Chemical Engineering Science 74 (2012) 1–8

Contents lists available at SciVerse ScienceDirect

Chemical Engineering Science

0009-25

doi:10.1

n Corr

E-m

journal homepage: www.elsevier.com/locate/ces

Hydrodeoxygenation of pyrolysis oil 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

a r t i c l e i n f o

Article history:

Received 22 April 2011

Received in revised form

26 January 2012

Accepted 27 January 2012Available online 4 February 2012

Keywords:

Microreactor

Hydrodeoxygenation

Multiphase

Mass transfer

Fuel

Energy

09/$ - see front matter & 2012 Elsevier Ltd. A

016/j.ces.2012.01.052

esponding author.

ail addresses: njoshi@stevens.edu, njoshipasa

a b s t r a c t

Use of a packed bed microreactor for the first stage hydrodeoxygenation (HDO) of pyrolysis oil was

investigated. The effects of various processing conditions such as temperature, hydrogen partial

pressure, and residence time on Extent of HDO, hydrogen consumption, and space-time-consumption

were investigated using reduced sulfided NiMo/Al2O3 catalyst. External and internal mass transfer

resistances were examined in the microreactor. High hydrogen consumption along with small oxygen

removal suggests that in hydrodeoxygenation of pyrolysis oil hydrogenation cannot be avoided. Reactor

plugging at 543 K due to coke formation/polymerization shows that first stage HDO at a temperature

below 543 K was required to convert highly reactive compounds so that second stage HDO at higher

temperature can be conducted to remove oxygen completely. Hydrogen consumption and percent

oxygen removed for this first stage HDO are comparable to literature values; however, it is found that in

a microreactor these values are attainable at much lower pressure and residence time.

& 2012 Elsevier Ltd. All rights reserved.

1. Introduction

The current research focuses on utilization of pyrolysis oil (PO)as a renewable source of energy for transportation fuels. Pyrolysisoil obtained from fast pyrolysis of lignocellulosic biomass con-tains 40–50% oxygen, including oxygen from water, as result ofwhich it is immiscible with crude oil and has a lower heatingvalue. The heating value of pyrolysis oil is close to 20 MJ/kg,which is half that of crude oil (Boateng et al., 2007). Removal ofoxygen from pyrolysis oil and molecular weight reduction areparamount for increasing heating value, thermal stability, volati-lity and miscibility with crude oil. The emphasis of this study is onthe hydrodeoxygenation of pyrolysis oil using microreactorsystem.

Pyrolysis oil is a very complex mixture of oxygenated hydro-carbons. The complexity in the pyrolysis oil arises due to disin-tegration of cellulose, hemicelluloses, and lignin and broadspectrum of phenolic compounds and their uncontrolled interac-tions (Bridgwater and Cottam, 1992). More than 400 compoundshave been identified in pyrolysis oil, and include acids, aldehydes,ketones, guaiacols, alcohols, esters, furans, phenols, sugars, andsyringols (Huber et al., 2006).

There are several routes to convert pyrolysis oil into usefultransportation fuels such as steam reforming (followed byFischer–Tropsch synthesis), and catalytic cracking over zeolites.But lately hydrodeoxygenation (HDO) of pyrolysis oil has received

ll rights reserved.

@yahoo.com (N. Joshi).

a lot of attention due to the fact that aromatic compounds such asbenzene, toluene, and xylene (BTX) can be preserved for highoctane hydrocarbon fuels. Conversion of pyrolysis oil to a con-ventional transportation fuel can be accomplished by hydrotreat-ing it at high temperature, high hydrogen pressure, and in thepresence of catalysts which results in elimination of oxygen aswater (Czernik and Bridgwater, 2004). Therefore, conceptually(ideal) HDO can be characterized as follows:

C6H8O4þ6H2-6CH2þ4H2O (1)

which gives a maximum stoichiometric yield of 58 wt% on liquidin a carbon limited system (Bridgwater and Cottam, 1992).However, in many cases, model compounds are chosen insteadof pyrolysis oil for better understanding and control of thereaction process. Nimmanwudipong et al. (2011) used Guaiacolas a model compound of lignin-derived pyrolysis oil in thepresence of hydrogen to postulate the reaction network and topredict oxygen removal.

But the problem with the pyrolysis oil processing is that whenit is heated to temperatures above 333–373 K, the oil becomes avery hard, coke-like material. This chemical instability is attrib-uted to condensation reaction of unsaturated double bonds suchas olefins, aldehydes, and ketones similar to the phenol–formaldehyde polymerization. This is the reason why it is highlydesirable to eliminate these functions as quickly as possiblebefore they react to high molecular weight compounds (Grangeet al., 1996). In this reaction, the phenolic compounds areparticularly active coke precursors (Sharma and Bakhshi, 1993).

N. Joshi, A. Lawal / Chemical Engineering Science 74 (2012) 1–82

To avoid the above problem, Elliott and Neuenschwander(1996) have proposed a two-stage hydrotreating process forupgrading pyrolysis oil to transportation fuel. The first stage orstabilization stage is performed below 573 K to convert readilypolymerizing O-compounds such as methoxyphenols, biphenols,and ethers to phenols as well as to saturate double bonds such asolefins, aldehydes, and ketones. The second stage of hydrodeox-ygenation is performed above 623 K to obtain oxygen freeproduct (Furimsky, 2000). In this study, first stage HDO wasinvestigated in the microreactor system.

