Microreactor Approaches for Liquid Fuel Production from Bioderived

Article pubs.acs.org/IECR

Microreactor Approaches for Liquid Fuel Production from Bioderived Syngas −5 m3/h Prototype Development for HTHP Water Gas Shift Paolo Piermartini* and Peter Pfeifer Institute for Micro Process Engineering (IMVT), Karlsruhe Institute of Technology (KIT), Hermann-von-Helmholtz-Platz 1, 76344 Eggenstein-Leopoldshafen, Germany ABSTRACT: A lab-scale microstructured reactor was used for investigations on enhancing the H2/CO ratio in synthesis gas from biomass feedstocks via the water gas shift reaction at elevated pressure and temperature. A model mixture of carbon monoxide, carbon dioxide, water, and hydrogen was used to reproduce the typical synthesis gas composition from dry biomass gasification. Catalyst powders were prepared and characterized with regard to a simplified catalyst exchange for scale-up; an incipient wetness impregnation of commercial ceria was applied to compare the performance and characteristics to previously investigated catalyst layers produced by a combined incipient wetness impregnation and sol−gel technology. The catalytic activities of these Pt/CeO2 powders were further compared to a commercial high temperature iron water gas shift catalyst at temperatures of 400 to 600 °C and pressures of up to 45 bar. Increased pressure led again to higher values of CO conversion and to increased formation of hydrocarbons (CH4, C2H6, etc.) and coke. Catalyst coatings revealed a higher catalytic activity than the powders due to their higher surface area. A power law model from literature was adapted to fit the performance of the catalyst layers. This model was used to evaluate the scale-up of a microreactor in the scale of 5 m3/h STP throughput of raw synthesis gas and to potentially confirm the idea of a forced falling temperature gradient for the exothermic shift reaction with the selected catalyst coating.

1. INTRODUCTION Biomass is the only renewable energy source containing carbon that can be used in the long term as a feedstock for the generation of chemicals and fuels.1 The most prominent conversion route of biomass is via synthesis gas (“syngas”) by gasification and subsequent fuel synthesis. In this circumstance microstructured reactors may be used in the catalytic reaction steps. They offer unique options for temperature control including user-defined temperature gradients2 and efficient heat transfer3,4 and offer the potential of modular design.5 For applications where an entrained flow gasifier is applied to convert a pyrolysis product from biomass into syngas, the high temperature and high pressure conditions in the gasifier produce a tar-free gas and avoid compression before synthesis. Even when a hot gas cleaning is performed,6 the H2/CO ratio in the syngas is typically too low to meet the requirements of a downstream synthesis plant.7 One way to overcome this issue is to increase the hydrogen content of syngas via the water gas shift (WGS) reaction at gasifier pressure and above the temperature of the synthesis step.8 According to eq 1 the water gas shift reaction is a weakly exothermic, reversible reaction that is nearly independent of pressure: H 2O + CO ↔ H 2 + CO2

ΔHR0 = − 41.3

kJ mol

the conversion at the reactor outlet. However, higher inlet temperature and pressure may also favor methane, carbon, and C2+ species formation in parallel reactions such as methane synthesis (eq 2), coke formation (eq 3), and hydrocarbon formation (eq 4) with consequently a large number of byproducts. An understanding of potential byproduct formation and the optimum temperature profile is thus necessary for catalyst selection and definition of optimal operating conditions. CO + 3H 2 ↔ CH4 + H 2O 2CO(g) ↔ C(s) + CO2(g)

ΔHR0 = − 172.45

CO + 2H 2 ↔ (−CH 2−) + H 2O

kJ mol

kJ mol

ΔHR0 = − 158.5

(2)

(3)

kJ mol (4)

To explore the advantages of a forced temperature profile concept we performed some preliminary calculations with well proven WGS kinetics of a copper catalyst9 which we only modified in the pre-exponential factor to meet the expected conversions. From these results it can be seen that the COconversion may be rate limited at temperatures of about 350 °C (Figure 1); at higher temperatures the reaction rate at the

(1)

As shown in our previous publication,8 a steam content of 28 mol % in the gasification product is sufficient according to equilibrium calculations and experimental data to establish H2/ CO ratios above 2 over a wide temperature range, i.e. below 500 °C. Generally, higher inlet temperatures offer higher initial rates and a forced temperature profile allows further increase of © XXXX American Chemical Society

ΔHR0 = − 205.8

Special Issue: Recent Advances in Natural Gas Conversion Received: July 31, 2013 Revised: October 15, 2013 Accepted: October 17, 2013

A

dx.doi.org/10.1021/ie402491c | Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

Industrial & Engineering Chemistry Research

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Figure 1. CO conversion as function of reactor length at two different isothermal conditions and with a linearly decreasing temperature profile at 1 bar, reactor length and feed composition according to previous investigations,8 i.e. 32 vol.% CO, 10 vol.% CO2, 18 vol.% H2, 10 vol.% N2, and 30 vol. % H2O; plug flow conditions assumed; kinetics according to shift reaction adapted from reverse water gas shift literature9 by arbitrary lowering of the pre-exponential factor by 1000 to lie in the desired conversion range.

