Artigo científico

1. Propylene glycol from glycerol: Process evaluation and break-even
price determination
Roberto X. Jimenez a
, Andre F. Young b, *
, Heloisa L.S. Fernandes a
a
Department of Chemical Engineering, Federal University of Rio de Janeiro, Av. Athos da Silveira Ramos, 149, Rio de Janeiro, Brazil
b
Department of Chemical and Petroleum Engineering, Federal Fluminense University, Rua Passo da Patria, 156, Niteroi, Brazil
a r t i c l e i n f o
Article history:
Received 22 November 2019
Received in revised form
27 April 2020
Accepted 23 May 2020
Available online 31 May 2020
Keywords:
Glycerol
Process simulation
Propylene glycol
Hydrogen
Reforming
a b s t r a c t
Propylene glycol can be produced from the hydrogenolysis of glycerol. If hydrogen comes from a
renewable source, the produced propylene glycol can be considered completely renewable. The objective
of this work is to perform a technical and economic evaluation of a propylene glycol production process
which uses glycerol as the main raw material. Two possibilities regarding the hydrogen source are
explored: it may come from an external source or be produced locally. It is shown that the break-even
price of the partially renewable propylene glycol is lower than the market price of propylene glycol,
which indicates that this process is economically viable. It is also shown that the price of hydrogen
produced via steam reforming is not competitive with its price if produced by other sources. Conse-
quently, the fully renewable propylene glycol production process yields a lower profit, but it is still viable
and depends on the market.
© 2020 Elsevier Ltd. All rights reserved.
1. Introduction
Glycerol is the main by-product of biodiesel production [1e3].
Along with the increased supply of biodiesel, glycerol production
has also increased. In 2017, there were produced about 374.5
thousand cubic meters of glycerol, which meant a 9.5% increase in
comparison with 2016 [4].
The increasing glycerol surplus has motivated the development
of chemical processes which employ it as the main raw material. As
a consequence, the number of published scientific articles on the
use of glycerol as the main raw material in chemical processes has
increased. For example, articles related to glycerol in the Scopus
database have increased by almost 7 times per year between 2008
and 2016 [3].
Among the products that can be obtained from glycerol there
are propanediols (1,2- and 1,3-propanediol), propene, epichloro-
hydrin, methanol, ethanol, lactic acid, citric acid, propanoic acid
and hydrogen [5e7]. Monteiro [3] and Freitas [6] identified that 1,3-
propanediol, hydrogen and 1,2-propanediol or propylene glycol
(PG) are the products formed from glycerol on which researchers
have published more articles. According to these authors, the
number of patents or patent applications on PG production from
glycerol was larger than those on 1,3-propanediol or hydrogen
production until 2013.
Propylene glycol can be used as a surfactant, a moisturizer, an
antifreeze agent, a solvent, a preservative and in detergent for-
mulations [6,8]. It is usually obtained by the hydrolysis of propylene
oxide [8,9]. Propylene oxide is produced from propylene hydro-
chloride, which comes from the petrochemical industry.
In Brazil, the only registered producer of propylene glycol is
Dow®. It has a chemical plant located in the industrial complex of
Aratu, in the state of Bahia. The company also produces propylene
oxide, so propylene glycol is still produced by the conventional
route from petrochemicals. It is noteworthy that this is the only
propylene glycol production plant in Latin America [10]. A simpli-
fied flowsheet of this process technology is shown in Fig. 1.
Alternatively, PG can be produced by a more ecologically sound
route, which employs glycerol as a carbon and hydrogen source,
and which relies less on fossil sources. Some scientific articles about
this route were published, but there are only a few patents (granted
or under application) on it. Currently, if USPTO and Lens databases
are searched for patents on processes involving the keywords
propylene glycol, glycerol/glycerin and 1,2-propanediol, glycerol/
glycerin, only 12 results are retrieved. Among them, 10 patents
describe processes with liquid-phase reactors, one that employs a
vapor-phase reactor and one with a two-phase reactor (reactions* Corresponding author.
E-mail address: ayoung@id.uff.br (A.F. Young).
Contents lists available at ScienceDirect
Renewable Energy
journal homepage: www.elsevier.com/locate/renene
https://doi.org/10.1016/j.renene.2020.05.126
0960-1481/© 2020 Elsevier Ltd. All rights reserved.
Renewable Energy 158 (2020) 181e191

2. start in the liquid phase and generate acetol, which is transferred to
the vapor phase to react). It is interesting to note that, although the
search was not limited by date, all the results date to the last 8
years.
These results indicate that this technology is new and may
require more studies until consolidation. Considering this, the
objective of this work is to perform a technical and economic
assessment of the partially renewable and fully renewable PG
production processes from glycerol, to find out if these processes
are economically viable in the current Brazilian economic scenario.
Equipment design and process simulations were performed with
the aid of the Aspen HYSYS® v8.8 process simulator.
2. Process simulation and equipment design
In order to proceed to the economic analysis, the reaction ki-
netics and vapor-liquid equilibria must be correctly reproduced in
the simulations so that the equipment can be properly designed.
The properties of all components involved in the system are
available in the Aspen HYSYS® simulator database.
