Progress toward Biomass and Coal-Derived Syngas Warm Cleanup

Article pubs.acs.org/IECR

Progress toward Biomass and Coal-Derived Syngas Warm Cleanup: Proof-of-Concept Process Demonstration of Multicontaminant Removal for Biomass Application Christopher J. Howard, Robert A. Dagle,* Vanessa M. Lebarbier, James E. Rainbolt, Liyu Li, and Dave L. King* Energy and Environmental Directorate, Institute for Integrated Catalysis, Pacific Northwest National Laboratory, Richland, Washington 99352, United States ABSTRACT: Systems comprising multiple sorbent and catalytic beds have been developed for the warm cleanup of coal- and biomass-derived syngas. Tailored specifically for biomass application, the process described here consists of six primary unit operations: (1) a Na2CO3 bed for HCl removal, (2) two regenerable ZnO beds in parallel for bulk H2S removal, (3) a ZnO bed for H2S polishing, (4) a NiCu/SBA-16 sorbent for trace metal (e.g., AsH3) removal, (5) a steam reforming catalyst bed for tars and light hydrocarbon reformation and NH3 decomposition, and (6) a Cu-based LT-WGS catalyst bed. Simulated biomassderived syngas containing possible inorganic contaminants (H2S, AsH3, HCl, and NH3) and hydrocarbons (methane, ethylene, benzene, and naphthalene) was used to demonstrate process effectiveness. The efficiency of the process was demonstrated for a period of 175 h, during which time no signs of deactivation were observed. However, postrun analysis revealed that small levels of sulfur slipped through the sorbent bed train to the two downstream catalytic beds. Future improvements will be made to the trace metal polishing sorbent to ensure complete inorganic contaminant removal (to low parts per billion level) prior to the catalytic steps. However, dual regenerating ZnO beds were effective for continuous removal for the vast majority of the sulfur present in the feed gas. The process was effective for complete AsH3 and HCl removal. The steam reforming catalyst completely reformed all the hydrocarbons present in the feed (methane, ethylene, benzene, and naphthalene) to additional syngas. However, postrun evaluation, under kinetically controlled conditions, indicates some deactivation of the steam reforming catalyst occurred. Spent catalyst characterization suggests this can be attributed, in part, to coke formation, likely due to the presence of benzene and/or naphthalene in the feed. Future adaptation of this technology may require dual, regenerable steam reformers. The process and materials described in this report hold promise for the warm cleanup of a variety of contaminant species within warm syngas.

1.0. INTRODUCTION Gasifier-derived syngas from both biomass1 and coal2 has many applications in the area of catalytic transformation to fuels and chemicals. A significant amount of research has recently taken place examining the feasibility of producing biomass-derived synthesis gas (syngas) for subsequent conversion to mixed alcohols3,4 or other fuels.1 In all cases, the gasifier-derived syngas must be treated to remove a number of impurities which would otherwise poison the processing catalysts. Inorganic impurities include alkali salts, chloride, sulfur compounds, heavy metals, ammonia, and various P, As, Sb, and Secontaining compounds.5 Many of these must be removed to part per billion levels due to their strong, detrimental interaction with downstream water−gas shift and synthesis catalysts. Product gas from a biomass gasifier also contains small quantities of hydrocarbon gases, such as methane and ethane, and organic liquids broadly classified as tars. These tars (e.g., aromatic hydrocarbons) are notorious for condensing and subsequently polymerizing on downstream equipment such as compressor and gas turbine surfaces. Tars and hydrocarbon gases also contribute to significant carbon deposition on water−gas shift catalyst, which is necessary for adjusting the raw gas composition to a synthesis gas suitable for the subsequent synthesis reaction. Tars and hydrocarbons also contain a significant fraction of the biomass carbon that could be converted into additional syngas. © XXXX American Chemical Society

Liquid phase gas cleaning approaches currently exist for removal of the contaminant species and are proven technologies.6 However, condensation is required for removal of the tars and water in the gas. Solvents are required at ambient or lower temperature to remove the inorganic contaminants, also producing a liquid waste stream. Reheating the gas to temperatures necessary for light hydrocarbon reformation and then adjustment of the hydrogen/carbon monoxide ratio, necessary for the downstream synthesis, leads to thermal inefficiencies. The efficiency of syngas utilization would be significantly improved if all the contaminants could be removed at temperatures higher than that associated with the condensation of tars (greater than approximately 300 °C) and prior to downstream synthesis of fuels and chemicals. A major impurity of concern in syngas is H2S. For coal warm syngas cleanup application, a regenerable adsorbent based on ZnO at the pilot scale was described by RTI. This reduces H2S to 0.5−5 ppm, which is sufficient for integrated gasification combined cycle (IGCC), a technology which offers direct combustion for power generation.7 Syngas for synthesis application needs to be taken to much lower levels, ∼100 Received: February 13, 2013 Revised: May 8, 2013 Accepted: May 23, 2013

