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Hydrodeoxygenation of acetic acid in a microreactor
Chemical Engineering Science 84 (2012) 761–771
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Hydrodeoxygenation of acetic acid in a microreactor Narendra Joshi n, Adeniyi Lawal New Jersey Center for Micro-Chemical Systems, Department of Chemical Engineering and Materials Science, Stevens Institute of Technology, Hoboken, NJ 07030, USA
H I G H L I G H T S c c c c c
Acetic acid can be converted by HDO below 300 1C at atmospheric pressure. Reaction pathways have been constructed based on literature reviews and product analysis. External mass transfer resistance was negligible at overall flow velocity of 2.54 m/s. Internal mass transfer resistance was negligible at an average particle size of 113 mm. The conversion of acetic acid decreases as internal reactor diameter increases.
a r t i c l e i n f o
abstract
Article history: Received 5 May 2012 Received in revised form 22 September 2012 Accepted 24 September 2012 Available online 1 October 2012
Acetic acid was used as a model compound for pyrolysis oil in a hydrodeoxygenation (HDO) study. HDO of acetic acid was performed in a packed bed microreactor. The catalyst was reduced sulfided NiMo/ Al2O3. The effects of state of aggregation of acetic acid, temperature, hydrogen partial pressure, liquid flow rate, reactor diameter, and residence time on conversion, yield, space-time consumption, and space-time yield were investigated. External and internal mass transfer and heat transfer resistances were also examined in the microreactor. Temperature was a major factor in HDO of acetic acid. Many consider hydrodeoxygenation as an unattractive process due to high pressure requirements (1050–3000 psig). In this work, attempt has been made to show that HDO of acetic acid can be conducted at atmospheric pressure with a significant conversion achieved. More acetic acid was converted during HDO as temperature was increased at constant pressure of 300 psig. Conversion was much higher for vapor phase acetic acid at atmospheric pressure than liquid phase acetic acid. HDO of gas phase acetic acid in a blank reactor compared to a catalytic HDO showed that thermal decomposition of acetic acid did not occur appreciably. Partial pressure of hydrogen above 240 psig had no effect on the conversion of liquid phase acetic acid. Conversion of vapor phase acetic acid increased as the partial pressure of hydrogen increased from 3 psig to 15 psig. Residence time was 0.06 s for a maximum conversion of liquid phase acetic acid, whereas it was 0.03 s for a maximum conversion of vapor phase acetic acid. The conversion of acetic acid for both liquid and vapor phases decreases significantly as the flow rate of acetic acid increases. As reactor diameter increases beyond 0.8 mm, the conversion reduces significantly. Mass transfer resistance was negligible at the superficial velocity of 2.54 m/s and at an average catalyst particle size of 113 mm. Radial temperature difference in the microreactor was less than 5%. & 2012 Elsevier Ltd. All rights reserved.
Keywords: Microreactor Heat transfer Fuel Energy Mass transfer Hydrodeoxygenation
1. Introduction Depleting petroleum reserves, rising prices, and environmental and political concerns have made renewable resources for transportation fuel attractive. Though there are other renewable sources of energy such as wind, solar, and hydroelectric, biomass is the only renewable source of energy that can be used for liquid fuel, fitting into the current infrastructure of transportation fuel utilization.
n
Corresponding author. Tel.: þ1 201 216 8314; fax: þ1 201 216 8306. E-mail addresses: [email protected], [email protected] (N. Joshi).
0009-2509/$ - see front matter & 2012 Elsevier Ltd. All rights reserved. http://dx.doi.org/10.1016/j.ces.2012.09.018
Effective utilization of biomass as the transportation fuel may reduce the world’s dependency on non-renewable fossil fuel. Another important advantage of the biomass is that it contributes no new carbon dioxide to the atmosphere (Peter, 2002). Raw pyrolysis oil from biomass can be further processed to obtain transportation fuel. Processing alternatives include hydrodeoxygenation, catalytic cracking, and steam reforming followed by Fischer Tropsch synthesis. Lately more attention has focused on hydrodeoxygenation as a means of increasing the energy density of the product fuel. Hydrodeoxygenation of pyrolysis oil requires about 450 1C to reduce most of the oxygen content in the oil. But at this temperature an extreme degradation of oil takes
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place and it becomes a very hard, coke like material. This chemical instability is attributed to unsaturated double bonds such as acetic acid, aldehyde, ketone, and so forth which react through condensation reaction similar to the phenol-formal polymerization (Grange et al., 1996). Therefore, it is very crucial to eliminate these functions before they react to high molecular weight compounds. Experiments with oxygenated model compounds show that it can be accomplished by a low temperature (below 300 1C) hydrotreatment (Elliott and Neuenschwander, 1996; Elliott et al., 1988; Elliott and Oasmaa, 1991; Grange et al., 1996). High acetic acid content (upto 25%) provides corrosive nature to pyrolysis oil (Milne et al., 1997; Xu et al., 2009) which needs to be reduced for long term storage. It also contains unsaturated double bond functional group which leads to polymerization of pyrolysis oil at higher temperature (above 300 1C) during hydrodeoxygenation. Therefore, converting acetic acid by hydrodeoxygenation at temperatures below 300 1C will increase the stability of pyrolysis oil and in turn, it is expected to reduce polymerization of pyrolysis oil significantly when hydrotreated at temperatures above 300 1C. By conducting hydrodeoxygenation of acetic acid we hope that it will provide an opportunity to investigate upgrading mechanism. Using acetic acid as a model compound we also hope to study reaction pathways of HDO of acetic acid. An understanding of reaction pathways and kinetics is expected to be of value in process design and modeling (LaVopa and Satterfield, 1987). However, studying reaction pathways and kinetics of HDO of PO is very difficult as it contains more than 300 oxygenated compounds. Therefore, acetic acid is selected as our model compound to investigate the hydrodeoxygenation process and reaction pathways. Biomass is a bulky material with a very low density and scattered across geographical areas. Research has shown that transporting biomass to a central location and processing for biofuel is not economically feasible due to very high cost of transportation. The research suggest that biofuel production facilities will have to be built close to the sources of feedstock, probably within 50 miles (David, 2008); therefore, a large scale processing plant will not be profitable. Even for a medium size biomass conversion facility, the amount of biomass required is overwhelming. It is calculated that a facility producing 50 million gallon biofuel per year would require a truck loaded with biomass to arrive every six minutes around the clock (David, 2008). The requirements of high production cost and very large quantity of biomass discourages people to use the biomass conversion facility with conventional macroreactor system. However, on-site and on-demand distributed system could be an alternative solution. In this regard, microreactor system could be a suitable alternative. Microreactor system consists of devices that are miniature in sizes and scaling up is done by increasing number of channels within a reactor unit as well as by increasing number of reactor units. The conventional method of commercializing a new process starts at the laboratory scale where results are collected which are then used in a pilot plant to obtain critical parameters. The information collected from pilot plant is used to design and build the larger production unit. In microreactor technology, scale-up can be achieved by numbering up straight from laboratory scale bypassing a costly pilot plant (Jenck, 2009). There are number of other benefits of using microreactor system. Study showed that heat transfer coefficient reached upto 25,000 W/m2 K in microdevices exceeding those of conventional heat exchangers by at least one order of magnitude (Schubert et al., 1998). In micromixers typical fluid layer thickness can be set to a few tens of micrometers, consequently, mixing time in micromixers amount to milliseconds which is hardly achievable using stirring equipment and other conventional mixers (Branebjerg et al., 1996; Knight et al., 1998). A comparison of microreactor with
conventional trickle-bed reactor for the hydrogenation of styrene showed that the throughput of the liquid for microreactor is eight times higher than that of the trickle-bed reactor (Nijhuis et al., 2003). As a result of enhanced mass transfer, the hydrogen concentration on catalyst is significantly higher than on the trickle-bet catalyst (34 mol/m3 compared to 22 mol/m3). Microreactors possess ultra-low transport resistances, therefore mass diffusion and heat transfer are extremely fast, resulting in rapid thermal and reaction equilibrium (Besser et al., 2003). High yield, improved product quality, better selectivity and safe operation are attainable due to very high heat and mass transfer (Halder et al., 2007; Okafor et al., 2010; Tadepalli et al., 2007; Voloshin and Lawal, 2010). Mixing in the microchannels is attainable only by inter-diffusion of reactants due to laminar flow (Adeosun and Lawal, 2005), however due to short transverse diffusional distance, rapid and effective mixing is attainable in the microreactor which can quickly bring reactants in contact with catalyst in a heterogeneous reaction (Hessel et al., 2001). Enhancement of mixing in two-phase flows can also be achieved by selecting appropriate inlet T-orientations to provide a short slug length. Studies have shown that introduction of gas and liquid feeds head to head or perpendicular to each other with the liquid stream parallel to the microchannel markedly improves the mixing (Qian and Lawal, 2006). There are other potential benefits of microreactors regarding application which are listed below (Ehrfeld et al., 2000) 1. 2. 3. 4. 5.
Earlier start of production at lower costs. Easier scale-up of production capacity. Smaller plant size for distributed production. Lower costs for transportation, materials and energy. More flexible response to market demands.
Many researches have considered microreactor technology as a potential for process intensification. Process intensification involves the development of innovative methods and devices that, in comparison to existing approaches, offer the chance of a dramatic improvement in the quality of production, substantial reduction in the ratio of equipment size to production capacity, and significant drop in the consumption of energy and production of waste. For these reasons, microstructured devices and components have an important role to play (Matlosz et al., 2009). Commercially available catalysts such as CoO/MoO3 and NiO/MoO3 on Al2O3 support used for removing sulfur, nitrogen, and oxygen from petrochemical feedstock are generally selected for hydrodeoxygenation process as well. According to research conducted at Pacific Northwest National Laboratory, the sulfided form of CoO/MoO3 and NiO/MoO3 are much more active for hydrodeoxygenation than the oxide form (Elliott, 2007). Sulfidation creates active sites that can play a role in the rupture of carbon-heteroatom bond (Senol, 2007). In this study we used the sulfided form of NiO/MoO3 on Al2O3. Model compounds are chosen instead of pyrolysis oil for better understanding and control of the HDO reaction process. Nimmanwudipong et al. used 2-methoxyphenol as a model compound of lignin-derived pyrolysis oil in the presence of hydrogen to elucidate the reaction network and to predict oxygen removal (Nimmanwudipong et al., 2011). The authors have shown that the catalytic conversion of 2-methoxyphenol in the presence of hydrogen catalyzed by Pt/Al2O3 involves three major classes of reactions: hydrogenation, hydrodeoxygenation, and transalkylation. Xu et al. has demonstrated that acetic acid can be converted to ethyl acetate using reduced Mo-10Ni/g-Al2O3 via hydrodeoxygenation at 473 K and 3 MPa hydrogen pressure (Xu et al., 2009). They have proposed that ethyl acetate was produced from acetic acid via a three step reaction in which produced aldehyde was
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converted to ethyl alcohol by hydrogenation which then reacted with acetic acid to form ethyl acetate. Similarly, LaVopa and Satterfield studied the hydrodeoxygenation of dibenzofuran on a sulfided NiMo/Al2O3 at 360 1C and 7.0 MPa and found that singlering cyclohexane predominated (LaVopa and Satterfield, 1987). They also found that the catalyst in the oxide form had lower activity and double-ring products predominated over single-ring compounds. Also, Li et al. concluded that during hydrodeoxygenation of 1-Naphthol catalyzed by sulfided Ni–Mo/g-Al2O3 aromatic ring hydrogenation and direct oxygen extrusion (HDO) occur in parallel (Li et al., 1985). The objective of the research presented here was to evaluate hydrodeoxygenation of acetic acid in a packed bed microreactor. Effects of various reaction variables such as liquid flow velocity, residence time, hydrogen partial pressure, temperature, state of aggregation of acetic acid, and reactor diameter on conversion, yield, and space-time yield were studied.
