Production of liquid hydrocarbon fuels by catalytic conversion of biomass-derived levulinic acid

Article abstract of DOI:10.1039/c1gc15047b

Levulinic acid derived from ligno-cellulosic biomass has the potential to be utilized as a platform intermediate molecule in the production of renewable liquid fuels for the transportation sector. Herein we report a catalytic process for the conversion of levulinic acid to γ-valerolactone (GVL) using a RuRe/C catalyst that is significantly more active than a traditional Ru/C catalyst. The bimetallic catalyst is active for the reduction of levulinic acid and simultaneous decomposition of formic acid with good stability in the presence of sulfuric acid, the homogeneous catalyst commonly used in the production of levulinic acid from carbohydrates. Results from techno-economic analyses show that the integration of this new process with catalytic decarboxylation of GVL to butene followed by alkene oligomerization could provide a cost-effective route for the conversion of ligno-cellulosic biomass to liquid hydrocarbon fuels.

Full text of DOI:10.1039/c1gc15047b

View Online / Journal Homepage / Table of Contents for this issue
Dynamic Article Links
Green Chemistry
Cite this: Green Chem., 2011, 13, 1755
www.rsc.org/greenchem
PAPER
Production of liquid hydrocarbon fuels by catalytic conversion of
biomass-derived levulinic acid†
Drew J. Braden, Carlos A. Henao, Jacob Heltzel, Christos C. Maravelias and James A. Dumesic*
Received 13th January 2011, Accepted 16th March 2011
DOI: 10.1039/c1gc15047b
Levulinic acid derived from ligno-cellulosic biomass has the potential to be utilized as a platform
intermediate molecule in the production of renewable liquid fuels for the transportation sector.
Herein we report a catalytic process for the conversion of levulinic acid to g-valerolactone (GVL)
using a RuRe/C catalyst that is significantly more active than a traditional Ru/C catalyst. The
bimetallic catalyst is active for the reduction of levulinic acid and simultaneous decomposition of
formic acid with good stability in the presence of sulfuric acid, the homogeneous catalyst
commonly used in the production of levulinic acid from carbohydrates. Results from
techno-economic analyses show that the integration of this new process with catalytic
decarboxylation of GVL to butene followed by alkene oligomerization could provide a
cost-effective route for the conversion of ligno-cellulosic biomass to liquid hydrocarbon fuels.
provide a cost-effective route for the conversion of ligno-
Introduction
cellulosic biomass to liquid hydrocarbon fuels.
The production of liquid hydrocarbon fuels from renewably
Previous work has shown that levulinic acid has the potential
grown ligno-cellulosic biomass is a challenge that has been
to be produced at cost effective margins3 by the hydrolysis and
the central focus of recent studies.1 In this respect, strategies
subsequent dehydration of hexose sugars, the most prevalent
for using chemical catalysts to produce liquid hydrocarbon
carbohydrate component of ligno-cellulosic biomass, at molar
biofuels have focused on identifying chemical transformations
yields higher than 60%,4–6 accompanied by an equal molar
of carbohydrate-based molecules that accomplish two general
amount of formic acid. Thus, catalyst systems have been
goals: (i) selective deoxygenation of carbohydrates for the
developed for the conversion of levulinic acid into fuel additives
production of functional intermediates, and (ii) C–C coupling
such as GVL, ethyl levulinate (EL), and methyltetrahydrofuran
between these intermediates for the production of higher
(MTHF).3,7,8 However, because of the inherent blending limits
molecular weight compounds for use in transportation fuels.
of these fuel additives in hydrocarbon fuels, studies have also
In the present paper, we address the role of levulinic acid as
addressed catalytic reactions for the conversion of GVL to
a functional intermediate for the conversion of cellulose to
higher molecular weight hydrocarbons that are fully compatible
liquid hydrocarbon fuels. More generally, levulinic acid has
with gasoline, diesel and jet fuels used today.6,9,10 Serrano–
been identified by Bozell and Petersen2 as one of the top 10
Ruiz and co-workers reported the production of 5-nonanone-
platform molecules for the production in biorefineries of value-
based diesel fuel components by the conversion of GVL to
added chemicals and liquid transportation fuels. We describe
valeric acid (VA) combined with ketonization.6 Lange and co-
herein a catalytic process for the conversion of levulinic acid to
workers reported a strategy for the production of VA-based
g-valerolactone (GVL) in the presence of sulfuric acid, thereby
providing a new route for the recovery and recycling of the
fuel components, such as alkyl (mono/di) valeric esters, by
combination of an intermediate VA production step followed by
sulfuric acid used in the cellulose deconstruction step; and, we
esterification with small poly-alcohols.9 More recently, Bond,
present results from techno-economic analyses to show that the
et al. reported a dual reactor approach for the production of
integration of this new process with catalytic decarboxylation
C8–C24 alkenes from an aqueous solution of GVL, involving
of GVL to butene followed by alkene oligomerization could
decarboxylation of GVL to butene over an acidic silica-alumina
catalyst followed by oligomerization of butene isomers over an
acidic ion-exchange catalyst (Amberlyst-70).10
While the above catalytic processes are promising approaches
for the conversion of GVL to liquid transportation fuels, they
all rely on the cost effective production of GVL from biomass-
derived levulinic acid. Although several catalyst systems have
Department of Chemical Engineering, University of Wisconsin-Madison,
Madison, Wisconsin, 53706, USA
† Electronic supplementary information (ESI) available: Details on
techno-economic analyses, extraction of GVL using butyl acetate, and
the reactor metal leaching study. See DOI: 10.1039/c1gc15047b
This journal is
The Royal Society of Chemistry 2011
Green Chem., 2011, 13, 1755–1765 | 1755
©

View Online
been reported for the reduction of levulinic acid to GVL,11–14
most studies have used idealistic feeds of levulinic acid and
processing conditions that are not amenable for use in a cascade
process operating with streams of levulinic acid derived from
ligno-cellulosic biomass. Thus, we have studied heterogeneous
catalysts for the production of GVL from levulinic acid
feedstocks produced by the deconstruction of ligno-cellulosic
biomass in aqueous solutions of sulfuric acid, and we have
developed an integrated strategy for the conversion of biomass
to alkene oligomers. Importantly, we have carried out techno-
economic studies and sensitivity analyses that provide insights
into the parameters that govern this chemical-catalytic strategy.
Fig. 2 Turnover frequency versus time-on-stream for flow reactor
studies of the conversion of formic acid (triangles) and levulinic acid
(squares) at 423 K and 35 bar over a Ru/C catalyst (aqueous feed
composition: [formic acid] = [levulinic acid] = 0.3 mol L-1. Sulfuric acid
was added to the aqueous feed at 50 h time-on-stream (dashed line)
[H2SO4] = 0.5 mol L-1).
Results
Levulinic acid reduction over a Ru/C catalyst
The production of levulinic acid from cellulose is most
commonly accomplished using aqueous sulfuric acid as the
hydrolysis catalyst,5,15 and thus, we aimed at achieving the
catalytic reduction of levulinic acid to GVL, combined with
decomposition of the formic acid co-product to H2 and CO2,
both in the presence of sulfuric acid, as shown in Fig. 1.