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dx.doi.org/10.1021/ie4004927 | Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

Industrial & Engineering Chemistry Research

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alkali carbonates.6 Filtration technologies have been widely reported6,12 and therefore are not discussed here. In this study we provide discussion for choice of catalytic and sorbent materials utilized for cleanup of a range of contaminants that are found in both biomass and coal derived syngas. Preliminary work on the individual unit operations has been done. We have evaluated them in an integrated fashion in order to determine how they perform in tandem. Performance data for some of the individual unit operations are provided here; for the chloride removal, trace metal polishing, and NH3 decomposition. However, the emphasis of this report is to describe the demonstration for the integrated process and report lessons learned. Separate reports are planned that will further detail development specifics for the sulfur removal (both bulk and polishing operations) and for the tar reforming catalyst. The integrated process was operated continuously for approximately 175 h under model biogas conditions. Catalytic and sorbent performance and postrun material characterizations are discussed.

ppb or lower, and other impurities in the gas also need to be removed such as chloride, phosphine and arsine, heavy metals, and others. The approach we have taken in the cleanup of impurities is to use a series of adsorption beds that operate at successively lower temperatures. This includes Na2CO3 for chloride removal, both regenerable and nonregenerable ZnO beds for bulk and trace sulfur removal, and a novel absorbent comprising Ni−Cu particles trapped within a mesoporous silica framework for removal of other trace contaminants.8 We describe the operation of this system in this paper. The authors believe this to be one of the first if not the first open report to describe results for a process designed to reduce a multitude of contaminants, in the integrated fashion as described here, for either coal or biomass application. The total treatment process for warm gas cleanup was broken down into six key stages: chloride removal, bulk sulfur removal, sulfur polishing, removal of trace contaminants, steam reformation of tars and light hydrocarbons, and low-temperature water−gas shift. The integrated process was demonstrated using simulated syngas. The model gas included inorganic impurities that could be found in both coal and biomass: HCl, H2S, AsH3, and NH3.5,9,10 Tar additives that represent biomassderived components consisted of benzene and naphthalene. Methane and ethylene were added as additional hydrocarbon components. While biomass-derived syngas does not typically contain many of these inorganic impurities in significant quantity, inorganic species such as AsH3 and relatively large quantity of H2S were included so as to demonstrate applicability of this process to either coal or biomass application. Biomass syngas conditioning requires an additional step for tar and hydrocarbon treatment/conversion. Therefore, a steam reforming catalyst was added to reform the tars and hydrocarbons into additional useful syngas. The materials for the six major unit operations include: (1) Na2CO3 absorption bed to remove HCl. At 450 °C, Na2CO3 absorbent can remove HCl in syngas from 100 ppm down to less than 1 ppm, with 50 wt % capacity; (2) two parallel ZnO beds to remove the sulfur from >1000 ppm to a few ppm (both regenerable) for bulk sulfur removal; (3) ZnO bed for high capacity sulfur polishing, downstream of bulk sulfur removal (single use); (4) metal-based absorption bed to capture a low parts per million (ppm) level of sulfur and AsH3; (5) tar and light hydrocarbon steam reforming unit operation, which can also promote NH3 decomposition; and (6) commercial lowtemperature water−gas shift catalyst (LT-WGS) to tailor the final H2/CO syngas composition and as an indirect means to assess contaminant removal efficiency. The Cu-based LT-WGS catalyst is known to be extremely sensitive to impurities.11 It should be noted that for hydrogen and synthetic fuel/ chemical production separation of CO2 is likely necessary.10 Unfortunately, suitable CO2 sorption materials for warm temperature operation are not yet proven technologies. Warm gas CO2 sorption material research development is currently underway in our laboratory. Promising materials should be integrated into future warm syngas cleanup strategies. Furthermore, water removal would also be necessary in order to provide a high quality syngas. Associated technologies for warm temperature water removal (e.g., membrane separation) are not discussed here. The process reported here focuses on contaminant removal without CO2 capture or water removal. Finally, in actual application, hot gas filters would be implemented upstream of the sorbent and catalyst beds in order to capture particulates as well as alkali salts in the form of