2.1. Propylene glycol production from glycerol
Glycerol (or glycerin) obtained from biodiesel is available in
different purity levels. Crude glycerin is defined as the glycerol-rich
mixture (around 75% glycerol) obtained after the separation of
biodiesel. After neutralization, glycerin is usually called blonde; its
weight percent composition is 85.0% glycerol, 11.7% water, 3.2%
methanol, and 0.2% biodiesel. After distillation, USP grade glycerol
or bidistilled glycerol (99.5% glycerol) is obtained [11]. The advan-
tage of using blonde glycerin as a raw material is that, as the puri-
fication process has a cost, it ends up being cheaper. However,
bidistilled glycerol is regarded as the process raw material in this
work, since kinetic data for the hydrogenolysis and reforming re-
actions refer to glycerol-water mixtures only [12].
Furthermore, according to Resolution no. 30 published by ANP
(Brazilian National Agency of Petroleum, Natural Gas and Biofuels),
which regulates biodiesel production activities in Brazil [4],
including the construction stage, the glycerol purification stage
must be present in every biodiesel production unit. Thus, in all
plants currently operating in Brazil, a stream of composition close
to that of bidistilled glycerol is available, or at least its composition
can be adjusted for this purpose.
In literature, a wide variety of metal catalysts has been used to
promote the CeO bond cleavage reaction of the glycerol molecule.
Cu-based catalysts are widely studied as they offer a high degree of
conversion and selectivity of PG [13e20].
Two mechanisms have been proposed for the production of
propylene glycol from glycerol: according to the dehydration-
hydrogenation mechanism, glycerol is initially dehydrated at
catalyst acid sites to acetol, which is hydrogenated to 1,2-
propanediol [5,13,21]. A three-phase dehydrogenation-dehydra-
tion-hydrogenation mechanism is also suggested [15,17,22,23].
The literature recommends a Langmuir-Hinshelwood-type
expression to model the heterogeneous kinetics of formation of
propylene glycol and by-products (1,3-propanediol, ethylene glycol,
ethanol, methanol, propanol and propanoic acid) through the
dehydration-hydrogenation mechanism [12,19,24,25]. Two types of
catalyst active sites are considered, as observed in Table 1 [12,26].
Zhou et al. [12] employed a CueZnOeAl2O3 catalyst with a
molar ratio of 1:1:0.5 of Cu/Zn/Al, with a particle diameter of
0.17 mm (80e100 mesh), and reported a conversion of 81.5% of the
glycerol and selectivity of 93.4% in terms of propylene glycol. Also,
the authors proposed the rate expressions of the two reaction steps
presented in Table 1, which can be found in Eqs. (1) and (2).
r1 ¼
k1bGCG
ð1 þ bGCG þ bACA þ bPCPÞ
(1)
r2 ¼
k2bGCG
ð1 þ bGCG þ bACA þ bPCPÞ
1 þ
ffiffiffiffiffiffiffiffiffiffiffiffi
bHPH
p 2
(2)
In Eqs. (1) and (2), Ci is the molar concentration of component i,
PH is the partial pressure of hydrogen, ki is the rate constant of
reaction i and bi is the absorption rate constant of component i [12].
The numerical values of these constants are described in Table 2. A
preliminary simulation of the reaction in laboratory conditions was
carried out in Aspen HYSYS®; simulated and experimental data
[12] were compared and relative errors of 3.2% in the glycerol
conversion and 7.0% in the acetol to propylene glycol selectivity
ratio were found. Since average relative errors between experi-
mental and theoretical glycerol and propylene glycol flowrates at
Fig. 1. Simplified propylene oxide hydration process flowsheet.
Table 1
Steps involved in the reaction of propylene glycol formation over a Cu/ZnO/Al2O3
catalyst. G ¼ glycerol, A ¼ acetol, PG ¼ propylene glycol, Q1 type 1 active site, Q2
type 2 active site [12].
Stage Kinetic Adsorption/Desorption
H2 þ 2Q1 #2HQ1 e bH
G þ Q2 #GQ2 e bG
GQ2/AQ2 þ H2O k1 e
2HQ1 þ AQ2 /PGQ2 þ 2Q1 k2 e
AQ2 #A þ Q2 e bÀ1
A
PGQ2 #P þ Q2 e bÀ1
PG
R.X. Jimenez et al. / Renewable Energy 158 (2020) 181e191182

3. the reactor output are 6.3% and 7.6%, respectively [12], deviations
obtained in the preliminary simulation can be considered accept-
able. Thus, the rate expressions in Eqs. (1) and (2) are used in this
work.
The UNIQUAC model was used for phase equilibria calculations
for the system containing the components involved in the pro-
duction of propylene glycol, since it provides a satisfactory
description of phase equilibria for similar systems [27].
The propylene glycol process flowsheet can be seen in Fig. 2. The
stream properties are presented as Supporting Information,
Table S3. The raw material is fed into the process and is put under
operating conditions before entering the reactor. Excess hydrogen
is used to facilitate mass transfer in the three-phase system within
the reactor. The global reaction is shown in Eq. (3).