Noble metal catalysts have already been incorporated in microreactors and applied to the WGS reaction in literature, but these microreactor systems have been exclusively applied to ambient pressures for fuel cell applications.18,19 These studies showed that the Pt/CeO2/Al2O3 catalyst is one of the best candidates with respect to selectivity and activity. Generally, the catalysts showed a high CO2-productivity.19 In a previous study20 Ru catalysts supported on ZrO2 coated microstructured foils were applied to the water gas shift reaction at temperatures of 250−300 °C and at ambient pressure. Neither the typical low-temperature catalysts such as copper−zinc or rutheniumzirconia nor the iron-chrome catalysts seemed to be adequate for our selected temperature range (400−600 °C). Only the Pt/CeO2/Al2O3-based catalyst promised a remarkable activity at the desired temperatures17 for the starting point of our previous study.8 This led to the development of a highly active Pt/CeO2 coating prepared via a combined ceria sol−gel and Pt impregnation method with productivity in the range of 25 g/ (g*h) of carbon monoxide conversion on catalyst mass (platinum+ceria) basis. The effects of pressure, temperature, and residence time on the product distribution were tested, and particular attention was given to the selectivity of secondary reactions and byproduct generation. The main byproducts were methane and coke; coke was identified as a result of the reaction of the carbon monoxide rich feed in the inlet of the reactor due to welding material. In the current publication we focus on a comparison of the Pt/CeO2 coating8 with a conventional Pt/CeO2 catalyst powder and a commercial iron catalyst, packed in different microstructure geometry and the selection between those for the scale-up to a reactor in the size range 5 m3/h of raw syngas feed. By applying an adapted power law rate equation, we fitted the reaction data on the coatings and compared this to the scale-up design, which was done by assuming a forced falling temperature profile.

entrance of a reactor is high but the CO conversion is limited due to thermodynamic equilibrium.10 If a decreasing temperature profile along the reactor length is defined, the calculations indicate that the overall conversion at the reactor outlet can be increased. With the knowledge that a linearly decreasing temperature profile of up to 150 K can be defined for a microreactor within only 8 cm reactor length,11 this may allow the assumption that a microreactor with a decreasing temperature profile could be part of a compact and easy-to-operate reactor concept for water gas shift reaction. In the literature, many catalysts have been reported to facilitate the water gas shift reaction at different temperatures. For the low temperature WGS reaction (200−300 °C) copper−zinc catalysts are normally used, but novel catalytic materials have also been investigated. For example, gold nanoparticles supported on TiO2 and CeO2 have been successfully used over a wide range of temperatures, between 50 and 400 °C.12 Au/TiO2 exhibits good CO-conversion at temperatures below 300 °C, while Au/CeO2 was most active at temperatures above 300 °C. Furthermore, coprecipitated bimetallic gold catalysts have been studied. Gold catalysts containing ceria as support showed a good activity in the low temperature range of 200−350 °C, especially an Au−Pt/ceria catalyst.13 An increase in the conversion of the carbon monoxide and a reduction of surface oxygen at lower temperatures was observed for coprecipitation of gold and metal catalysts.13 Iron catalysts have conventionally been applied for the high temperature WGS reaction (400−500 °C) and have recently been improved.14−16 Noble metals (Rh, Pt, Pd, and Ru) supported on CeO2−Al2O3 were investigated for the water gas shift at relatively high temperatures (500−700 °C) in order to find an active catalyst that could be directly coupled with the biomass gasification process.17 The Pt-CeO2/Al2O3 catalyst showed the highest CO-conversion at temperatures of 700 °C, and the ceria support seemed to improve the catalyst performance. B



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2. EXPERIMENTAL SECTION 2.1. Microreactor. The lab reactor was built of Nicrofer 3220H/Alloy 800H. This material was chosen for its satisfying pressure resistance at temperatures of up to 600 °C and low coke formation compared to standard stainless steel.8 Special graphite seals were used to prevent leaks at high pressure under the low total flow rates applied to the reactor. Two sides of the reactor could be opened, one for the foil integration (i.e., a full opening) and one side for inserting powder catalyst from the top while having a fixed foil stack. The housing of the reactor is 225 mm long and fits microstructured foils with 25 mm width. The high length of the channels/slits was chosen to avoid back mixing phenomena, i.e. axial mass transport, which could occur at the low gas velocities used in these investigations. In the case of the catalyst coating fourteen stainless steel foils structured with microchannels (Figure 2) were successively

Figure 3. Reactor arrangement.