Hydrodeoxygenation of pyrolysis oil performed in a conven-tional macroreactor system could have severe heat and masstransfer limitations as well as concerns of high operating cost andsafety. Moreover, as biomass has very low density and is scatteredacross geographical areas, macroreactor system would not beeconomically viable since profitability requires such a system toprocess a huge quantity of biomass (David, 2008). According toWright et al. (2010) one of the reasons which make transportationfuel production from pyrolysis oil uneconomical is the infancy ofupgrading technologies and they suggested that various alter-native technologies need to be explored for possible improve-ments. In this regard, on-site, on-demand distributed system suchas microreactor system could be a suitable alternative solution forcentralized processing. In a microreactor, diffusion time is shortand influence of mass transfer on rate of reaction is greatlyreduced. Heat transfer is also greatly improved compared toconventional macroreactor. As heat and mass transfer is max-imized, global apparent kinetics approach intrinsic chemicalkinetics (Drinkenburg, 2003). Risk associated with high pressureand temperature processing will be minimal in the microreactorsystem compared to conventional macroreactor system (Hendershot,2003). Hence, a microreactor system could be a solution to on-siteon-demand production of fuels which will reduce inventory andtransportation costs of low density biomass. Moreover, becausemicroreactor possesses ultra-low transport resistances, mass diffu-sion and heat transfer are extremely quick providing agility ofoperation such that thermal and composition reach equilibrationinstantaneously (Besser et al., 2003). High yield, improved productquality, better selectivity and safe operation are attainable due toinstantaneous heat and mass transfer (Halder et al., 2007; Okaforet al., 2010; Tadepalli et al., 2007b; Voloshin and Lawal, 2010).Mixing in the microchannels is attainable only by inter-diffusions ofreactants due to laminar flow (Adeosun and Lawal, 2005), howeverdue to short transverse diffusional distance rapid and effectivemixing is attainable in the microreactor which can quickly bringreactants in contact with catalyst in a heterogeneous reaction (Hesseland Lowe, 2001).

Enhancement of mixing in two-phase flows can also beincreased by selecting appropriate inlet T-orientations to providea short slug length. Studies have shown that introduction of gasand liquid feeds head to head or perpendicular to each other withthe liquid stream parallel to the microchannel markedly improvethe mixing (Qian and Lawal, 2006).

Hydrodeoxygenation of pyrolysis oil is a heterogeneous reac-tion involving hydrogen gas, liquid pyrolysis oil and solid catalyst.In a microreactor system, the micro-mixer produces Taylor flow(slug flow) consisting of alternating gas bubbles of hydrogen andpyrolysis oil slugs before entering the packed bed microreactor.However, flow visualization experiments conducted by Tadepalliet al. (2007a) demonstrated a development of transitional Taylorflow due to distortion of slug boundaries by catalyst particleswithin the packed bed microreactor.

Catalysts generally used for hydrodeoxygenation such as CoO/MoO3 and NiO/MoO3 on Al2O3 support are commercially availablecatalysts used for removing sulfur, nitrogen, and oxygen frompetrochemical feedstocks. According to research conducted at

Pacific Northwest National Laboratory, sulfided form of CoO/MoO3 and NiO/MoO3 are much more active for hydrodeoxygena-tion than the oxide form (Elliott, 2007). The sulfidation createsactive sites that can play a role in the rupture of carbon–heteroatom bond (Senol, 2007). In this study we used sulfidedform of NiO/MoO3 with Al2O3 supported catalyst.

Elliott (2007) and Elliott et al. (2009) have reported hydrogenconsumption and oxygen removal during first stage HDO to be inthe average of 0.15 m3/m3 and 7% using various catalysts; and theoperating pressure used for this HDO was about 13.79 MPa.Another work by Elliott and Oasmaa (1991) on catalytic hydro-treating of black liquor oils reported that oxygen content wasreduced from 9 wt% (dry basis) to 4 wt% in an average, over manyexperimental runs in operating conditions of temperature at553 K, around 6.89 MPa pressure and 25–35 min of residencetime. Similarly, Venderbosch et al. (2009) were able to removeabout 5% of oxygen as water, with hydrogen consumption of0.078 m3/m3 during hydrodeoxygenation of pyrolysis oil con-ducted at 448 K and 20.68 MPa. In the hydrodeoxygenation oftar from gasification of biomass, Gevert and Hernelind (2002)shows that when tar was heated from 473 K to 553 K, oxygencontent decreased from 28 wt% to 22 wt%, but when pressure wasincreased from 6.89 MPa to 11.03 MPa at temperature of 553 K,the reduction of oxygen was from 24 wt% to 21 wt% only andconcluded that the role of the pressure was not as strong as thatof temperature. Gevert et al. also reported that the reactor wasclogged at the temperature above 573 K, possibly due to coking.

The objective of the research presented here was to evaluatethe performance of a packed bed microreactor for hydrodeox-ygenation of pyrolysis oil by studying the effects of variousreaction variables such as mass transfer limitations, liquid flowvelocity, residence time, hydrogen partial pressure, temperature,and reactor diameter on the Extent of HDO, space-time consump-tion, and hydrogen consumption.

2. Experimental

2.1. Materials

Presulfided NiO/MoO3/Al2O3 catalyst was obtained from Albe-marle (presulfied and supplied by Eurecat, USA), Houston, Texas.The catalyst was ground and sieved to obtain particles withdiameters in a range of 75–150 mm. Surface area of the catalystwas 164 m2/g and average pore diameter was 106 A. The surfacearea and pore diameter were obtained by using multipoint BETtechnique and the instrument used was Quantochrome Autosorb-1.The catalyst was reduced with 5.0 sccm of hydrogen at 593 K and3.45 MPa for two hours. The surface area of the reduced catalyst was209.0 m2/g and average pore diameter was 92.0 A. Pyrolysis oilprepared from Sawdust was obtained from Dynamotive EnergyCorporation, British Columbia, Canada. The pyrolysis oil was vacuumfiltered using number 410 filter paper (1�10�6 m retention) fromVWR to remove residual particles. Filtration was enhanced byadding Standard Super-Cel (a diatomite filtration aid) obtained fromCelite Corporation, Lompoc, California. The gas used was extra dryhydrogen from Praxair and inert nitrogen was used to performmaterial balance.

