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Carbon dioxide and ethanol from sugarcane biorefinery as renewable feedstocks to environment-oriented integrated chemical plants
Camila Fernandes Ribeiro Machadoa, Ofelia de Queiroz Fernandes Araujoa, José Luiz
de Medeirosa, Rita Maria de Brito Alvesb,1
aFederal University of Rio de Janeiro, Rio de Janeiro, BrazilbUniversidade de São Paulo, São Paulo, Brazil
emails: [email protected] (C.F.R. Machado), [email protected] (O.Q.F. Araújo), [email protected] (J.L. de Medeiros), [email protected] (R.M.B. Alves)
Supplementary Material
APPENDIX A. Thermodynamic Basis for Process Simulations
Tables A.1 and A.2, Figures A.1 and A.2, References
APPENDIX B. Kinetic Basis for Process Simulations
Tables B.1, B.2 and B.3, and References
APPENDIX C. Process Flowsheets as Simulated in ASPEN-HYSYS
Figures C.1, C.2, C.3, C.4, C.5, C.6, C.7 and C.8
APPENDIX D. Mass Balance for the Eco-Pole Units as Simulated in ASPEN
HYSYS
Figures D.1, D.2, D.3, D.4, D.5, D.6 and D.7
1 Corresponding author: email: [email protected], Tel. (+55-11) 3091-2265.
Appendix A. Thermodynamic Basis for Process Simulations
The selected models used in the simulations are presented in Table A.1. For units U1,
U2 and U3, UNIQUAC activity coefficient is used for the liquid phase and Redlich–
Kwong Equation of State (RK) for the vapor phase. In the absence of binary interaction
parameters in literature, these are calculated using the Coeff Estimation UNIFAC VLE
tool of HYSYS Process Simulator. The depropanizer column from Unit 1 was simulated
with Peng-Robinson Equation of States (PR), because it comprises only propane and
propylene. For the simulation of units U4, U6 and U7 (Fig. 1), Wilson’s activity
coefficient model is employed due to satisfactory results from comparison between
theoretical and experimental data available in literature (Mello, 2010). PR was used to
model the high-pressure gas phase of the reactors.
Table A.1. Thermodynamic models
Process Unit Thermodynamic model
U1UNIQUAC + Redlich–Kwong Equation of State
Peng-Robinson (depropanizer column)
U2 UNIQUAC + Redlich–Kwong Equation of State
U3 UNIQUAC + Redlich–Kwong Equation of State
U4 Wilson model + Peng-Robinson Equation of State
U5 UNIQUAC + Peng-Robinson Equation of State
U6Wilson model + Peng-Robinson Equation of State
Peng-Robinson Equation of State (ethylene column)
U7 Wilson model + Peng-Robinson Equation of State
Fig. A.1 presents the equilibrium curves for the experimental data found in literature,
where PB is the bubble point and DP is the dew point. From the curves, it can be
ascertained that the thermodynamic packages selected for simulation properly
represents these binaries.
In the production of ethylene carbonate (EC), ethylene glycol (EG) and dimethyl
carbonate (DMC), Unit 5, the formation of the DMC-Methanol azeotrope takes place.
Hsu et al., 2010 presented the binary interaction parameters between four components:
methanol (MeOH), DMC, EG, EC. The liquid phase is represented by the UNIQUAC
activity model and the RK model for the vapor phase. The experimental data for the
MeOH-EG binary are available from Gmehling and Onken (1977).
Figure A.1. Vapor-liquid equilibrium of: a) methanol and propylene; (b) propylene and PO; (c) ethylene oxide and water; and (d) ethanol and water binaries. Graph built with ASPEN HYSYS – (a) and (b) UNIQUAC - Ohgaki et al., 1988, (c) Wilson - Coles and Popper, 1950 and (d) Wilson - David and Dodge, 1959. X stands for the mole fraction of a component in the mixture. BP = Boiling Point, DP = Dew Point, Exp-X=experimental value of x (liquid phase molar fraction), Exp-Y= experimental value of y (vapor phase molar fraction).
For the MeOH-EC, DMC-EC, DMC-EG and EG-EC binaries, the experimental data
are from Fang and Qian (2005). The MeOH-DMC interaction parameters are from
Rodriguez et al. (2002). Camy et al. (2003) present phase equilibrium calculations for
the quaternary methanol/CO2/DMC/water mixture, using Soave-Redlich-Kwong
Equation of State (SRK) with MHV2 mixture rules. These data are used to adjust the
binary parameters Aij and Aji of the UNIQUAC model, which are presented in Table
A.2. Fig. A.2 presents equilibrium curves at suitable temperatures for representing the
operational conditions of the simulated process in ASPEN HYSYS (ASPENTECH Inc)
with UNIQUAC-RK models and the experimental data presented by Hsu et al. (2010)
and Camy et al. (2003).