2. Experimental 2.1. Materials Presulfided NiO/MoO3/Al2O3 catalyst obtained from Albemarle (sulfided and supplied by Eurecat USA), Houston, Texas was ground and sieved to obtain particles with diameters in the range of 75–150 mm. An average surface area of the sieved catalysts was ˚ The surface area 164 m2/g and average pore diameter was 106 A. and pore diameter were obtained by using multipoint BET technique and the instrument used was Quantochrome Autosorb-1. The catalyst was reduced with 5.0 sccm of hydrogen at 593 K and 3.45 MPa for 2 h. The average surface area of the reduced catalysts ˚ ACS regent was 209.0 m2/g and average pore diameter was 92.0 A. grade (conc. 99.7%) acetic acid was purchased from Pharmco Inc. The gas used was extra dry hydrogen from Praxair. Nitrogen was used as tracer to perform a material balance.
2.2. Experimental setup A HPLC pump (Laboratory Alliance Series III) was used to control the flow rate of acetic acid. Mass flow controllers (Porter Model 201) were used to control the flow rates of hydrogen and nitrogen. Ranges of superficial velocities of acetic acid, hydrogen and nitrogen gases used were 0.0011–0.0065 m/s, 0.54–4.71 m/s, and 0.36–2.75 m/s respectively. The liquid and gas phases were combined in a T-junction mixer (Upchurch) with 508 10 6 m through-hole. Reynolds number for the combined flow was less than 100 for all experiments indicating laminar flow. The fluids exiting from the T-junction exhibited a Taylor flow pattern with a liquid slug length in the range 0.001–0.003 m, whereas gas bubble length varied from 0.001 to 0.005 m. A microreactor was prepared from a 0.0016 m (1/16 in.) 316 stainless steel tubing with 762 10 6 m internal diameter, and was gravity filled with catalyst. The total length of the packed bed microreactor varied from 0.025 to 0.18 m. Hastelloy micron filter-cloth (200 1150 mesh, Unique Wire Weaving Co., Hillside, New Jersey) was placed at the ends of the reactor to retain the catalyst. The reactor system was pressurized using a back pressure regulator (GO Regulator Co.). The entrance and the exit pressures of the fluids (liquid and gas combined) in the reactor were measured. The pressure drop along the reactor varied from 0.07 MPa to 0.2 MPa depending upon reactor length. Acetic acid was vaporized prior to entering to a reactor using a vaporizer heated to a temperature above the boiling point (118 1C) of acetic acid. The product stream was condensed using a chiller (model NESLAB RTE 7, Thermo Fisher Scientific).
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2.3. Analysis The water content in the HDO product was analyzed using Volumetric Karl Fischer (KF) Titration Workstation (Model 375, Denver Instrument) with the use of hydranal reagents obtained from Sigma-Aldrich. The titrant used was hydranal titrant-2 while the working medium was hydranal solvent. Before the actual titration, the hydranal solvent was titrated to dryness in a drift determination step. This step removes moisture from solvent, electrode, and titration vessel. The HDO product was added using a syringe through the septum on the KF cell. Analysis of liquid product was conducted in a HPLC (Shimadzu series: mobile phase degasser [DGU_20A5], pump station [LC-20AT], auto-sampler [SIL-20AC], and reflective index detector [RID-10A]) equipped with BioRad Aminex HPX-87H column. The mobile phase consisted of 0.007 N aqueous H3PO4 with an isocratic flow (flow rate of 6 10 7 m3/min). The gas phase was analyzed for CO2, CO, and CH4 using Varian 450GC with Hayesep N and Mol sieve 5A columns in series. HDO of acetic acid involves complex reactions consisting of a series of reactions forming acetaldehyde, ethanol, and ethyl acetate and parallel reactions forming acetone, carbon dioxide, carbon monoxide, and methane. Acetic acid reacts with hydrogen forming acetaldehyde and H2O. A model for the formation of acetaldehyde is suggested by (Grootendorst et al., 1994) with following pathway: CH3COOH þH2 ¼ 4CH3CHOþH2O, DH1¼1.27 kJ/mol
(Reaction 1)
and (Xu et al., 2009) suggested the formation of ethyl acetate with following reaction mechanism: CH3CHO þH2 ¼ 4CH3CH2OH, DH1 ¼ 80.98 kJ/mol
(Reaction 2)
CH3CH2OH þCH3COOH ¼ 4CH3COOCH2CH3 þH2O, DH1¼ 4.95 kJ/mol
(Reaction 3)
Acetone, CO2, CH4, and CO are formed according to assumptions of following reaction mechanisms as provided by (Blake and Jackson, 1968; Nguyen et al., 1995; Pestman et al., 1997) 2CH3COOH¼ 4CH3COCH3 þCO2 þH2O, DH1¼37.64 kJ/mol
(Reaction 4)
CH3COOH ¼ 4CH4 þCO2, DH1¼15.13 kJ/mol
(Reaction 5)
CH3COOH þH2 ¼ 4CH4 þCOþH2O,
DH1¼12.3 kJ/mol
(Reaction 6)
Based on the analysis of the product and the literature reviews, reaction pathways of HDO of acetic acid are proposed as shown in Fig. 1. The reaction of acetic acid was characterized in terms of conversion, space-time yield (STY), space-time consumption (STC), and yield which are defined as follows: Conversion ð%Þ ¼
Yield ð%Þ ¼
Amount of acetic acid reacted 100% Amount of acetic acid fed
ð1Þ
Amount of products formed 100% Theoretical amount of product that could be formed
ð2Þ Spacetimeyield ðSTYÞ rate of formation of product, g=g cat h ¼
Amount of product formed Amount of catalyst time
ð3Þ
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Fig. 1. Reaction pathways of HDO of acetic acid (Based on literature review and analysis of products).
Fig. 2. Dependence of conversion and yield on temperature. Reaction conditions: 2.07 MPa (300 psig); gas phase: hydrogen (80%) and nitrogen; liquid phase: acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst; liquid flow rate: 8.33 10 10 m3/s (0.05 ml/min).
3. Results and discussions 3.1. Dependence of conversion and yield on temperature In hydrodeoxygenation of acetic acid, temperature was found to be the most important parameter in removing oxygen. A series of experiments was conducted to study the dependence of conversion of liquid phase acetic acid and yield of products on temperature at constant pressure of 300 psig. Residence time was kept constant by varying reactor length (catalyst loading) to compensate the change in gas velocity. The results shown in Fig. 2 indicate that the conversion of acetic acid and yields of acetaldehyde and ethyl acetate increase as temperature increases. It was observed that upto 210 1C only acetaldehyde was formed but as the temperature increased above 210 1C formations of ethyl acetate and ethanol were detected. A significant amount of acetone was detected at 400 1C and 450 1C. A series of experiments was conducted with acetic acid in a vapor phase at atmospheric pressure. Acetic acid was vaporized by heating above its boiling point (118 1C) prior to entering a reactor. The results shown in Fig. 3 indicate that the conversion of vapor phase acetic acid increases to 60 percent as temperature increases from 200 1C to 450 1C compared to 47 percent conversion for liquid phase acetic acid.
A work by Xu et al. on hydrodeoxygenation of acetic acid showed that ethyl acetate was the final product at 200 1C and proposed that ethyl acetate was produced via three step reactions in which acetaldehyde and ethanol were the intermediates (Xu et al., 2009). From analysis of the product we have confirmed that HDO of acetic acid produced all these compounds. Our result also showed the formation of acetone which was not mentioned in the literature by Xu et al. It is also found that the rate of conversion of vapor phase acetic acid at 200 1C is comparable to the literature value mentioned by Xu et al. Analysis of gaseous products showed that methane, carbon dioxide, and carbon monoxide were formed as the temperature increased to 450 1C for catalytic HDO of vapor phase acetic acid as shown in Fig. 4. Another set of experiments was conducted with acetic acid in the vapor phase without catalyst in the reactor. As the temperature increased from 200 1C to 450 1C formation of carbon dioxide, carbon monoxide, and methane was not observed except for a trace amount of carbon dioxide at 400 1C and 450 1C. Therefore, thermal decomposition did not play a role in the formation of carbon dioxide, carbon monoxide, and methane when experiments were conducted with catalytic hydrodeoxygenation upto 450 1C. An increase in conversion as the temperature increases is indicative of HDO of acetic acid being highly influenced by kinetics. As temperature increases more reactants
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Fig. 3. Dependence of conversion and yield on temperature. Reaction conditions: 15 psig; gas phase: hydrogen (80%), nitrogen, and acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst; acetic acid flow rate: 8.33 10 10 m3/s (0.05 ml/min).