2.2. Experimental setup

In this section, the experimental setup and procedure isdescribed. A HPLC pump (Laboratory Alliance Series III) was usedto control the flow rate of pyrolysis oil and mass flow controllers(Porter Model 201) were used to control the flow rates of com-pressed hydrogen and nitrogen. Ranges of superficial velocities of

0

10

20

30

Wt %

Hemicelluloseand Cellulose Lignin

Fig. 1. Typical chemical composition of pyrolysis oil (PO) (Milne et al., 1997).

N. Joshi, A. Lawal / Chemical Engineering Science 74 (2012) 1–8 3

liquid oil, 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. Theliquid and gas phases were combined in a T-junction mixer(Upchurch) with 508�10�6 m through-hole. As the Reynoldsnumber for the combined flow was less than 100 for all experi-ments, all the flows were in the laminar regime. The fluids exitingfrom the T-junction exhibited the Taylor flow pattern with the liquidslug length in the range 0.001–0.004 m, whereas gas bubble lengthvaried from 0.001 to 0.005 m. A microreactor was prepared from a0.0016 m 316 stainless steel tubing with 762�10�6 m internaldiameter, and was gravity filled with catalyst. The total length of thepacked bed microreactor varied from 0.025 to 0.18 m. Hastelloymicron filter-cloth (200�1150 meshes, Unique Wire Weaving Co.,Hillside, New Jersey) was placed at the ends of the reactor to retainthe catalyst. The reactor was submerged in a constant temperaturefluidized bed (model SBL-2D, Barloworld Scientific). The bed mate-rial was sand (nominal particle size of 120�10�6 m) and it wasfluidized with compressed air at 69 kPa. The reactor system waspressurized using the Back Pressure Regulator (GO Regulator Co.).The entrance and the exit pressures of the fluids (liquid and gascombined) in the reactor were measured by using pressure trans-ducers. The pressure drop in the reactor varied from 0.1 MPa to0.3 MPa depending upon reactor length.

2.3. Analysis

Composition of gas phase was measured by a Varian CP-3800gas chromatograph equipped with a Mole Sieve 5 A column at308 K, thermal conductivity detector, and argon as carrier gas. Themethod used for the GC was internal normalization as describedby Grob (1995). Nitrogen as an inert gas was used to determineunreacted hydrogen in the product. Inlet hydrogen and nitrogenflow rates were set using the mass flow controllers. As a mixtureof hydrogen and nitrogen entered the reactor with known flowrates, hydrogen was consumed due to the reaction whereas theamount of nitrogen remained the same since it is an inert.Therefore the flow rate of the nitrogen was used as a basis tocalculate the unreacted hydrogen at outlet. A difference betweenthe inlet and outlet hydrogen was calculated as hydrogenconsumption.

Water content in the filtered pyrolysis oil was measured usingVolumetric Karl Fischer (KF) Titration Workstation (Model 375,Denver Instrument) with the use of hydranal reagents obtainedfrom Sigma-Aldrich. Titrant used was hydranal titrant-2 while theworking medium was hydranal solvent. Before the actual titra-tion, the hydranal solvent (4.0�10�5 m3 added to the titrationvessel) was titrated to dryness in a drift determination step. Thisstep removes moistures from solvent, electrode, and titrationvessel. Filtered pyrolysis oil 5.0�10�8 m3 was added using asyringe through septum on the KF cell.

Liquid phase was analyzed using GC/MS, and High Perfor-mance Liquid Chromatography (HPLC). GC/MS analysis was per-formed using Varian Instrument (GC 3900, equipped with aVarian CP-1177 split/splitless injector, and a Varian CP-8410autosampler); and for detection the ion trap mass spectrometer(Varian Saturn 2100T) was used. Capillary column used for gaschromatograph was Factor Four, VF-5ms (30 m long with2.5�10�4 m diameter and 2.5�10�7 m film thickness). Analysisof light non-volatile organic compounds of pyrolysis oil wasconducted in a HPLC (Shimadzu series: mobile phase degasser[DGU_20A5], pump station [LC-20AT], autosampler [SIL-20AC],and reflective index detector [RID-10A]) equipped with BioRadAminex HPX-87H column. The mobile phase consisted of 0.007 Naqueous H3PO4 with an isocratic flow (flow rate of 6�10�7 m3/min).Hydrodeoxygenation reaction involves reaction of H2 with oxygen inpyrolysis oil forming H2O. The theoretical removal of oxygen can be

calculated based on the amount of hydrogen reacted. Hydrogenreacted was expressed in terms of Extent of HDO, H2 consumption,and space–time-consumption (reaction rate of disappearance) whichwere defined as follows:

H2 consumption¼ ðflow rate of H2 in2flow rate of H2 outÞ ð2Þ

Extent of HDO¼amount of hydrogen consumed

amount of hydrogen that would have been consumed

for a complete removal of oxygen from pyrolysis oil

� 100

ð3Þ

space2time consumption ðSTCÞ ðrate of disappearance of H2Þ

¼amount of H2 consumed

amount of catalyst x timeð4Þ

3. Results and discussion

3.1. Analysis of pyrolysis oil

A typical composition of pyrolysis oil produced by fast pyr-olysis is shown in Fig. 1.

Some properties of pyrolysis oil used in this study are shownin Table 1. Water soluble and insoluble fractions of PO wereobtained using method described by Brown et al. (2009). Volatilesand non-volatiles of water soluble fraction were determined byusing thermogravimetric analysis (TGA Q50 V6.7 Build 203). Theanalysis was performed at 523 K for 77 min to simulate thetemperature in the GC/MS injector.