Table 1. Geometry Data of the Microstructured Foil Stacks Figure 2. Photo of one foil with pillar structure (left) and catalyst coated microchannels (right).

stacked in the reactor. Figure 3 provides a description of the arrangement of the microstructured foils in the reactor. The dimensions of the microchannels are given in Table 1. The geometric surface area is the sum of the channel bottom, the side channel walls, and the top of the fin. This is the area coated by the catalyst. The dimensions of the microstructured foils for the fixed bed investigations are also presented in Table 1 and Figure 2. One pair of foils with micropillars are stacked face-toface with the pillar side into the reactor for powder packing; when they were placed inside the reactor and the stack was pressed from the back side the powder was filled from the foil entrance thanks to a movable flange (Figure 3). 2.2. Catalyst Preparation. The foils made from 1.4404/ 316L (stainless steel) were first cleaned in water as well as isopropyl alcohol for 30 min in an ultrasonic bath and then annealed in air at 500 °C for 5 h. Annealing stainless steel foils generates a chromium oxide layer, which improves the adhesion of the oxide catalyst to the foil surface20 and also avoids formation of carbon on the stainless steel foils; this has been observed in the case of the micropacked bed where no catalyst support prevents the contact to the steel. Pt(5 wt.%)/CeO2 coated catalyst was freshly prepared for the investigations in variation of the CO/H2O ratio through the

geometric dimensions of the microstructured foils

micro channel

micro pillars

channel width/pillar distance (μm) channel/pillars height (μm) fin width/pillars diameter (μm) channels/pillars per foil length of foil (mm) width of foil (mm) geometric surface area of one foil (mm2) number of foils height of stack (mm) channels/pillars volume of a foil (mm3)

200 200 100 50 150 25 5235 14 4.2 300

800 400 800 1029 150 25 2171.5 2 0.8 650

combined sol−gel process and incipient wetness impregnation, described in our previous publication.8 Manufacturing of the Pt/CeO2 powder catalyst occurred through a cerium oxide powder with a level of purity of 96% (Alfa Aesar, CAS 1306-38-3). Five g of the 5 wt.% Pt/95 wt.% CeO2 catalyst were prepared through impregnation of ca. 4.75 g of cerium oxide powder with a 0.02 mol/L platinum nitrate aqueous solution. The solution was mixed and heated in a flask under continuous stirring, and vacuum (200−300 mbar) was applied to slowly evaporate the solvent at temperatures about 70−100 °C. The complete impregnated powder was carefully recovered from the wall of the flask, dried for 24 h at 80 °C, and calcined at 500 °C for 5 h. After the calcination the particles were crushed in a mortar and sieved. A particle fraction with a diameter between 50 and 100 μm was chosen for the reaction testing in order to keep the 1:10 ratio between particle C



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mass fraction of catalyst powder (see section 2.2) in order to increase τmod.. For both powders, Pt/CeO2 and Fe catalyst, the catalyst mass thus was 0.9 g and τmod. was always 118000 g*s/ m3 (STP). For the experimental basis to fit the power law rate equation the CO to H2O ratio was further varied from 0.5 to 1.25 with an overall constant flow rate and constant component concentrations of CO2, H2, and N2 on the catalyst coating.

diameter and dimensions of the microstructures. The obtained apparent density was between 1 and 2 g/cm3. To establish a reproducible catalyst mass in the micropacked bed design and to compare catalyst activity that is not too far different in the overall active metal surface area of Pt, a mixture with pure ceria powder (weight basis) was produced, and approximately 2 g of the powder mixture was employed into the microstructure; the Pt/CeO2 catalyst amount (metal+support) in this mixture was 0.9 g. The overall powder amount (2 g) was proven to fill the applied one pair of foils in filling tests. The filling tests with pure ceria powder were also used to check negligible blind activity of the original ceria. A commercial high temperature iron catalyst was also obtained in the same size range of 50 and 100 μm diameter by sieving. The total amount of catalyst powder introduced into the one foil pair was 0.9 g due to the lower packing density. No inert powder was added here. Before the catalytic testing all catalysts, except the iron system, were reduced in situ in 1 L/min STP flow of 5 vol.% hydrogen/nitrogen mixture for 4 h at 500 °C. For the iron system a hydrogen atmosphere was avoided due to the possible formation of Fe(0). 2.3. Catalyst Characterization. Sorption measurements were conducted in an Autosorb 1C machine from Quantachrome. The total catalyst surface area was determined through nitrogen adsorption using the BET method for the calculation of the specific (metal + support) surface area (SBET). The active metal surface area per mass or volume of catalyst (metal + support) and the dispersion of the active metal particles on the surface were determined from chemical adsorption measurement in the Autosorb 1C machine, considering the active metal loading for each catalyst. The gases used for the adsorption on platinum and iron were hydrogen and ammonia, respectively. The chemisorption was carried out at 40 °C (hydrogen) and 100 °C (ammonia). A stoichiometry of 2 was used for the hydrogen (dissociative adsorption of H2 on the metal), while a stoichiometry of 1 was chosen for the ammonia adsorption. All the samples were outgassed at 120 °C prior to measurement. The platinum samples were furthermore reduced with H2 at 400 °C. The “Extrapolation to P=0 Method” was used for all samples applying the data points from strong adsorption. A pressure between 40 and 300 mTorr hydrogen was used for the isotherm. 2.4. Reaction Conditions. The following experimental conditions were used for investigating the water gas shift reaction in the microstructured lab reactor: Each catalyst was tested for more than 100 h without observable deactivation (see also ref 8). A standard feed gas of 32 vol.% CO, 10 vol.% CO2, 18 vol.% H2, 10 vol.% N2, and 30 vol.% H2O was applied to the catalysts at a modified residence time for the catalyst coatings of τmod. 26200 g*s/m3 (STP), see eq 5. For an overall catalyst coating mass (Pt+ceria) of 0.2 g distributed on 14 microchannel foils this τmod. corresponds to a total gas flow of 0.46 L/min STP. The reactor wall temperature ranged from 400 to 550 °C, and the pressure was increased from 1 bar (atmospheric conditions) up to 45 bar. For the catalyst packing the same overall flow rate was applied. Since a higher powder mass is easily packed by filling of one slit obtained from a face-to-face arrangement of two foils with pillars, the catalyst powder was diluted with inert ceria in the experiments. From preliminary experiments (not reported) we explored a significant lower catalytic activity and increased the