Fig. 4. Dependence of gaseous products formation on temperature. Reaction conditions: 15 psig; gas phase: hydrogen (80%), nitrogen, and acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst; acetic acid flow rate: 0.05 ml/min.
collide with high enough energy to overcome activation energy barrier to form products. Furthermore, high conversion of acetic acid can be achieved at atmospheric pressure in contrast to high pressure (70–200 atm.) used by other groups for DHO of bio-oil (Bulushev and Ross, 2011; Furimsky, 2000; Huber et al., 2006). A carbon balance around the reactor showed that 14% of carbon was unaccounted for which was assumed lost as uncondensed acetaldehyde, ethanol, ethyl acetate, and acetone. From an oxygen balance, 97% of the oxygen is accounted for, of which 47% is in carbon dioxide, carbon monoxide, and water. As a result of HDO of acetic acid, oxygen to carbon ratio reduced from 1.33 in the feed to 1.0 in the product. 3.2. Dependence of conversion and yield on hydrogen partial pressure Very high hydrogen pressure (1050 psig–3000 psig) requirement due to low reactivity of some of the oxygenates in pyrolysis oil discourages many to consider hydrodeoxygenation as a viable alternative for biofuel production. In this work, hydrodeoxygenation of acetic acid at low pressure was examined for the conversion of acetic acid. A set of experiments was carried out at 150 1C with liquid phase acetic acid in the range of 300–600 psig total pressures to study the effect of inlet hydrogen partial pressure on conversion, yield and STY. Reaction temperature and residence time were kept constant. The residence time was kept constant by varying reactor length (catalyst loading). The H2 partial pressure was varied by changing total pressure. The results in Fig. 5
indicate that increasing hydrogen partial pressure had no effect on conversion and yield for the selected pressure range, indicating that adsorbed hydrogen on the catalyst surface had reached a maximum (saturated) value. The increase in STY with increase in hydrogen partial pressure was due to an increase in hydrogen concentration (due to constant residence time) resulting in a higher reaction rate. Another set of experiments was carried out at 450 1C with vapor phase acetic acid in the range of 3–15 psig to study the effect of inlet hydrogen partial pressure on conversion and yield. The reactions were carried out at constant temperature and residence time with acetic acid in vapor phase. Partial pressure of hydrogen was varied by changing the composition of hydrogen with nitrogen. The results in Fig. 6 indicate that conversion of acetic acid increases as the partial pressure of H2 increases and reaches a maximum value at a pressure close to 15 psig. The increase in conversion could be due to increase in concentration of hydrogen on the catalyst surface. 3.3. Dependence of conversion and yield on flow rate of acetic acid Concentration of hydrogen on the catalyst surface affects the conversion of acetic acid during HDO. As mass transfer of hydrogen to catalyst surface occurs via liquid slugs, dependence of conversion and yield on flow rate of acetic acid was examined. Experiments were conducted to study the dependence of yield and conversion on flow rate of acetic acid by varying the liquid flow rate from 0.03–0.15 ml/min. Other operating conditions such
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Fig. 5. Dependence of conversion yield and STY on inlet hydrogen partial pressure. Reaction conditions: 150 1C; gas phase: hydrogen (80%) and nitrogen; liquid phase: acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst; liquid flow rate: 0.05 ml/min.
Fig. 6. Dependence of conversion and yield on inlet hydrogen partial pressure. Reaction conditions: 450 1C; gas phase: hydrogen (80%), nitrogen, and acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst; acetic acid flow rate: 0.05 ml/min.
as temperature, pressure and gas flow rate were kept constant. In this experiment, the actual volumetric gas flow rate was at least one order of magnitude higher than the liquid flow rate; hence the residence time was essentially constant. The results shown in Fig. 7 indicate that as the liquid flow rate increases, conversion and yield decrease. During the experiment, a steady increase in liquid slug length was observed as liquid flow rate was increased from 0.03 to 0.15 ml/min. As mass transfer from gas to liquid slugs through hemispherical caps of gas bubble was a strong function of liquid slug length, the convective mass transfer rate of hydrogen to catalyst surface through acetic acid decreased when liquid flow rate was increased (Kreutzer et al., 2001) which was consistent with our findings. Similar experiments were repeated to study the dependence of conversion and yield on flow rate of vapor phase acetic acid. The results shown in Fig. 8 indicate that as the flow rate of acetic acid increases from 0.05 m/s to 0.2 ml/min, the conversion decreases from 60% to 7% which could be due to decrease in concentration of hydrogen on the catalyst surface and decrease in residence time caused by vaporization of acetic acid. 3.4. Dependence of conversion and yield on residence time As reactants must come in contact with catalyst surface simultaneously for reaction to occur, the amount of time spent on the catalyst surface by the reactants affects the conversion. Residence time was varied by increasing reactor length (catalyst
loading) while keeping flow rate, temperature and pressure constant to study the dependence of conversion, yield and STY on it. In microreactors, the residence time is very small compared to conventional macroreactors; however for maximum conversion, an optimum residence time is required. This can be done by increasing reactor length which in turn increases catalyst loading. The residence time was calculated based on empty reactor volume. The result in Fig. 9 shows that the conversion and yield increase as residence time increases. The conversion approaches close to the maximum value at a residence time of 0.06 s for the given reaction conditions. STY decreases as residence time increases due to increase in catalyst loading causing decrease in average reactant concentration per gram of catalyst. Similar set of experiments was conducted with vapor phase acetic acid. The results in Fig. 10 show that the conversion increases and reaches a maximum value of 60% at a residence time of 0.03 s. 3.5. External mass transfer limitation The heterogeneous reaction of acetic acid and hydrogen on the sulfided NiMo/Al2O3 catalyst surface involves transfer of hydrogen into the liquid phase, and diffusion through liquid phase to the catalyst through a boundary layer surrounding the catalyst surface. As mass transfer rates through both gas–liquid and liquid–solid interfaces are affected by flow velocity, the rate of product formation was measured as a function of total (gas and
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Fig. 7. Dependence of yield, STY and conversion on flow rate of acetic acid. Reaction conditions: 2.07 MPa (300 psig), 150 1C; gas phase: hydrogen (80%) and nitrogen; liquid phase: acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst.
Fig. 8. Dependence of yield and conversion on flow rate of acetic acid. Reaction conditions: 15 psig, 450 1C; gas phase: hydrogen (80%), nitrogen, and acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst.
Fig. 9. Dependence of STY, conversion, and yield on residence time. Reaction conditions: 2.07 MPa (300 psig), 150 1C; gas phase: hydrogen (80%) and nitrogen; liquid phase: acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst; liquid flow rate: 8.33 10–10 m3/s (0.05 ml/min).
liquid) superficial velocity, while keeping the residence time constant by varying catalyst bed volume. Fig. 11 shows that, for the range of velocity studied (keeping gas–liquid ratio constant), the superficial velocity had no effect on the space-time yield (STY). This indicates that at the lowest velocity selected for this study, the boundary layer is so thin that it no longer offers any significant resistance to diffusion; hence hydrodeoxygenation of acetic acid is not limited by the external mass transfer for the selected velocities.
3.6. Internal mass transfer limitation As more active metal catalysts remain inside the pores, the rate of mass transfer of reactants from catalyst surface to interiors of the catalyst pores affects the conversion. The effect of internal mass transfer limitation was studied by varying catalyst particle size. Two different particle size ranges of 38–45 mm and 75–150 mm were selected in order to study the effect of catalyst particle size on reaction rate of acetic acid. The reaction rates of
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Fig. 10. Dependence of conversion and yield on residence time. Reaction conditions: 15 psig, 450 1C; gas phase: hydrogen (80%), nitrogen and acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst; acetic acid flow rate: 8.33 10 10 m3/s (0.05 ml/min).
Fig. 11. Dependence of space-time yield (STY) on total fluid flow velocity. Reaction conditions: 2.07 MPa (300 psig), 150 1C; gas phase: hydrogen (80%) and nitrogen; liquid phase: acetic acid (Conc. 99.7%).
acetic acid for particle size ranges of 38–45 mm and 75–150 mm were 0.000184 and 0.000192 g of acetic acid/g catalyst/s respectively which indicate that there was no diffusional mass transfer limitation for the particle size range of 75–150 mm. Internal mass transfer limitation can also be estimated by calculating Thiele modulus for the particle size range of 75–150 mm assuming a pseudo-first-order reaction with respect to hydrogen and acetic acid (Furimsky, 2000; LaVopa and Satterfield, 1987) according to Eq. (4) (Fogler, 2005) where –r0 HAc is the reaction rate, rp is the catalyst particle density, De is the effective diffusivity, which was estimated using the equation De ¼(DAB jp sc)/t where DAB is the binary diffusivity of hydrogen in the liquid reactant. DAB is estimated to be 5.7 10 9 m2/s according to Wilke–Chang equation(Perry and Green, 1997), and the typical values of porosity (jp), constriction factor (sc), and tortuosity (t) used were: 0.4, 0.8, and 3 respectively. The Thiele modulus was estimated to be 0.15 which corresponds to an internal effectiveness factor of unity,(Fogler, 2005) indicating that actual overall rate of reaction was equal to the rate of reaction that would result if entire interior surface were exposed to the external catalyst surface conditions (CAs, Ts). Therefore, it can be concluded that the reaction rate of acetic acid was not limited by internal mass transfer within the catalyst particles. 0:5 d r nr 0 fexp ðThiele ModulusÞ ¼ p p HAc ð4Þ 6 De n C H 2 3.7. Heat transfer limitation A qualitative analysis of radial heat transfer limitation was conducted for liquid phase acetic acid by calculating Damkohler
number (Da) for heat transfer and comparing it to 0:4 RT w =Ea (Mears, 1971) which is defined as: DHðr Þð1 A ÞR2 RT w obs 0 Da ¼ ð5Þ o0:4 lT w ð1 þ bÞ Ea If the above relation holds, the radial temperature difference in the reactor will be less than 5%. In the above equation, the heat of reaction (DH) was 6.06 kJ/mole; the observed reaction rate ðr obs Þ was 10.66 mol/m3 s at temperature (Tw) of 423 K; the radius of the tubular reactor (Ro) was 0.00038 m; the bed porosity (A) was assumed to be 0.3; effective thermal conductivity of porous catalyst (l) was in the order of 0.1 W/m K (Ajmera et al., 2002); and the ratio of diluent to catalyst volume (b) was 0. The activation energy (Ea) for the reaction was 83.74 KJ/mole (Rachmady and Vannice, 2000) and R was the gas constant. Calculation showed that the left hand side in the above equation was two orders of magnitude smaller than the right hand side. Therefore, the radial heat transfer effect in a microreactor is not a factor. 3.8. Dependence of conversion and space-time-yield on reactor diameter A set of experiments was conducted to study the effect of reactor diameter on conversion and STY using liquid phase and gas phase acetic acid as mixing of gas and liquid reactants depends upon the reactor diameter. In this experiment, internal reactor diameters of 0.8 mm, 3.2 mm, and 6.4 mm were used while temperature, pressure, residence time, and superficial velocity were kept constant. The results shown in Figs. 12 and 13 indicate that conversion and STY decreases as the reactor
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769
Fig. 12. Dependence of conversion and STY on reactor diameter. Reaction conditions: 2.07 MPa (300 psig), 150 1C; gas phase: hydrogen (80%) and nitrogen; liquid phase: acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst.