Table A.2. UNIQUAC model parameters for the EC/MeOH/DMC/EG binaries.
i=DMC
j=EC
i=MeOH
j=EC
i=MeOH
j=EG
i=CO2
j=MeOH
i=DMC
j=MeOH
Aij
Aji
2.5273
-6.7598
-0.54094
15.892
-32.587
2.2712
131.7089
502.1199
636.888
36.688 Source: Hsu, et al.(2010) and Camy et al. (2003)
Figure A.2. Liquid-vapor equilibrium: (a) binary Methanol and water, (b) Methanol and DMC, (c) Methanol and EG in UNIQUAC-RK and Hsu et al. (2010), (d) Methanol and CO2 (e) DMC and EC in UNIQUAC-RK and Schwinghammer et al. (2006).
References
Camy, S., Pic, J.S, Badens, E., Condoret, J.S., 2003. Fluid phase equilibria of the reacting
mixture in the dimethyl carbonate synthesis from supercritical CO2. J. Supercrit. Fluids, 25
(1), 19-32.
Coles, IC. F., Popper, F., 1950. Ethylene Oxide-Acetaldehyde and Ethylene Oxide-Water
Systems. Ind. Eng. Chem. 42 (7), 14344-1438.
David, F., Dodge, B.F., 1959. Vapor-Liquid Equilibrium at High Pressures. The Systems
Ethanol-Water and 2-Propanol-Water. J. Chem. Eng. Data, 4 (2), 107-121.
Fang, Y.J., Qian, J.M., 2005. Isobaric vapor-liquid equlibria of binary mixtures containing
the carbonate group -OCOO-. J. Chem. Eng. Data 50, 340–343.
Gmehling, J., Onken, U., 1977. Vapor-Liquid Equilibrium Data Collection, Bd. I, Bd. IIa.
DECHEMA Chemistry Data Series, Frankfurt, 1977-1978.
Hsu, K., Hsiao, Y., Chien, I., 2010. Design and Control of Dimethyl Carbonate-Methanol
Separation via Extractive Distillation in the Dimethyl Carbonate Reactive-Distillation
Process. Ind. Eng. Chem. Res. 49 (2), 735–749.
Mello, F.H., 2010. Simulação de Sistema de Absorção, Stripping e Reabsorção de Óxido
de Etileno. MSc Thesis, UNICAMP, Dept. Chem. Eng., Brazil.
Ohgaki, K., Takata, H., Washida, T., Katayama, T., 1988. Phase equilibria for four binary
systems containing propylene. Fluid Phase Equilib. 43(1), 105-113.
Rodriguez, A., Canosa, J., Dominguez, A., Tojo, J., 2002. Vapour–liquid equilibria of
dimethyl carbonate with linear alcohols and estimation of interaction parameters for the
UNIFAC and ASOG method. Fluid Phase Equili. 201 (1), 187-201.
Schwinghammer, S., Siebenhofer, M., Marr, R., 2006. Determination and modelling of the
high-pressure vapour–liquid equilibrium carbon dioxide–methyl acetate. J. Supercrit.
Fluids 38 (1), 1-6.
APPENDIX B. Kinetic Basis for Process Simulations
The main features related to reaction simulations are described for each unit. For those
modeled using kinetic rates, the respective kinetic expressions were modified to meet
the ASPEN HYSYS requirements. The equation for heterogeneous catalysis as given in
ASPEN HYSYS is shown below
rate=Numerator /Deno min ator(B.1)
Numerator=k∗f (Basis )−k '* f ' (Basis )(B.2)
Deno min ator=[1+K 1∗f 1 ( Basis )+1+K 2∗f 2 (Basis )+. .. ]n(B.3)
k=A *exp (−E/ RT )∗T β
(B.4)
k '=A '*exp (−E ' / RT )∗T β '
(B.5)
K 1=A 1 *exp (−E 1 /RT )(B.6)
K 2=A 2 *exp (−E 2/RT )(B.7)
The functions of the Basis (f, f’, f1, f2, ...) are the product of “concentrations” (in the
Basis units) to the power of the specified exponents.