C3H8O3 þ H2/C3H8O2 þ H2O
Glycerol PG
(3)
The reaction is carried out at T ¼ 493 K, PH ¼ 4 MPa and a
hydrogen to glycerol molar ratio of 5:1. The catalyst is assumed to
be CueZnOeAl2O3 and the residence time is adjusted to 85
kgCat:h:kg:molÀ1
Glycerol, in order to reproduce the results obtained by
Zhou et al. [12]. Moreover, the authors compared their results at
493 K with others available in the literature [21], and showed
consistency in the values that were reported. Zhou et al. [12] proved
that glycerol conversion increases with higher pressures (testing
3 MPa, 4 MPa and 5 MPa), but at 493 K it remains almost the same,
independent on pressure. A conservative approach was used to
choose a 4 MPa pressure for the reaction simulation. The plant was
designed to operate with a flowrate of 1070 kg/h of glycerol. This
was the average glycerol production in biodiesel plants in Brazil in
2018 [4].
The reactor effluent is mainly composed by propylene glycol,
hydrogen, unreacted glycerol and by-products. The product stream
exchanges heat with the glycerol stream to minimize utility ex-
penses. Then it follows to the separation section, in which excess
hydrogen is recovered with a flash vessel, recycled and mixed with
pure hydrogen, thereby ensuring a 5:1 M ratio of hydrogen to
glycerol at the process inlet. Finally, two distillation columns adjust
the propylene glycol purity to the required value (USP grade,
99.5%).
The literature suggests the use of a trickle-bed-reactor for
glycerol hydrogenolysis in industrial-scale processes [25,28,29].
Inside the reactor, liquid and gas flow downward through the cat-
alytic bed. The net flowrate must be low enough so that the liquid
phase forms a thin film over the fixed bed and descends by gravity
and gas-phase drag [30]. Thus, the contact among the three phases
is maximized. Trickle-bed-reactors have been widely used in other
processes involving hydrogen, mainly in the hydrodesulfurization
of petroleum fractions [30,31].
The catalyst density was not mentioned in the articles reviewed
but is required for the calculation of the fixed-bed volume. The
fixed-bed volume includes the volume occupied by the catalyst
solid mass (30%) and the volume of pore voids and voids among the
catalyst particles (70%), which are generally in a proportion of 3:3:4
[32]. The solid density was calculated from the molar fraction of its
components. Assuming the void fraction in the bed to be equal to
0.4, the densities of solid, pellets and bulk can be estimated. For the
purpose of reactor sizing, the trickle-bed model was simplified to a
plug-flow fixed bed model, as available in Aspen HYSYS®.
The required catalyst mass was obtained by solving the design
equations for a plug-flow reactor. A plug-flow reactor which
operates at conditions similar to those in the work by Zhou et al.
[12] requires 992.8 kg of catalyst. Likewise, a conversion of 98.2%
requires 1489 kg of catalyst, while a conversion of 99.99% requires
an additional 993 kg of catalyst [12]. Consequently, it is not
convenient to try to achieve such high glycerol conversions. Thus,
the plant was designed to operate with an overall glycerol
Table 2
Parameters for the kinetic model of the glycerol hydrogenolysis reaction [12].
Kinetic and absorption constants Parameters
Pre-exponential Factor
mol
kg s
Activation Energy
kJ
mol
k1 1.54 Â 107
86.56
k2 7.16 Â 106
57.8
bG 2.22 36.42
bA 8.73 25.94
bPG 5.80 25.77
bH 1.86 Â 10À5
MPa 36.24
Fig. 2. Propylene glycol production flowsheet.
R.X. Jimenez et al. / Renewable Energy 158 (2020) 181e191 183

4. conversion of 98.2%. Finally, the reactor volume can be calculated
by the ratio of catalyst mass and bulk density.
The reaction must occur under isothermal conditions. To meet
this specification, a cooling jacket was applied, with water at 298 K
as coolant. The heat exchange area was assumed to be approxi-
mately equal to the lateral area of the plug-flow reactor.
In addition to acetol and propylene glycol, the authors identified
the presence of by-products such as ethylene glycol, methanol,
ethanol, propanol and propanoic acid in the product stream, among
which ethylene glycol is the most expressive. Therefore, it is
necessary to simulate the formation of these products so that the
simulation is in agreement with the experimental study.
Ethylene glycol and methanol modeling was performed by
specifying the conversion value found experimentally [12]. Due to
the lack of information on the proportion of the remaining by-
products, their simulation was not feasible. Since the by-products
are lower alcohols (except for propanoic acid), they were regar-
ded as a single representative compound (methanol) in the
simulations.
In a preliminary analysis of the separation system, it was found
that for a better recovery of unreacted hydrogen it is necessary to
cool the product stream from the reactor. To save resources and
accomplish energy integration, it was decided to implement a
shell-and-tube heat exchanger through which this stream and the
glycerol and water feed stream are passed. The heat transfer rate,
the heat exchange area and the global heat transfer coefficient were
obtained with the aid of the simulator, from the values of the inlet
and outlet streams’ temperatures, pressures and flowrates.