In hydrodeoxygenation of pyrolysis oil, high temperatures aredetrimental as it may lead to plugging of reactor channels due toformation of coke-like material. Venderbosch et al. (2009)reported that oxygen containing components of pyrolysis oilespecially the carbonyl compounds such as aldehydes, ketones,and carboxylic acids are responsible for the instability. Thereforefirst stage of HDO was performed at temperatures below 573 K toconvert highly reactive compounds from pyrolysis oil. HPLCanalyses of the hydrodeoxygenated pyrolysis oil compared tofeed oil were conducted by preparing solutions of feed andproduct in acetone with ethanol as an internal standard. Calibra-tion curves of pure compounds were used for individual com-pounds in pyrolysis oil. The percent conversions of somecompounds of PO are shown in Table 2. Table 2 shows thathydroxy-aldehyde, xylose, and acetic acid were reduced by morethan 15%; whereas hydroxy-acetone, glyoxal, and levoglucosanwere reduced by less than 10% by first stage (mild) HDO whichwas conducted at 453 K and 2.07 MPa.

Similarly, phenolic compounds in pyrolysis oil were preparedin acetone with Fluoranthene as an internal standard and quanti-fied using GC/MS which are shown in Table 3. The GC/MS analysisof hydrodeoxygenated pyrolysis oil conducted at 453 K and2.07 MPa showed that percent conversions of furfural and iso-eugenol were 4.49 and 1.87, respectively.

Table 1Pyrolysis oil properties.

Density (298 K) (g/ml) 1.19

Viscosity (296 K) (cSt) 55.47

Elemental analysis (wt%)

C 40.88

H 7.66

O 50.69

N 0.6

Ash 0.17

Pyrolysis oil fractionMoisture (wt%) 26.7

Water insoluble (wt%) 24.7

Water soluble volatiles at 523 K (wt%) 41.7

Water soluble non-volatiles at 523 K (wt%) 6.9

Surface tension (m N/m) 35–39a

a Dynamotive (2006).

Table 2HPLC analyses of feed and product pyrolysis oil. Reaction conditions: 2.07 MPa,

453 K; gas phase: hydrogen and nitrogen, liquid phase: whole pyrolysis oil, 20 mg

of sulfided NiMo/Al2O3 catalyst.

Oxygenates in PO wt% in Feed PO % Conversion

Xylose 0.0192 25.66

Glyoxal 0.0160 6.77

Acetic acid 0.2679 15.13

Hydroxy-acetone 0.0487 3.89

Hydroxy-acetaldehyde 0.0253 59.04

Levoglucosan 0.1194 0.52

Table 3GC/MS analysis of pyrolysis oil.

Oxygenates in PO wt% in Feed PO

Furfural 0.0843

Furfuryl alcohol 0.1014

Phenol 0.0814

p-Cresol 0.0470

Guaiacol 0.3595

Methyl guaiacol 0.2353

Methyl syringol 0.3653

Syringol 1.5890

Iso-eugenol 0.0446

N. Joshi, A. Lawal / Chemical Engineering Science 74 (2012) 1–84

A first stage HDO in a microreactor with run time of 85 min atresidence time of 0.033 s converted 1.5% of feed into the coke;however at residence time of 0.017 s, the coke formation was only0.63% after 4 h of the HDO. Therefore, it appears that in additionto temperature, residence time may have an influence on cokeformation during the HDO of PO. Analyses of gaseous productsshowed formation of iso-butane, carbon dioxide, iso-pentane,nonane, and other unknown gaseous products. Up to 10% of iso-butane and 1.2% of carbon dioxide were detected as well as traceamounts of iso-pentane, nonane. Methane and carbon monoxidewere not detected. An elemental carbon balance around thereactor after the hydrodeoxygenation of pyrolysis oil provided27% loss of carbon of which 4% was accounted for by cokeformation, 21% was accounted for by carbon from iso-butaneand carbon dioxide, and 2% loss of carbon was assumed to be inthe trace amounts of iso-pentane, nonane, and other unknowngaseous products.

The first stage hydrodeoxygenation in a microreactor removed6% of theoretically removable oxygen from feed pyrolysis oil at theoperating condition of 453 K and 2.07 MPa. A material balanceprovided the hydrogen consumption of 12 sccm. The high hydrogen

consumption indicated that hydrogenation was unavoidable as thesulfided NiMo catalyst was not optimized for the removal of oxygenonly. Similar occurrence of hydrogenation and HDO in parallel wasalso reported in the HDO of 1-naphthol by Li et al. (1985).

3.2. Reaction and catalyst activity

During hydroprocessing of pyrolysis oil hydrodeoxygenation(HDO), hydrogenation (HYD), hydrodenitrogenation (HDN),hydrodesulphurization (HDS), and hydrodemetallization (HDM)occur simultaneously (Furimsky, 2000). Since nitrogen, sulfur, andash contents are less than one percent in pyrolysis oil, HDO andHYD are the dominant reactions during hydroprocessing. Attemperatures below 573 K, hydroprocessing of pyrolysis oilcauses partial reduction and partial deoxygenation of oil.

Catalyst deactivation is a major concern in the chemicalindustry where mass production is required to be profitable.Frequent regeneration of deactivated catalyst is required if thelife time of catalyst is very short as in the catalytic cracking ofpetroleum naphthas. Catalyst deactivation could be caused bygradual change in surface crystal structure, irreversible depositionof a substance on the active site or a deposit of carbonaceousmaterial on the catalyst surface (Fogler, 2005). During hydro-deoxygenation of pyrolysis oil, Ferrari et al. (2001) found thatsulfided NiMo/Al2O3 catalyst was rapidly deactivated by cokedeposition due to the acidity of the alumina support. Theirfindings concluded that oxygen substituted phenolic compoundssuch as guaiacol and catechol were the cause of coke deposition.Similarly, Laurent and Delmon (1994) concluded that the pre-sence of water during hydroprocessing of pyrolysis oil caused theloss of two third of the initial activity of a sulfided NiMo/Al2O3

catalyst in less than 60 h. The decrease in catalytic activity wasdue to a partial oxidation of the nickel sulfide phase into oxidizednickel species caused by water. Reduced sulfided NiMo/Al2O3

used in the present study did not show any sign of catalystdeactivation in terms of hydrogen consumption for at least85 min, although there is a sign of gradual increase in cokeformation as pressure drop in the reactor increased over time.Hence, prolonged exposure of catalyst to hydrotreating is neededto study the life time of the catalyst.