While reports in literature suggest that ruthenium should be
an active metal for the heterogeneously catalyzed reduction
of levulinic acid to GVL,12,13 previous studies have reported
that ruthenium may not be suitable for use in the presence of
sulfuric acid. Osada, et al.16 reported that a Ru/TiO2 catalyst,
used for the gasification of lignin to methane, deactivated
significantly after soaking the catalyst in a sulfuric acid solution.
Similarly, Miyazawa, et al.17 reported that Ru/C, used in
combination with Amberlyst-70 for the conversion of glycerol to
1,2-propanediol, was deactivated by sulfur and SO2 evolved from
Amberlyst-70 during reaction. Conversely, ruthenium has been
identified as a catalytic component that possesses high activity
for hydrodesulfurization and hydrogenation reactions used in
hydrotreating schemes.18,19 Accordingly, we started our studies
using a 5 wt% Ru/C catalyst to carry out the decomposition of
formic acid to H2 and CO2 and the simultaneous reduction of
levulinic acid to produce GVL.
the feed at a level of 0.5 mol L-1. It can be seen that the
catalytic activity decreased slowly from the beginning of the
experiment until the addition of sulfuric acid (dashed line),
followed by a rapid decrease in catalytic activity to a significantly
lower level (approximately 3% of the original activity) upon
addition of sulfuric acid. Although operating at a low turnover
frequency, the results in Fig. 2 demonstrate for the first time
that a Ru/C catalyst can operate with stable activity for the
production of GVL from levulinic acid (and decomposition of
formic acid) in the presence of sulfuric acid. The weight hourly
space velocity (WHSV) was decreased by a factor of 3.3 after
sulfuric acid addition (at 100 h time-on-stream) to compensate
for low levulinic and formic acid conversions (<5% conversion).
During these experiments, we observed the formation of an
anhydride species comprised of levulinic acid and sulfuric acid,
formed through a reversible and catalyst-independent reaction
at low WHSV. Similar sulfuric acid anhydride species have
been reported in literature20 and the levulinic acid-sulfuric acid
anhydride species formed in the current study were confirmed
using liquid chromatography coupled with mass spectrometry.
High selectivities for production of GVL (>98%) were observed
before addition of sulfuric acid, which decreased (to 60–
70%) after addition of sulfuric acid due to formation of the
aforementioned anhydride species. Measurable quantities were
not detected in the reactor effluent of 4-hydroxyvaleric acid,
resulting from the partial hydrogenation of levulinic acid, or
angelica lactones, resulting from the cyclization and dehydration
of levulinic acid.21 After taking into consideration the formation
of the anhydride species, the material balance for sulfuric acid
closed to within 99%; no measurable amounts of H2S were
observed. Analysis of the gas phase showed that the formic acid
was almost completely converted to CO2 (>98%), and only trace
amounts of methane were observed. Following the addition of
sulfuric acid, the material balance for levulinic acid closed to
within 95%. The levulinic and formic acid conversions for all
points measured were approximately between 5–35% and 15–
50%, respectively. No measurable quantities (>1 ppm) of catalyst
metals were observed in the liquid reactor effluent, as measured
by inductively coupled plasma (ICP) analysis.
Fig. 1 Reaction scheme for formic acid decomposition and levulinic
acid reduction for the production of GVL.
Fig. 2 presents the catalytic turnover frequencies (moles
species converted per mole of surface metal sites per time) versus
time-on-stream for the conversion of levulinic acid and formic
acid over a 5 wt% Ru/C catalyst in a flow reactor. Initially, a
feed of levulinic and formic acid (0.3 mol L-1 of each acid) was
passed over the 5% Ru/C catalyst at 423 K and 35 bar without
sulfuric acid in the feed. A minimal co-feed of H2 (5 cm3 min-1
(STP)) was used to ensure that the catalyst remained reduced.
Then, after 50 h time-on-stream, sulfuric acid was added to
1756 | Green Chem., 2011, 13, 1755–1765
This journal is
The Royal Society of Chemistry 2011
©

View Online
Levulinic acid reduction over a RuRe/C catalyst
promoter of Ru for conversion of levulinic and formic acids
in the presence of sulfuric acid.
Based on the stable catalytic activity measured for the conversion
of levulinic acid and formic acid in the presence of sulfuric acid
over the Ru/C catalyst, we worked towards the identification
of other Ru-based materials that might show higher catalytic
activity. Bimetallic catalysts that couple typical hydrogenation
metals (e.g., Pt, Pd, Ru) with more oxophilic metals (e.g., Re,
Mo) have proven to be advantageous for various conversion
processes.22 For example, researchers at DuPont demonstrated
that Pd–Re and Ru–Re based catalysts are active for the conver-
sion of maleic acid to tetrahydrofuran,23–25 and Pd-Re particles
have also been demonstrated to be active for the hydrogenolysis
of other carboxylic acids.26 According to the phase diagram for
a Ru–Re mixture, the two metals are completely miscible at all
temperatures of interest in the current study. Accordingly, we
prepared a bimetallic 15 wt% RuRe(3 : 4)/C catalyst for use in
the levulinic reduction reaction.
The turnover frequencies versus time-on-stream for conver-
sion of levulinic and formic acids (2.2 mol L-1 of each acid) over
the RuRe/C catalyst are presented in Fig. 3, in the absence as
well as in the presence of sulfuric acid. It can be seen that stable
catalytic activity was achieved in the presence of sulfuric acid.
In fact, the catalytic activity in the presence of sulfuric acid was
higher by a factor of two than in the absence of sulfuric acid.
Fig. 3 also shows the rate of GVL formation, from which it
can be seen that the RuRe/C catalyst achieves high selectivity
for the formation of GVL (>95%) in the presence and in the
absence of sulfuric acid. The levulinic acid conversions for all
points measured were between 15–40%. Because of the increased
activity of the RuRe/C catalyst, reaction kinetics studies of this
catalyst were carried out at higher WHSV, and measurable levels
of anhydride species comprised of levulinic acid and sulfuric
acid were not observed. Formic acid was completely converted
for all samples with a high selectivity to CO2 (>99%). The
material balance for levulinic acid conversion over the RuRe/C
catalyst closed to within 95%, and no measurable quantities of
byproducts were observed in the reactor effluent. In addition, no
measurable quantities (>1 ppm) of catalyst metals were observed
in the liquid effluent streams from the flow reactor. Comparison
of the results in Fig. 2 and 3 shows that Re is an effective
High conversion catalyst comparison
For the GVL production step to be cost effective on an industrial
scale, a high selectivity to GVL must be achieved at a high
levulinic acid conversion. Thus, we carried out flow reaction
experiments at high levulinic acid conversions (>80%) for both
the 5 wt% Ru/C and 15 wt% RuRe(3 : 4)/C catalysts using a
concentrated levulinic and formic acid feed (2.2 mol L-1 of each
acid) in the presence of sulfuric acid (0.5 mol L-1) at 423 K. (As
discussed below, this feed composition is typical of that expected
from the deconstruction of cellulose.) For the high conversion
experiments, the steady state GVL production rate for the Ru/C
catalyst (0.0028 mmol min-1 g-cat-1) is more than 10 times lower
than the steady state GVL production rate for the RuRe/C
catalyst (0.034 mmol min-1 g-cat-1). We also note that the steady-
state rate for the Ru/C catalyst is significantly lower than the
rate measured using the lower concentration feed (0.3 mol L-1)
at a lower conversion (approximately 0.02 mmol min-1 g-cat-1).
