realistic permeance obtained would be for instance half of that from the reference case, the membrane area required for the same would be doubled if all other design conditions would remain invariant, hence doubling the total cost of the membrane. This is true only if the hydrogen mass transfer across the membrane is the ratelimiting step of the separation process. However, this can be determined only through a detailed membrane model, which is beyond the scope of this work. An analogue situation would result from a system having the reference permeance but with double the price of the membrane. Both of these effects are studied in this section with a sensitivity analysis on the membrane cost. Recent studies reported production costs between 7500 and $10 800/m2;9,26−29 however, uncertainty remains while the technology is not commercial. To consider the effect of the membrane cost on the hydrogen, a sensitivity analysis was conducted, showing the results in terms of the total cost of hydrogen in Figure 9. At $33 560/m2, the total cost of

FINAL REMARKS AND CONCLUSIONS

This work performed a comparative analysis of membranebased configuration for distributed production of hydrogen for filling stations using membrane reactor technology with a conventional system based on steam reformer and PSA. Within the membrane solution, three different layouts were evaluated in the analysis, discussing results obtained for steam reforming membrane reactor, autothermal reforming membrane reactor, and water-gas shift membrane reactor. In this study, the reference characteristics of the permselective membrane were taken from the U.S. DOE target for the hydrogen and fuel cell program. As described above, this target implies an optimistic performance of the membrane as well as for the membrane production process in terms of cost. To take into consideration feasible deviations from that target (at least in the middle term period), a sensitivity analysis was performed in terms of membrane cost. The results show that the SR-MR system could be competitive with the SR-PSA technology with approximately a 6-fold increase in the cost of the membrane, as a result of either a higher price per m2 or a lower permeance of the membrane. Another comment regarding the permeance assumed for the membrane is that the target addressed is set for a membrane designed for the water-gas shift reaction (up to 400 °C). The experience shows that, to achieve the level of permeance defined in such a target, the thickness of the membrane should be around 2−3 μm. However, the membranes designed for operating in a reforming reactor at around 600 °C tend to have slightly thicker layers, not under 4 μm. This means that, if two membranes are prepared by the same technique giving the same permeability, the one with the thicker layer would result in having a lower permeance. On the other hand, a given membrane has a higher permeability at higher temperatures giving the opposite condition. In the comparison between the systems addressed in this study, the permeance of the membrane was taken as constant for both the water-gas shift membrane reactor and the reforming membrane reactors. This approach was defined because we are not addressing a proven membrane with a specified characterization of the permeability at different temperatures. Furthermore, we believe that considering a permeance for the reforming reactor even higher than that in the U.S. DOE target would be an unrealistic assumption. Results show how the SR-MR and ATR-MR configurations can perform at higher thermal efficiency than the conventional system. This can make them very attractive with elevated costs of NG. The SR-MR and ATR-MR configurations also show

Figure 9. Change in the total cost of hydrogen production with membrane price.

hydrogen for the SR-MR solution levels up to that of SR-PSA, and then at $47 860/m2, the capital costs between these two systems becomes equal, leaving the difference to the cost of production. H