Fig. 13. Dependence of conversion and STY on reactor diameter. Reaction conditions: 15 psig, 450 1C; gas phase: hydrogen (80%), nitrogen, and acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst.
diameter increases for both liquid phase and gas phase acetic acid. Various researches on reactor diameters have also shown similar results. According to Liu et al. experiments conducted on hydrogenations of styrene, 1-octene, and toluene in microreactor channel sized of 1 mm, 1.5 mm, and 2.0 mm showed that the apparent rate constant (hence conversion and rate) increased dramatically with decreasing channel size (Liu et al., 2005). A work by Vandu et al. also showed that mass transfer decreased drastically with increased channel dimension (Vandu et al., 2005). It has been suggested that the increase in mass transfer and hence increase in conversion as reactor diameter decreases is because, decreasing channel size disproportionately enhances the concentration of hydrogen on the catalyst surface. This observation can be explained by intensified gas/liquid mixing inside the smaller channel due to recirculation (Kashid et al., 2007; Liu et al., 2005; Thulasidas et al., 1997; Vandu et al., 2005; Zaloha et al., 2012). Therefore the higher conversion and STY for microreactor with diameter of 0.0008 m ( o1 mm) could be attributed to the higher concentration of hydrogen on the catalyst surface due to intensified recirculation of liquid slugs. Based on the results presented in Figs. 12 and 13, it can be deduced that as reactor diameter is increased, the overall performance of reactor will diminish; hence comparatively, microreactor (IDo1 mm) performs better than larger diameter reactors in terms of conversion and reaction rate.
4. Conclusions The hydrodeoxygenation reaction of acetic acid in a microreactor was studied by evaluating the dependences of various
operating conditions using reduced sulfided NiMo/Al2O3 catalyst. Experimental results indicate that as the temperature increases, more acetic acid is converted to acetaldehyde, ethanol, ethyl acetate, acetone, carbon dioxide, carbon monoxide, and methane. Conversion of vapor phase acetic acid is higher than that of liquid phase acetic acid as temperature increases due to higher rate of diffusion of hydrogen to vapor phase acetic acid. A high conversion of vapor phase acetic acid at atmospheric pressure shows that HDO of acetic acid can be performed without applying extreme pressure, although this may not be realizable for pyrolysis oil. A set of experiments conducted with a blank reactor indicates that thermal decomposition does not occur in the temperature range 200 1C–450 1C. In conclusion, the temperature effect has a major contribution in converting acetic acid to products that are not corrosive and possibly less susceptible to polymerization in the presence of phenolic compounds. Experiment on the effect of hydrogen partial pressure indicates that conversion of liquid phase acetic acid remains constant as the inlet partial pressure of hydrogen increases from 240 psig to 480 psig during hydrodeoxygenation reaction The conversion of vapor phase acetic acid during hydrodeoxygenation increases and reaches a maximum as the partial pressure of hydrogen increases from 3 psig to 15 psig. Studies of the dependence on the flow rate of liquid phase acetic acid show that increasing the flow rate of acetic acid decreases conversion and yield because of decrease in convective mass transfer rate of hydrogen from gas bubble to catalyst surface through acetic acid slug. In the case of vapor phase acetic acid, conversion and yield decrease as the flow rate of acetic acid increases due to decrease in residence time as a result of
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vaporization of acetic acid. The dependence of conversion and yield on the residence time was studied by increasing reactor length. The maximum values of conversion and yield reach at residence time of 0.06 s for liquid acetic acid. At this residence time, the STY obtained was the lowest due to decrease in average reactant concentration per gram of catalyst. Similarly, a maximum conversion reaches at 0.03 s for the vapor phase acetic acid. External mass transfer resistance was found to be negligible in microreactor at overall flow velocity of 2.54 m/s. Similarly, internal mass transfer was found to be negligible at an average particle size of 113 mm under the operating conditions selected in the microreactor. The heat transfer limitation in the microreactor was not a factor as the radial temperature difference in the microreactor was less than 5%. Hence, HDO reaction could be considered as kinetically controlled in the microreactor under the reaction conditions selected which is consistent with the strong temperature effect. The conversion and STY decrease as internal reactor diameter increases because of decreased concentration of hydrogen on catalyst surface due to diminishing rate of recirculation. Based on the products formed during low severity ( o300 1C) HDO of acetic acid, it can be concluded that acetic acid can be converted to products that are less corrosive. This means that by conducting HDO of pyrolysis oil at temperatures below 300 1C, corrosive nature of pyrolysis oil can be reduced. Based on the water formed during HDO of acetic acid, it can be said that oxygen content of acetic acid can be reduced. This implies that low severity (first stage) HDO at temperatures below 300 1C can be accomplished so that condensation reaction can be avoided in turn when HDO of pyrolysis oil is conducted at temperatures above 300 1C. However, low conversion of acetic acid and formations of acetaldehyde, acetone, and ethyl acetate during HDO, specially, at temperatures below 300 1C indicates that pyrolysis oil will still have unsaturated functions after HDO which could still cause condensation reaction during HDO at temperatures above 300 1C. As mentioned in the introduction section, microreactor system could be a suitable alternative in biofuel processing due to bulky nature of biomass which are scattered across the geographical area. However, due to low conversion of functions at low temperatures ( o300 1C), separation and recycling cost could be very high which could potentially upset the benefits of microreactor system. Therefore, improved catalysts need to be screened for conducting DHO for higher conversion of functions which have capability of rupturing carbon–oxygen bonds at temperatures below 300 1C. Identifying and quantifying all the products formed during HDO of pyrolysis oil will be very difficult. Using acetic acid as a model compound, HDO products were analyzed and quantified. Based on the products formed and the literature reviews, reaction pathways have been constructed. These reaction pathways will be valuable for kinetics study. The disadvantage of hydrodeoxygenation of pyrolysis oil is that it requires very high hydrogen pressure (1050–3000 psig). Therefore, many consider hydrodeoxygenation as an unattractive process. In this work, attempt has been made to show that HDO of acetic acid as a model compound of pyrolysis soil can be conducted at atmospheric pressure with a significant conversion achieved.
Nomenclature De
jp sc
Effective diffusivity of hydrogen in the bulk liquid (m2/s) Particle porosity Constriction factor
t jexp dp
rp r0 HAc CH2 Da DH robs Tw Ro A b
l Ea
Tortuosity Thiele modulus Catalyst particle diameter (m) Particle density (kg/m3) Reaction rate of H2 (kmol/kg s) H2 concentration in liquid (kmol/m3) Damkohler number Heat of reaction (kJ/mol) Observed reaction rate (mol/m3 s) Reactor wall temperature (K) Radius of a tubular reactor (m) Bed porosity Ratio of diluent to catalyst volume Effective thermal conductivity (W/m-K) Activation energy (kJ/mol)
Acknowledgments Narendra Joshi gratefully acknowledges support from the Robert C. Stanley Graduate Fellowship Program of Stevens Institute of Technology and GK-12 program of National Science Foundation under grant DGE-0740462. In addition, we are grateful to Dr. James Manganaro for his technical contribution to this work and Steve Mayo (Albemarle) for providing catalysts. References Adeosun, J.T., Lawal, A., 2005. Mass transfer enhancement in microchannel reactors by reorientation of fluid interfaces and stretching. Sens. Actuators B 110, 101–111. Ajmera, S.K., Delattre, C., Schmidt, M.A., Jensen, K.F., 2002. Microreactors for measuring catalyst activity and determining reaction kinetics. Stud. Surf. Sci. Catal., 97–102. Besser, R.S., Ouyang, X., Surangalikar, H., 2003. Hydrocarbon hydrogenation and dehydrogenation reactions in microfabricated catalytic reactors. Chem. Eng. Sci. 58, 19–26. Blake, P.G., Jackson, G.E., 1968. The thermal decomposition of acetic acid. J. Chem. Soc. B. Branebjerg, J., Gravesen, P., Krog, J.P., Nielsen, C.R., 1996. Fast Mixing by Lamination, Micro-Electro Mechanical Systems. MEMS 96 Proceedings. An Investigation of Micro-Structures, Sensors, Actuators, Machines and Systems. IEEE The Ninth Annual International Workshop, pp. 441–446. Bulushev, D.A., Ross, J.R.H., 2011. Catalysis for conversion of biomass to fuels via pyrolysis and gasification: a review. Catal. Today 171, 1–13. David, R., 2008. The price of biofuels. Technol. Rev. 111, 42. Ehrfeld, W., Hessel, V., Lowe, H., 2000. Microreactors: New Technology for Modern Chemistry. John Wiley & Sons. Elliott, D.C., 2007. Historical developments in hydroprocessing bio-oils. Energy Fuels 21, 1792–1815. Elliott, D.C., Baker, E.G., Piskorz, J., Scott, D.S., Solantausta, Y., 1988. Production of liquid hydrocarbon fuels from peat. Energy Fuels 2, 234–235. Elliott, D.C., Neuenschwander, G.G., 1996. Liquid fuels by low-severity hydrotreating of biocruide. Dev. Thermochem. Biomass Convers. 1, 611–621. Elliott, D.C., Oasmaa, A., 1991. Catalytic hydrotreating of black liquor oils. Energy Fuels 5, 102–109. Fogler, H.S., 2005. Elements of chemical reaction engineering, Fourth ed. Prentice Hall, Professional Technical Reference, Massachusetts, USA, pp 830. Furimsky, E., 2000. Catalytic hydrodeoxygenation. Appl. Catal. A 199, 147–190. Grange, P., Laurent, E., Maggi, R., Centeno, A., Delmon, B., 1996. Hydrotreatment of pyrolysis oils from biomass: reactivity of the various categories of oxygenated compounds and preliminary techno-economical study. Catal. Today 29, 297–301. Grootendorst, E.J., Pestman, R., Koster, R.M., Ponec, V., 1994. Selective reduction of acetic acid to acetaldehyde on iron oxides. J. Catal. 148, 261–269. Halder, R., Lawal, A., Damavarapu, R., 2007. Nitration of toluene in a microreactor. Catal. Today 125, 74–80. Hessel, V.E., Ehrfeld, W., Lowe, H., 2001. Microreactors. Wiley-VCH, Weinheim. Huber, G.W., Iborra, S., Corma, A., 2006. Synthesis of transportation fuels from biomass: chemistry, catalysts, and engineering. Chem. Rev. 106, 4044–4098. Jenck, J.F., 2009. Impact of Microtechnologies on Chemical Processing. Wiley New Jersey, USA. Kashid, M.N., Platte, F., Agar, D.W., Turek, S., 2007. Computational modelling of slug flow in a capillary microreactor. J. Comput. Appl. Math. 203, 487–497. Knight, J.B., Vishwanath, A., Brody, J.P., Austin, R.H., 1998. Hydrodynamic focusing on a silicon chip: mixing nanoliters in microseconds. Phys. Rev. Lett. 80, 3863–3866.