The indexes 1, 2, ... in the constants K, A and E indicate the row number in the matrix of
denominator terms.
n is the denominator exponent
Propylene Unit
In this work, Propylene (C3H6) is produced by the catalytic dehydrogenation of
propane (C3H8) in the presence of CO2. In this reaction, CO2 acts as an oxidant to
combine the dehydrogenation of C3H8 with a reverse water gas shift, and hence, the
equilibrium of the dehydrogenation of C3H8 can be shifted to the product side (Du et al.,
2015). The presence of CO2 increases the propylene yield and reduces catalyst
deactivation (Du et al., 2015; Takahara et al., 1998).
Simulation of the propylene unit is based on Wu et al. (2013a), which employs a
Cr2O3-ZrO2 catalyzed reaction in a microreactor. The reactor is simulated as a
conversion reactor model for the reported reactions (Eqs. B.8-B.11).
C3H8 + CO2 → C3H6 + CO + H2O (B.8)
C2H6 → C2H4 + H2 (B.9)
C2H6 + H2 → 2CH4 (B.10)
C3H8 + CO2 → C2H6 + 2CO + H2 (B.11)
It worth to notice that Wu et al (2013) employ a microreactor with asmall gas flow
rate and low reactant concentrations, which could be an issue concerning scale-up
However, since it is assumed that the reactor is limited by kinetics and not by mass
transfer, and hence the same conversion could be achieved upon scaling up, the
propylene unit may be simulated as a conversion reactor using the data provided in the
reference. . Other authors also report similar conversion, as Cheng et al. (2015), Dinse
et al. (2009), Croppi et al. (2014), Liu et al. (2011), Nowicka et al. (2014), Tóth et al.
(2016), Du et al. (2015), Michorczyk et al. (2012).
The dehydrogenation of propane to propylene using CO2 occurs at 550 ºC, with
propane conversion of 53.3%. Under these conditions, selectivity is 79% for propylene,
20.5% for CH4, and 0.5% for C2H4. A molar ratio CO2/propylene of 3 is assumed.
Propylene Oxide Unit
In this work, propylene oxide (PO) is produced by epoxidation of propylene with
H2O2 (Eq. B.12), which occurs with excess of methanol and in the presence of
TS-1/SiO2 catalyst in a tubular reactor. The reaction rate is given in Eq. B 13, where α =
0.32, β = 0.68, A = 1.3x103 and Ea = 46.8 kJ/mol (Wu et al., 2013b). The reaction rate
parameters were obtained based on experimental data from the reaction between 35 and
50 ºC.
C3H8 + H2O2 → C2H4O + H2O (B.12)
r=A .exp (−Ea
R .T ) .CH 2 O2
α .CC3 H6
β (B.13)
The increase in concentration of H2O2 reduces the selectivity towards propylene
oxide, while increasing methanol concentration improves the selectivity. Pressure and
temperature increases promote enhanced PO selectivity (Wu et al., 2013b). The reactor
is simulated at 40 ºC and 4 bar.
Propylene Carbonate Unit
Propylene carbonate (PC), in this work, is synthesized by the reaction between
propylene oxide (PO) and CO2 (Eq. B.14) (Li et al, 2012). Thus, the unit consumes CO2
directly and indirectly through the use of PO from the Propylene Oxide Unit. Kinetics
follows Arrhenius expression with A=8.47 x 10-2 and Ea = 11.035 kJ/mol at the
operating temperature (Eq. B.14) (Li et al, 2012).
C3H6O + CO2 → C4H6O3 (B.14)
r=A .exp (−Ea
R . T ) .CC3 H 6 O (B.15)
Methanol Unit
The synthesis reaction of methanol by hydrogenation of CO2 is divided into three
reversible and independent reactions as shown in Eqs. B.16, B.17 and B.18. The global
reaction is favored by low temperatures and high pressures, which increase methanol
selectivity and yield. Despite the fact that CO2 conversion increases with the increase in
temperature, yield and selectivity are reduced (Graaf et al., 1988). The reaction takes
place at high pressure with low conversion, approximately 20%, recirculating the non-
reacted gases to increase the process yield (Mota et al., 2014).