τmod. =

mcat. catalyst mass = ̇ V(STP) volume flow rate(STP)

(5)

From the measured inlet concentration [CO]in and end concentration [CO]end, the conversion of carbon monoxide X(CO) could be determined: X(CO) =

[CO]in − [CO]end [CO]in

(6)

From the measured outlet concentrations of all the products of the water gas shift reaction, the concentration ratio K can be defined: K=

[CO2 ]vol ∗[H 2]vol [CO]vol ∗[H 2O]vol

(7)

For the determination of the WGS equilibrium constant Kp (additionally applied in the graphs) the program HSCChemistry 5.0 was applied. The program minimizes the Gibbs free energy for a set of species with given initial composition of these species. This initial composition corresponds to the measured outlet gas composition in the experiments. Nitrogen and all byproducts were considered as inert in these calculations (allowed species were H2O, H2, CO, and CO2) to judge only the shift activity of the catalysts. Dividing the mole flow of converted CO ΔṅCO by the overall molecules of active platinum sites, the mean turnover frequency TOF (s−1) can be defined: TOF =

ΔnCO ̇ nactivePt

(8) 8

The test setup is described in the previous publication, and isothermal conditions inside the microreactor have been validated previously.

3. EXPERIMENTAL RESULTS AND DISCUSSION 3.1. Results from Pt/CeO2 Powder Catalyst. As mentioned in section 2.2 the evaluation of the Pt/CeO2 powder was performed under a comparable ratio of total flow to overall active metal surface area as in the previous study. The physical characterization will be shown in the comparison to the catalyst layers. Figure 4 compares the equilibrium constant Kp of the shift reaction with the experimental concentration ratio K for the different pressures 1, 15, and 45 bar as a function of wall temperature of the reactor. It is obvious that the higher pressure and a higher temperature allow higher conversion of CO to form hydrogen. This is not due to any influence of the thermodynamics but due to higher pressure of the reactants which have a more severe influence than the products. When comparing the K values to CO conversion it can be stated that the highest K represents approximately 30% conversion; this is 2/3rd the value which should be reached according to the equilibrium conditions. The equilibrium was never reached in these experiments. D



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powder manufacture with high surface area was not successful. The precursor for ceria is an ammonium nitrate compound, Ce(NH4)2(NO3)6. Larger amounts of gel get converted explosively into the oxide powder during calcination. Thus such preparation was currently ruled out for ceria powder synthesis. As the total amount of active Pt surface area was not exactly the same in the comparison between the powder and the coated catalyst, the mean turnover frequency (TOF) of CO over the catalyst bed (see eq 8) was calculated for the different pressures and temperatures for both systems. The value of TOF increases up to 6 s−1 for reaction temperatures of about 550 °C at a pressure of 45 bar (Figure 5). Good agreement among the Figure 4. Experimental concentration ratio K at the reactor outlet compared to the WGS equilibrium constant as function of the reactor wall temperature for the Pt/CeO2 powder catalyst; different pressures (⧫ = 1 bar, ■ = 15 bar, ▲ = 45 bar) and 1.18E+05 s*g/m3 STP, inlet composition 32 vol.% CO, 10 vol.% CO2, 18 vol.% H2, 10 vol.% N2, 30 vol.% H2O.