2.0. EXPERIMENTAL SECTION 2.1. Catalytic and Sorbent Materials. 2.1.1. HCl Removal by Na2CO3. HCl is present in the syngas stream produced from gasification of coal13 and biomass5 and the exact amount varies with feedstock. In one study, a coal-derived syngas was reported to contain 120 ppm HCl.13 HCl can be removed from the process stream by reaction within a fixed bed of alkaline material, such as calcium carbonate or sodium carbonate.14 Despite its higher price, sodium carbonate was utilized due to its higher efficiency. The reaction of HCl with Na2CO3 results in the formation of sodium chloride (NaCl(s)) in the fixed bed and the release of H2O and CO2 into the process stream. 2.1.2. H2S Removal by ZnO. Syngas can contain H2S concentrations of up to approximately 1.5 wt % from coal13 and is typically 100−200 ppm from biomass.5 It is necessary to eliminate H2S from the syngas prior to reaction with the LTWGS and reforming catalysts as they would suffer from deactivation by exposure to sulfur.15 Removal of sulfur can be done using ZnO. ZnO can be reduced to ZnS according to eq 1 and is thermodynamically favorable in the low temperature range (350−400 °C):16 ZnO(s) + H 2S(g) → ZnS(s) + H 2O(g)

(1)

To remove H2S from the process stream, the syngas was passed through: first, a regenerable ZnO bed lowering the H2S level to ∼5 ppm and then, a disposable ZnO polishing bed lowering the H2S concentration to less than 100 ppb.17 Under these conditions, parts per billion (ppb) levels of H2S observed at the outlet of the polishing bed are expected to be captured by the trace metal polishing step discussed in the next section. Regeneration of the ZnS was performed under air according to eq 2. An unwanted side reaction is the formation of sulfate (eq 3) which can be mitigated if the regeneration is performed at temperatures higher than where sulfurization is performed (e.g., 760 °C).18 The regeneration protocol utilized in this study is detailed below. ZnS(s) + O2 (g) → ZnO(s) + SO2 (g)

(2)

ZnS(s) + 2O2 (g) → ZnSO4 (s)

(3)

2.1.3. Trace Metal Contaminant Removal by CuNi/SBA-16 Sorbent. Initial screening of potential materials for the removal B



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of trace level inorganic compounds stemmed from investigations into alternate candidate materials for sulfur removal. One promising material for adsorption under the conditions of warm syngas cleanup studied was Ni,19 and as such, it was considered for the removal of other impurities as well. Two primary concerns in the utilization of nickel were particle sintering and methanation activity. These two effects were respectively lessened by loading and trapping the Ni particles into the cage structure of the SBA-16 and alloying with copper.8 Ni and Cu were also considered as promising sorbent materials since downstream synthesis catalysts frequently contain those metals. Thus, contaminants that could poison the downstream catalysts would instead be trapped by the upstream sacrificial sorbent bed. The combined loading of copper and nickel for this sorbent material was 28.8 wt %, with a copper to nickel molar ratio of 1 to 9. Complete preparation details are described in previous work.8 2.1.4. Tar and Light Hydrocarbon Steam Reforming and NH3 Decomposition Catalyst. The primary purpose of the tar reformer catalyst is to transform the tars and light hydrocarbons into additional syngas (i.e., H2 and CO). Additionally, steam reforming catalysts are known to be active for ammonia decomposition.5,20 Traditional Ni-based steam reforming catalysts are prone to severe deactivation by coking.21 We have thus developed precious metal based catalysts that offer better resistance against coking and sintering. For this study, a 5 wt % Ir/MgAl2O4 (60−100 mesh) steam reforming catalyst was chosen. This particular Ir-based catalyst was chosen for both its favorable activity and stability, after extensive catalyst screening (details to be reported elsewhere). The Ir catalyst was prepared by incipient wetness impregnation of MgAl2O4 (Sasol Puralox 30/140) calcined at 500 °C with iridium nitrate solution (19.3 wt % Ir in nitric acid) dissolved in deionized water. After impregnation, the catalyst was dried at 110 °C for 8 h and calcined under air at 500 °C for 3 h. 2.1.5. Low-Temperature Water−Gas Shift (LT-WGS) Catalyst. After eliminating all the impurities from the process stream, the “clean” syngas was passed through a water−gas shift unit to adjust the H2/CO ratio to the desired ratio. A commercial Cu/ZnO/Al2O3 catalyst highly active for the WGS reaction at 200−280 °C was utilized.11 Cu/ZnO/Al2O3 catalysts are susceptible to poisoning by sulfur and other inorganic impurities.11 Observed deactivation of the WGS catalyst provides thus indication for potential slippage of inorganic impurities or tars through the upstream stages of the process. 2.2. Process Description and Material Placement Rationale. Lab-scale proof-of-concept demonstration for the process and materials employed therein utilized multiple fixed bed reactors. Ultimately a system such as this may have to operate adiabatically; hence the cleanup generally proceeds from higher temperature beds to lower. Although it should be pointed out that in this particular case the high temperature tar reforming unit is in the middle of the process thus contributing a thermal inefficiency. Nonetheless, for proof-of-concept purposes it was utilized as such in this demonstration. A block flow diagram of the lab-scale proof of concept demonstration’s process, depicting contaminant’s target reduction levels, is shown in Figure 1. The first stage of the process was the removal of HCl by Na2CO3 in reactor one (R1) operating at 450 °C, an optimal temperature as determined by separate investigation as discussed below. The placement of HCl removal first was