CO + 2H2 → CH3OH (B.16)
CO2 + H2 → CO + H2O (B.17)
CO2 +3H2 → CH3OH + H2O (B.18)
The kinetics expressions (B.19, B.20 and B.21) proposed by Graaf et al. (1988) is
adopted, however they have to be rearranged to comply with the reaction format of the
simulation environment (HYSYS) and complemented with data from the kinetic
parameters from Graaf et al. (1988).
r 'CH 3 OH , A 2=K A2 KCO K H 2[ f CO f H 2−
f CH 3 OH
f H 2 K °P1]
(1+KCO f CO+ KCO2 f CO 2)(1+K H 2
12 f H 2
12 +K H2O f H 2 O)
(B.19)
r ' H 2O ,B 2=KB 2 KCO 2 K H 2( f CO 2 f H 2− f H20 f CO/ K °P2)
(1+KCO f CO+KCO 2 f CO 2)(1+K H 2
12 f H 2
12 + KH 2O f H 2O)
(B.20)
r 'CH 3 OH ,C 2=KC 2 KCO2 KH 2[ f CO2 f H 2−
f CH 3 OH f H2 O
f H 22 K °P 3
]
(1+KCO f CO+KCO 2 f CO 2)(1+K H 2
12 f H 2
12 +K H 2 O f H 2O)
(B.21)
Table B.1 presents the kinetic parameters used in the simulation.
Table B.1. Kinetic parameters for reaction rates Eq. B.19 Eq. B.20 Eq. B.21
A 2.33E+01 8.07E+01 4.82E-05
E 51800 5.50E+05 -2.20E+03
A'
1.02E+14 1.89E-01 1.98E+06
E'
1.50E+05 5.90E+05 5.64E+04
DMC, EC and EG Unit
The syntheses of DMC, EC and EG analyzed in this work are described by Eqs.
B.22 and B.23 (Han et al., 2001; Wang et al., 2011). The reaction in Eq. B.22 involves
the formation of EC from CO2 and EO. In Eq. B.23, the intermediary EC is converted
into DMC and EG with excess of methanol.
C2H4O + CO2 →C3H4O3 (B.22)
C3H4O3+2CH3OH→C3H6O3+C2H6O2 (B.23)
The reactions present superior yield and selectivity in homogeneous catalysis
(Peppel, 1958; Sakakura et al., 2007; Wang et al., 2011), but it would require catalyst
separation. The reaction is discussed elsewhere (Souza et al., 2013a, 2013b). According
to Peppel (1958), reaction of Eq. B.22 is considered fast, irreversible, with conversion
and selectivity close to 100%. Therefore, Souza (2013a, 2013b) proposed hypothetical
rate parameters considering activation energy equal to the activation energy of reaction
in Eq. B.23, with pre-exponential constant 30 times greater than the one in Eq. B.23,
herein adopted. For simulation, the reactions follow kinetic expressions proposed by
Hsu et al. (2010) and Fang and Xiao (2004).
rec=k+¿ CEO CCO2 −K−¿CEC¿ ¿ (B.24)
rec=k+¿ CEC CMeOH−K−¿
CEG CDMC
CMeOH
¿¿ (B.25)
Ethylene Unit
In the context of the proposed Eco-Pole, operating in synergy with the bioethanol
industry, the production of green ethylene is highly attractive. The bioethanol from
biomass is considered a potential precursor of the green ethylene, increasing its
technical and economic potential. It is worth noting that, before the advent of the steam
cracking technology, ethylene was produced by ethanol dehydration.
Kagyrmanova et al. (2011) present a catalytic dehydrogenation study for ethylene
through pilot scale experiments and simulation results to estimate kinetic model
parameters. Dehydrogenation occurs in the presence of an aluminum based catalyst in a
multi-tubular reactor at normal pressure (1atm) and temperature between 370 – 400 ºC.
The reaction network for formation of ethylene comprises consecutive parallel reactions
with diethyl ether (DEE, (C2H5)2) as an intermediary. The experiments revealed that the
main products of ethanol (C2H5OH) dehydrogenation were ethylene, DEE, acetaldehyde
and butenes, according to the reactions in Equations B.26 to B.30.
C2H5OH →C2H4 + H2O (B.26)
2C2H5OH → (C2H5)2 + H2O (B.27)
(C2H5)2 →2C2H4 + H2O (B.28)
C2H5OH →C2H4O + H2 (B.29)
2C2H4 →C4H8 (B.30)
Highest selectivity of ethylene is achieved at 450oC, when ethanol concentration is
greater than 94% (mass base) (Kagyrmanova ey al., 2011). It is worth to notice two
important points: (1) with the conversion of ethanol above 98% and temperature lower
than 440 ºC, the lowest rate of ethanol/ethylene consumed is obtained, 1.7 kg/kg; (2) the
catalyst presents satisfactory performance and stability under these conditions.