Catalyst characterization
The Ru/C and RuRe/C catalysts used for the production of
GVL were analyzed using various characterization techniques.
The number of surface metal sites was determined for reduced
Ru/C and RuRe/C catalyst samples by CO chemisorption.
The total number of surface metal sites increased from 150 to
247 mmol sites g-1 following the impregnation of Re on the Ru/C
catalyst (4 : 3 molar ratio of Re to Ru).
Temperature programmed reduction (TPR) experiments were
carried out on fresh Re/C, Ru/C and RuRe/C catalyst samples.
Results for the H2 uptake as a function of temperature are shown
in Fig. 4. The H2 uptake for the Ru/C catalyst is centered
at approximately 400 K, whereas the Re/C catalyst showed a
broader H2 uptake that was centered at a higher temperature
(approximately 650–700 K). Interestingly, the RuRe/C catalyst
showed a H2 uptake that was centered at approximately 450 K,
a temperature that is slightly higher than for Ru/C but much
lower than Re/C. This finding suggests that Ru catalyzes the
reduction of rhenium oxide. Thus, Ru and Re are suggested to
Fig. 3 Turnover frequency versus time-on-stream for the conversion
of an equal molar (2.2 mol L-1) solution of levulinic acid (squares)
and formic acid (not shown) and the production of GVL (circles) over
RuRe/C at 423 K and 35 bar in the presence (solid symbols) and absence
(open symbols) of sulfuric acid (0.5 mol L-1).
Fig. 4 TPR profiles for fresh Re/C, Ru/C and RuRe/C catalyst
samples.
This journal is
The Royal Society of Chemistry 2011
Green Chem., 2011, 13, 1755–1765 | 1757
©

View Online
be in intimate contact on the carbon support for the RuRe/C
catalyst under reducing conditions.
X-Ray diffraction (XRD) experiments were carried out on
Ru/C and RuRe/C catalyst samples before and after use in flow
reaction kinetics experiments, both in the presence and absence
of sulfuric acid. Diffraction peaks corresponding to Ru and Re
particles were not observed for all catalyst samples measured.
Thus, the Ru and Re particles are well-dispersed on the carbon
support before and after the reaction.
Integrated biofuel production strategy
Using experimental results for the production of GVL over
a RuRe/C catalyst in the presence of sulfuric acid at high
conversions of levulinic acid, we developed an integrated
biofuel production strategy that involves the complete utilization
of ligno-cellulosic biomass. The framework of our chemical
catalytic approach is based on the biochemical process model
originally developed by the National Renewable Energy Labora-
tory (NREL) for the conversion of corn stover to ethanol.27–30
A
schematic representation for our integrated conversion strategy
is presented in Fig. 5, and includes processing steps required
to handle each component of biomass (e.g., hemi-cellulose,
cellulose and lignin). The hemi-cellulose fraction of the biomass
is first removed through dilute acid pretreatment that produces
an aqueous stream of hemi-cellulose monomers (mainly xylose
when starting with corn stover).31,32 We then pass the solid
biomass containing the cellulose and lignin fractions to a
reactor for cellulose deconstruction to levulinic and formic
acids in aqueous sulfuric acid, followed by our cascade catalytic
process for the production of liquid hydrocarbon transporta-
tion fuels. The recovered xylose fraction, the insoluble lignin
fraction and the solid humins (i.e., insoluble polymerization
products formed during cellulose deconstruction) are sent to
a boiler/turbogenerator for the production of process heat and
electricity.
Fig. 6 Schematic flowsheet for the cascade conversion of cellulose to
liquid alkene oligomers.
423 K to an aqueous solution of sulfuric acid (0.5 mol L-1), as
recently reported,6 achieving high levels of levulinic and formic
acids (>2 mol L-1). The yield of levulinic and formic acids from
the cellulose deconstruction is approximately 55%, with the re-
mainder of the partially deoxygenated carbohydrate monomers
forming humins.6 Following the cellulose deconstruction step,
the insoluble lignin and humins formed during the reaction
are separated from the aqueous stream using a settling tank.
The aqueous stream of levulinic and formic acids, containing
sulfuric acid, can be converted to GVL at essentially quantitative
yields over the 15% RuRe(3 : 4)/C catalyst (catalytic reactor 1)
at 423 K. The increased hydrophobic nature of GVL enables its
separation from the aqueous layer by liquid–liquid extraction
with an acetate solvent, as reported previously.6 We selected
butyl acetate over other acetate solvents (e.g., propyl, isobutyl
and isoamyl acetates), because it is relatively inexpensive and has
adequate selectivity for this extraction process. Results for GVL
extraction using butyl acetate as the solvent are reported in the
supporting information in Table S4.† Following the extraction
of GVL, the sulfuric acid aqueous solution is recycled back to
the deconstruction reactor. Essentially all of the butyl acetate
(>99%) can be recovered by simple distillation. The GVL
product can then be upgraded to alkene oligomers using the
dual reactor catalytic conversion process developed by Bond,
et al.,10 in which an aqueous solution of GVL (60 wt%) is
Fig. 5 Schematic diagram for the integrated production of liquid alkene
oligomers from ligno-cellulosic biomass.
Fig. 6 shows a schematic flowsheet of our cascade process
to convert the cellulose portion of ligno-cellulosic biomass to
liquid hydrocarbon fuels. The conversion of cellulose to an equal
molar mixture of levulinic and formic acids can be carried out
(in a batch reactor) by the progressive addition of cellulose at
1758 | Green Chem., 2011, 13, 1755–1765
This journal is
The Royal Society of Chemistry 2011
©

View Online
first decarboxylated to form a mixture of butene isomers over a
silica-alumina catalyst (catalytic reactor 2) at 648 K and 36 bar.
Essentially all of the water is condensed in a separator at 373–
383 K and 36 bar and recycled to be mixed with more GVL. The
butene, in the presence of CO2, can then be oligomerized to
larger alkene oligomers over an Amberlyst-70 catalyst (catalytic
reactor 3) at 443 K and 36 bar. The oligomers can be condensed
at ambient conditions leaving a relatively pure steam of CO2
at 36 bar. Based on the study by Bond, et al., we assume that
99% of the GVL produced from levulinic acid is converted to
alkene oligomers (of carbon length C8–C24). The distribution of
oligomers produced, as reported by Bond et al.,10 is included in
the supporting information (Table S3†).
boiler/turbo generation unit to produce heat and electricity. In
scenario 1, we assume that 85% of the C5 fraction of the hemi-
cellulose is recovered, concentrated (using excess heat), and
sent to the reboiler/turbogenerator. Using catalytic activities
reported in literature, we determined the catalyst amounts
required in the batch reactor (cellulose deconstruction reactor),6
reactor 2 (GVL decarboxylation reactor)10 and reactor 3 (butene
oligomerization reactor).10 Experimental results for the high
conversion study discussed above were used to determine the
required amount of RuRe/C catalyst for reactor 1 (levulinic acid
reduction reactor) which used inlet feed concentrations that are
similar to those used in the experimental investigation above.
Finally, to facilitate the comparison between the two scenarios,
the processing capacity in scenario 2 was adjusted so that the
two scenarios have the same C6 sugar throughput and therefore
the same oligomer production capacity.