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Contents lists available at SciVerse ScienceDirect
Chemical Engineering Science journal homepage: www.elsevier.com/locate/ces
Hydrodeoxygenation of acetic acid in a microreactor Narendra Joshi n, Adeniyi Lawal New Jersey Center for Micro-Chemical Systems, Department of Chemical Engineering and Materials Science, Stevens Institute of Technology, Hoboken, NJ 07030, USA
H I G H L I G H T S c c c c c
Acetic acid can be converted by HDO below 300 1C at atmospheric pressure. Reaction pathways have been constructed based on literature reviews and product analysis. External mass transfer resistance was negligible at overall flow velocity of 2.54 m/s. Internal mass transfer resistance was negligible at an average particle size of 113 mm. The conversion of acetic acid decreases as internal reactor diameter increases.
a r t i c l e i n f o
abstract
Article history: Received 5 May 2012 Received in revised form 22 September 2012 Accepted 24 September 2012 Available online 1 October 2012
Acetic acid was used as a model compound for pyrolysis oil in a hydrodeoxygenation (HDO) study. HDO of acetic acid was performed in a packed bed microreactor. The catalyst was reduced sulfided NiMo/ Al2O3. The effects of state of aggregation of acetic acid, temperature, hydrogen partial pressure, liquid flow rate, reactor diameter, and residence time on conversion, yield, space-time consumption, and space-time yield were investigated. External and internal mass transfer and heat transfer resistances were also examined in the microreactor. Temperature was a major factor in HDO of acetic acid. Many consider hydrodeoxygenation as an unattractive process due to high pressure requirements (1050–3000 psig). In this work, attempt has been made to show that HDO of acetic acid can be conducted at atmospheric pressure with a significant conversion achieved. More acetic acid was converted during HDO as temperature was increased at constant pressure of 300 psig. Conversion was much higher for vapor phase acetic acid at atmospheric pressure than liquid phase acetic acid. HDO of gas phase acetic acid in a blank reactor compared to a catalytic HDO showed that thermal decomposition of acetic acid did not occur appreciably. Partial pressure of hydrogen above 240 psig had no effect on the conversion of liquid phase acetic acid. Conversion of vapor phase acetic acid increased as the partial pressure of hydrogen increased from 3 psig to 15 psig. Residence time was 0.06 s for a maximum conversion of liquid phase acetic acid, whereas it was 0.03 s for a maximum conversion of vapor phase acetic acid. The conversion of acetic acid for both liquid and vapor phases decreases significantly as the flow rate of acetic acid increases. As reactor diameter increases beyond 0.8 mm, the conversion reduces significantly. Mass transfer resistance was negligible at the superficial velocity of 2.54 m/s and at an average catalyst particle size of 113 mm. Radial temperature difference in the microreactor was less than 5%. & 2012 Elsevier Ltd. All rights reserved.
Keywords: Microreactor Heat transfer Fuel Energy Mass transfer Hydrodeoxygenation
1. Introduction Depleting petroleum reserves, rising prices, and environmental and political concerns have made renewable resources for transportation fuel attractive. Though there are other renewable sources of energy such as wind, solar, and hydroelectric, biomass is the only renewable source of energy that can be used for liquid fuel, fitting into the current infrastructure of transportation fuel utilization.
n
Corresponding author. Tel.: þ1 201 216 8314; fax: þ1 201 216 8306. E-mail addresses: [email protected], [email protected] (N. Joshi).
0009-2509/$ - see front matter & 2012 Elsevier Ltd. All rights reserved. http://dx.doi.org/10.1016/j.ces.2012.09.018
Effective utilization of biomass as the transportation fuel may reduce the world’s dependency on non-renewable fossil fuel. Another important advantage of the biomass is that it contributes no new carbon dioxide to the atmosphere (Peter, 2002). Raw pyrolysis oil from biomass can be further processed to obtain transportation fuel. Processing alternatives include hydrodeoxygenation, catalytic cracking, and steam reforming followed by Fischer Tropsch synthesis. Lately more attention has focused on hydrodeoxygenation as a means of increasing the energy density of the product fuel. Hydrodeoxygenation of pyrolysis oil requires about 450 1C to reduce most of the oxygen content in the oil. But at this temperature an extreme degradation of oil takes
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place and it becomes a very hard, coke like material. This chemical instability is attributed to unsaturated double bonds such as acetic acid, aldehyde, ketone, and so forth which react through condensation reaction similar to the phenol-formal polymerization (Grange et al., 1996). Therefore, it is very crucial to eliminate these functions before they react to high molecular weight compounds. Experiments with oxygenated model compounds show that it can be accomplished by a low temperature (below 300 1C) hydrotreatment (Elliott and Neuenschwander, 1996; Elliott et al., 1988; Elliott and Oasmaa, 1991; Grange et al., 1996). High acetic acid content (upto 25%) provides corrosive nature to pyrolysis oil (Milne et al., 1997; Xu et al., 2009) which needs to be reduced for long term storage. It also contains unsaturated double bond functional group which leads to polymerization of pyrolysis oil at higher temperature (above 300 1C) during hydrodeoxygenation. Therefore, converting acetic acid by hydrodeoxygenation at temperatures below 300 1C will increase the stability of pyrolysis oil and in turn, it is expected to reduce polymerization of pyrolysis oil significantly when hydrotreated at temperatures above 300 1C. By conducting hydrodeoxygenation of acetic acid we hope that it will provide an opportunity to investigate upgrading mechanism. Using acetic acid as a model compound we also hope to study reaction pathways of HDO of acetic acid. An understanding of reaction pathways and kinetics is expected to be of value in process design and modeling (LaVopa and Satterfield, 1987). However, studying reaction pathways and kinetics of HDO of PO is very difficult as it contains more than 300 oxygenated compounds. Therefore, acetic acid is selected as our model compound to investigate the hydrodeoxygenation process and reaction pathways. Biomass is a bulky material with a very low density and scattered across geographical areas. Research has shown that transporting biomass to a central location and processing for biofuel is not economically feasible due to very high cost of transportation. The research suggest that biofuel production facilities will have to be built close to the sources of feedstock, probably within 50 miles (David, 2008); therefore, a large scale processing plant will not be profitable. Even for a medium size biomass conversion facility, the amount of biomass required is overwhelming. It is calculated that a facility producing 50 million gallon biofuel per year would require a truck loaded with biomass to arrive every six minutes around the clock (David, 2008). The requirements of high production cost and very large quantity of biomass discourages people to use the biomass conversion facility with conventional macroreactor system. However, on-site and on-demand distributed system could be an alternative solution. In this regard, microreactor system could be a suitable alternative. Microreactor system consists of devices that are miniature in sizes and scaling up is done by increasing number of channels within a reactor unit as well as by increasing number of reactor units. The conventional method of commercializing a new process starts at the laboratory scale where results are collected which are then used in a pilot plant to obtain critical parameters. The information collected from pilot plant is used to design and build the larger production unit. In microreactor technology, scale-up can be achieved by numbering up straight from laboratory scale bypassing a costly pilot plant (Jenck, 2009). There are number of other benefits of using microreactor system. Study showed that heat transfer coefficient reached upto 25,000 W/m2 K in microdevices exceeding those of conventional heat exchangers by at least one order of magnitude (Schubert et al., 1998). In micromixers typical fluid layer thickness can be set to a few tens of micrometers, consequently, mixing time in micromixers amount to milliseconds which is hardly achievable using stirring equipment and other conventional mixers (Branebjerg et al., 1996; Knight et al., 1998). A comparison of microreactor with
conventional trickle-bed reactor for the hydrogenation of styrene showed that the throughput of the liquid for microreactor is eight times higher than that of the trickle-bed reactor (Nijhuis et al., 2003). As a result of enhanced mass transfer, the hydrogen concentration on catalyst is significantly higher than on the trickle-bet catalyst (34 mol/m3 compared to 22 mol/m3). Microreactors possess ultra-low transport resistances, therefore mass diffusion and heat transfer are extremely fast, resulting in rapid thermal and reaction equilibrium (Besser et al., 2003). High yield, improved product quality, better selectivity and safe operation are attainable due to very high heat and mass transfer (Halder et al., 2007; Okafor et al., 2010; Tadepalli et al., 2007; Voloshin and Lawal, 2010). Mixing in the microchannels is attainable only by inter-diffusion of reactants due to laminar flow (Adeosun and Lawal, 2005), however due to short transverse diffusional distance, rapid and effective mixing is attainable in the microreactor which can quickly bring reactants in contact with catalyst in a heterogeneous reaction (Hessel et al., 2001). Enhancement of mixing in two-phase flows can also be achieved by selecting appropriate inlet T-orientations to provide a short slug length. Studies have shown that introduction of gas and liquid feeds head to head or perpendicular to each other with the liquid stream parallel to the microchannel markedly improves the mixing (Qian and Lawal, 2006). There are other potential benefits of microreactors regarding application which are listed below (Ehrfeld et al., 2000) 1. 2. 3. 4. 5.
Earlier start of production at lower costs. Easier scale-up of production capacity. Smaller plant size for distributed production. Lower costs for transportation, materials and energy. More flexible response to market demands.
Many researches have considered microreactor technology as a potential for process intensification. Process intensification involves the development of innovative methods and devices that, in comparison to existing approaches, offer the chance of a dramatic improvement in the quality of production, substantial reduction in the ratio of equipment size to production capacity, and significant drop in the consumption of energy and production of waste. For these reasons, microstructured devices and components have an important role to play (Matlosz et al., 2009). Commercially available catalysts such as CoO/MoO3 and NiO/MoO3 on Al2O3 support used for removing sulfur, nitrogen, and oxygen from petrochemical feedstock are generally selected for hydrodeoxygenation process as well. According to research conducted at Pacific Northwest National Laboratory, the sulfided form of CoO/MoO3 and NiO/MoO3 are much more active for hydrodeoxygenation than the oxide form (Elliott, 2007). Sulfidation creates active sites that can play a role in the rupture of carbon-heteroatom bond (Senol, 2007). In this study we used the sulfided form of NiO/MoO3 on Al2O3. Model compounds are chosen instead of pyrolysis oil for better understanding and control of the HDO reaction process. Nimmanwudipong et al. used 2-methoxyphenol as a model compound of lignin-derived pyrolysis oil in the presence of hydrogen to elucidate the reaction network and to predict oxygen removal (Nimmanwudipong et al., 2011). The authors have shown that the catalytic conversion of 2-methoxyphenol in the presence of hydrogen catalyzed by Pt/Al2O3 involves three major classes of reactions: hydrogenation, hydrodeoxygenation, and transalkylation. Xu et al. has demonstrated that acetic acid can be converted to ethyl acetate using reduced Mo-10Ni/g-Al2O3 via hydrodeoxygenation at 473 K and 3 MPa hydrogen pressure (Xu et al., 2009). They have proposed that ethyl acetate was produced from acetic acid via a three step reaction in which produced aldehyde was
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converted to ethyl alcohol by hydrogenation which then reacted with acetic acid to form ethyl acetate. Similarly, LaVopa and Satterfield studied the hydrodeoxygenation of dibenzofuran on a sulfided NiMo/Al2O3 at 360 1C and 7.0 MPa and found that singlering cyclohexane predominated (LaVopa and Satterfield, 1987). They also found that the catalyst in the oxide form had lower activity and double-ring products predominated over single-ring compounds. Also, Li et al. concluded that during hydrodeoxygenation of 1-Naphthol catalyzed by sulfided Ni–Mo/g-Al2O3 aromatic ring hydrogenation and direct oxygen extrusion (HDO) occur in parallel (Li et al., 1985). The objective of the research presented here was to evaluate hydrodeoxygenation of acetic acid in a packed bed microreactor. Effects of various reaction variables such as liquid flow velocity, residence time, hydrogen partial pressure, temperature, state of aggregation of acetic acid, and reactor diameter on conversion, yield, and space-time yield were studied.