A minimum temperature approach of 10 K was used to avoid an
excessively large heat exchange area. The initial estimate of pres-
sure drop on both sides of the exchanger was 10 kPa [33]. The heat
exchanger is designed with floating heads to minimize thermal
expansion effects that may occur for a temperature difference
greater than 30 K between the tube fluid (reactor effluent) and the
shell fluid [34].
The methodology described by Seider et al. [33], which is based
on the relative volatility of the compounds, was adopted for the
synthesis of the separation system. The reactor outlet is a two-
phase stream, which consists of hydrogen, methanol, water, ace-
tol, propylene glycol, ethylene glycol and glycerol (from the most
volatile to the least volatile). Three cuts were required to achieve
the specified purity: Hydrogen/Methanol, Water/Propylene Glycol
and Propylene Glycol/Ethylene Glycol.
Hydrogen is the first component that must be removed, as it is
the excess reagent and can be recycled into the process, thereby
reducing raw materials expenditures. Thus, for the separation of
hydrogen from the reactor effluent stream, a flash vessel was used.
The liquid-gas separator should be such that the vessel diameter
must be large enough so that the gas velocity is less than the liquid
droplet separation velocity and that the vessel height is equal to its
diameter, or 1 m if the diameter is less than 1 m. It is recommended
that hold-up time is about 10 min so that the liquid level inside the
equipment is appropriate. Finally, the feed inlet must be at least
60 cm above the liquid level [35].
To maximize the recovery of pure hydrogen, it is necessary that
the flash vessel also removes heat from the streams, to promote
condensation of the less volatile components. Such heat exchange
was performed by implementing a coil inside the vessel with
cooling water at 298 K. Aspen HYSYS® Adjust block was used so that
the cooling rate varies to ensure that the purity of the hydrogen
obtained is at least 99.5%.
The next two cuts were made through distillation columns. The
first column cuts between water (light key) and propylene glycol
(heavy key), and the second column adjusts the purity of the final
product to the specified value by cutting between propylene glycol
(light key) and ethylene glycol (heavy key). The columns were
simulated rigorously and its input parameters were estimated
based on typical methods.
The minimum number of stages was calculated using the Fenske
equation. The number of theoretical stages was obtained by the
Gilland-Molokanov correlation. The minimum reflux ratio was
estimated from the Underwood equation. Finally, the Kirkbride
equation was used to determine the optimum feed stage [35]. It is
noteworthy that the calculations of these estimates were per-
formed in the simulator itself, using the Shortcut Distillation block.
From the obtained values, it was possible to start the rigorous
design of the distillation columns.
In a first analysis, it was found that the operation under atmo-
spheric pressure demands extreme temperatures in condensers
and reboilers. As a consequence, common utilities such as steam
and cooling water could not be used. Also, degradation or poly-
merization reactions of residual glycerol may occur at elevated
temperatures (about 473 K) [36].
Thus, to avoid such problems, it was decided that the equipment
should operate under vacuum. The operating pressure was set as
close as possible to atmospheric pressure, but the reboilers tem-
peratures were kept at least 15 K below the glycerol degradation
temperature. The literature suggests the use of packed columns for
vacuum distillation and high viscosity mixtures, as in this case [35].
Structured packings offer better separation efficiencies and pres-
sure drops, but are more expensive. Thus, the columns were
designed with random packings of 1 in Pall metal rings, due to the
high number of theoretical stages needed [33].
For the calculation of the packed bed height, the Height Equiv-
alent to a Theoretical Plate (HETP) method was used. For vacuum
distillation, the required column diameter can be estimated from
the flooding condition. This condition should be avoided and can be
estimated from the flooding velocity through the Leva correlation
[33]. The results obtained in the simulation are in agreement with
the work of Gandarias et al. [37].
2.2. Hydrogen production: glycerol steam reforming
Reforming of oxygenated hydrocarbons consists of the cleavage
of the CeC bond to form CO and H2. Then CO and water may react in
a water-gas shift reaction, releasing more H2 and CO2 [38]. Glycerol
reforming needs a catalyst that promotes cleavage of the CeC, OeH
and CeH bonds, and if possible also promote the shift reaction
[38,39]. Reforming can be conducted in both liquid and gaseous
phases. In the latter case, it is necessary to vaporize glycerol, what
requires higher temperatures [40]. The water-gas shift reaction
occurs only in the gas phase.
Glycerol aqueous phase reforming occurs at higher pressures
and lower temperatures, if compared to steam reforming, and of-
fers greater possibilities for using crude glycerin as a raw material
[41]. The major disadvantage of this route is its low reaction rates
[42]. Vapor-phase reforming, in turn, can be conducted at atmo-
spheric pressures and has presented energetic advantages and
higher yield in terms of H2 [43]. That is why this was the route
chosen in this work.
In literature, there are studies about the activity of different
catalysts and operating conditions that maximize the glycerol
reforming and shift reactions, to the detriment of methanation and
coke formation. Such reactions decrease the efficiency of the pro-
cess in terms of H2 generation and impair the catalytic activity by
carbon deposition at the active catalyst sites [44,45].
Catalysts involving noble metals, such as Pt, Ru and Pd, and non-
noble metals, such as Ni and Co, have been successfully tested in
several studies for both aqueous- and vapor-phase reforming
[45e47]. The main disadvantage of nickel-based catalysts is coke
R.X. Jimenez et al. / Renewable Energy 158 (2020) 181e191184

5. formation, whereas for noble metal catalysts the high cost is an
important drawback.