Techno-economic analysis
We performed techno-economic analyses based on ASPEN
simulation models. The process model specifications were based
in part on values reported by NREL for the conversion of corn
stover to ethanol.27–30 Our goal was to establish a common
basis for the comparison of alternative biofuel production
strategies. To that end, we initially used the same corn stover
feedstock and processing rate (2000 MT dry corn stover per day),
similar process equipment whenever possible, the same costs
for feedstock, chemicals and utilities, and the same economic
parameters (e.g., reference index year, income tax rate, return
on investment, depreciation). Due to the large plant capacity,
the plant design consists of three identical processing trains
operating in parallel and each train has multiple parallel unit
operations (e.g., reactors, distillation columns). The details of
our process model are included in the supporting information.†
In addition to examining the overall feasibility of our pro-
posed chemical catalytic strategy, we also studied the significance
of using an alternative raw material. To that end, we developed
two processing scenarios. In scenario 1, corn stover is used as
raw material, while in scenario 2, loblolly pine, an important
and prevalent softwood that grows readily in the southern
United States, is used. It is important to note that the hemi-
cellulose fraction of loblolly pine is rich in C6 sugars,33 instead
of C5 sugars which are abundant in the hemi-cellulose of corn
stover (see Table 1). Thus, for the case of loblolly pine, a higher
fraction of the original biomass can be transformed into valuable
liquid fuels using the proposed route. Furthermore, the use
of loblolly pine as a feedstock leads to a simplified flowsheet
in which the dilute acid pretreatment step used to extract C5
sugars from the hemi-cellulose fraction of the corn stover is no
longer needed. Additionally, the small fraction of C5 sugars in
the raw material leaves the process as humins along with the
lignin. In both cases the unused materials (humins, lignin, and
recovered C5 sugars in the case of corn stover) are sent to the
Capital cost comparison
The capital costs for the two aforementioned oligomer pro-
duction scenarios as well as the ethanol production process
analyzed by NREL are shown in Table 2. The same capital
costs were used for processing steps that are similar to both the
production of alkene oligomers and the production of ethanol
(e.g., feedstock handling, pretreatment, storage, utilities). The
cost of the boiler/turbogenerator for the oligomer production
process is estimated using the standard cost scaling equation
(eqn (1)),34 which relates the unit capacity (e.g. kmol h-1 carbon
sent to boiler/turbogenerator) to its cost, using reference values.
0.62
⎛
⎜
⎜
⎞
⎟
⎟
⎟
⎟
⎠
Capacity
Capacity
(1)
Cost = Costref
⎜
⎜
⎝
ref
The capital costs for all unit operations are given in Table 2.
First, we note that the cost of the boiler/turbogenerator unit
in the three processes ranges between 34 and 49% of the total
capital cost, making it the most expensive unit. Thus, an accurate
estimation of its capital cost as a function of its capacity, using
eqn (1), is important. Second, the results in Table 2 show that the
total installed equipment cost for scenario 1 (corn stover) is 23%
($38.3 million) higher than the ethanol process. This difference
is due mainly to the larger boiler/turbogenerator ($33.8 million
of difference). Third, compared to the ligno-cellulosic ethanol
strategy, the cost of scenario 2 (loblolly pine) is approximately
9% lower ($15.5 million). This difference is because the process
in scenario 2 has no preprocessing section and a smaller turbo
generation unit due to the higher loblolly pine mass fraction used
to generate fuels (i.e., a lower biomass fraction sent to the boiler).
Finally, we note that the cost of the equipment required to
catalytically convert cellulose into alkene oligomers in both sce-
narios ($53 million) is similar to the cost of the equipment used
for the saccharification/fermentation and distillation/solids
recovery sections in the ethanol process ($48 million).
Table 1 General feedstock compositions and costs used in techno-
economic analyses
Component
Corn Stover
Loblolly Pine
C6 carbohydrates
C5 carbohydrates
Lignin
41%
24%
18%
17%
83
55%
9%
32%
4%
Operating cost comparison
A comparison of the operating costs for the biochemical
production of ethanol and the catalytic production of alkene
oligomers is shown in Fig. 7. We assumed a price of $83 per
Other
Cost (US$ dry-MT-1)
57
This journal is
The Royal Society of Chemistry 2011
Green Chem., 2011, 13, 1755–1765 | 1759
©

View Online
Table 2 Capital costs for the ethanol production process and for alkene oligomer production process scenario 1 (corn stover processing) and scenario
2 (loblolly pine processing)
Ethanol process
(MM$)
Oligomer process scenario 1
(MM$)
Oligomer process scenario 2
(MM$)
Process Sections
(%)
(%)
(%)
7.3
Feedstock handlinga
Pretreatmenta
10.9
36.2
21.8
26.1
3.5
3.2
56.1
6.3
6.6
10.9
36.2
5.4
17.9
10.9
22.1
13.3
15.9
2.1
2.0
34.2
3.8
Saccharification/fermentationa
Distillation/solids recoverya
Wastewater treatmentb
Storagea
2.8
3.2
89.9
6.3
12.3
7.8
1.5
8.4
12.8
0.9
8.8
1.4
1.6
44.4
3.1
6.1
3.9
0.8
4.1
6.3
0.4
4.4
0.4
2.9
3.2
72.1
6.3
12.7
7.7
1.5
8.2
12.8
0.9
8.8
1.9
2.2
48.5
4.2
8.5
5.1
1.0
5.5
8.6
0.6
5.9
0.5
Boiler/turbogeneratorb
Utilitiesa
Cellulose deconstructionc
GVL Productionc
Cooling/CO2 separationc
GVL extraction/H2SO4 recyclec
Butene productionc
Butene & CO2 separationc
Butene oligomerizationc
Oligomer condensationc
Total installed equip. cost
Total project investmentd
0.7
202.4
465.5
0.7
148.6
341.7
164.1
377.4
a Taken from study by Kazi, et al.30 b Estimated from value reported by Kazi, et al.30 c Determined using Aspen Icarus. d Includes other direct (e.g.,
instrumentation, insulation) and indirect (e.g., engineering, contingency) costs.
solid waste disposal (e.g., ash). Using current market prices,
the purchase cost of precious catalyst metals (i.e. Ru and Re)
required for scenarios 1 and 2 is $33 million. We assume that
the catalyst is leased and has a salvage value at the end of
the life of the project. Thus, the catalyst cost is included as an
operating cost for the refurbishing and continued maintenance
of the GVL production catalyst (see supporting information
for more details†). It should be noted that the costs for
the SiO2/Al2O3 catalyst (used in GVL decarboxylation) and
Amberlyst-70 catalyst (used in butene oligomerization) are
negligible compared to the RuRe/C catalyst, and hence they
are not considered in this analysis.
As shown in Fig. 7, the feedstock cost is the most significant
fraction of operating costs in all strategies. Nevertheless, the
feedstock cost as well as the cost of “other raw materials”
in scenario 2 are significantly lower than the corresponding
costs in scenario 1 and the bioethanol strategy. This difference
is due to a lower raw material price (in US$ dry-MT-1) and
Fig. 7 Annual operating costs for the biochemical production of
the higher C6 sugar content, which means that less feedstock
is required to produce the same amount of fuel. The cost
of “other raw materials” is lower because scenario 2 has no
preprocessing stage. Also, we note that the cost of the enzyme
for the biochemical production of ethanol is considerable
and surpasses any variable cost for the production of alkene
oligomers.