2. Experimental 2.1. Materials Presulfided NiO/MoO3/Al2O3 catalyst obtained from Albemarle (sulfided and supplied by Eurecat USA), Houston, Texas was ground and sieved to obtain particles with diameters in the range of 75–150 mm. An average surface area of the sieved catalysts was ˚ The surface area 164 m2/g and average pore diameter was 106 A. and pore diameter were obtained by using multipoint BET technique and the instrument used was Quantochrome Autosorb-1. The catalyst was reduced with 5.0 sccm of hydrogen at 593 K and 3.45 MPa for 2 h. The average surface area of the reduced catalysts ˚ ACS regent was 209.0 m2/g and average pore diameter was 92.0 A. grade (conc. 99.7%) acetic acid was purchased from Pharmco Inc. The gas used was extra dry hydrogen from Praxair. Nitrogen was used as tracer to perform a material balance.
2.2. Experimental setup A HPLC pump (Laboratory Alliance Series III) was used to control the flow rate of acetic acid. Mass flow controllers (Porter Model 201) were used to control the flow rates of hydrogen and nitrogen. Ranges of superficial velocities of acetic acid, hydrogen and nitrogen gases used were 0.0011–0.0065 m/s, 0.54–4.71 m/s, and 0.36–2.75 m/s respectively. The liquid and gas phases were combined in a T-junction mixer (Upchurch) with 508 10 6 m through-hole. Reynolds number for the combined flow was less than 100 for all experiments indicating laminar flow. The fluids exiting from the T-junction exhibited a Taylor flow pattern with a liquid slug length in the range 0.001–0.003 m, whereas gas bubble length varied from 0.001 to 0.005 m. A microreactor was prepared from a 0.0016 m (1/16 in.) 316 stainless steel tubing with 762 10 6 m internal diameter, and was gravity filled with catalyst. The total length of the packed bed microreactor varied from 0.025 to 0.18 m. Hastelloy micron filter-cloth (200 1150 mesh, Unique Wire Weaving Co., Hillside, New Jersey) was placed at the ends of the reactor to retain the catalyst. The reactor system was pressurized using a back pressure regulator (GO Regulator Co.). The entrance and the exit pressures of the fluids (liquid and gas combined) in the reactor were measured. The pressure drop along the reactor varied from 0.07 MPa to 0.2 MPa depending upon reactor length. Acetic acid was vaporized prior to entering to a reactor using a vaporizer heated to a temperature above the boiling point (118 1C) of acetic acid. The product stream was condensed using a chiller (model NESLAB RTE 7, Thermo Fisher Scientific).
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2.3. Analysis The water content in the HDO product was analyzed using Volumetric Karl Fischer (KF) Titration Workstation (Model 375, Denver Instrument) with the use of hydranal reagents obtained from Sigma-Aldrich. The titrant used was hydranal titrant-2 while the working medium was hydranal solvent. Before the actual titration, the hydranal solvent was titrated to dryness in a drift determination step. This step removes moisture from solvent, electrode, and titration vessel. The HDO product was added using a syringe through the septum on the KF cell. Analysis of liquid product was conducted in a HPLC (Shimadzu series: mobile phase degasser [DGU_20A5], pump station [LC-20AT], auto-sampler [SIL-20AC], and reflective index detector [RID-10A]) equipped with BioRad Aminex HPX-87H column. The mobile phase consisted of 0.007 N aqueous H3PO4 with an isocratic flow (flow rate of 6 10 7 m3/min). The gas phase was analyzed for CO2, CO, and CH4 using Varian 450GC with Hayesep N and Mol sieve 5A columns in series. HDO of acetic acid involves complex reactions consisting of a series of reactions forming acetaldehyde, ethanol, and ethyl acetate and parallel reactions forming acetone, carbon dioxide, carbon monoxide, and methane. Acetic acid reacts with hydrogen forming acetaldehyde and H2O. A model for the formation of acetaldehyde is suggested by (Grootendorst et al., 1994) with following pathway: CH3COOH þH2 ¼ 4CH3CHOþH2O, DH1¼1.27 kJ/mol
(Reaction 1)
and (Xu et al., 2009) suggested the formation of ethyl acetate with following reaction mechanism: CH3CHO þH2 ¼ 4CH3CH2OH, DH1 ¼ 80.98 kJ/mol
(Reaction 2)
CH3CH2OH þCH3COOH ¼ 4CH3COOCH2CH3 þH2O, DH1¼ 4.95 kJ/mol
(Reaction 3)
Acetone, CO2, CH4, and CO are formed according to assumptions of following reaction mechanisms as provided by (Blake and Jackson, 1968; Nguyen et al., 1995; Pestman et al., 1997) 2CH3COOH¼ 4CH3COCH3 þCO2 þH2O, DH1¼37.64 kJ/mol
(Reaction 4)
CH3COOH ¼ 4CH4 þCO2, DH1¼15.13 kJ/mol
(Reaction 5)
CH3COOH þH2 ¼ 4CH4 þCOþH2O,
DH1¼12.3 kJ/mol
(Reaction 6)
Based on the analysis of the product and the literature reviews, reaction pathways of HDO of acetic acid are proposed as shown in Fig. 1. The reaction of acetic acid was characterized in terms of conversion, space-time yield (STY), space-time consumption (STC), and yield which are defined as follows: Conversion ð%Þ ¼
Yield ð%Þ ¼
Amount of acetic acid reacted 100% Amount of acetic acid fed
ð1Þ
Amount of products formed 100% Theoretical amount of product that could be formed
ð2Þ Spacetimeyield ðSTYÞ rate of formation of product, g=g cat h ¼
Amount of product formed Amount of catalyst time
ð3Þ
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Fig. 1. Reaction pathways of HDO of acetic acid (Based on literature review and analysis of products).
Fig. 2. Dependence of conversion and yield on temperature. Reaction conditions: 2.07 MPa (300 psig); gas phase: hydrogen (80%) and nitrogen; liquid phase: acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst; liquid flow rate: 8.33 10 10 m3/s (0.05 ml/min).
3. Results and discussions 3.1. Dependence of conversion and yield on temperature In hydrodeoxygenation of acetic acid, temperature was found to be the most important parameter in removing oxygen. A series of experiments was conducted to study the dependence of conversion of liquid phase acetic acid and yield of products on temperature at constant pressure of 300 psig. Residence time was kept constant by varying reactor length (catalyst loading) to compensate the change in gas velocity. The results shown in Fig. 2 indicate that the conversion of acetic acid and yields of acetaldehyde and ethyl acetate increase as temperature increases. It was observed that upto 210 1C only acetaldehyde was formed but as the temperature increased above 210 1C formations of ethyl acetate and ethanol were detected. A significant amount of acetone was detected at 400 1C and 450 1C. A series of experiments was conducted with acetic acid in a vapor phase at atmospheric pressure. Acetic acid was vaporized by heating above its boiling point (118 1C) prior to entering a reactor. The results shown in Fig. 3 indicate that the conversion of vapor phase acetic acid increases to 60 percent as temperature increases from 200 1C to 450 1C compared to 47 percent conversion for liquid phase acetic acid.
A work by Xu et al. on hydrodeoxygenation of acetic acid showed that ethyl acetate was the final product at 200 1C and proposed that ethyl acetate was produced via three step reactions in which acetaldehyde and ethanol were the intermediates (Xu et al., 2009). From analysis of the product we have confirmed that HDO of acetic acid produced all these compounds. Our result also showed the formation of acetone which was not mentioned in the literature by Xu et al. It is also found that the rate of conversion of vapor phase acetic acid at 200 1C is comparable to the literature value mentioned by Xu et al. Analysis of gaseous products showed that methane, carbon dioxide, and carbon monoxide were formed as the temperature increased to 450 1C for catalytic HDO of vapor phase acetic acid as shown in Fig. 4. Another set of experiments was conducted with acetic acid in the vapor phase without catalyst in the reactor. As the temperature increased from 200 1C to 450 1C formation of carbon dioxide, carbon monoxide, and methane was not observed except for a trace amount of carbon dioxide at 400 1C and 450 1C. Therefore, thermal decomposition did not play a role in the formation of carbon dioxide, carbon monoxide, and methane when experiments were conducted with catalytic hydrodeoxygenation upto 450 1C. An increase in conversion as the temperature increases is indicative of HDO of acetic acid being highly influenced by kinetics. As temperature increases more reactants
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Fig. 3. Dependence of conversion and yield on temperature. Reaction conditions: 15 psig; gas phase: hydrogen (80%), nitrogen, and acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst; acetic acid flow rate: 8.33 10 10 m3/s (0.05 ml/min).