dx.doi.org/10.1021/ef301960e | Energy Fuels XXXX, XXX, XXX−XXX

Energy & Fuels

Article

(13) Campanari, S.; Manzolini, G.; Garcia de la Iglesia, F. J. Power Sources 2009, 186, 464−477. (14) McKinsey & Company The Role of Battery Electric Vehicles, Plugin Hybrids and Fuel Cell Electric Vehicles; McKinsey & Company: New York, 2010. (15) Hsu, C. S.; Robinson, P. R. Practical Advances in Petroleum Processing; Hsu, C. S., Robinson, P. R., Eds.; Springer: New York, 2006; Vol. 1, p 866. (16) Kohl, A.; Nielsen, R. Gas Purification, 5th ed.; Gulf Professional Publishing: Houston, TX, 1997. (17) Peng, D. Y.; Robinson, D. B. Ind. Eng. Chem. Fundam. 1976, 15, 59−64. (18) Rackett, H. G. J. Chem. Eng. Data 1970, 15, 514−517. (19) Spencer, C. F.; Danner, R. P. J. Chem. Eng. Data 1972, 17, 236− 241. (20) Haar, L.; Gallagher, J. S.; Kell, J. H. NBS/NRC Steam Tables; Hemisphere Publishing Corporation: Washington, D.C., 1984. (21) U.S. Energy Information Administration, U.S. DOE. U.S. Natural Gas Prices; www.eia.gov/dnav/ng/ng_pri_sum_dcu_nus_m. htm (accessed Sept 10, 2012). (22) U.S. Energy Information Administration, U.S. DOE. U.S. Electricity Data; www.eia.gov/electricity/monthly/epm_table_grapher. cfm?t=epmt_5_3 (accessed Sept 10, 2012). (23) Directed Technologies, Inc. Cost and Performance Comparison of Stationary Hydrogen Fueling Appliances; Directed Technologies, Inc.: Arlington, VA, 2002. (24) Access Intelligence. Chemical Engineering Plant Cost Index (CEPCI) (Annual Index); Access Intelligence: New York, 2013. (25) United States Department of Energy (U.S. DOE). Hydrogen from Coal Program Research, Development, and Demonstration Plan; U.S. DOE: Washington, D.C., 2009. (26) Schwartz, J. II.D.1 Advanced Hydrogen Transport Membranes for Coal Gasification. DOE Hydrogen and Fuel Cell Program. FY 2011 Annual Progress Report; United States Department of Energy (U.S. DOE): Washington, D.C., 2011. (27) Emerson, S. C. II.D.4 Experimental Demonstration of Advanced Palladium Membrane Separators for Central High-Purity Hydrogen Production. DOE Hydrogen and Fuel Cell Program. FY 2009 Annual Progress Report; United States Department of Energy (U.S. DOE): Washington, D.C., 2010; pp 114−118. (28) Manzolini, G.; Gazzani, M.; Turi, D. M.; Macchi, E. Proceedings of the International Conference on Greenhouse Gas Control Technologies (GHGT-11); Kyoto, Japan, Nov 18−22, 2012. (29) Atsonios, K.; Koumanakos, A.; Panopoulos, K. D.; Doukelis, A.; Kakaras, E. Proceedings of the ASME Turbo Expo 2013; San Antonio, TX, June 3−7, 2013.

higher total efficiency (80.2 and 81.0%, respectively) than the SR-PSA system (75.6%); however, because of the higher fraction of electric consumption, the hydrogen production costs equal to 1.33 and $1.41/kgH2 for SR-MR and ATR-MR, respectively, remain higher than the $1.27/kgH2 of the reference case. The study reflects that, with a proper design of the membrane reactor, the capital cost of the system could actually be reduced in comparison to the traditional solution, even considering a specific cost of the perm-selective membranes much higher than the expected threshold of $5382/m2. This reduction in the investment cost takes place thanks to the exclusion of the PSA unit for purification of hydrogen. As a result, the calculation of the total cost of hydrogen considering a 10 year lifetime of the system places the SR-MR solution as more competitive than the SR-PSA, with $4.62/kgH2 rather than $4.73/kgH2, respectively. The ATR-MR and WGS-MR solutions result in a slightly higher cost of $4.78/kgH2. The estimations show that, up to $33 560/m2, the total cost of hydrogen for the SR-PSA and SR-MR systems would be leveled. The use of membrane reactor technology presents a promising outlook in the field of distributed production of hydrogen. Although the solution is both technically and economically feasible, the design of the membrane reactor unit requires careful attention and expertise.

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AUTHOR INFORMATION

Corresponding Author

*Telephone: +39-022399-3817. E-mail: leonardo.roses@mail. polimi.it. Notes

The authors declare no competing financial interest.

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REFERENCES

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dx.doi.org/10.1021/ef301960e | Energy Fuels XXXX, XXX, XXX−XXX