Simulating the process at temperatures higher than 460 ºC, the reaction achieves 98% of
ethanol conversion. With temperatures above 420 ºC, ethylene yield diminishes, due to
the increase in the formation of butenes and other sub-products. At 420 ºC, a greater
yield of ethylene (97%) is observed, corresponding to 98.7% conversion of ethanol and
98.2% ethylene selectivity.
In this study, the reaction is simulated in a conversion reactor model, based on the
conversion of ethanol and selectivity of products, from which the conversion of each
reaction is estimated (Table B.2).
Table B.2. Reaction conversion for ethylene synthesis.
Ethylene Oxide Unit
Ethylene oxide production is achieved by direct catalytic oxidation of ethylene with
oxygen (O2) as given in Eqs. B.31 and B.32. The reactions were simulated in a tubular
reactor.
C2H4 + 1/2O2→ C2H4O (B.31)
C2H4 + 3O2 → 2CO2 + 2H2O (B.32)
The simulated tubular reactor operates at 220 oC and 20 bar. Under these conditions,
conversion of ethylene is around 10% to 12%, with ethylene oxide selectivity between
78% and 82%. The formation of CO2 occurs in parallel with the formation of ethylene
oxide. The reaction rates are described by Bingchen et al. (1999).
Table B.3 shows the adjusted kinetic parameters to the equivalent model available in
ASPEN HYSYS
Table B.3. Constants of the reaction rate used in ASPEN HYSYS.
Reaction Base Component Conversion
(Eq. B.26) Ethanol 74.2 %
(Eq. B.27) Ethanol 4.63 %
(Eq. B.28) Diethyl ether 20 %
(Eq. B.29) Ethanol 0.25 %
(Eq. B.30) Ethylene 11.1 %
A E (J/mol-1)
1st term 5.913x10-2 -18322
2nd term 0.3998 36828
References
Bingchen, Z., Yuanshun, W., Hongshi, W., 1999. Modelling Catalytic Reaction-Absorption Coupling Process of Ethylene Oxide Synthesis in Slurry Reactor, Chem. Eng. Sci. 54 (10), 1531-1534.
Cheng,Y., Zhang, F., Zhang, Y., Miao, C., Hua, W., Yue, Y., Gao, Z. Oxidative dehydrogenation of ethane with CO2 over Cr supported on submicron ZSM-5 zeolite. Chinese Journal of Catalysis, 36(8), 1242-1248, 2015.
Croppi, C., Iaquaniello, G., Palo, E., Salladini, A. Novel process scheme for selective propane. Dehydrogenation. CAtalytic REactors based on New mAterials for C1-C4 valorization (CARENA). Funding scheme: FP7‐NMP‐2010‐LARGE‐4. Grant Agreement: N° 263007, Available at :https://www.carenafp7.eu/index.php?option=com_docman&task=doc_download&gid=108
Dinse, A., Khennache, S., Frank, B., Hess, C., Herbert, R., Wrabetz, S., Schlögl, R., Schomäcker, R. Oxidative dehydrogenation of propane on silica (SBA-15) supported vanadia catalysts: Akinetic investigation. Journal of Molecular Catalysis A: Chemical, 307(1-2), 43–50 (2009).
Du, X. Yao, B., Gonzalez-Cortes, S. Kuznetsov, V.L., AlMegren, H., Xiao, T., Edwards, P.P. Catalytic dehydrogenation of propane by carbon dioxide: a medium-temperature thermochemical process for carbon dioxide utilization. Faraday Discuss., 183, 161-176, 2015. DOI: 10.1039/C5FD00062A.
Fang, Y.J., Xiao, W.D., 2004. Experimental and Modeling Studies on a Homogeneous Reactive Distillation System for Dimethyl Carbonate Synthesis by Transesterification. Sep. Purifi. Technol. 34 (1-3), 255–263.
Graaf, G.H., Stamhuis, E.J., Beenackers, A.A.C.M., 1988. Kinetics of low-pressure methanol synthesis. Chem. Eng. Sci. 43 (12), 3185-3195.
Han, M.S., Lee, B.G., Ahn, B.S., Park, K.Y., Hong, S.I., 2001. Kinetics of Dimethyl Carbonate Synthesis From Ethylene Carbonate and Methanol Using Alkali-Metal Compounds as Catalysts. React. Kinet. Catal. Lett. 73 (1), 33-38.