Kinetic modeling of reforming of oxygenated compounds such
as ethanol and methanol has been satisfactorily done through
power-law expressions [48]. The activation energy depends on the
type of catalyst employed in the reactions. For noble metals (Pt and
Ru), the reaction rate expression depends linearly on glycerol
concentration due to the high activity of such catalysts, while for
other metals it is a function of a fractional power of the glycerol
concentration. The non-dependence of the rate on water concen-
tration is because it is used in excess [49]. Table 3 presents glycerol
reforming kinetic parameters for power-law expressions obtained
for different catalysts.
The Soave-Redlich-Kwong (SRK) equation was chosen to
calculate the thermodynamic properties in this simulation. This
choice is supported by glycerol reforming studies available in the
literature [50e54].
The glycerol steam reforming process for hydrogen production
is designed to meet the hydrogen demand required for hydro-
genolysis and is shown in Fig. 3. The stream properties are pre-
sented as Supporting Information, Table S4. Glycerol and water
enter the process at a 1:9 ratio. The reforming process involves
three reaction steps, carried out in three different reactors. Such
reactors operate under different conditions and employ different
catalysts [55]. The first reactor performs reforming itself, as
described in Eq. (4):
C3H8O3#3 CO þ 4 H2 (4)
The water-gas shift reaction occurs in the second and third re-
actors, as described in Eq. (5):
CO þ H2O#CO2 þ H2 (5)
It is recommended to perform this reaction in two steps because
it is exothermic. So, the first reactor, which is called the high-
temperature shift reactor (HTS), operates at 648 K and atmo-
spheric pressure. The second reactor, which is called the low-
temperature shift reactor (LTS), operates at 498 K and at the same
pressure, in order to shift equilibrium in the products direction
[52].
A Fe2O3 catalyst was used for the HTS reactor. This catalyst has a
low thermal deactivation rate and allows to promote the kinetics by
increasing temperaure. For the LTS reactor, a high-activity CuO
catalyst was used, which increases the products formation rate
under lower temperatures [52].
The last reactor effluent consists mainly of H2, CO2, CO traces
and water vapor, which is condensed, separated and recycled. The
surplus stream enters a Pressure Swing Adsorption (PSA) separa-
tion system, which produces hydrogen at a purity of 99.99%.
The reforming reactor and the two shift reactors are of the fixed-
bed PFR type. Thus, for design, it is first necessary to evaluate the
amount of catalyst to be employed and then calculate the reactor
volume. For the first reactor, the raw material must be heated to the
operational temperature of 973 K and the effluent must be cooled
down to 648 K, at which the HTS reactor operates, and then finally
cooled to 498 K for the LTS [55].
Due to the lack of information on all reactions in a same study
and under the same conditions, the conversion values and opera-
tional conditions adopted by Villaça [52] were used. The conversion
values obtained for methane generation and carbon formation are
much lower than that of the main by-products, so it was decided to
neglect these by-products. The carbon formed will accumulate at
the active catalyst sites over time, thereby decreasing the efficiency
of the process. This issue was considered in the economic evalua-
tion. Methane molar fraction decreases with increasing tempera-
ture, and for the temperature adopted in this work (973 K), no
methane is detected in the reactor output stream. It can also be
observed that the amount of H2 produced almost remains constant
with increasing temperature, which is in agreement with experi-
ments [41].
As the temperature increases, the molar fraction of carbon
monoxide increases. This is because the shift reaction (which
transforms CO to CO2) is exothermic and its yield is higher at low
temperatures. To model the two shift reactors (HTS and LTS), the
Equilibrium Reactor block from Aspen HYSYS® was used to calculate
the output streams’ compositions. This can be justified by the fact
that this reaction indeed tends to equilibrium in the reactor. This
stage of the simulation is different from that presented by Villaça
[52], but is in agreement with the considerations exposed earlier.
Aspen HYSYS® contains in its database the equilibrium constant
value of this reaction as a function of temperature.
Excess water is added to the process. Thus, after the LTS reactor,
water must be condensed in a flash vessel and recirculated.
Aspen HYSYS® does not allow the simulation of adsorption
processes with conventional blocks. Thus, the PSA unit is simplified
to a Splitter block, in which a process stream is separated into
several streams with defined compositions. In the present work,
purity and separation data reported in the work of Villaça [52] and a
pressure of 0.7 MPa were considered.
For sizing, two adsorption towers were considered. The columns
operate in cycles, thus ensuring that the process operates contin-
uously. The adsorption columns are filled with a type 5A alumi-
nosilicate molecular sieve, indicated to separate carbon dioxide
from hydrogen-rich streams [30]. The quantity of adsorbent can be
calculated considering an average adsorption capacity value of
0.255 kg of adsorbent per kg of adsorbate. With the bulk density,
the bed volume is obtained [30,56].
2.3. Economic analysis
The total investment cost was estimated by means of the Lang
method. The FOB price of each equipment was estimated from
correlations available in the literature [33]. These correlations were
calculated for a specific date, so updating is required. It can be done
Table 3
Kinetic parameters of power-law models for glycerol steam reforming [49].