3.3. External mass transfer limitation

External mass transfer involves diffusion of reactants from bulkfluid to the external surface of the catalyst where the reactionbetween gaseous H2 and liquid PO takes place. The heterogeneousreaction of pyrolysis oil 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 aboundary layer surrounding the catalyst surface. As external masstransfer rates through both gas–liquid and liquid–solid interfacesare affected by flow velocity, the rate of hydrogen consumption wasmeasured as the superficial flow velocity was varied, while keepingthe residence time constant by varying catalyst loading. Fig. 2 showsthat, for the range of velocity indicated, the superficial velocity hasno effect on the space-time consumption (STC). This indicates that atthe lowest velocity selected for this study, the boundary layer is sothin that it no longer offers any significant resistance to the diffusionacross the boundary layer; hence the hydrodeoxygenation is notlimited by the external mass transfer.

3.4. Internal mass transfer limitation

The effect of internal mass transfer limitation was studied byvarying catalyst particle size. Two different particle size ranges of38–45�10�6 m and 75–150�10�6 m were selected in order to

0

2

4

6

0.0

STC

(kg

H/k

g ca

t./ s

x

0.00

028)

Superficial flow velocity (m/s)

1.0 2.0 3.0 4.0 5.0

Fig. 2. Effect of fluid flow velocity on space–time consumption (STC). Reaction

conditions: 2.07 MPa, 453 K; gas phase: hydrogen and nitrogen, liquid phase:

whole pyrolysis oil.

0

2

4

6

0

10

20

30

40

50

0.0 0 0.01 0.02 0.03 0.04 0.05

H2

Con

sum

ptio

n (s

ccm

)

Ext

ent o

f H

DO

(%

)

Superficial PO Velocity (m/s)

Extent of HDO (%)

H2 Consumption

STC

STC

(kg

H2/

kg c

at./s

x 0.

0002

8)

Fig. 3. Effect of superficial pyrolysis oil velocity on STC, Extent of HDO and H2

consumption. Reaction conditions: 2.07 MPa, 453 K; gas phase: hydrogen and

nitrogen, liquid phase: whole pyrolysis oil, sulfided NiMo/Al2O3 catalyst.

Fig. 4. Effect of residence time on STC, and Extent of HDO. Reaction conditions:

2.07 MPa, 453 K; gas phase: hydrogen and nitrogen, liquid phase: whole pyrolysis

oil, sulfided NiMo/Al2O3 catalyst, liquid flow rate: 8.33�10�10 m3/s.

N. Joshi, A. Lawal / Chemical Engineering Science 74 (2012) 1–8 5

study their effect on reaction rate (STC). The STC of hydrogen forparticle ranges of 38�10�6–45�10�6 m and 75�10�6–150�10�6 m were 0.00097 and 0.00092 kg H2/kg catalyst/s, respec-tively which indicate that there was no diffusional mass transferlimitation for the particle size range of 75�10�6–150�10-6 m.

To substantiate the above result, the internal mass transferlimitation was numerically evaluated for the particle size range of75�10�6–150�10�6 m (mean diameter of 113�10�6 m)according to the Weisz–Prater criterion (Fogler, 2005) given inEq. (5): where �r0

ðobsÞ) is the reaction rate (STC), rc is the catalystparticle density, De is the effective diffusivity, which was esti-mated using the equation De¼(DABjpsc)/t where DAB is the binarydiffusivity of hydrogen in the liquid reactant. DAB is estimatedto be 4.74�10�6 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 Weisz–Prater criterion (Cwp) was calculated to be0.0005 which was considerably smaller than 1 indicating thatdiffusion rate for the catalyst particle size range of 75�10�6–150�10�6 m was much higher compared to actual reaction rate.Hence there was no internal diffusion limitation, and consequentlythe concentration gradient within the catalyst particle was negligible

Cwp ¼�r0ðobsÞrcR2

DeCAs

ð5Þ

Internal mass transfer limitation could also be estimated bycalculating Thiele modulus for the particle size range of 75�10�6–150�10�6 m assuming a pseudo-first-order reaction with respect tohydrogen and representative compounds of pyrolysis oil (Furimsky,2000; LaVopa and Satterfield, 1987) according to Eq. (6) (Fogler,2005). The Thiele modulus was estimated to be 0.01 which corre-sponds to an effectiveness factor of unity (Fogler, 2005), indicatingthat actual overall rate of reaction was equal to the rate of reactionthat would result if entire interior surface were exposed to theexternal catalyst surface conditions (CAs

, Ts). Therefore, it can beconcluded that the rate of hydrogen consumption (STC) was notlimited by internal mass transfer

fexp ¼dp

6

rpr0H2

DeCH2

� �0:5

ð6Þ

3.5. Effect of superficial velocity of pyrolysis oil

Experiments were conducted to study the effect of superficialPO velocity on STC, Extent of HDO and hydrogen consumption byvarying the liquid velocity from 0.0065 to 0.0391 m/s. Otheroperating conditions such as temperature, pressure, and gas flowrate are kept constant. In this experiment, the actual volumetricgas flow rate was at least one order of magnitude higher than the

liquid flow rate; hence the residence time was essentially con-stant. The results shown in Fig. 3 indicate that as the superficialliquid velocity increases, the H2 consumption, Extent of HDO, andSTC decrease. During the experiment, a steady increase in liquidslug length was observed as liquid velocity was increased from0.0065 to 0.0391 m/s. As mass transfer from gas to liquid slugsthrough hemispherical caps of gas bubble was a strong function ofliquid slug length, the convective mass transfer rate of hydrogento liquid pyrolysis oil decreased when liquid velocity wasincreased (Kreutzer et al., 2001). Hence, the decrease in liquidvelocity will enhance the mass transfer rate of gas to liquid andcatalyst surface, and therefore resulting in increase in Extent ofHDO, STC, and H2 consumption.