The analysis of the gaseous byproduct species according to Table 2 at the highest pressure of 45 bar shows that the amount Table 2. Methane and C2+ Concentration on the Pt/CeO2 Powder Catalyst at Different Conditions As a Function of the CO Conversiona pressure 45 bar

a

temp 400 450 500 550

°C °C °C °C

CH4/vol.%

ΣC2+/vol.%

CO conversion/%

0.06 0.15 0.06 0.4

0.68 0.58 -

2 13 29 36

Figure 5. Turnover frequency (TOF) as a function of the reactor wall temperature for Pt/CeO2 coating (closed symbols) and Pt/CeO2 powder (empty symbols) at different pressure (diamonds 1 bar, triangles 45 bar); overall gas flow of 0.46 L/min STP, gas composition 32 vol.% CO, 10 vol.% CO2, 18 vol.% H2, 10 vol.% N2, 30 vol.% H2O.

τmod. = 1.18E+05 s*g/m3 STP.

values of the TOF between the coating and the powder exists only for the higher pressures (shown explicitly for 45 bar). However, for 1 bar the turnover frequency of the powder catalyst is only 40% of the TOF of the coating (Table 4). Comparing also the ratio of the TOF values at 15 bar, which is between those of 1 and 45 bar, we believe that this lowering effect is caused by some bypass inside the reactor which is induced by the higher relative pressure drop of the packed bed. As the flow velocity is low, laminar conditions can be assumed; in such circumstances the relative pressure drop is only

of methane and C2+ hydrocarbons is negligible. At the lowest temperature there is a tendency to form C2+, whereas methane is more prominent at higher temperature. 3.2. Discussion on Differences between Pt/CeO2 Powder and Pt/CeO2 Coating. As stated in our previous publication,8 the catalyst coating easily reached the equilibrium conversion under the applied total flow rate at 45 bar. This better performance can be explained by a two times higher overall Pt surface area loaded into the reactor in the case of the coating. In general, the coating exhibits a much better Pt utilization according to Table 3. Pt dispersion is much higher mainly due to the superior BET surface area of the coated support which provides much better initial condition for the formation of small Pt crystals than the commercial ceria powder. However, a trial applying the same sol−gel route for

Table 4. Turnover Frequency (TOF) and Ratio α for Pt/ CeO2 Coating and for Pt/CeO2 Powder at Different Pressure and Reactor Temperature for a Total Flow Rate of 0.46 L/min STP temp [°C]

Table 3. Comparison of Physical Data from Pt/CeO2 Coating and Pt/CeO2 Powder Pt/CeO2 coating BET surface area (m2/g) volume specific BET surface area (m2/m3) active metal surface area (m2/g) vol.-spec. active metal surface area (m2/m3) metal dispersion (%)

pressure [bar] TOFcoating [s−1]

Pt/CeO2 powder

58 3 × 1008

1.30 1.06 × 1008

2,171 1.1 × 1007

0.308 2.5 × 1007

17.6

2.5

TOFpowder [s−1]

α = TOFpowder/ TOFcoating

E

1 15 45 1 15 45 1 15 45

400 °C 450 °C 500 °C 550 °C 1.85 2.19 2.19 0.43 1.49 3.02 0.23 0.68 1.38

2.85 4.05 4.52 0.97 2.09 4.52 0.34 0.52 1.00

3.65 4.93 5.44 1.59 2.54 5.69 0.44 0.52 1.05

3.84 4.93 6.06 2.23 4.67 6.35 0.58 0.95 1.05



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commercial iron powder catalyst and the Pt powder catalyst. Even though different adsorption methods were used for the determination of the active metal surface area, it may be stated that a much smaller site density is obvious for the noble-metal free system. Considering the low site density and the minor initial CO conversion at temperatures about 400 °C together with the knowledge that the catalyst was prepared for high temperature range, the temperature window was expanded in the catalytic tests to 600 °C maximum. Higher temperatures could not be used due to the graphite sealing. Figure 6 compares the equilibrium constant of the shift reaction with the experimental concentration ratio K on the

dependent on the gas velocity and independent of the gas density according to the Ergun equation. Thus, the higher the pressure the lower the relative pressure drop will be (calculated values of pressure drop for the 15 cm long packed bed range from 600 mbar at 1 bar to 10 mbar at 45 bar at an assumed bed porosity of 60%). In consequence this may reduce the importance of fabrication tolerances between the foil outer geometry and reactor cavity which embeds the foil and a lower bypass around the packed bed inside or outside the pressed foil stack may occur. At the lowest pressure the average bypass should consequently be in the order of 60% of the total flow to explain the difference in TOF. Microchannel arrangements can be assumed to operate at near-zero bypass as the overall pressure drop is generally 2 orders of magnitude lower than in the micropacked bed, and this practically rules out the importance of small gaps between reactor and foils (for comparison: calculated values of the pressure drop for the 15 cm long microchannels are around 10 mbar at 1 bar to 0.2 mbar at 45 bar). However, one contradiction still exists: At the lowest pressure the bypass in the packed bed increases with decreasing reaction temperature. A slightly increasing viscosity of gas with increasing temperature and the proportional dependence on the viscosity in the Ergun equation in the laminar flow regime would explain a higher deviation at high temperature. However, the opposite effect is observed. Hence other effects can also be assumed. The rate increase due to strong interaction of the ceria support with platinum8 may also decline with decreasing Pt dispersion and increasing dimension of the platinum particles for the catalyst powder, due to a larger number of Pt−Pt bonds. This leads to an inhibited interaction between the ceria support and the platinum metal and consequently to lower carbon monoxide conversions of the active sites at low pressure and/or temperature. 3.3. Results from Commercial Iron Catalyst. The applied iron catalyst is based on pure Fe2O3 (iron(III) oxide)/Fe3O4 (iron(II−III) oxide) as determined from X-ray diffraction (not shown). A slight tendency of phase change was observed between the two species before and after reaction; the Fe3O4 content slightly increased and the peak height to width ratio was reduced after the reaction tests. This may indicate that a slight sintering occurred and the catalyst became more reduced. However, these data may also be influenced by oxidation upon air exposure while unloading the catalyst and ex situ analysis. Table 5 shows a general comparison of the