Hsu, K., Hsiao, Y., Chien, I. Design and Control of Dimethyl Carbonate-Methanol Separation via Extractive Distillation in the Dimethyl Carbonate Reactive-Distillation Process. Ind. Eng. Chem. Res., Vol 49, No. 2, pp 735–749, 2010.
Kagyrmanova, A.P., Chumachenko, V.A., Korotkikh, V.N., Kashkin, V. N., Noskov, A.S., 2011. Catalytic dehydration of bioethanol to ethylene: Pilot-scale studies and process simulation, Chem. Eng. J. 176-177, 188-194.
Li, B., Zhang, L., Song Y., Bai, D., Jing H., 2012. Brønsted acid improved cycloaddition of carbon dioxide to propylene oxide. J. Mol. Catal. A: Chem. 363– 364, 26– 30.
Liu, l., Deng, Q.F., Agula, B., Zhao, X., Renb, T.Z., Yuan, Z.Y. Ordered mesoporous carbon catalyst for dehydrogenation of propane to propylene. Chem. Commun., 47, 8334-8336, DOI: 2011. 10.1039/C1CC12806J
Michorczyk, P., Zeńczak, K., Niekurzak, R., Ogonowski, J. Dehydrogenation of propane with CO2 - a new green process for propene and synthesis gas production. Polish Journal of Chemical Technology, 14(4), 77-82, 2012.
Mota, C.J.A, Monteiro, R.S., Maia, E.B.V., Pimentel, A.F., Miranda, J.L., Alves, R.M.B., Coutinho, P.L.A., 2014. O Dióxido de Carbono como Matéria-Prima para a Indútria Química. Produção do Metanol Verde, Rev. Virtual Quim. 6(1), 44-59.
Nowicka, E., Reece, C., Willock, D., Golunski, S., Hutchings, G.J. Utilization of CO2 in the process of propane oxidative dehydrogenation. ICEC 2014, 8th International Conference on Environmental Catalysis, Asheville, North Carolina, USA, 24 - 27 August 2014. Available at: https://www.icec2014.us/pdf/GC-P-18.pdf.
Peppel, W.J., 1958. Preparation and properties of the alkylene carbonates. Ind. Eng. Chem. 50 (5), 767–770.
Sakakura, T., Choi, J., Yasuda, H., 2007. Transformation of carbon dioxide. National
Institute of Advanced Industrial Science and Technology (AIST), Chem. Rev. 107, 2365–
2387.
Souza, L.F.S., Ferreira, P.R.R., Medeiros, J.L., Alves, R.M.B., Araújo, O.Q.F., 2013a.
Production of DMC from CO2 via Indirect Route: Process Synthesis and Analysis. In:
Third International Conference on Sustainable Chemical Product and Process Engineering,
Dalian. Proceedings of SCPPE-2013.
Souza, L.F.S., Ferreira, P.R.R., Medeiros, J.L., Alves, R.M.B., Araújo, O.Q.F., 2013b.
Production of DMC from CO2 via Indirect Route: Technical–Economical–Environmental
Assessment and Analysis. ACS Sustainable Chem. Eng. 2 (1), 62-69.
Tóth, A., Halasi, G., Bánsági, T. Solymosi, F. Reactions of propane with CO 2 over Au catalysts. Journal of Catalysis, 337, 57-64, 2016.
Wang, J., Sun, J., Shi, C., Cheng, W., Zhang, X., Zhang, S., 2011. Synthesis of dimethyl carbonate from CO2 and ethylene oxide catalyzed by K2CO3-based binary salts in the presence of H2O. Green Chem. 13, 3213-3217.
Wu, R., Xie, P., Cheng, Y., Yue, Y., Gu, S., Yang, W., Miao, C., Hua, W., Gao, Z., 2013a. Hydrothermally prepared Cr2O3–ZrO2 as a novel efficient catalyst for dehydrogenation of propane with CO2. Catal. Commun. 39 (50), 20-23.