Figure 1. Block flow diagram (BFD) of warm syngas cleanup lab-scale proof of concept demonstration.

due to the moderate temperature at which it is conducted, the high potential for chloride to negatively impact other sorbents and catalysts, and the relative inertness of Na2CO3 for activity other than the formation of NaCl. It has also been advised that the most corrosive impurities should be removed first without allowing their interaction with any of the process equipment or active materials.6 The second stage was configured for bulk removal of sulfur through adsorption by ZnO. Similar to HCl removal, bulk sulfur removal was conducted early due to sulfur’s potency as a catalyst poison, the moderate temperature at which the adsorption occurs and ZnO’s inertness to unintended reactions. The system was designed such that two regenerable parallel beds (R2a and R2b) would be used at 450 °C to remove the bulk sulfur from the gas phase. The regeneration process consisted of an 8 h high temperature oxidation stage at 650 °C under 2% O2 at GHSV = 8200 h−1, followed by a 6 h lower temperature reduction stage at 450 °C under 8% H2 at GHSV = 8200 h−1. Subsequently, the ZnO beds of R2a and R2b experienced alternating cycles of 16 h of operation and 16 h of regeneration over the duration of the test. Both of the beds were loaded with what calculated to be twice the amount of ZnO necessary for 16 h of bulk H2S removal at 450 °C (calculated from the ZnO sulfur weight capacity which is approximately 25 wt % at this temperature). While the bulk removal beds are known to effectively remove H2S to as low as 5 ppm, this concentration is still high enough to poison the downstream catalysts, so a polishing bed operating at 300 °C was placed immediately after the bulk sulfur removal beds. The additional bed of ZnO (R3) acts as the sulfur polishing stage, further lowering the H2S levels close to the thermodynamic limit near 100 ppb. The fourth reactor (R4) contained the CuNi-SBA16 material for removal of trace metal contaminants and ppb level H2S. The CuNi/SBA-16 sorbent was loaded in a fixed bed reactor, reduced at 300 °C for 8 h under 10% H2 in N2, and operated at 260 °C. The fifth reactor (R5) was used for the tar reforming stage, holding the 5 wt % Ir/MgAl2O4 steam, reforming catalyst. The tar reforming bed was placed after the ZnO and CuNi beds due to the catalyst’s proclivity to deactivate over time with even slight exposure to sulfur. The hydrocarbons within the process C



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stream must be fully converted to CO2; otherwise they will flow through the LT-WGS bed and likely induce coke formation. Additionally, the high temperature tar reformer can be utilized for NH3 decomposition.5 The 5 wt % Ir/MgAl2O4 catalyst was reduced under 10% H2/ N2 for 1 h at 500 °C, then 6 h at 850 °C, and operated at 850 °C. The sixth and final reactor contained the low-temperature Cu/ZnO/Al2O3 water−gas shift catalyst and was mixed with SiC diluent (1:1 by weight) in order to facilitate isothermal operation. The catalyst was reduced under 10% H2/N2 for 8 h at 230 °C and operated at 230 °C. 2.3. Process Evaluation Methodology. Material loading details for each reactor are presented in Table 1. The feed gas