Another aspect in the production of alkene oligomers is the co-
production of electricity and process heat. To estimate the output
of the larger boiler/turbogenerator, reference values previously
reported for ethanol production29 were scaled using simple
mass and energy balances. Our analysis indicates that oligomer
production scenarios 1 and 2 consume approximately 30% of
the boiler/turbogenerator heating, and provide an electricity
surplus which translates into an additional annual revenue of
approximately 24 and 20 MM$ yr-1, respectively.
ethanol (reported by Kazi, et al.,30 blue), alkene oligomer production
scenario 1 (corn stover, red), and scenario 2 (loblolly pine, green). Total
annual costs: ethanol process = 140 MM$, oligomer scenario 1 = 90
MM$, oligomer scenario 2 = 46 MM$.
dry MT for corn stover28 and $57 per dry ton for loblolly pine
($26 per green short ton-1 assuming 50% moisture) based on
a 50 mile collection radius specific to south eastern Georgia.35
The “other raw materials” includes the sulfuric acid and lime
used in the pretreatment of corn stover (to release the hemi-
cellulose), wastewater treatment chemicals and polymers, and
the boiler and cooling tower chemicals. Fixed costs include
employee salaries, insurance, overhead and maintenance. Waste
disposal costs include gypsum disposal following the sulfuric
acid neutralization during biomass pretreatment and other
1760 | Green Chem., 2011, 13, 1755–1765
This journal is
The Royal Society of Chemistry 2011
©

View Online
Revenue-cost balance
2 is lower ($4.31 GGE-1 vs. $7.78 GGE-1), which leads to a lower
revenue from sales at the break-even point.
Using the calculated capital and operating costs, a minimum
selling price (MSP) for the produced fuels can be determined. To
make the two oligomer routes comparable to the ethanol process,
fuel production (i.e., ethanol and oligomers) was expressed in
terms of gallons of gasoline equivalent (GGE). The calculated
MSP value for the NREL ethanol process30 is $5.13 GGE-1,
while the values for oligomer production scenarios 1 and
2 are equal to $7.78 GGE-1 and $4.31 GGE-1, respectively.
Fig. 8 shows how the total production cost (based on the
aforementioned capital and operational cost calculations) is
balanced by the sales of electricity and fuels. First, we note
that the total production cost of the ligno-cellulosic ethanol
route is higher than in both oligomer scenarios. This difference
is due mainly to the larger operating costs (e.g., cost of cellulases
– see Fig. 7). However, the NREL strategy leads to higher
fuel production and sales. In particular, it leads to higher fuel
sales than scenario 1 (although the two routes process the same
amount of feedstock) since its biomass-to-fuels yield is higher.
The NREL strategy also leads to higher production than in
scenario 2, despite the high biomass-to-fuels yield of the latter,
because it processes 2000 MT of corn stover per day versus 1365
MT of loblolly pine per day. Second, we note that the production
cost in scenario 2 is substantially lower than in scenario 1, due
to lower capital (Table 2) and operational (Fig. 7) costs. It is
important to stress here that the amounts of fuel produced in
the two alkene oligomer scenarios are identical (recall that they
convert the same amount of C6 sugars), but the MSP for scenario
Sensitivity analysis
We conducted sensitivity analysis studies to determine how vari-
ations in key economic and processing parameters impact the
MSP in the two oligomer strategies. We investigated the impact
of feedstock price, total installed equipment cost, levulinic acid
reduction catalyst refurbishing cost, and electricity price. To
calculate the corresponding sensitivities, we varied the nominal
(base case) values by 20%. In addition, the impact of selling the
generated CO2 instead of venting it into the atmosphere (base
case) was also considered. The results are shown in Fig. 9 and
10.
Fig. 9 Sensitivity analysis for the production of alkene oligomers –
scenario 1 (corn stover) (high levels = red bars, low levels = blue bars).
Fig. 10 Sensitivity analysis for the production of alkene oligomers –
scenario 2 (loblolly pine) (high levels = red bars, low levels = blue bars).
For scenario 1, the feedstock price and the installed equipment
cost have the largest impact; a 20% variation in the former leads
to a 9.5% change in the MSP, while a similar variation in the
latter leads to a 9.0% change. We note that the feedstock cost
accounts for 65% of the total operating cost, or 40% of the
total production cost (see Fig. 7). For scenario 2, we note that
the two aforementioned parameters remain the most important,
but the impact of the equipment cost is higher; increasing it
by 20% leads to a 12% change in the MSP, while a similar
increase in feedstock price results in a MSP that is only 8%
higher. This difference is because the capital cost for scenario
2, although lower than in scenario 1, is a larger fraction of
the total production cost, as shown in Fig. 8b. At the same
time, the feedstock cost in scenario 2 is 32% of the production
cost (60% of the total operating cost) as opposed to 40% in
scenario 1. The sensitivity of the MSP with respect to the price
Fig. 8 Revenue-cost balance for the biochemical production of ethanol,
and alkene oligomer production scenario 1 (corn stover) and scenario
2 (loblolly pine). The calculation of the return on investment (ROI)
and operating costs as well as electricity sales is based on rigorous
ASPEN models, nominal prices for raw materials and utilities, and the
methodology followed by Kazi, et al.30 Fuel price and hence fuel sales
are adjusted so that the total production cost of each strategy (ROI +
income tax + operating costs) is equal to the total revenues (electricity
sales + fuel sales). The MSP is the price that leads to this “break even”
point. a) Absolute values; b) Percentages of the total.
This journal is
The Royal Society of Chemistry 2011
Green Chem., 2011, 13, 1755–1765 | 1761
©

View Online
of electricity in scenario 1 is 3% and in scenario 2 is 5%, despite
the fact that more electricity is produced and sold in scenario 1.
This result can be explained if we recognize that the lower MSP
of oligomers in scenario 2 leads to substantially smaller fuel
sales, which means that electricity sales in scenario 2 are a larger
fraction of total sales than in scenario 1 (see Fig. 8b). The cost of
catalyst refurbishing had essentially negligible effect (<1% for
both scenarios). Finally, using a CO2 price of $36 ton-1, based on
an amine recovery system,36 we calculated that selling the CO2
in the high pressure gas stream coming from the CO2/butene
separation had a relatively small effect (2% and 4% variation for
scenarios 1 and 2 respectively).
in all strategies is the boiler/turbogenerator. In this respect,
it would be advantageous to couple a biofuels facility with
an existing plant that utilizes similar unit operations (i.e.,
boiler/turbogenerator, wastewater treatment) or to retrofit an
existing facility. In addition to equipment cost considerations,
the facility should be located in close proximity to a supply
of feedstock. For these reasons, a paper mill would appear to
be potentially suitable for effective implementation of a biofuels
process. Since 1970, the number of paper mills has decreased and
the average mill production capacity has increased (from 93,000
to 216,000 MT yr-1).37 This decrease is caused by the closure
of smaller mills (<100,000 MT yr-1 production capacity) and
their replacement with larger mills. Hence, the integration of a
biofuels production process with a relatively smaller scale paper
mill may have promising potential.