Figure 6. Experimental concentration ratio K at the reactor outlet compared to the WGS equilibrium constant as function of the reactor wall temperature for the Fe2O3/Fe3O4 powder catalyst; different pressures (⧫ = 1 bar, ■ = 15 bar, ▲ = 45 bar) and 1.18E+05 s*g/m3 STP, inlet composition 32 vol.% CO, 10 vol.% CO2, 18 vol.% H2, 10 vol.% N2, 30 vol.% H2O.

iron catalyst for the different pressures 1, 15, and 45 bar as a function of wall temperature of the reactor. The experimental data approach gradually the equilibrium for increasing temperature, but the pressure effect is negligible for this catalyst. This may lead to the assumption that either the concentration dependence of the reactants is zero or the products inhibit a higher conversion at higher partial pressures. The formation of methane and C2+ hydrocarbons was even lower as compared to the Pt/CeO2 catalyst which may lead to the assumption that the iron catalyst is favorable in terms of selectivity. However, the total conversions are different from the platinum catalysts. The byproduct formation could also increase when the conversion increases; however, this comparison would only possible by further increase of residence time. 3.4. Choice between Catalysts for Scale-Up. The activity results are shown in Figure 7 for the three different catalyst types at a pressure of 45 bar; the figure again compares the equilibrium constant of the shift reaction with the experimental concentration ratio K. As can be seen, the catalytic coating clearly outperforms the powder catalysts in terms of adjusting the concentration ratio toward the desired equilibrium, equivalent to a H2/CO ratio of 2. It means that a 5 times higher catalyst mass is needed in the case of the Pt powder system to obtain the plotted curve; for the iron catalyst it is the same increase. As mentioned in section 2.4 each catalyst was tested for more than 100 h, and in this time no deactivation by sintering or coking of the catalysts was observed (see also ref 8).

Table 5. Comparison of Physical Data from Pt/CeO2 Powder and Commercial Iron Catalyst (No Prereduction before Ammonia Sorption Measurement)a Pt/CeO2 powder

Fe2O3/Fe3O4 catalyst

BET surface area (m2/g) volume specific BET surface area (m2/m3) active metal surface area (m2/g)

1.30 1.06 × 1008

73.7 2.58 × 1008

0.308 (H2)

vol.-spec. active metal surface area (m2/m3) metal dispersion (%)

2.5 × 1007 (H2) 2.5 (H2)

7.04 × 10−03 (NH3) 2.46 × 1004 (NH3)

mean pore diameter (nm)

2.18

1.06 × 10−3 (NH3) 12.4

a

Mean pore diameter calculated according to the Barret-JoynerHalenda method. F



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Figure 8. Experimental and simulated concentration ratio K at the reactor outlet compared to the WGS equilibrium constant as function of the reactor wall temperature for the Pt/CeO2 coating; feed flow 0.46 L/min STP at CO/H2O ratio of 1.25 and 1 bar.

Figure 7. Experimental concentration ratio K at the reactor outlet compared to the WGS equilibrium constant as function of the reactor wall temperature for the Pt/CeO2 coating (■) with τmod. = 26200 g*s/m3 (STP), the Pt/CeO2 powder (□) and the Fe2O3/Fe3O4 powder (■)with τmod. = 118000 g*s/m3 (STP) at 45 bar; constant overall throughput of 0.46 L/min STP, concentrations 32 vol.% CO, 10 vol.% CO2, 18 vol.% H2, 10 vol.% N2, 30 vol.% H2O.