by using a multichannel vaporizer. The resulting feed had a steam to carbon molar ratio of 1.73. Primary syngas and hydrocarbon components within the product gas were analyzed with an Agilent gas chromatograph (3000A) equipped with mol sieve 5A, PPU, and OV-1 columns. Detection and quantification for the inorganic components H2S and AsH3 was performed using an HP 6890N GC equipped with an Rtx-Q Bond porous layer open tubular (PLOT) capillary column (30m length, 0.53 mm inner diameter, d.f. = 0.20 μm). Detection was performed by a Vicci pulsed He discharge ionization detector (PHDID). The PHDID lower detection limits for H2S and AsH3 were 100 and 15 ppb, respectively. Quantification of HCl and NH3 was not available in the integrated process evaluation study. Concentrations of gaseous components at thermodynamic equilibrium were determined using a Gibbs Reactor with ChemCad (version 6.4) simulation software. At the end of the long-term evaluation test period, the spent tar reformer and LT-WGS catalysts were separately tested to identify possible loss of activity. The two catalysts remained loaded in the lab-scale proof of concept demonstration system and were tested at increased space velocities so as to compare activities when under kinetic control. The tar reforming catalyst was tested at 550−750 °C and GHSV = 126 000 h−1 using a molar feed composition of 4.3% H2, 78.3% N2, 1.2% CH4, 2.1% CO, 2.1% CO2, 0.2% ethylene, and 11.8% H2O. The LT-WGS catalyst was tested at 150−230 °C, GHSV = 36 000 h−1 using a molar feed composition of 1.0% H2, 74.3% N2, 0.7% CO, 0.1% CO2, and 20.5% H2O. Following these tests, fresh catalysts were loaded, reduced, and tested individually in the same manner as their spent equivalents. These data sets provide a comparison of the fresh and spent versions of both of these catalysts so that possible loss of activity over time could be identified. Separate HCl removal experiments were performed at 300− 550 °C and GHSV = 80 000 h−1 using NaCO3 (50−60 mesh). A premixed cylinder containing 50% H2O, 13% CO, 10% CO2, 20% H2, 7% CH4, and 103 ppm HCl was introduced to the experimental system at ambient pressure and in the same method as described above. Separate AsH3 decomposition experiments were performed at 150−600 °C and GHSV = 5000 h−1 using NaCO3 (35−60 mesh), glass beads (0.5 mm diameter), NiCu/SBA-16 (35−60 mesh), and a blank reactor with no material present. A premixed cylinder containing 9 ppm AsH3 in helium was introduced to the experimental system at ambient pressure and in the same method as described above. Separate NH3 decomposition experiments were performed at 700−900 °C and GHSV = 12 000−36 000 h−1 using a commercial Ni-based catalyst (G90B United Catalyst Inc.). An analytical method using NH3 ion selective electrode was developed which can accurately measure a broad range of NH3 concentration (from 1 ppm up to 1 wt %) in steam-containing warm syngas. A premixed feed gas and water were introduced to the experimental system at ambient pressure and in the same method as described above. The resulting molar feed composition contained 9.8% CO, 15.4% H2, 8.9% CO2, 4.9% CH4, 50% H2O, 10% He, 1% N2 and either 1000 or 100 ppm NH3. 2.4. Material Characterization. The spent sorbents and catalysts were characterized by means of X-ray diffraction (XRD), inductively coupled plasma−optical emission spectroscopy (ICP-OES), ion chromatography (IC), scanning trans-

Table 1. Reactor Loadings particle size reactor

material

R1 R2a R2b R3 R4

Na2CO3 ZnO ZnO ZnO CuNi/ SBA16 5% Ir/ MgAl2O4 a CuZnAl2O3 a SiC

R5 R6 R6 a

(mesh #)

(μm)

mass (g)

Sigma-Aldrich Süd-Chemie Süd-Chemie Süd-Chemie PNNL

35−60 35−60 35−60 35−60 35−60

250−500 250−500 250−500 250−500 250−500

4.80 2.88 2.88 2.88 1.80

PNNL

60−100

150−250

0.20

Süd-Chemie

60−80

177−250

1.62

Atlantic Equipment Engineers

60−100

150−250

1.62

source

Physically mixed together prior to loading into reactor.

composition simulated that of the effluent from a biomass-fed gasifier (Table 2). Syngas, light hydrocarbons, trace inorganic Table 2. Feed Gas Composition syngas and carriers 16.5% 8.3% 8.3% 5.7% 0.8% 46.1% bal 0.8%

H2 CO CO2 CH4 C2H4 H2O He N2 tar simulants benzene naphthalene

4006 ppm 492 ppm contaminants

HCl H2S AsH3 NH3

51 ppm 1065 ppm 6 ppm 232 ppm

species, and inert gases were blended from compressed gas cylinders using mass flow controllers. Liquid feeds of water and tar simulant (86.2 wt % benzene mixed with 13.8 wt % naphthalene) were fed by means of a Lab Alliance HPLC pump and Teledyne ISCO high-pressure syringe pump, respectively. The liquids were vaporized prior to being fed into the system D