Discussion
Impact of newly developed levulinic acid reduction catalyst
systems
To assess this potential, we estimated how the total installed
equipment cost for the loblolly pine processing strategy varies
as a function of the softwood processing capacity (see Fig. 11).
The results presented above were based on a total processing
capacity of 1365 dry MT of loblolly pine per day (478,000
MT yr-1), corresponding to a total installed equipment cost
of $149 million. However, if an alkene oligomer production
process was developed in a retrofitted paper mill for a lower
processing capacity of 200 MT dry softwood per day (70,000 MT
yr-1), then the total installed equipment cost would decrease to
approximately $44 million. Furthermore, if the facility already
included a boiler/turbogenerator and wastewater treatment
equipment, then the additional required installed equipment
cost could potentially be decreased to as low as $22 million. At
the same time, the projected MSP of $4.31 GGE-1 should still
be a valid approximation for plant capacities as low as 200 dry
MT of loblolly pine per day, because the design for our process
is composed of 3 identical processing trains each containing a
number of parallel units.
Due to their inherently high value, the stability of precious
metal catalysts is an important concern. The experimental
results described above demonstrate that the RuRe/C catalyst
maintains stable activity in the presence of sulfuric acid. The
TPR study revealed that the Ru and Re metals are in intimate
contact on the RuRe/C catalyst under reducing conditions. The
XRD study demonstrated that the Ru and Re particles are well-
dispersed on the carbon support even after reaction testing
in the presence of sulfuric acid. In addition, ICP analyses of
reactor effluent streams (see supporting information†) revealed
that under the acidic conditions used during reaction kinetics
studies in the presence of sulfuric acid, the catalyst is exposed
to relatively high concentrations (e.g., 100 ppm) of metal
ions that are leached from the reactor walls (i.e., Fe, Ni, Cr,
Co, Mo). While the reaction kinetics results presented here
demonstrate that the RuRe/C bimetallic catalyst shows good
stability and selectivity in the presence of these metal ions from
the reactor walls, it is recommended that additional reaction
kinetic studies be conducted using various biomass-synthesized
feeds to determine the long term effects of soluble humins or
other soluble residues on the performance of this catalyst.
The application of the newly developed bimetallic RuRe/C
catalyst for the reduction of levulinic acid in the presence of
sulfuric acid could lead to several advantages for an integrated
biofuel production scheme. We have demonstrated that the
levulinic acid reduction step can be operated at high conversion
with good selectivity to GVL, thus minimizing concerns related
to levulinic acid recycle. The proposed GVL extraction process,
utilizing butyl acetate, enables the recycling of most of the sulfu-
ric acid catalyst, eliminating costly neutralization and disposal
steps. The simultaneous conversion of formic acid provides an
internal hydrogen source that can be used to minimize external
hydrogen requirements. In addition, the carbon removed from
the carbohydrate starting material as CO2 could potentially be
sequestered in relatively pure streams from the GVL production
and the alkene oligomer production steps.
Fig. 11 Total capital cost as a function of biomass processing capacity
for the conversion of a softwood (loblolly pine) feedstock into liquid
alkene oligomers (calculations according to eqn (1)).
Potential for future improvements
Process integration
The focus of the techno-economic analysis in this study was
centered on assessing the feasibility of the alkene oligomer
production strategy in comparison with the production of
ethanol, the most prevalent biomass-derived fuel in the world
Large capital investment is potentially a barrier for the imple-
mentation of new technologies such as those presented here.
As discussed above, the single largest capital cost component
1762 | Green Chem., 2011, 13, 1755–1765
This journal is
The Royal Society of Chemistry 2011
©

View Online
today. This analysis revealed that the proposed chemical-
catalytic strategy compares favorably, especially when feedstocks
rich in C6 sugars are used. Importantly, the MSP value of
$4.31 GGE-1 can be achieved using the proven experimental
yields for levulinic acid production from ligno-cellulose (55%
yield), GVL production from levulinic acid (>99% yield), and
the production of alkene oligomers (C8–C24) from GVL (99%
yield). It should be noted that, in comparison with the ethanol-
based process (85% yield of ethanol from cellulose, 76% yield
of ethanol from xylose), the overall yield of butene oligomers is
relatively low for the base case studied (~50% yield of oligomer
monomers per cellulose monomer processed) with the least
selective conversion step being the production of levulinic acid
(55% yield). However, the production of levulinic acid from
cellulose (and glucose) is not thermodynamically limited, and
preliminary results have suggested that levulinic acid yields in
excess of 70% may be attainable, leaving significant room for
improvement, and ultimately leading to a further significant
reduction in the fuel MSP.
presence of sulfuric acid, a common homogeneous catalyst used
in the conversion of cellulose into levulinic acid.
Techno-economic analyses reveal that the production of liquid
hydrocarbons via our levulinic-acid-based approach is econom-
ically competitive with a state-of-the-art ligno-cellulosic ethanol
production strategy. The most sensitive process parameters are
the feedstock and total installed equipment costs. In addition
to cost, it was determined that the feedstock composition plays
an important role in the process economics, such that biomass
sources rich in C6 sugars allow for more cost effective process
designs. While the proposed process appears to be economically
competitive, there exists ample room for improvement in product
yields (i.e., levulinic acid production) and catalyst activities (i.e.,
levulinic acid conversion) that could further improve the process
economics.
Materials and Methods
Ru/C and RuRe/C catalyst preparation and pretreatment
The present study has compared the production of alkene
oligomers with the production of ethanol in terms of the MSP
based on GGE (gallon gasoline equivalent). However, MSP
should not be the only comparison criterion. First, the proposed
strategy results in a high pressure stream of highly concentrated
CO2, which is ideal for carbon capture and sequestration. In
addition to leading to a carbon negative process, this also means
that if emission regulations or emission trading schemes are in-
troduced, then CO2 consumers will receive credits, which would
further improve the economics of the proposed route. Second,
a mix of alkene oligomers is a better liquid fuel than ethanol,
because it has higher energy density, it is easier to handle, and
it does not require modifications of the existing infrastructure.
If desired, the alkenes can be converted to alkanes by reduction
over a metal catalyst, and only one equivalent of H2 is required
to convert one alkene molecule to its corresponding alkane. In
this respect, the final fuel components produced from the alkene
oligomer strategy do not suffer from blending limitations, in
contrast to ethanol and other oxygenated fuel additives. Also,
the proposed process has the potential to produce oligomer
products that contain more than 93% of the energy content
contained within the C6 sugar monomers and only 31% of the
initial carbohydrate weight (for the stoichiometric conversion of
glucose to alkenes plus CO2 and H2O). Thus, the production
of alkene oligomers from ligno-cellulosic biomass appears to
be an advantageous and economically attractive approach for
producing renewable liquid hydrocarbon transportation fuels
from non-fossil fuel sources.
The 5 wt% Ru/C catalyst used in the levulinic acid reduc-
tion studies was purchased from Sigma-Aldrich. The 15 wt%
RuRe(3 : 4)/C catalyst was prepared by incipient wetness im-
pregnation of the 5 wt% Ru/C catalyst using an aqueous
solution of perrhenic acid (HReO4, 53 wt% Re). Prior to reaction
kinetics measurements, each catalyst was reduced at 673 K
(0.5 K min-1 ramp, 2 h hold) under flowing H2 (100 cm3 min-1
(STP)). For reactions using feeds containing sulfuric acid, the
catalyst was pretreated in situ for 24 h with a 0.5 mol L-1 solution
of sulfuric acid at a flow rate of 0.1 cm3 min-1 at the reaction
temperature in the presence of flowing H2 (25 cm3 min-1 (STP)).