Catalyst Temperature (K) Glycerol Order Steam Order Activation Energy
kJ
mol
Ru= Al2O3 623e773 1 e 21.2
Pt= C 623e673 1 e e
Co= Al2O3 723e823 0.1 0.4 67.2
Ni= Al2O3 723e823 0.48 0.34 60.0
Co À Ni= Al2O3 773e823 0.25 0.36 63.3
Ni= CeO2 873e923 0.233 e 103.4
Ni À ZrO2=
CeO2
973 0.3 e 43.4
R.X. Jimenez et al. / Renewable Energy 158 (2020) 181e191 185

6. with the Plant Cost Index, published periodically by Chemical En-
gineering Magazine. The value of the Plant Cost Index for the year
on which the correlations were based is 394. For 2018 it is 603.1
[57].
For non-corrosive environments with hydrogen, Seider et al.
[33] recommend the use of a 1% Cr and 0.5% Mo steel alloy (ASME
code SA-387B). This material was applied only in large equipment
used in operations involving the presence of hydrogen.
For the calculation of production costs, it was considered that
the plant operates continuously 24 h a day for 330 days a year.
Prices for raw materials, utilities and catalysts are summarized in
Table 4.
It was not possible to stablish an average price for hydrogen,
because several production processes employ a variety of raw
materials and energy sources. For instance, for hydrogen produced
via water electrolysis the cost may vary from 1.28 to 4.14 US$/kg,
these values being the most recent that could be found in the
literature [58]. Negro et al. [59] presented the different costs of
hydrogen in Brazil if it was produced via electrolysis, fossil fuel
reforming, or renewable fuel reforming. However, these values are
out of date and may affect the veracity of the economic analysis. For
this reason, these costs were updated using the Brazilian National
Consumer Price Index [60], that represents the accumulated infla-
tion in the country over time. The results can be observed in Table 5.
Thus, in the present work, the hydrogen price range is 0.23e4.14
US$/kg, considering the lowest price of hydrogen if it is obtained by
natural gas reforming and the highest price if it is obtained by water
electrolysis. The price of hydrogen produced by electrolysis using
solar energy was disregarded, since its value is very different from
others and the price range obtained would be too wide [59].
Rajkhowa et al. [25] quantified the dominant factors in the loss
of catalytic activity for a Cu-based catalyst during the glycerol
hydrogenolysis reaction, concluding that for this copper catalyst,
sintering occurs due to the agglomeration of Cu atoms, decreasing
the effective area of the catalytic surface. The authors evaluated the
stability of the catalyst for periods of 80e90 h under maximum
operability conditions and different raw material compositions. For
a feed of pure glycerol, the catalyst showed high stability
throughout the experiments. However, for a feed stream containing
pollutants such as sulfur, chlorine, and glycerides from biodiesel
production, its activity decreases. This is one more reason to
consider bidistilled glycerol as the raw material. Thus, the loss of
catalytic activity was neglected.
In the hydrogen production process, Levalley et al. [61] indicate
that copper and iron catalysts for the shift reaction keep a good
performance from 2 to 4 years. An average time of 3 years was
assumed for the economic analysis. Berman and Epstein [62]
proved that RueAl2O3 catalysts suffer significant thermal decom-
position above 1373 K. As the reforming reactor operates at 973 K, it
was considered that thermal deactivation is not critical, and the
catalyst may be replaced together with the other catalysts.
The annual consumption of raw materials and utilities was ob-
tained directly from Aspen HYSYS®. For the calculation of
manpower, the Brazilian average wage for industrial operators was
considered [63]. The number of workers per shift was estimated
from the number of equipment units, considering five shifts as
recommended by Seider et al. [33]. The remaining production costs
were estimated from correlations available in the literature.
3. Results and discussion
A brief description of the process equipment size and cost can be
found in Table 6, while the relative costs of equipment pieces are
shown in Fig. 4. To produce 821 kg/h of propylene glycol, 22 kg/h of
hydrogen, 1070 kg/h of glycerol, and 267 kg/h of water are needed.
Producing 23 kg/h of hydrogen requires 174 kg/h of glycerol and
127 kg/h of water. It is clear that the distillation columns are the
most expensive equipment in the propylene glycol process. This is
due to the high purity specification of PG, which leads to larger
distillation columns. Due to the high steam demand, the steam
generation unit costs represent more than half of the equipment
costs for the steam reforming process, as can be seen in Fig. 5.
Figs. 6 and 7 show the costs distribution of the utilities
consumed in each process. Cooling water and steam are the main
utilities used in the PG process because of the distillation columns’
Fig. 3. Glycerol steam reforming flowsheet.
Table 4
Prices of raw materials, utilities and products involved in the glycerol hydro-
genolysis and glycerol steam reforming processes.