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catalysts in order to ascertain if metal sintering occurred during testing. Quantification for the inorganic species on spent materials was conducted by means of ICP-OES using a Perkin-Elmer Optima 7300 DV ICP-OES. Chloride quantification was obtained through analysis of spent samples using ionchromatography (IC), wherein each sample was added to water and analyzed using a Dionex ICS-3000. Total carbon analysis was performed by ALS Environmental in Tucson, AZ. STEM analysis was performed with FEI Titan 80-300 operated at 300 kV. The FEI Titan is equipped with CEOS GmbH double-hexapole aberration corrector for the probeforming lens, which allows imaging with ∼0.1 nm resolution in (STEM) mode. The STEM images were acquired on a high angle annular dark field (HAADF) with an inner collection angle of 52 mrad. In general, the sample preparation involved mounting of powder samples on copper grids covered with lacey carbon support films and immediately loading them into the TEM airlock to minimize an exposure to atmospheric O2. Ir metal particle size distribution was determined counting approximately 150 particles, for the fresh and spent reforming catalysts.

mission electron microscopy−energy dispersive spectrometry (STEM-EDS), and total carbon analysis. Note that prior to characterization each of the fixed bed reactors was cut such that the contained bed was separated into multiple sections, as shown in Figure 2. For each analytical technique, a sample representative of each material’s starting condition was analyzed for reference.

3.0. RESULTS 3.1. Individual Unit Evaluations. Gas composition data through each individual process unit was collected before and immediately following the long-term stability evaluation, as each of the process units were successively either added or removed from the process stream. These data sets are provided in Table 3. The passing of the feed gas through R1 containing Na2CO3 had no appreciable effect on the majority of the components within the feed, except for the diminution of the AsH3 below detection limits. However, as noted above, analytical for both HCl and NH3 was not available during the integrated process evaluation. Characterization of the spent Na2CO3 material, as discussed below, revealed the presence of Cl, confirming its ability to capture HCl. While analytical for the HCl was unavailable for the integrated process evaluation, a separate single bed investigation further verified reports by others6,22 for Na2CO3 to be effective for HCl removal. Na2CO3 was

Figure 2. Bed unloading diagram.

It should also be noted that the conditions under which the bulk ZnO beds were removed were not identical. Upon completion of the long-term stability test, bed R2a had reached the end of 16 h in operation and R2b had just completed the regeneration process. The R2a bed did not undergo regeneration prior to unloading, so it represents a ZnO bed that has just undergone a normal adsorption cycle. The material unloaded from R2b is representative of a freshly regenerated ZnO bed. Over the duration of the long-term stability evaluation, beds R2a and R2b were alternated such that each bed treated the process stream for six adsorption stages. Analysis by XRD was conducted using a Rigaku MiniFlex II (θ-2θ, 30 kV, 15 mA), equipped with a Cu anode and monochromator to filter out Kβ peaks. XRD was also used to measure the Cu particle size for the LT-WGS fresh and spent Table 3. Gas Compositions for Unit Process Tests stage

He %

system startup feed bal R1 outlet bal R2a outlet bal R3 outlet bal a R4 outlet R5 outlet bal R6 outlet bal system shutdown R6 outlet bal R5 outlet bal R4 outlet bal R3 outlet bal R2a outlet bal R1 outlet bal feed bal a

H2 %

N2 %

CH4 %

CO %

CO2 %

50.6 50.2 50.4 50.8

1.6 1.5 1.4 1.5

10.3 10.2 9.9 9.7

14.7 14.6 11.9 11.2

15.2 15.3 17.2 17.6

49.8 52.6

0.7 0.5

0.0

13.6 0.4

12.7 25.4

63.4 59.0 49.0 48.1 48.9 48.8 49.0

0.8 0.9 1.4 1.5 1.4 1.5 1.4

0.0 0.0 10.2 10.2 10.1 10.2 10.2

0.6 16.6 14.3 14.3 14.3 14.5 14.5

25.4 13.4 15.4 15.4 15.4 15.2 15.2

C2H4 % 1.3 1.3 1.3 1.3

1.5 1.5 1.5 1.5 1.5

ethane %

H2S %

benzene %

naphthalene %

H2S ppm

AsH3 ppm

0.1 0.1

0.68 0.57 0.64 0.44

N/A

2117 2244 2

18

0.1 0.1

0.62 0.65 0.66 0.68 0.70

2 2244 2117

18

Leak in condensor observed for this data set, thus, results not available. E



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Figure 3. HCl removal over 50−60 mesh Na2CO3 as a function of temperature (simulated syngas molar feed composition: 50% H2O, 13% CO, 10% CO2, 20% H2, 7% CH4, 103 ppm HCl; 80 000 h−1 GHSV; atmospheric pressure).