Levulinic acid reduction experiments
Reaction kinetics measurements for the reduction of levulinic
acid were carried out using a packed bed reactor. A schematic
of the flow reactor system used for reaction kinetics studies is
shown in the supporting information in Fig. S1.† A packed-bed
reactor consisting of a half-inch or quarter-inch tube (Hastelloy
C327) was used for the levulinic acid reduction experiments.
Quartz wool was used above and below the catalyst bed to keep
the catalyst in place. The reactor was heated with an aluminium
block heated externally by a well-insulated furnace (Applied
Test Systems). Type-K thermocouples (Omega) were used to
measure the reaction temperature outside the reactor wall.
The reaction temperature was maintained by PID temperature
controllers (Love controls) that were connected to a variable
transformer (Tesco). Mass flow controllers (Brooks 5850E) were
used to regulate the flow of H2, He and other gases during the
experiments. The liquid feed was pumped from a graduated
cylinder by an HPLC pump (Lab Alliance series 1) to a needle
located at the entrance of the catalyst bed. A back-pressure
regulator (GO model BP-60), located downstream from the
reactor, was used to control the total pressure, which was
measured by three gauges located downstream from the liquid
pump, at the inlet of the reactor and downstream of the reactor.
A gas-liquid separator (Pemburthy) at room temperature was
used to collect the liquid phase for analysis. A bubble flow meter,
located downstream from the back pressure regulator, measured
the gas flow rate out of the reaction system.
Conclusions
GVL has been identified as a viable biomass-derived platform
chemical from which liquid hydrocarbon fuels can be derived;
however, its efficient and cost effective production remains a
challenge. Our studies of the conversion of levulinic acid into
GVL over a traditional ruthenium-based catalyst have identified
a more active RuRe/C catalyst. This RuRe/C catalyst is selective
for the production of GVL from levulinic acid while exhibiting a
high activity for the simultaneous decomposition of formic acid.
Additionally, this bimetallic catalyst shows good stability in the
This journal is
The Royal Society of Chemistry 2011
Green Chem., 2011, 13, 1755–1765 | 1763
©

View Online
The aqueous product distributions were measured by HPLC
analysis (Waters, using a Biorad Reztec column) of the liquid
effluent drained from the separator located downstream of the
flow reactor. Non-UV active species (e.g., sugars) were quantified
using a refractive index detector (Waters 410), and UV-active
species were quantified using a UV detector (Waters PDA 990)
at appropriate absorbance wavelengths. Total organic carbon
analysis (TOC) was used to determine the total organic carbon
in solution. An HPLC-MS (Agilent), equipped with an ion trap,
was used to qualitatively identify species in solution.
overall reactor conversions and yields, separation recoveries,
etc. Conversion reactor models coupled with rigorous separation
models (e.g., multi-stage distillation and liquid–liquid extraction
units) were used to obtain accurate values. The solution of
the corresponding material and energy balances provided the
necessary information to conduct equipment sizing.38 Finally,
detailed capital cost calculations based on unit capacities, oper-
ational conditions, and selected materials of construction were
performed using ASPEN Process Economic Analyzer. In order
to have a common basis, we adopted the methodology which was
used for the economic analysis by NREL for the evaluation of
their ethanol process, accounting for all sources of operational
revenue (e.g., liquid fuels and electricity), operational costs
(e.g. raw material and utility consumption) capital depreciation,
taxes, internal rate of return, etc.
Catalyst characterization
Adsorption isotherms for CO on Ru/C and RuRe/C catalysts
were collected using a standard volumetric gas handling system
employing capacitance manometers for precision pressure mea-
surement ( 0.5 ¥ 10-5 mbar). Following reduction pretreatment
at 673 K (0.5 K min-1 ramp, 2 h hold), adsorbed hydrogen was
desorbed by vacuum treatment (1 ¥ 10-7 mbar) for 1 h, and the
sample was subsequently cooled to 300 K. For chemisorption
studies, small amounts of CO (1–10 mmol) were dosed onto the
catalyst until saturation of the surface sites was achieved. The
amount of gas adsorbed was determined volumetrically from
the dose and equilibrium pressures and system volumes and
temperatures. Times of 20–30 min were allowed for each dose to
reach equilibrium. After saturation, the catalyst was evacuated
at 300 K for 1 h to remove weakly bound probe gas molecules
and a second isotherm was collected. The irreversible CO uptake
was calculated by subtracting the two isotherms.
TPR studies were carried out using a continuous flow appa-
ratus consisting of a mass flow controller (Teledyne-Hastings)
and tube furnace connected to a variable power-supply and
PID temperature controller (Love Controls) with a K-type
thermocouple (Omega), shown in the supporting information
in Fig. S2.† The effluent was monitored by a mass spectrometer
system consisting of a quadruple residual gas analyzer (Stanford
Instruments RGA 200) inside a vacuum chamber. Vacuum was
provided by a diffusion pump connected in series to a rotary
pump. The effluent was introduced into the vacuum chamber
via a constricted quartz capillary, resulting in a pressure of 5 ¥
10-5 mbar inside the chamber. Dried, unreduced catalyst samples
(100–300 mg) were loaded into a 12.6 mm (0.5 inch) outer
diameter, fritted quartz tube and reduced in situ. A temperature
ramp of 10 K min-1 was used to heat the catalyst from room
temperature to 873 K where it was held for 20 min.
Acknowledgements
This work was supported by the Department of Energy through
the Great Lakes Bioenergy Research Center.
Notes and references
1 D. Martin-Alonzo, J. Q. Bond and J. A. Dumesic, Catalytic conver-
sion of biomass to biofuels, Green Chem., 2010, 12, 1493–1513.
2 J. J. Bozell and G. R. Petersen, Technology development for the
production of biobased products from biorefinery carbohydrates -
the US Department of Energy’s “Top 10” revisited, Green Chem.,
2010, 12, 539–554.
3 D. J. Hayes, F. Fitzpatrick, M. H. B. Hayes, J. R. H. Ross,
Biorefineries: Industrial Processes and Products, Vol. 1 (Wiley-VCH,
Weinheim), (2006).
4 B. Girisuta, L. P. B. M. Janssen and H. J. Heeres, A kinetic study
on the conversion of glucose to levulinic acid, Chem. Eng. Res. Des.,
2006, 84, 339–349.
5 B. Girisuta, L. P. B. M. Janssen and H. J. Heeres, Kinetic study on
the acid-catalyzed hydrolysis of cellulose to levulinic acid, Ind. Eng.
Chem. Res., 2007, 46, 1696–1708.
6 J. C. Serrano-Ruiz, D. J. Braden, R. M. West and J. A. Dumesic,
Conversion of cellulose to hydrocarbon fuels by progressive removal
of oxygen, Appl. Catal., B, 2010, 100, 184–189.