3.6. Effect of residence time

Residence time was varied by increasing reactor length (cat-alyst loading) while keeping flow rate, temperature and pressureconstant to study the effect on Extent of HDO and STC. Inmicroreactors, the residence time is very small compared toconventional macroreactors; however, it requires an optimumresidence time for maximum hydrogen consumption. This can bedone by increasing reactor length which in turn increases catalystloading. The result in Fig. 4 shows that Extent of HDO increases asresidence time increases. The highest residence time of 1.23 sapproached close to maximum residence time for the givenreaction condition. STC decreases as residence time increasesdue to increase in catalyst loading causing decrease in averagereactant concentration per gram of catalyst at higher Extent ofHDO. Hence higher STC could be achieved at lower residencetime.

N. Joshi, A. Lawal / Chemical Engineering Science 74 (2012) 1–86

3.7. Effect of hydrogen partial pressure

A set of experiments was carried out in the range of 1.24–3.10 MPa to study the effect of hydrogen partial pressure onExtent of HDO and STC. Reaction temperature and residence timewere kept constant for all experimental runs. The residence timewas kept constant by varying reactor length (catalyst loading).The H2 partial pressure was varied two ways: by changinghydrogen flow rate at constant total pressure and by changingtotal pressure at constant H2 flow rate. The results in Fig. 5 showthat increasing H2 partial pressure by increasing hydrogen flowrate increases the Extent of HDO. The increase in Extent of HDOwas due to increase in dissolved hydrogen on catalyst surface asH2 partial pressure was increased. The constant STC observed isdue to constant hydrogen concentration resulting in a constantreaction rate. The results in Fig. 6 indicate that increasinghydrogen partial pressure by increasing total pressure had noeffect on the Extent of HDO. The constant Extent of HDO at thetotal pressure range selected indicated that dissolved hydrogenon the catalyst surface reached a maximum (saturated) value. Theincrease in STC with increase in hydrogen partial pressure wasdue to an increase in hydrogen concentration (due to constantresidence time) resulting in a higher reaction rate.

3.8. Effect of temperature

In hydrodeoxygenation reaction of pyrolysis oil, temperature isthe most important parameter in removing oxygen from oxygenated

Fig. 5. Effect of hydrogen partial pressure on STC, and Extent of HDO. Reaction

conditions: 453 K; total pressure: 2.07 MPa; gas phase: hydrogen and nitrogen,

liquid phase: whole pyrolysis oil, sulfided NiMo/Al2O3 catalyst, liquid flow rate:

8.33�10�10 m3/s.

Fig. 6. Effect of hydrogen partial pressure on STC, and Extent of HDO. Reaction

conditions: 453 K; gas phase: hydrogen (90%) and nitrogen, liquid phase: whole

pyrolysis oil, sulfided NiMo/Al2O3 catalyst, liquid flow rate: 8.33�10�10 m3/s.

compounds. A series of experiments was conducted by varyingtemperature to study its effect on Extent of HDO, hydrogenconsumption, and space–time consumption. In this study pressureand residence time were kept constant. Residence time was keptconstant by varying reactor length (catalyst loading) to compensatefor the change ingas velocity. The results shown in Fig. 7 indicatethat there was an increase in the space-time consumption (STC) astemperature was increased. This increase in STC with temperaturecould be attributed primarily to the influence of reaction kineticsand secondarily to diffusional mass transfer at the temperaturerange utilized (Fogler, 2005). Another phenomenon observed in thisstudy was that as temperature was increased, Extent of HDO andhydrogen consumption were increased; however the reactorplugged when the temperature was increased to 543 K. This reactorplugging was attributed to condensation reaction of unsaturateddouble bonds with phenolic compounds similar to phenol–formaldehyde polymerization mentioned in the introduction sectionto produce tar-like materials (Elliott, 1996; Grange et al., 1996;Sharma and Bakhshi, 1993).

3.9. Effect of reactor diameter

A set of experiments was conducted to study the effect ofreactor diameter on Extent of HDO, and STC. In these experi-ments, reactor internal diameters of 0.0008 m, 0.0032 m, and0.0064 m were compared while temperature, pressure, residencetime, and superficial velocity were kept constant. The resultsshown in Fig. 8 indicate that Extent of HDO and STC decreasedsignificantly as the reactor diameter was increased. The higherExtent of HDO and STC for microreactor with diameter of0.0008 m (o1 mm) could be attributed to enhanced heat andmass transfer rates while for higher diameter reactors, resistancesto heat and mass transfers appeared to be significant. Since thefluid flow velocity was kept constant, the convective masstransfer would remain constant for all reactors. Therefore, thehigher Extent of HDO and STC for microreactor with diameter of0.0008 m must be attributed to the higher rate of diffusioncausing faster mixing due to shorter diffusional path lengthcompared to larger diameter reactors. Also surface area to volumeratio decreased as reactor diameter increased causing decrease inheat transfer rate. Based on the results presented in Fig. 8 it couldbe deduced that as reactor diameter is increased, the overallperformance of reactor will diminish; hence comparativelymicroreactor performs better than macroreactors in terms ofExtent of HDO and reaction rate.

0

2

4

6

0

10

20

30

40

273 373 473 573 673 773

HC

onsu

mpt

ion

(scc

m)

Ext

ent o

f H

DO

(%

)

Temperature K

Extent of HDO

H2 Consumption

STCReactor plugged

STC

(kg

H/K

g ca

t./s

x0.