Figure 4. Decomposition of AsH3 through an empty reactor, over Na2CO3, CuNi/SBA-16, and glass beads (9 ppm AsH3 in helium; 5000 h−1 GHSV; atmospheric pressure).

disappearance of AsH3 over the Na2CO3 bed (R1), and this prompted additional further investigation. In another separate investigation, Na2CO3 was evaluated specifically for AsH3 decomposition over a wide temperature range using a feed containing 9 ppm AsH3 (balance helium) and at GHSV = 5000 h−1. For comparison purpose, the CuNi/ SBA-16 material, glass beads, and a blank with no material present were similarly evaluated. AsH3 decomposition conversion for these tests is shown in Figure 4. CuNi/SBA-16 provided complete removal from the gas stream at all temperatures tested (200−475 °C). Na2CO3 completely removed AsH3 above 350 °C. Both glass beads and a blank reactor exhibited complete decomposition of the AsH3 at 550− 600 °C. This provides further evidence that the Na2CO3 operating at 450 °C in the R1 unit of the process evaluation

separately evaluated specifically for HCl decomposition over a wide temperature range using a simulated syngas feed containing 103 ppm HCl and at GHSV = 80 000 h−1. HCl removal results are described in terms of sorbent weight capacity and shown in Figure 3 as a function of temperature. Optimal sorbent removal was observed at 450−500 °C with approximately a 30 wt % weight capacity. It should be noted that in these separate studies performance of K2CO3 was found to be substantially inferior to Na2CO3. This separate investigation provides complementary information for HCl removal over Na2CO3 that was not provided during the integrated process evaluation due to the analytical constraints. Further analysis of the gas phase results obtained during the process testing, as shown in Table 3, shows the unexpected F



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Figure 5. Decomposition of NH3 through a Ni-based catalyst (feed gas: 9.8% CO, 15.4% H2, 8.9% CO2, 4.9% CH4, 50% H2O, 10% He, 1% N2, and either 1000 or 100 ppm NH3; atmospheric pressure).

Figure 6. Conversion of target species during long-term process evaluation, downstream from the LT-WGS reactor (230 °C, 8000 h−1 GHSV, atmospheric pressure).

reduced to levels below detectable limits, it was difficult to ascertain from the gas phase analysis whether the CuNi/SBA16 provided any benefit. However, as discussed below, trace levels of sulfur were observed in the spent material indicating that the CuNi/SBA16 sorbent was able to eliminate trace level of H2S. It should be noted that there was an unintended leak at the exit condenser, specific to the time of sampling R4; thus, this particular data set for the system startup (R4 outlet) is not available. However, as shown in the system shutdown there was no meaningful change in concentrations when comparing R3 and R4 outlet concentrations. This is important to note as the CuNi/SBA-16 material can catalyze methanation if operated at too high temperatures. Thus, methanation over CuNi/SBA-16

apparently decomposed a substantial portion of the AsH3, in addition to capturing Cl as originally intended. The addition of R2a to the process stream introduced the ZnO bed designed for bulk removal of H2S. The gas composition data (Table 3) clearly shows that the ZnO effectively removes the vast majority of H2S; with the measured content falling from 2244 ppm at R1’s outlet to 2 ppm at the outlet of R2a. Similarly, the ZnO polishing bed in R3 further removed H2S from the gas stream while leaving other components unchanged. The ZnO in R3 effectively lowered H2S content to levels below the detectable limit of the instrumentation of 100 ppb. The process stream was next directed through the CuNi/ SBA-16 bed in R4. As both the AsH3 and H2S were already G



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Figure 7. Activity testing of the 5% Ir/MgAl2O4 catalyst (4.3% H2, 78.3% N2, 1.2% CH4, 2.1% CO, 2.1% CO2, 0.2% ethylene, 11.8% H2O; 126 600 h−1 GHSV; atmospheric pressure).

was not observed when operated at 260 °C during the course of this study. Addition of the tar reformer catalyst housed in R5 resulted in gas composition changes. Hydrocarbon content became undetectable as methane, ethylene, and benzene were all reformed. While naphthalene was not directly analyzed, we believed it was fully reformed. Indeed, separate experiments not presented here have shown that it is easier to reform naphthalene than benzene under the present conditions. Hence, the steam reformer appears to have correctly functioned in completing converting the hydrocarbon feed to syngas. Additionally, while analytical for the detection of NH3 was not available during the course of the 175 h testing, separate tests support the likelihood of decomposition of NH3 to N2 in the steam reformer. Indeed, in a separate investigation, NH3 decomposition was performed over a commercial Ni steam reforming catalyst using a feed similar to that used for the process evaluation. By feeding syngas with no N2, over the catalyst at 850 °C at a GHSV of 36 000 h−1, more than 95% of NH3 was decomposed to N2 and H2, decreasing the concentration of NH3 from 100 ppm down to less than 5 ppm. These results presented in Figure 5 were obtained using a Ni-based catalyst that is typically much less active than those that are precious metal-based. Thus, we believe that the NH3 was likely reduced to very low levels through the R5 bed, containing Ir-based reforming catalyst, during the process testing. However, further investigation utilizing online analysis for NH3 is certainly warranted using this Ir-based reforming catalyst. Finally, addition of the LT-WGS catalyst in R6 resulted in increased levels of H2 and CO2, as well as a decrease in CO, as a result of water−gas shift reaction occurring. Small levels of CO were still present at the outlet of R6 as a result of thermodynamic constraint. 3.2. Long-Term Process Evaluation. While the efficiency of the sorbents and catalysts was evaluated by determining the contaminant species level of the spent materials, the primary method of evaluating the system’s performance in real time was