Wu, G., Wang, Y., Wang, L., Feng, W., Shi, H., Lin, Y., Zhang, T., Jin, X., Wang, S., Wu, X., Yao, P., 2013b. Epoxidation of propylene with H2O2 catalyzed by supported TS-1 catalyst in a fixed-bed reactor: Experiments and kinetics. Chem. Eng. J. 215-216, 306-314
Appendix C. Process Flowsheets as Simulated in ASPEN-HYSYS
Figure C.1. Process Flowsheet for Propylene Unit (U1)
Figure C.2. Process Flowsheet for Propylene Oxide Unit (U2)
Figure C.3. Process Flowsheet for Propylene Carbonate Unit (U3)
Etapa de Reação
Etapa de Separação
Figure C.4. Process Flowsheet for Methanol Unit (U4)
Figure C.5. Process Flowsheet for Dimethyl Carbonate, Ethylene Carbonate and Ethylene Glycol Unit (U5)
Figure C.6. Process Flowsheet for Ethylene Unit (U6)
Figure C.7. Process Flowsheet for Ethylene Oxide Unit (U7)
Figure C.8. Process Flowsheet for the integrated chemical complex – Eco-Pole
APPENDIX D. Mass Balance for the Eco-Pole Units as Simulated in ASPEN HYSYS
Table D.1: Mass balance for Unit 1 (Propylene Production) - ASPEN HYSYS
Unit LP STEAM PROPYLENE96% TOP T-121 BOTTOM
T-118 TOP T-118 C3=/C3
Vapor Fraction 1.00 0.00 1.00 0.00 0.97 0.00
Temperature oC 159.00 56.00 49.00 278.00 47.00 73.00
Pressure bar 6.00 20.20 20.00 20.20 20.00 30.10
Molar Flow kgmol/h 1723.80 248.90 140.80 1699.30 599.70 389.60
Mass Flow kg/h 31054.10 10914.40 5923.30 193685.10 25274.40 16837.80
Unit BOTTOM T-104 TOP T-104 PROPYLENE-
001 TOP T-115 MAKE UPn-OCTANE PROPANE
Vapor Fraction 8.05E-4 1.00 0.88 0.98 0.00 1.00
Temperature oC 55.00 41.00 -26.00 38.00 50.00 25.00
Pressure bar 20.20 20.00 2.00 30.00 20.00 2.00
Molar Flow kgmol/h 2299.00 1241.10 406.20 194.90 0.80 203.20
Mass Flow kg/h 218959.50 38200.60 17902.50 7930.30 94.80 8962.30
Table D.2. Mass Balance for the unit U2 (Propyle Oxide Production) - ASPEN HYSYS
Unit OP-4 METHANOL 99.5% H2O2 99.9% VENT 3 H2O2*Vapor Fraction 0.00 0.00 0.00 1.00 0.00
Temperature oC 40.00 83.00 156.00 53.00 25.00
Pressure bar 4.00 2.00 1.20 2.00 1.00
Molar Flow kgmol/h 1293.40 964.30 161.80 0.00 79.50
Mass Flow kg/h 42781.50 30944.80 5538.00 0.00 2702.70
Unit OP-25 H2O+ H2O2 OP-METHANOL VENT-5 Vapor Fraction 0.00 3.0 E-06 0.00 0.49
Temperature oC 53.00 125.00 53.00 53.00
Pressure bar 2.00 1.2.00 1.00 1.00
Molar Flow kgmol/h 83.60 245.30 1048.00 0.10
Mass Flow kg/h 4707.60 7122.80 35652.80 3.50
Table D.3. Mass Balance for unit U3 (Propylene Carbonate Production) - ASPEN HYSYS
Unit BOTTOM T-107
PROPYLENE OXIDE METHANOL 1 PROPYLENE
OXIDE-1 PURGE 4
Vapor Fraction 0.00 1.00 0.00 1.00 1.00
Temperature oC 410.00 45.00 64.00 31.00 31.00
Pressure bar 1.00 1.00 1.00 1.00 1.00
Molar Flow kgmol/h 36.70 7.90 2.70 5.30 0.50
Mass Flow kg/h 3750.30 375.30 86.10 296.80 29.70
Unit LP STEAM 5
PROPYLENE CARBONATE OP-13 CO2-1
Vapor Fraction 1.00 0.00 0.00 1.00
Temperature oC 159.00 170.00 53.00 30.00
Pressure bar 6.00 0.50 2.00 5.00
Molar Flow kgmol/h 33.70 36.70 40.00 36.70
Mass Flow kg/h 607.60 3750.30 2250.00 1616.30
Table D.4. Mass Balance for unit U4 – Methanol Production - ASPEN HYSYS