Fig. 4. Dependence of gaseous products formation on temperature. Reaction conditions: 15 psig; gas phase: hydrogen (80%), nitrogen, and acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst; acetic acid flow rate: 0.05 ml/min.
collide with high enough energy to overcome activation energy barrier to form products. Furthermore, high conversion of acetic acid can be achieved at atmospheric pressure in contrast to high pressure (70–200 atm.) used by other groups for DHO of bio-oil (Bulushev and Ross, 2011; Furimsky, 2000; Huber et al., 2006). A carbon balance around the reactor showed that 14% of carbon was unaccounted for which was assumed lost as uncondensed acetaldehyde, ethanol, ethyl acetate, and acetone. From an oxygen balance, 97% of the oxygen is accounted for, of which 47% is in carbon dioxide, carbon monoxide, and water. As a result of HDO of acetic acid, oxygen to carbon ratio reduced from 1.33 in the feed to 1.0 in the product. 3.2. Dependence of conversion and yield on hydrogen partial pressure Very high hydrogen pressure (1050 psig–3000 psig) requirement due to low reactivity of some of the oxygenates in pyrolysis oil discourages many to consider hydrodeoxygenation as a viable alternative for biofuel production. In this work, hydrodeoxygenation of acetic acid at low pressure was examined for the conversion of acetic acid. A set of experiments was carried out at 150 1C with liquid phase acetic acid in the range of 300–600 psig total pressures to study the effect of inlet hydrogen partial pressure on conversion, yield and STY. Reaction temperature and residence time were kept constant. The residence time was kept constant by varying reactor length (catalyst loading). The H2 partial pressure was varied by changing total pressure. The results in Fig. 5
indicate that increasing hydrogen partial pressure had no effect on conversion and yield for the selected pressure range, indicating that adsorbed hydrogen on the catalyst surface had reached a maximum (saturated) value. The increase in STY with increase in hydrogen partial pressure was due to an increase in hydrogen concentration (due to constant residence time) resulting in a higher reaction rate. Another set of experiments was carried out at 450 1C with vapor phase acetic acid in the range of 3–15 psig to study the effect of inlet hydrogen partial pressure on conversion and yield. The reactions were carried out at constant temperature and residence time with acetic acid in vapor phase. Partial pressure of hydrogen was varied by changing the composition of hydrogen with nitrogen. The results in Fig. 6 indicate that conversion of acetic acid increases as the partial pressure of H2 increases and reaches a maximum value at a pressure close to 15 psig. The increase in conversion could be due to increase in concentration of hydrogen on the catalyst surface. 3.3. Dependence of conversion and yield on flow rate of acetic acid Concentration of hydrogen on the catalyst surface affects the conversion of acetic acid during HDO. As mass transfer of hydrogen to catalyst surface occurs via liquid slugs, dependence of conversion and yield on flow rate of acetic acid was examined. Experiments were conducted to study the dependence of yield and conversion on flow rate of acetic acid by varying the liquid flow rate from 0.03–0.15 ml/min. Other operating conditions such
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Fig. 5. Dependence of conversion yield and STY on inlet hydrogen partial pressure. Reaction conditions: 150 1C; gas phase: hydrogen (80%) and nitrogen; liquid phase: acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst; liquid flow rate: 0.05 ml/min.
Fig. 6. Dependence of conversion and yield on inlet hydrogen partial pressure. Reaction conditions: 450 1C; gas phase: hydrogen (80%), nitrogen, and acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst; acetic acid flow rate: 0.05 ml/min.
as temperature, pressure and gas flow rate were kept constant. In this experiment, the actual volumetric gas flow rate was at least one order of magnitude higher than the liquid flow rate; hence the residence time was essentially constant. The results shown in Fig. 7 indicate that as the liquid flow rate increases, conversion and yield decrease. During the experiment, a steady increase in liquid slug length was observed as liquid flow rate was increased from 0.03 to 0.15 ml/min. As mass transfer from gas to liquid slugs through hemispherical caps of gas bubble was a strong function of liquid slug length, the convective mass transfer rate of hydrogen to catalyst surface through acetic acid decreased when liquid flow rate was increased (Kreutzer et al., 2001) which was consistent with our findings. Similar experiments were repeated to study the dependence of conversion and yield on flow rate of vapor phase acetic acid. The results shown in Fig. 8 indicate that as the flow rate of acetic acid increases from 0.05 m/s to 0.2 ml/min, the conversion decreases from 60% to 7% which could be due to decrease in concentration of hydrogen on the catalyst surface and decrease in residence time caused by vaporization of acetic acid. 3.4. Dependence of conversion and yield on residence time As reactants must come in contact with catalyst surface simultaneously for reaction to occur, the amount of time spent on the catalyst surface by the reactants affects the conversion. Residence time was varied by increasing reactor length (catalyst
loading) while keeping flow rate, temperature and pressure constant to study the dependence of conversion, yield and STY on it. In microreactors, the residence time is very small compared to conventional macroreactors; however for maximum conversion, an optimum residence time is required. This can be done by increasing reactor length which in turn increases catalyst loading. The residence time was calculated based on empty reactor volume. The result in Fig. 9 shows that the conversion and yield increase as residence time increases. The conversion approaches close to the maximum value at a residence time of 0.06 s for the given reaction conditions. STY decreases as residence time increases due to increase in catalyst loading causing decrease in average reactant concentration per gram of catalyst. Similar set of experiments was conducted with vapor phase acetic acid. The results in Fig. 10 show that the conversion increases and reaches a maximum value of 60% at a residence time of 0.03 s. 3.5. External mass transfer limitation The heterogeneous reaction of acetic acid and hydrogen on the sulfided NiMo/Al2O3 catalyst surface involves transfer of hydrogen into the liquid phase, and diffusion through liquid phase to the catalyst through a boundary layer surrounding the catalyst surface. As mass transfer rates through both gas–liquid and liquid–solid interfaces are affected by flow velocity, the rate of product formation was measured as a function of total (gas and
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Fig. 7. Dependence of yield, STY and conversion on flow rate of acetic acid. Reaction conditions: 2.07 MPa (300 psig), 150 1C; gas phase: hydrogen (80%) and nitrogen; liquid phase: acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst.
Fig. 8. Dependence of yield and conversion on flow rate of acetic acid. Reaction conditions: 15 psig, 450 1C; gas phase: hydrogen (80%), nitrogen, and acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst.
Fig. 9. Dependence of STY, conversion, and yield on residence time. Reaction conditions: 2.07 MPa (300 psig), 150 1C; gas phase: hydrogen (80%) and nitrogen; liquid phase: acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst; liquid flow rate: 8.33 10–10 m3/s (0.05 ml/min).
liquid) superficial velocity, while keeping the residence time constant by varying catalyst bed volume. Fig. 11 shows that, for the range of velocity studied (keeping gas–liquid ratio constant), the superficial velocity had no effect on the space-time yield (STY). This indicates that at the lowest velocity selected for this study, the boundary layer is so thin that it no longer offers any significant resistance to diffusion; hence hydrodeoxygenation of acetic acid is not limited by the external mass transfer for the selected velocities.
3.6. Internal mass transfer limitation As more active metal catalysts remain inside the pores, the rate of mass transfer of reactants from catalyst surface to interiors of the catalyst pores affects the conversion. The effect of internal mass transfer limitation was studied by varying catalyst particle size. Two different particle size ranges of 38–45 mm and 75–150 mm were selected in order to study the effect of catalyst particle size on reaction rate of acetic acid. The reaction rates of
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Fig. 10. Dependence of conversion and yield on residence time. Reaction conditions: 15 psig, 450 1C; gas phase: hydrogen (80%), nitrogen and acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst; acetic acid flow rate: 8.33 10 10 m3/s (0.05 ml/min).
Fig. 11. Dependence of space-time yield (STY) on total fluid flow velocity. Reaction conditions: 2.07 MPa (300 psig), 150 1C; gas phase: hydrogen (80%) and nitrogen; liquid phase: acetic acid (Conc. 99.7%).
acetic acid for particle size ranges of 38–45 mm and 75–150 mm were 0.000184 and 0.000192 g of acetic acid/g catalyst/s respectively which indicate that there was no diffusional mass transfer limitation for the particle size range of 75–150 mm. Internal mass transfer limitation can also be estimated by calculating Thiele modulus for the particle size range of 75–150 mm assuming a pseudo-first-order reaction with respect to hydrogen and acetic acid (Furimsky, 2000; LaVopa and Satterfield, 1987) according to Eq. (4) (Fogler, 2005) where –r0 HAc is the reaction rate, rp is the catalyst particle density, De is the effective diffusivity, which was estimated using the equation De ¼(DAB jp sc)/t where DAB is the binary diffusivity of hydrogen in the liquid reactant. DAB is estimated to be 5.7 10 9 m2/s according to Wilke–Chang equation(Perry and Green, 1997), and the typical values of porosity (jp), constriction factor (sc), and tortuosity (t) used were: 0.4, 0.8, and 3 respectively. The Thiele modulus was estimated to be 0.15 which corresponds to an internal effectiveness factor of unity,(Fogler, 2005) indicating that actual overall rate of reaction was equal to the rate of reaction that would result if entire interior surface were exposed to the external catalyst surface conditions (CAs, Ts). Therefore, it can be concluded that the reaction rate of acetic acid was not limited by internal mass transfer within the catalyst particles. 0:5 d r nr 0 fexp ðThiele ModulusÞ ¼ p p HAc ð4Þ 6 De n C H 2 3.7. Heat transfer limitation A qualitative analysis of radial heat transfer limitation was conducted for liquid phase acetic acid by calculating Damkohler
number (Da) for heat transfer and comparing it to 0:4 RT w =Ea (Mears, 1971) which is defined as: DHðr Þð1 A ÞR2 RT w obs 0 Da ¼ ð5Þ o0:4 lT w ð1 þ bÞ Ea If the above relation holds, the radial temperature difference in the reactor will be less than 5%. In the above equation, the heat of reaction (DH) was 6.06 kJ/mole; the observed reaction rate ðr obs Þ was 10.66 mol/m3 s at temperature (Tw) of 423 K; the radius of the tubular reactor (Ro) was 0.00038 m; the bed porosity (A) was assumed to be 0.3; effective thermal conductivity of porous catalyst (l) was in the order of 0.1 W/m K (Ajmera et al., 2002); and the ratio of diluent to catalyst volume (b) was 0. The activation energy (Ea) for the reaction was 83.74 KJ/mole (Rachmady and Vannice, 2000) and R was the gas constant. Calculation showed that the left hand side in the above equation was two orders of magnitude smaller than the right hand side. Therefore, the radial heat transfer effect in a microreactor is not a factor. 3.8. Dependence of conversion and space-time-yield on reactor diameter A set of experiments was conducted to study the effect of reactor diameter on conversion and STY using liquid phase and gas phase acetic acid as mixing of gas and liquid reactants depends upon the reactor diameter. In this experiment, internal reactor diameters of 0.8 mm, 3.2 mm, and 6.4 mm were used while temperature, pressure, residence time, and superficial velocity were kept constant. The results shown in Figs. 12 and 13 indicate that conversion and STY decreases as the reactor
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Fig. 12. Dependence of conversion and STY on reactor diameter. Reaction conditions: 2.07 MPa (300 psig), 150 1C; gas phase: hydrogen (80%) and nitrogen; liquid phase: acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst.