monitoring stability of the LT-WGS catalyst. This was achieved by gas phase analysis at the outlet of the LT-WGS reactor (R6) as a function of time-on-stream. The only gas components observed over the course of the test were H2, CO, CO2, He, and N2 (H2O was condensed out prior to GC sampling). Figure 6 shows the CO conversion during the stability test. CO conversion maintained an approximate steady value of 95% over the entire 175 h time-on-stream. The CO conversion approaches that as predicted by thermodynamics, as depicted in Figure 6. This value was calculated using ChemCad software utilizing the effluent compositions obtained post-tar reformer (R5). Complete methane, ethylene, and benzene conversion was observed over the entire duration of the test. Thus, no significant deactivation was observed for the process. However, separate activity performance evaluation for the LT-WGS catalyst when under kinetic control was necessary for a more accurate assessment, which is discussed below. 3.3. Postrun Activity Evaluations. The sequential removal of each unit operation from the process stream upon completion of the long-term stability test was performed in order to systematically assess performance for each unit operation after the 175 h test duration, and yielded data that generally coincided with that collected upon startup (shown in Table 3). The removal of R6 showed loss of water−gas shift effects by way of the lower H2 and CO2 content at the outlet of R5. When the tar reformer catalyst in R5 was taken offline and the gas composition at the outlet of R4 analyzed, the concentrations of hydrocarbons (methane, ethylene, benzene) matched those of the feed stock observed upon startup. The dilution effect of the reformed hydrocarbons, as observed upon startup, was also reversed and N2 and CO2 levels returned to their respective nominal values. When R4 was removed, no change in the gas composition was observed. Removal of the ZnO beds R3 then R2a shows the levels of H2S increasing to 2 then 2244 ppm, respectively, while all other components in the gas stream remained unaffected. This directly mirrors the results observed upon startup. Similarly, the removal of R1 H



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Figure 8. Activity testing of the Cu/ZnO/Al2O3 catalyst (1.0% H2, 74.3% N2, 0.7% CO, 0.1% CO2, 20.5% H2O; 36 000 h−1 GHSV; atmospheric pressure).

Figure 9. XRD patterns for R1 (Na2CO3).

shows the return of AsH3 in the feed but no effect on the other species. Further elucidation of potential catalyst deactivation was provided by a set of activity tests performed on the tar reformer and LT-WGS catalysts upon completion of the demonstration test. Activity test results for the spent tar reformer catalyst as compared to that of a fresh catalyst, evaluated separately under identical conditions, are presented in Figure 7. This comparison illustrates that a marked loss of methane conversion was observed for the spent 5 wt % Ir/MgAl2O4 catalyst. As discussed below, this is primarily the result of coking of the Ir nanoparticles likely induced by the benzene and naphthalene present in the feed, as evidenced by the presence of carbon on the spent materials. These results indicate that the steam

reforming catalysts did in fact deactivate, observable when tested under kinetic-control. This suggests that real industrial applications utilizing fixed-bed steam reforming reactors may require duel regenerating beds. For purposes of the process evaluation, the steam reformer was oversized thus explaining why loss of activity was not observed during the 175 h test duration. It should be noted that the concentration of tar stimulants feed were intentionally 2−5 times greater than what is typically observed in real biomass-derived syngas5 in order to exaggerate any coking effects. Analogous to the testing of the tar reformer catalyst, activity test results for the LT-WGS and its fresh counterpart are provided in Figure 8. The spent Cu/ZnO/Al2O3 catalyst was less active than the fresh catalyst. CO conversion was 39% and I



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Table 4. ICP and IC Results R1

a

R2a

section (no.)

sulfur (ppm)

arsenic (ppm)

chloride (ppm)

fresh inlet middle outlet

126 619 155

Progress toward Biomass and Coal-Derived Syngas Warm Cleanup

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