Unit H-07 H2 H-01 H-11 C-01 VENT-1Vapor Fraction 1.00 1.00 1.00 0.88 1.00 1.00
Temperature oC 220.00 40.00 170.00 10.00 0.00 78.00
Pressure bar 46.00 20.00 51.00 44.00 4.00 10.00
Molar Flow kgmol/h 11413.00 2015.00 2015.00 11413.00 796.00 3.00
Mass Flow kg/h 121802.00 4062.00 4062.00 121802.00 35086.00 120.00
Unit H-14 H-03 H-16 H-17 RESIDUALWATER2 VENT-1
Vapor Fraction 1.00 1.00 0.00 0.00 0.00 1.00
Temperature C 44.00 44.00 10.00 78.00 180.00 78.00
Pressure bar 50.00 50.00 10.00 10.00 10.00 10.00
Molar Flow kgmol/h 10069.00 10070.00 1344.00 676.00 665.00 3.00
Mass Flow kg/h 88083.00 88342.00 33718.00 21617.00 11981.00 120.00
Table D.5. Mass Balance for unit U5 (Methanol DMC, EC and EG Production) - ASPEN HYSYS
UnitBOTTOM
TOP T-110 CO2-1 CO2-5 CO2-4 CO2rec C3T-110
Vapor Fraction 0.00 1.00 1.00 1.00 1.00 1.00 0.00
Temperature oC 152.00 128.00 30.00 80.00 160.00 100.00 100.00
Pressure Bar 0.10 0.08 5.00 13.00 39.50 39.50 39.50
Molar Flow kgmol/h 49.20 85.50 129.20 129.20 129.20 43.70 128.70
Mass Flow kg/h 4064.70 5307.90 5683.90 5683.90 5683.90 1924.00 11287.20
Unit C8 C11 EGrec EGfinal TOP T-106 BOTTOM T-106 VENT 6
Vapor Fraction 1.00 0.00 0.00 0.00 1.00 0.00 1.00
Temperature oC 100.00 40.00 166.00 40.00 38.00 166.00 41.00
Pressure bar 39.50 1.50 0.36 1.00 0.29 0.29 0.58
Molar Flow kgmol/h 43.30 183.90 1508.40 85.50 668.10 134.70 5.50
Mass Flow kg/h 1905.80 5885.20 93702.20 5307.90 27694.50 9372.60 214.20
Unit TOP T-116BOTTOM
TOP T-117 BOTTOM T-117 DMC EC
T-116Vapor Fraction 0.00 0.00 0.00 0.00 0.00 0.00
Temperature oC 41.00 162.00 53.00 166.00 40.00 40.00
Pressure bar 0.58 0.59 0.30 0.36 1.00 1.00
Molar Flow kgmol/h 580.00 1604.80 89.90 1514.90 89.90 49.20
Mass Flow kg/h 19936.20 102097.20 7991.90 94105.30 7991.90 4064.70
Table D.6. Mass Balance for unit U6 (Ethylene Production) - ASPEN HYSYS
Unit BIOETHANOL B-09 B-04 ETHANOL 70% B-14
Vapor Fraction 0.00 0.00 1.00 0.00 0.0002
Temperature oC 25.00 40.00 138.00 56.00 40.00
Pressure bar 1.00 7.00 8.00 3.00 7.00
Molar Flow kgmol/h 2200.00 973.00 3294.00 455.00 995.00
Mass Flow kg/h 93450.00 26964.00 88245.00 17957.00 27731.00
Unit VENT 4 1-BUTENE + DIETIEL ÉTER B-15 B-12
Vapor Fraction 0.0272 0.00 0.9971 1.00
Temperature oC 56.00 144.00 -12.00 40.00
Pressure bar 3.00 30.00 30.00 30.00
Molar Flow kgmol/h 0.00 106.00 1287.00 1392.00
Mass Flow kg/h 1.00 6153.00 35938.00 42091.00
Table D.7. Mass Balance for unit U6 (Ethylene Oxide Production) - ASPEN HYSYS
Unit OXYGEN PURGE-1 B-16
BOTTOM
T-112
BOTTOM
T-109
Vapor Fraction 1.00 1.00 0.9971 0.00 0.00
Temperature oC 30.00 41.00 -12.00 29.00 114.00
Pressure bar 5.07 18.25 20.27 2.00 2.00
Molar Flow kgmol/h 199.90 18.10 421.00 376.40 499.70
Mass Flow kg/h 6396.80 514.60 11756.50 16582.60 9069.10
Unit TOP T-109 BOTTOM T-108 TOP T-108 TOP T-112
Vapor Fraction 1.00 0.00 1.00 1.00
Temperature oC 22.00 49.00 41.00 11.00
Pressure bar 2.00 18.30 18.25 2.00
Molar Flow kgmol/h 642.60 902.70 181.40 266.30
Mass Flow kg/h 26348.60 26628.30 5147.90 9766.00