Discussing the different catalyst results with regard to the choice for a scale-up, different arguments were raised: On the one hand exchange of catalyst is feasible with powder, but on the other hand a much higher amount of Pt metal is currently needed for the desired conversion when comparing the Pt powder catalyst with the coating results; byproduct formation is comparable between the different catalysts at least when taking into account the overall higher CO conversion on the coated catalyst. As we suppose that a multiple coating would be feasible when deactivation occurs, we decided for the smaller amount of Pt and thus the coating for conducting further work. 3.5. Model for Reactor Scale-up with Pt/CeO2 Coating. Based on the variation of the CO/H2O concentration in the range 0.5 to 1.25, a plug flow model with 100 intervals for the integration of the conversion was used to fit a power law rate model. The applied kinetic model and the initial constants originated from investigations on a Pt/CeO2/Al2O3 catalyst.21 By simple manual modification of the pre-exponential factor and the activation energy, an engineering description for the performance of the Pt/CeO2 coating was derived. The final model is shown in eq 9 for the reaction rate r, together with a reversibility factor β in eq 10: This simple approach, without analysis of statistical evaluation, allowed describing the experimental values quite well (Figures 8 and 9).

Figure 9. Experimental and simulated concentration ratio K at the reactor outlet compared to the WGS equilibrium constant as function of the reactor wall temperature for the Pt/CeO2 coating; feed flow 0.46 L/min STP at CO/H2O ratio of 0.5 and 1 bar.

4. SCALE-UP TO 5 M3/H THROUGHPUT 4.1. Considerations for Scale-up. Several considerations have been done within the scale-up into the range of 5 m3/h throughput. Some of the issues arose from the experimental results, and some were subject of manufacture compromises. The following list should provide an overview over the steps performed: • Adjusted stacking: Since the overall heat transfer capabilities of the microchannel reactors is usually much greater than required in an alternating stack of plates for reaction and cooling, a number of plates for reaction can be stacked above each other without intermediate cooling. In this case the heat flux through the stack is determined by the channel walls or more specifically by the fins between the channels on the individual plates. According to the solution of the heat balance by a porous body approach,23 10 coated catalyst plates were stacked in between the cooling plates for the technical scale (Figure 10). • Required inspection of the catalyst: Flanges were proposed on the inlet and outlet of the reaction gas for inspection and flow coating issues. • Optimized pressure drop and temperature profile: In order to establish a forced falling temperature profile for optimization of the local kinetics inside the reactor2 a counter current cooling with air flow was designed for the first prototype (steam cooling is envisaged later on). In order to meet the

⎛ −83.5*103J/mole ⎞ rCO = 1*105mole/(kg*s) exp⎜ ⎟ RT ⎝ ⎠ 0.13 0.49 −0.45 −0.12 × pCO pH O pH pCO (1 − β) 2

2

2

(9)

with β=

pCO pH 2

2

KeqpCO pH O 2

(10)

Keq is the equilibrium constant for the water gas shift, which is calculated in this case through a correlation from the literature:22 ⎛ 4577.8 ⎞ Keq = exp⎜ − 4.33⎟ ⎝ T ⎠

(11) G



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• Reduction of dead volume: Since the residence time in the inlet of the lab reactor also influenced the coke formation, the inlet volume of the reactor was kept minimal. In order to ensure no flow maldistribution a CFD study was conducted while using mean values for the properties of the reactant flow at the desired operation conditions range. Figure 11 shows the contour plot of the velocity magnitude inside the reactor, and the homogeneous color in the channel region indicates that both functionalities were combined successfully. • Pressure resistance: The body of the reactor, i.e. the welding seam, the wall thickness of the flanges, and the flow chambers formed from the diffusion bonding of the designed plates for the cooling side were dimensioned by help of finite element methods. 50 bar and 550 °C was the defined operation condition. • Material and construction suitable for diffusion bonding: Not only the material decides for the diffusion bonding properties but also the geometry of the channels has to be considered. This has been included in the design, and, while trying to increase the volume specific surface area of the reactor core, the best compromise for the etched channel geometry was chosen to be 300 × 150 μm (width × height). • Safety regulations: Regarding the fact that diffusion welding is not yet a validated method with regard to pressure certification, a pressure vessel was constructed around the reactor. Together with the specification of the reaction volume enclosed in the reactor plus the connection tubing (toward a safety valve), the pressure limit of the surrounding vessel was set to 8 bar taking into account the expansion. The reactor is shown in Figure 12. With regard to the flow coating, experiments with test stacks of identical structure are currently conducted. These stacks consist of 10 foils with 67 channels per foil. Figure 13 shows test stacks in the as-fabricated stage and after the coating procedure as well as a picture of the front inlets of a coated stack lighted from the back side. This stack has been coated six times with an amount of 2.5 mL sol−gel solution per step. This gives an overall ceria deposition of 100 mg. The intended catalyst amount is 200 mg taking into account the desired catalyst amount per wall surface area. As can be seen from Figure 13b, there are small ceria residues located on some channel outlets, but the predominant number of channels are free for flow. A detailed investigation of the catalyst distribution is currently performed by help of hot-wire anemometry24 and cross cuts at different position. However, this will be detailed in another study. 4.2. Evaluation of the Scale-up. With regard to the initial idea of a forced falling temperature profile the chosen design can be evaluated by help of a rate plot with the derived kinetic equation and the ability of the reactor design for process intensification can be checked. Figure 14 shows such rate plot as a function of CO conversion and local reactor temperature in a classical way. The regions of constant rate are highlighted in different color and separated by lines of constant rate. With a specified inlet temperature of 530 °C the ideal situation would be an immediate temperature jump to much higher temperature (this is technically not feasible); thus a more practical “optimum” may consist of the attempt to keep the initial rate constant as long as possible (tangent trajectory to 20 mol/kg*s rate line). At a conversion of around 35% at 580 °C the temperature should be lowered in the following according to the lines of constant rate to follow the path