0002

8)

Fig. 7. Effect of temperature on H2 consumption, and Extent of HDO. Reaction

conditions: 2.07 MPa; gas phase: hydrogen and nitrogen, liquid phase: whole

pyrolysis oil, sulfided NiMo/Al2O3 catalyst, liquid flow rate: 8.33�10�10 m3/s.

0.0

1.0

2.0

3.0

4.0

0

20

40

60

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8

STC

(Kg

H2/

Kg

cat/s

x0.0

0028

)

Ext

ent o

f H

DO

(%

)

Reactor Diameter ( m x 0.01)

Extent of HDO STC

Fig. 8. Effect of reactor diameter on Extent of HDO and STC. Reaction conditions:

2.07 MPa, 453 K; gas phase: hydrogen and nitrogen, liquid phase: whole pyrolysis

oil, sulfided NiMo/Al2O3 catalyst.

N. Joshi, A. Lawal / Chemical Engineering Science 74 (2012) 1–8 7

4. Conclusions

A first stage hydrodeoxygenation of whole pyrolysis oil in apacked bed microreactor at lower pressure and residence timecompared to macroreactors has been demonstrated. Microreactorperformance study was conducted based on the effect of variousoperating conditions using reduced sulfided NiMo/Al2O3 catalyst.The experimental results indicate that external mass transferresistances were found to be negligible in microreactor at overallflow velocity of 0.91 m/s. Similarly, internal mass transfer wasfound to be negligible under the operating conditions selected inthe microreactor. Hence, reaction could be considered as kineticallycontrolled in the microreactor under the reaction conditionsselected. A study on the effect of pyrolysis oil flow rate shows thatincreasing flow rate of pyrolysis oil decreases Extent of HDO, H2

consumption, and STC because of decrease in convective masstransfer rate of hydrogen from gas bubble to pyrolysis oil slug.The effect of residence time on Extent of HDO was studied byincreasing reactor volume. Close to maximum Extent of HDO wasachieved at residence time of 1.23 s. At this residence time the STCobtained was the lowest due to decrease in average reactantconcentration per gram of catalyst. Therefore, higher STC could beachieved by decreasing residence time. The Extent of HDO wasincreased when hydrogen partial pressure was increased up to2.07 MPa because of increase in dissolved hydrogen on catalystsurface when hydrogen flow rate was increased at constant totalpressure, whereas STC remained constant due to constant hydrogenconcentration. The Extent of HDO and STC decrease as internalreactor diameter increases because of increased heat and masstransfer resistances due to longer diffusional path length. Therefore,microreactor with internal diameter of less than 1 mm performedbetter than larger (3.2 mm, and 6.4 mm) internal diameter reactors.

High hydrogen consumption along with small oxygen removalindicates that simultaneous hydrogenation is unavoidable whileperforming hydrodeoxygenation of pyrolysis oil which suggeststhat catalyst for HDO-only needs to be formulated and optimizedfor maximum removal of oxygen from pyrolysis oil. The increasein STC as temperature increases could be attributed primarily toincreased reaction kinetics. However, reactor plugging at 543 Kshows that first stage HDO at temperature below 543 K wasrequired to convert highly reactive compounds.

Hydrogen consumption and percent removed oxygen of firststage HDO product are comparable to literature values; however,it is found that in microreactor these values are attainable atmuch lower pressure and residence time than in a macroreactor.

Nomenclature

Cwp Weisz–Prater criterionr0ðobsÞ observed reaction rate (mol/g h)

rc catalyst particle density (g/cm3)R catalyst radius (mm)De effective diffusivity of hydrogen in the particle (cm2/s)CAs

H2 concentration on catalyst surface (mol/L)jp porositysc constriction factort tortuosityfexp Thiele modulusdp catalyst particle diameter (mm)rp particle density (g/cm3)r0H2

reaction rate of H2 (mol/g h)CH2

H2 concentration in liquid (mol/L)

Acknowledgments

Narendra Joshi gratefully acknowledges support from theRobert C. Stanley Graduate Fellowship Program of Stevens Insti-tute of Technology and GK-12 program of National ScienceFoundation under Grant DGE-0740462. In addition, we are grate-ful to Dr. Jim Manganaro for his technical contribution to thiswork and Steve Mayo (Albemarle) for providing catalysts. Wewould also like to acknowledge Dr. Akwasi A. Boateng (EasternRegional Research Center, USDA) for providing pyrolysis oil.

References

Adeosun, J.T., Lawal, A., 2005. Mass transfer enhancement in microchannelreactors by reorientation of fluid interfaces and stretching. Sens. Actuators,B: Chem. 110, 101–111.

Besser, R.S., Ouyang, X., Surangalikar, H., 2003. Hydrocarbon hydrogenation anddehydrogenation reactions in microfabricated catalytic reactors. Chem. Eng.Sci. 58, 19–26.

Boateng, A.A., Daugaard, D.E., Goldberg, N.M., Hicks, K.B., 2007. Bench-scalefluidized-bed pyrolysis of switchgrass for bio-oil production. Ind. Eng. Chem.Res. 46, 1891–1897.

Bridgwater, A.V., Cottam, M.L., 1992. Opportunities for biomass pyrolysis liquidsproduction and upgrading. Energy Fuel 6, 113–120.

Brown, R.C., Li, M., Kuzhiyi, N., Johnston, P., Jones, S., 2009. What does It mean tocharacterize bio-oil? In: Proceedings of TC Biomass Conference. Chicago, IL.

Czernik, S., Bridgwater, A.V., 2004. Overview of applications of biomass fastpyrolysis oil. Energy Fuel 18, 590–598.