7 I. T. Horvath, H. Mehdi, V. Fabos, L. Boda and L. T. Mika, g-
Valerolactone - a sustainable liquid for energy and carbon-based
chemicals, Green Chem., 2008, 10, 238–242.
8 H. Mehdi, et al., Integration of homogeneous and heteroge-
neous catalytic processes for a multi-step conversion of biomass:
from sucrose to levulinic acid, g-valerolactone, 1,4-pentanediol, 2-
methyltetrahydrofuran, and alkanes, Top. Catal., 2008, 48, 49–54.
9 J. P. Lange, et al., Valeric Biofuels: A Platform of Cellulosic
Transportation Fuels, Angew Chem, 2010, 122, 4581–4585.
10 J. Q. Bond, D. M. Alonso, D. Wang, R. M. West and J. A. Dumesic,
Integrated Catalytic Conversion of gamma-Valerolactone to Liquid
Alkenes for Transportation Fuels, Science, 2010, 327, 1110–1114.
11 L. E. Manzer, Production of 5-methylbutyrolactone from levulinic
acid, Int. Patent 6617464 B2, (2003).
A Scintag PAD V X-ray diffractometer with monochromated
Cu-Ka X-ray tube was used in XRD studies. The tube voltage
and current were 45 kV and 40 mA, respectively. Diffraction
patterns for were collected in the 5◦ to 40◦ 2q range, with
0.02◦ intervals and a dwell time of 2 s. Prior to diffraction
studies, the catalyst samples were reduced in flowing hydrogen
(200 cm3 min-1) at 673 K and passivated in a flowing 2% O2/He
mixture (150 cm3 min-1).
12 L. E. Manzer, Catalytic synthesis of alpha-methylene-gamma-
valerolactone: a biomass-derived acrylic monomer, Appl. Catal., A,
2004, 272, 249–256.
13 Z.-p. Yan. and L. Lin, Synthesis of g-valerolactone from biomass-
based levulinic acid over Ru/C catalyst in, Energy Fuels, 2009, 23,
3853–3858.
14 R. J. Haan, J. P. Lange, L. Petrus, C. J. M. Petrus-Hoogenbosch,
Hydrogenation process for the conversion of a carboxylic acid or an
ester having a carbonyl group, US Patent 2007208183 (2007).
15 A. Corma, S. Iborra and A. Velty, Chemical routes for the transfor-
mation of biomass into chemicals, Chem. Rev., 2007, 107, 2411–2502.
16 M. Osada, N. Hiyoshi, O. Sato, K. Arai and M. Shirai, Subcritical
water regeneration of supported ruthenium catalyst poisoned by
sulfur, Energy Fuels, 2008, 22, 845–849.
Process synthesis and techno-economic analysis
The process synthesis studies were based on detailed process
simulation models developed in ASPEN Plus, incorporating
experimental data and other published information to specify
raw materials compositions, reactor operational conditions,
1764 | Green Chem., 2011, 13, 1755–1765
This journal is
The Royal Society of Chemistry 2011
©

View Online
17 T. Miyazawa, S. Koso, K. Kunimori and K. Tomishige, Glycerol
hydrogenolysis to 1,2-propanediol catalyzed by a heat-resistant ion-
exchange resin combined with Ru/C, Appl. Catal., A, 2007, 329,
30–35.
18 Y. J. Kuo, R. A. Cocco and B. J. Tatarchuk, Hydrogenation and
Hydrodesulfurization over Sulfided Ruthenium Catalysts: II. Impact
of Surface Phase-Behavior on Activity and Selectivity, J. Catal., 1988,
112, 250–266.
19 Y. J. Kuo and B. J. Tatarchuk, Hydrogenation and Hydrodesulfu-
rization over Sulfided Ruthenium Catalysts, J. Catal., 1988, 112,
229–249.
20 M. Zerella, S. Mukhopadhyay and A. Bell, T. Synthesis of mixed
acid anhydrides from methane and carbon dioxide in acid solvents
in, Org. Lett., 2003, 5, 3193–3196.
28 A. Aden and T. Foust, Technoeconomic analysis of the dilute sulfuric
acid and enzymatic hydrolysis process for the conversion of corn
stover to ethanol, Cellulose, 2009, 16, 535–545.
29 A. R. Aden, et al., Lignocellulosic biomass to ethanol process design
and economics utilizing co-current dillute acid prehydrolysis and
enzymatic hydrolysis for corn stover, NREL/TP-510-32438, (2002).
30 F. K. Kazi, et al., Techno-economic comparison of process technolo-
gies for biochemical ethanol production from corn stover, Fuel, 2010,
89, 520–528.
31 L. d. C. Sousa, S. P. S. Chundawat, V. Balan and B. E. Dale,
‘Cradle-to-grave’ assessment of existing lignocellulose pretreatment
technologies, Curr. Opin. Biotechnol., 2009, 20, 339–347.
32 M. P. Tucker, K. H. Kim, M. M. Newman and Q. A. Nguyen,
Effects of temperature and moisture on dilute-acid steam explosion
pretreatment of corn stover and cellulase enzyme digestibility, Appl.
Biochem. Biotechnol., 2003, 105, 165–178.
21 R. H. Leonard, Levulinic acid as a basic chemical raw material, Ind.
Eng. Chem., 1956, 48, 1330–1341.
22 E. L. Kunkes, et al., Catalytic conversion of biomass to monofunc-
tional hydrocarbons and targeted liquid-fuel classes, Science, 2008,
322, 417–421.
23 M. Mabry, W. Prichard and S. Ziemecki, in E. I. DuPont, de Nemours
and Company US Patent 4,550,185, (1985).
24 M. Mabry, W. Prichard and S. Ziemecki, in E. I. DuPont, de Nemours
and Company US Patent 4,609,636, (1986).
25 J. T. Schwartz, in E. I. DuPont, de Nemours and Company US Patent
5,478,952, (1995).
33 T. Marzialetti, et al., Dillute acid hydrolysis of loblolly pine: a
comprehensive approach, Ind. Eng. Chem. Res., 2008, 47, 7131–7140.
34 R. Smith, Chemical process design and integration, Wiley, Chichester,
West Sussex, England, Hoboken, NJ, 2005.
35 2008 Georgia Timber Report in Georgia Forestry Commision
(http://www.gfc.state.ga.us/ForestManagement/TimberSelling.cfm).
36 D. Mignard and C. Pritchard, Processes for the synthesis of liquid
fuels from CO2 and marine energy, Chem. Eng. Res. Des., 2006, 84,
828–836.
26 V. Pallassana and M. Neurock, Reaction paths in the hydrogenolysis
of acetic acid to ethanol over Pd(111), Re(0001), and PdRe alloys,
J. Catal., 2002, 209, 289–305.
27 A. Aden, Biochemical production of ethanol from corn stover:
2007 state of technology model, NREL/TP-510-43205, 1-10,
(2008).
37 P. J. Ince, X. Li, M. Zhou, J. Buongiorno and M. Reuter, United
States paper, paperboard, and market pulp capacity trends by process
and location, 1970-2000, Forest Products Laboratory, US Dept of
Agriculture, 2001, 1–36.
38 S. M. Walas, Chemical process equipment : selection and design
(Butterworths, Boston), (1988).
This journal is
The Royal Society of Chemistry 2011
Green Chem., 2011, 13, 1755–1765 | 1765
©