Figure 10. Stacking scheme of the 5 m3/h prototype reactor; straight channels represent the reaction zone, curved channels are dedicated to cooling.

requirement of low pressure drop of the cooling side, a face-toface arrangement of two cooling plates was projected. This allows the realization of larger channel diameters with the chosen microfabrication technique (see also Figure 10). • Changed channel geometry: Since it was decided to apply the etching technique for the large number of channels (roughly 100,000 channels on the reactive gas side) with regard to low fabrication time (milling is not yet available with multiple tools in our laboratories) the required geometric surface of the channels and thus the number of channels was determined via the surface specific catalyst loading which can be realized with the sol−gel technique. • Optimized welding material: In the previous investigation8 coke formation was mainly observed in the lab reactor on the welding seam and the stainless steel frit inside the reactor. The reactor material (alloy 800H) was not subject of coke formation. In order to solve this problem, a welding material made from DIN 1.4850 (33% Ni, 21% Cr, 4.5% Mn, 1.2% Nb, balance iron) was applied. H



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Figure 11. Velocity contours for the reactive gases inside the 5 m3/h-scale prototype reactor.

Figure 12. Technical-scale reactor prototype for 5 m3/h located on top of a dished boiler end for integration into the pressure vessel.

d r (X ) =0 dT

(12)

Figure 13. Photo of a test stack for flow coating with catalyst a) in the status as-fabricated (left) and after coating (right) as well as b) a coated stack with a spot light through the channels.

An adiabatic rise at the reactor entrance would provide a higher space time yield as the trajectory points into regions of even higher reaction rate, so a cooled system providing this “optimum” path seems feasible in terms of heat management. However, as the limit of coke formation in the specific catalystreactor design is set to 550 °C, the actually manufactured reactor with counter current cooling of the reaction will be a

compromise. The continuous line represents the educated guess of the temperature-conversion plot which is likely to occur inside the 5 m3/h prototype. However, such a process may, under the aspects of the temperature limit, still provide a I



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gas from entrained-flow gasification of pyrolysis product from solid biomass. The comparison of the different catalyst implementation techniques and catalyst types in lab investigations was used to choose the scale-up concept. The different considerations for establishing ideal situations in the prototype reactor are explained, and the final hardware is shown. Recent results from catalyst coating experiments with test pieces of identical microstructure as used in the technical-scale reactor are presented. A more detailed study will follow concerning the coating analysis. With regard to the idea of a forced temperature profile for reaction rate optimization at each point inside the reactor the design shows compromises due to the position of the lines of constant rate and as there were set temperature limits to avoid coke formation. In sum, however, the overall concept seems advantageous compared to the abilities of standard reactor equipment to establish better reaction pathways.

Figure 14. CO conversion − temperature plot with embedded lines of constant rate according the kinetics in section 3.5 (values at lines in unit mol/kg*s), with temperature limit to prevent coke formation, with “optimum” trajectory for optimized rate as function of conversion, trajectory for adiabatic temperature rise from the inlet reaction gas temperature and an educated guess of the most likely trajectory in the manufactured prototype.

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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higher space-time yield than if it is run in conventional equipment in undefined polytropic mode. In order to analyze the remaining advantage in more detail, a plot (Figure 15) similar to Figure 1 was generated by applying the derived kinetics obtained in section 3.5. The linear decrease from 530 to 480 °C provides still superior conversion over isothermal conditions at the individual temperatures. However, since there is a drastic reduction of the rate with decreasing temperature (see also conversion for isothermal operation at 350 °C), the effect of the temperature profile is not as pronounced as assumed from the preliminary calculations shown in Figure 1.

ACKNOWLEDGMENTS The financial support from the Baden-Württemberg Foundation under the acronym Bio11/ProShift and from KIC InnoEnergy for the project SynCon is gratefully acknowledged.

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REFERENCES

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5. CONCLUSION With regard to the development of a technical-scale reactor based on microchannel technology, the individual steps of the development are explained. A scale-up of a WGS unit was considered for the increase of the H2/CO ratio of raw synthesis

Figure 15. CO conversion as function of reactor length at four different isothermal conditions and with a linearly decreasing temperature profile at 45 bar for the Pt/CeO2 catalyst coating according to the derived kinetics in section 3.5; reactor length and feed composition corresponding to Figure 1, i.e. 32 vol.% CO, 10 vol.% CO2, 18 vol.% H2, 10 vol.% N2, and 30 vol.% H2O; plug flow conditions assumed. J



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