David, R., 2008. The price of biofuels. Technol. Rev.Drinkenburg, A., 2003. Process Intensification, Re-Engineering the Chemical

Processing Plant. CRC Press.Dynamotive, 2006. The Biooil Information Book.Elliott, D.C., Neuenschwander, G.G., 1996. Liquid fuels by low-severity hydrotreat-

ing of biocruide. Dev. Thermochem. Biomass Convers. 1, 611–621.Elliott, D.C., 2007. Historical developments in hydroprocessing bio-oils. Energy

Fuels 21, 1792–1815.Elliott, D.C., Hart, T.R., Neuenschwander, G.G., Rotness, L.J., Zacher, A.H., 2009.

Catalytic hydroprocessing of biomass fast pyrolysis bio-oil to produce hydro-carbon products. Environ. Prog. Sustainable Energy 28, 441–449.

Elliott, D.C., Neuenschwander, G.G., 1996. Liquid fuels by low-severity hydrotreat-ing of biocruide. Dev. Thermochem. Biomass Convers. 1, 611–621.

Elliott, D.C., Oasmaa, A., 1991. Catalytic hydrotreating of black liquor oils. EnergyFuel 5, 102–109.

Ferrari, M., Bosmans, S., Maggi, R., Delmon, B., Grange, P., 2001. CoMo/carbonhydrodeoxygenation catalysts: influence of the hydrogen sulfide partialpressure and of the sulfidation temperature. Catal. Today 65, 257–264.

Fogler, H.S., 2005. Elements of Chemical Reaction Engineering, fourth ed. PrenticeHall Professional Technical Reference.

Furimsky, E., 2000. Catalytic hydrodeoxygenation. Appl. Catal., A: Gen. 199,147–190.

Gevert, B., Hernelind, M., 2002. Hydrotreatment of tar formed in gasification ofbiomass. Fuel Chem. Div. 47, 171–172.

Grange, P., Laurent, E., Maggi, R., Centeno, A., Delmon, B., 1996. Hydrotreatment ofpyrolysis oils from biomass: reactivity of the various categories of oxygenatedcompounds and preliminary techno-economical study. Catal. Today 29,297–301.

Grob, R.L., 1995. Modern Practice of Gas Chromatography, third ed. Wiley-Interscience.

Halder, R., Lawal, A., Damavarapu, R., 2007. Nitration of toluene in a microreactor.Catal. Today 125, 74–80.

Hendershot, D., 2003. Process Intensification for Safety, Re-Engineering theChemical Processing Plant. CRC Press.

N. Joshi, A. Lawal / Chemical Engineering Science 74 (2012) 1–88

Hessel, V., 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.Kreutzer, M.T., Du, P., Heiszwolf, J.J., Kapteijn, F., Moulijn, J.A., 2001. Mass transfer

characteristics of three-phase monolith reactors. Chem. Eng. Sci. 56,

6015–6023.Laurent, E., Delmon, B., 1994. Influence of water in the deactivation of a sulfided

NiMo/[gamma]-Al2O3 catalyst during hydrodeoxygenation. J. Catal. 146,281–291.

LaVopa, V., Satterfield, C.N., 1987. Catalytic hydrodeoxygenation of dibenzofuran.Energy Fuel 1, 323–331.

Li, C.L., Xu, Z.R., Cao, Z.A., Gates, B.C., Petrakis, L., 1985. Hydrodeoxygenation of1-naphthol catalyzed by sulfided Ni–Mo/g-Al2O3: reaction network. AIChE. J.31, 170–174.

Milne, T.A., Agblevor, F., Davis, M., Deutch, S., Johnson, D., 1997. Developments inThermal Biomass Conversion. Academic and Professional, London, UK.

Nimmanwudipong, T., Runnebaum, R., Block, D., Gates, B., 2011. Catalytic reactionsof guaiacol: reaction network and evidence of oxygen removal in reactions

with hydrogen. Catal. Lett. 141, 779–783.Okafor, O.C., Tadepalli, S., Tampy, G., Lawal, A., 2010. Cycloaddition of isoamylene

and alfa-methylstyrene in a microreactor using Filtrol-24 catalyst: micro-reactor performance study and comparison with semi-batch reactor perfor-mance. Int. J. Chem. React. Eng. 8.

Perry, R.H., Green, Don W., 1997. Perry’s Chemical Engineering Handbook, seventhed. McGraw-Hill.

Qian, D., Lawal, A., 2006. Numerical study on gas and liquid slugs for Taylor flow ina T-junction microchannel. Chem. Eng. Sci. 61, 7609–7625.

Senol, O.l., 2007. Hydrodeoxygenation of Aliphatic and Aromatic Oxygenates onSulphided Catalysts for Production of Second Generation Biofuels. Departmentof Chemical Technology, Helsinki University of Technology.

Sharma, R.K., Bakhshi, N.N., 1993. Conversion of non-phenolic fraction of biomass-derived pyrolysis oil to hydrocarbon fuels over HZSM-5 using a dual reactorsystem. Bioresour. Technol. 45, 195–203.

Tadepalli, S., Halder, R., Lawal, A., 2007a. Catalytic hydrogenation of o-nitroanisolein a microreactor: reactor performance and kinetic studies. Chem. Eng. Sci. 62,2663–2678.

Tadepalli, S., Qian, D., Lawal, A., 2007b. Comparison of performance of micro-reactor and semi-batch reactor for catalytic hydrogenation of o-nitroanisole.Catal. Today 125, 64–73.

Venderbosch, R.H., Ardiyanti, A.R., Wildschut, J., Oasmaa, A., Heeres, H.J., 2009.Stabilization of biomass-derived pyrolysis oils. J. Chem. Technol. Biotechnol.85, 674–686.

Voloshin, Y., Lawal, A., 2010. Overall kinetics of hydrogen peroxide formation bydirect combination of H2 and O2 in a microreactor. Chem. Eng. Sci. 65,1028–1036.

Wright, M.M., Daugaard, D.E., Satrio, J.A., Brown, R.C., 2010. Techno-economicanalysis of biomass fast pyrolysis to transportation fuels. Fuel 89, S2–S10.