Raw Material Unit Price (US$/kg) Reference
Glycerol 0.4918 [64]
Hydrogen 1.28e4.14 [58]
Process Water 0.0008 [65]
Utilities
450 psi Vapor 0.0130 [66] a
Cooling Water 0.0008 [65]
Fuel (US$/m3
) 0.9768 [67]
Electricity (US$/kWh) 0.1032 [68]
Effluent Treatment 0.0560 [69]
Catalysts
Cu= ZnO= Al2O3 1.8000 [70]
Ru= Al2O3 21000 [70]
Fe2O3 13.900 [70]
CuO 6.7000 [70]
Products
Propylene Glycol 1.4500 [10]
a
Updated to 2018 with the Plant Cost Index.
R.X. Jimenez et al. / Renewable Energy 158 (2020) 181e191186

7. reboilers and condensers. Moreover, there is also a high electricity
demand due to the compressor.
Considering the production of partially renewable propylene
glycol (with hydrogen purchased from an external source) with an
average price of 2.185 US$/kg for H2, the break-even price for pro-
pylene glycol is 1.17 US$/kg, with a total investment of 4,722,678.07
US$. The break-even price of hydrogen production via glycerol
steam reforming is 9.01 US$/kg, with a total investment of
1,441,653.78 US$. The detailed cost statement of the processes is
presented as Supporting Information, Table S1.
It was found that the break-even price of propylene glycol is 1.36
US$/kg for a fully renewable PG production process (that is, a
propylene glycol production process integrated with a glycerol
reforming process to produce hydrogen). Therefore, the generation
of a 100% green PG resulted in a 16% increase in its final price.
Table 5
Prices of hydrogen from different sources in 2003 and 2018 [59].
Process Energy/Raw Material Source 2003 Price (US$/kg) 2018 Price (US$/kg)
Water
Electrolysis
Nuclear Energy 1.98 3.62
Hydraulic Energy 0.79 1.45
Natural Gas 1.58 2.90
Solar 20.76 38.10
Fossil Fuels Reforming Natural Gas 0.12 0.23
Gasoline 1.35 2.48
Methanol 1.19 2.19
Renewable Fuels Reforming Biogas 1.71 3.15
Ethanol 1.50 2.74
Table 6
Propylene glycol and glycerol reforming processes equipment sizing and cost estimates.
Equipment Equipment Size Size Variable Cost (US$)
Propylene Glycol Process
Heat Exchanger 58.81 Total Heat Transfer Area (m2
) 59,733.88
Reactor 1.20 Volume (m3
) 26,967.90
Distillation Colum 1 18 Theoretical Stages 16,450.77
Distillation Colum 2 60 Theoretical Stages 137,091.41
Flash Vessel 1.60 Height (m) 405,172.96
Glycerol Steam Reforming Process
Vapor Generation Unit e e 91,355.20
Heat Exchanger 1 2.86 Total Heat Transfer Area (m2
) 4403.80
Heat Exchanger 1 2.86 Total Heat Transfer Area (m2
) 4140.31
Reforming Reactor 0.0119 Volume (m3
) 7908.29
HTS Reactor 0.0055 Volume (m3
) 10,604.89
LTS Reactor 0.0052 Volume (m3
) 9592.25
Flash Vessel 1.60 Volume (m3
) 7908.29
PSA 0.1435 Volume (m3
) 13,404.60
Fig. 4. Propylene glycol process equipment cost distribution.
Fig. 5. Steam reforming process equipment cost distribution.
R.X. Jimenez et al. / Renewable Energy 158 (2020) 181e191 187

8. However, this additional expense can be offset when the expanding
market shares of more sustainable products are taken into account.
The costs distributions of the evaluated processes are shown in
Table 7. It can be seen that raw material and catalysts correspond to
55e60% of the total production cost. The renewable PG production
cost is higher in nearly US$ 1.2 million, due to the price of renew-
able hydrogen.
The main source of uncertainty in the economic analysis for the
propylene glycol production process considering that the hydrogen
is purchased from an external source is its purchase price, which
depends on the raw material used to produce hydrogen, the
quantity purchased and the distance between the PG plant and
hydrogen plant.
For this reason, a sensitivity analysis was performed to assess
the impact of the price of H2 on the PG production cost. It was found
that a 1% increase in the price of hydrogen implies a 0.045% increase
in the PG production cost. It follows that the price of propylene
glycol is not greatly influenced by small variations in the price of
hydrogen. Considering the limits of the hydrogen price range [59], a
price range for propylene glycol of 1.15e1.23 US$/kg is obtained.
Thus, it can be said that even for large variations in the price of
hydrogen, the propylene glycol production cost is not significantly
affected.
If a similar sensitivity analysis concerning the glycerol price is
performed, it is concluded that for a 1% increase in the price of
glycerol the glycerol reforming process cost increases by 0.41% and
the propylene glycol production process cost increases by 0.57%.
The selling price of PG and the quantity sold is subject to
changes due to market offer and demand. Fig. 8 shows the variation
of annual net profit with the variation of the PG price. This analysis
was performed for a constant propylene glycol production of
824 kg/h, which was the amount obtained by processing all the
glycerol generated by an average biodiesel plant in Brazil.
A similar analysis is shown in Fig. 9, but in this case the amount
of PG produced is varied and its price is kept constant. For the
renewable propylene glycol, the minimum quantity to be produced
to achieve a zero annual net profit is approximately 776 kg/h. For
the partially renewable propylene glycol with hydrogen obtained
Fig. 6. Propylene glycol process utilities cost distribution.