Fig. 13. Dependence of conversion and STY on reactor diameter. Reaction conditions: 15 psig, 450 1C; gas phase: hydrogen (80%), nitrogen, and acetic acid (Conc. 99.7%); sulfided NiMo/Al2O3 catalyst.
diameter increases for both liquid phase and gas phase acetic acid. Various researches on reactor diameters have also shown similar results. According to Liu et al. experiments conducted on hydrogenations of styrene, 1-octene, and toluene in microreactor channel sized of 1 mm, 1.5 mm, and 2.0 mm showed that the apparent rate constant (hence conversion and rate) increased dramatically with decreasing channel size (Liu et al., 2005). A work by Vandu et al. also showed that mass transfer decreased drastically with increased channel dimension (Vandu et al., 2005). It has been suggested that the increase in mass transfer and hence increase in conversion as reactor diameter decreases is because, decreasing channel size disproportionately enhances the concentration of hydrogen on the catalyst surface. This observation can be explained by intensified gas/liquid mixing inside the smaller channel due to recirculation (Kashid et al., 2007; Liu et al., 2005; Thulasidas et al., 1997; Vandu et al., 2005; Zaloha et al., 2012). Therefore the higher conversion and STY for microreactor with diameter of 0.0008 m ( o1 mm) could be attributed to the higher concentration of hydrogen on the catalyst surface due to intensified recirculation of liquid slugs. Based on the results presented in Figs. 12 and 13, it can be deduced that as reactor diameter is increased, the overall performance of reactor will diminish; hence comparatively, microreactor (IDo1 mm) performs better than larger diameter reactors in terms of conversion and reaction rate.
4. Conclusions The hydrodeoxygenation reaction of acetic acid in a microreactor was studied by evaluating the dependences of various
operating conditions using reduced sulfided NiMo/Al2O3 catalyst. Experimental results indicate that as the temperature increases, more acetic acid is converted to acetaldehyde, ethanol, ethyl acetate, acetone, carbon dioxide, carbon monoxide, and methane. Conversion of vapor phase acetic acid is higher than that of liquid phase acetic acid as temperature increases due to higher rate of diffusion of hydrogen to vapor phase acetic acid. A high conversion of vapor phase acetic acid at atmospheric pressure shows that HDO of acetic acid can be performed without applying extreme pressure, although this may not be realizable for pyrolysis oil. A set of experiments conducted with a blank reactor indicates that thermal decomposition does not occur in the temperature range 200 1C–450 1C. In conclusion, the temperature effect has a major contribution in converting acetic acid to products that are not corrosive and possibly less susceptible to polymerization in the presence of phenolic compounds. Experiment on the effect of hydrogen partial pressure indicates that conversion of liquid phase acetic acid remains constant as the inlet partial pressure of hydrogen increases from 240 psig to 480 psig during hydrodeoxygenation reaction The conversion of vapor phase acetic acid during hydrodeoxygenation increases and reaches a maximum as the partial pressure of hydrogen increases from 3 psig to 15 psig. Studies of the dependence on the flow rate of liquid phase acetic acid show that increasing the flow rate of acetic acid decreases conversion and yield because of decrease in convective mass transfer rate of hydrogen from gas bubble to catalyst surface through acetic acid slug. In the case of vapor phase acetic acid, conversion and yield decrease as the flow rate of acetic acid increases due to decrease in residence time as a result of
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vaporization of acetic acid. The dependence of conversion and yield on the residence time was studied by increasing reactor length. The maximum values of conversion and yield reach at residence time of 0.06 s for liquid acetic acid. At this residence time, the STY obtained was the lowest due to decrease in average reactant concentration per gram of catalyst. Similarly, a maximum conversion reaches at 0.03 s for the vapor phase acetic acid. External mass transfer resistance was found to be negligible in microreactor at overall flow velocity of 2.54 m/s. Similarly, internal mass transfer was found to be negligible at an average particle size of 113 mm under the operating conditions selected in the microreactor. The heat transfer limitation in the microreactor was not a factor as the radial temperature difference in the microreactor was less than 5%. Hence, HDO reaction could be considered as kinetically controlled in the microreactor under the reaction conditions selected which is consistent with the strong temperature effect. The conversion and STY decrease as internal reactor diameter increases because of decreased concentration of hydrogen on catalyst surface due to diminishing rate of recirculation. Based on the products formed during low severity ( o300 1C) HDO of acetic acid, it can be concluded that acetic acid can be converted to products that are less corrosive. This means that by conducting HDO of pyrolysis oil at temperatures below 300 1C, corrosive nature of pyrolysis oil can be reduced. Based on the water formed during HDO of acetic acid, it can be said that oxygen content of acetic acid can be reduced. This implies that low severity (first stage) HDO at temperatures below 300 1C can be accomplished so that condensation reaction can be avoided in turn when HDO of pyrolysis oil is conducted at temperatures above 300 1C. However, low conversion of acetic acid and formations of acetaldehyde, acetone, and ethyl acetate during HDO, specially, at temperatures below 300 1C indicates that pyrolysis oil will still have unsaturated functions after HDO which could still cause condensation reaction during HDO at temperatures above 300 1C. As mentioned in the introduction section, microreactor system could be a suitable alternative in biofuel processing due to bulky nature of biomass which are scattered across the geographical area. However, due to low conversion of functions at low temperatures ( o300 1C), separation and recycling cost could be very high which could potentially upset the benefits of microreactor system. Therefore, improved catalysts need to be screened for conducting DHO for higher conversion of functions which have capability of rupturing carbon–oxygen bonds at temperatures below 300 1C. Identifying and quantifying all the products formed during HDO of pyrolysis oil will be very difficult. Using acetic acid as a model compound, HDO products were analyzed and quantified. Based on the products formed and the literature reviews, reaction pathways have been constructed. These reaction pathways will be valuable for kinetics study. The disadvantage of hydrodeoxygenation of pyrolysis oil is that it requires very high hydrogen pressure (1050–3000 psig). Therefore, many consider hydrodeoxygenation as an unattractive process. In this work, attempt has been made to show that HDO of acetic acid as a model compound of pyrolysis soil can be conducted at atmospheric pressure with a significant conversion achieved.
Nomenclature De
jp sc
Effective diffusivity of hydrogen in the bulk liquid (m2/s) Particle porosity Constriction factor
t jexp dp
rp r0 HAc CH2 Da DH robs Tw Ro A b
l Ea
Tortuosity Thiele modulus Catalyst particle diameter (m) Particle density (kg/m3) Reaction rate of H2 (kmol/kg s) H2 concentration in liquid (kmol/m3) Damkohler number Heat of reaction (kJ/mol) Observed reaction rate (mol/m3 s) Reactor wall temperature (K) Radius of a tubular reactor (m) Bed porosity Ratio of diluent to catalyst volume Effective thermal conductivity (W/m-K) Activation energy (kJ/mol)
Acknowledgments Narendra Joshi gratefully acknowledges support from the Robert C. Stanley Graduate Fellowship Program of Stevens Institute of Technology and GK-12 program of National Science Foundation under grant DGE-0740462. In addition, we are grateful to Dr. James Manganaro for his technical contribution to this work and Steve Mayo (Albemarle) for providing catalysts. References Adeosun, J.T., Lawal, A., 2005. Mass transfer enhancement in microchannel reactors by reorientation of fluid interfaces and stretching. Sens. Actuators B 110, 101–111. Ajmera, S.K., Delattre, C., Schmidt, M.A., Jensen, K.F., 2002. Microreactors for measuring catalyst activity and determining reaction kinetics. Stud. Surf. Sci. Catal., 97–102. Besser, R.S., Ouyang, X., Surangalikar, H., 2003. Hydrocarbon hydrogenation and dehydrogenation reactions in microfabricated catalytic reactors. Chem. Eng. Sci. 58, 19–26. Blake, P.G., Jackson, G.E., 1968. The thermal decomposition of acetic acid. J. Chem. Soc. B. Branebjerg, J., Gravesen, P., Krog, J.P., Nielsen, C.R., 1996. Fast Mixing by Lamination, Micro-Electro Mechanical Systems. MEMS 96 Proceedings. An Investigation of Micro-Structures, Sensors, Actuators, Machines and Systems. IEEE The Ninth Annual International Workshop, pp. 441–446. Bulushev, D.A., Ross, J.R.H., 2011. Catalysis for conversion of biomass to fuels via pyrolysis and gasification: a review. Catal. Today 171, 1–13. David, R., 2008. The price of biofuels. Technol. Rev. 111, 42. Ehrfeld, W., Hessel, V., Lowe, H., 2000. Microreactors: New Technology for Modern Chemistry. John Wiley & Sons. Elliott, D.C., 2007. Historical developments in hydroprocessing bio-oils. Energy Fuels 21, 1792–1815. Elliott, D.C., Baker, E.G., Piskorz, J., Scott, D.S., Solantausta, Y., 1988. Production of liquid hydrocarbon fuels from peat. Energy Fuels 2, 234–235. Elliott, D.C., Neuenschwander, G.G., 1996. Liquid fuels by low-severity hydrotreating of biocruide. Dev. Thermochem. Biomass Convers. 1, 611–621. Elliott, D.C., Oasmaa, A., 1991. Catalytic hydrotreating of black liquor oils. Energy Fuels 5, 102–109. Fogler, H.S., 2005. Elements of chemical reaction engineering, Fourth ed. Prentice Hall, Professional Technical Reference, Massachusetts, USA, pp 830. Furimsky, E., 2000. Catalytic hydrodeoxygenation. Appl. Catal. A 199, 147–190. Grange, P., Laurent, E., Maggi, R., Centeno, A., Delmon, B., 1996. Hydrotreatment of pyrolysis oils from biomass: reactivity of the various categories of oxygenated compounds and preliminary techno-economical study. Catal. Today 29, 297–301. Grootendorst, E.J., Pestman, R., Koster, R.M., Ponec, V., 1994. Selective reduction of acetic acid to acetaldehyde on iron oxides. J. Catal. 148, 261–269. Halder, R., Lawal, A., Damavarapu, R., 2007. Nitration of toluene in a microreactor. Catal. Today 125, 74–80. Hessel, V.E., Ehrfeld, W., Lowe, H., 2001. Microreactors. Wiley-VCH, Weinheim. Huber, G.W., Iborra, S., Corma, A., 2006. Synthesis of transportation fuels from biomass: chemistry, catalysts, and engineering. Chem. Rev. 106, 4044–4098. Jenck, J.F., 2009. Impact of Microtechnologies on Chemical Processing. Wiley New Jersey, USA. Kashid, M.N., Platte, F., Agar, D.W., Turek, S., 2007. Computational modelling of slug flow in a capillary microreactor. J. Comput. Appl. Math. 203, 487–497. Knight, J.B., Vishwanath, A., Brody, J.P., Austin, R.H., 1998. Hydrodynamic focusing on a silicon chip: mixing nanoliters in microseconds. Phys. Rev. Lett. 80, 3863–3866.
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