Fig. 7. Steam reforming process utilities cost distribution.
Table 7
Production costs distribution for the partially renewable PG and renewable PG
processes.
Production Costs Distribution Partially Renewable PG Renewable PG
Raw Materials and Catalyst 59.7% 55.7%
Direct Costs 7.0% 10.7%
Indirect Costs 4.7% 6.8%
Utilities 18.7% 16.9%
Devaluation 5.0% 5.6%
Other Expenses 4.9% 4.3%
Fig. 8. Variation in the net annual profit as a function of PG selling price.
R.X. Jimenez et al. / Renewable Energy 158 (2020) 181e191188

9. from an external source (and a lower production cost), the mini-
mum quantity is 670 kg/h. So, it takes 106 kg/h more PG to make
green propylene glycol feasible, which represents a relatively
broader market share.
A comparison between the conventional process, which uses
propylene oxide as the main raw material, and the renewable and
partially renewable processes was carried out. Raw material and
utility consumption for the conventional process were estimated by
simplified mass and energy balances in Aspen HYSYS®, as shown in
Fig. 1. The stream properties are presented as Supporting Infor-
mation, Table S5. In this technology, the reaction consists of pro-
pylene oxide hydration in the presence of an acid catalyst (sulfuric
acid), and methanol is used as a solvent. The reaction is normally
carried out in a continuous stirred tank [71].
There are some clear disadvantages in the conventional process,
such as petroleum-based raw materials, toxic solvents and corro-
sion problems. Propylene oxide is commonly obtained from pro-
pylene (petroleum based), while methanol is produced from
natural gas. Corrosion problems must be taken in consideration,
because of the use of H2SO4, which is dissolved in water in 0,1% wt.
To produce the same amount of PG, 4000 kg/h of water, 40 kg/h of
H2SO4, 643 kg/h of methanol and 675 kg/h of propylene oxide are
needed. In comparison, the partially green process uses only 267 kg/
h of water and the fully green process consumes a total of 394 kg/h
of process water. Both processes require none of the other com-
ponents. Aspen HYSYS® also gives an estimation of the carbon
emissions. Its 2237 kg/h for the conventional process, against
467 kg/h for the propylene glycol production stage and 141 kg/h for
steam reforming.
Fig. 10 (A) shows that the conventional process’s raw material
cost is greater in nearly 25% than the partially renewable PG process
and 16% than the renewable PG process. This difference is caused by
the high demand of solvent, which is used in a propylene oxide-
methanol equivolume mixture. In Fig. 10 (B) the utilities demand
of each process can be compared. The propylene oxide process
utilities consumption is almost 3.5 times larger than in the
renewable PG process. This high consumption is due to the
requirement to keep the reactor temperature under 325 K and the
separation steps needed to achieve the same purity that has been
reached in the other processes. The reactor temperature is a critical
variable, because of propylene oxide’s low boiling point. Addi-
tionally, it is important to notice that the renewable PG process
consumes more utilities than the partially renewable PG process,
due to steam reforming.
Finally, these results suggest that both green and partially green
PG are economically feasible, because both break-even prices of
propylene glycol were lower than the reported market price, even
though its price depends on the existing demand. The environ-
mental advantages of these products seem clear, but, for future
work, it would be interesting to perform a formal environmental
analysis of the green propylene glycol production process, such as a
Life Cycle Assessment (LCA), and to compare it to the conventional
PG production process to have more elucidative results in this
sense.
4. Conclusion
The present work simulated the propylene glycol production
Fig. 9. Variation in the net annual profit as a function of PG production.
Fig. 10. Comparison of the PG production processes. (A) Raw material costs. (B) Util-
ities demands.
R.X. Jimenez et al. / Renewable Energy 158 (2020) 181e191 189

10. process using glycerol from biodiesel production as the main raw
material, in a Brazilian scenario. A total production cost of 1.17 US$/
kg was obtained when PG is produced using hydrogen from an
external source. When it is produced using hydrogen obtained from
glycerol steam reforming, its price increases to 1.36 US$/kg. A
hydrogen production cost of 9.01 US$/kg was calculated when it is
obtained from glycerol steam reforming, which is much higher
when compared to the market prices of hydrogen obtained by other
processes. It has caused a nearly 70% drop in the annual net profit of
propylene glycol production in the base scenario, although the
process remains economically viable. Fully renewable PG is
approximately 16% more expensive than partially renewable PG,
but it is possible, depending on the market.
Declaration of competing interest
The authors declare that they have no known competing
financial interests or personal relationships that could have
appeared to influence the work reported in this paper.
CRediT authorship contribution statement
Roberto X. Jimenez: Conceptualization, Investigation, Software,
Writing - original draft, Writing - review editing. Andre F. Young:
Conceptualization, Methodology, Supervision, Visualization,
Writing - original draft, Writing - review editing. Heloisa L.S.
Fernandes: Conceptualization, Supervision, Writing - original draft,
Writing - review editing.
Appendix A. Supplementary data
Supplementary data to this article can be found online at
https://doi.org/10.1016/j.renene.2020.05.126.
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