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Techno-economic analysis of the production of epichlorohydrin from glycerol. Alberto Almena, and Mariano Martín Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.5b02555 • Publication Date (Web): 28 Oct 2015 Downloaded from http://pubs.acs.org on November 3, 2015

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Industrial & Engineering Chemistry Research

Techno-economic analysis of the production of epichlorohydrin from glycerol. Alberto Almena, Mariano Martín1 Department of Chemical Engineering. University of Salamanca, Pza. Caídos 1-5, 37008 Salamanca (Spain) Abstract

In this work we evaluate the production of epichlorohydrin from glycerol, the main byproduct in the biodiesel industry. A process flowsheet is synthesized consisting of glycerol purification, reaction with HCl and final reaction with NaOH to produce epichlorohydrin. Both reactions are complex equilibriums and recovery of the reaction medium, hexanoic acid, and purification of the products is required. The process is selected based on the use of cheap and available secondary raw materials, namely HCl and NaOH, compared to other technical alternatives. We couple MATLAB with CHEMCAD to simulate the mass and energy balances. Finally, the economic analysis of the process shows that the process is profitable due to the high selling cost of the epichlorohydrin. For the production of 26.5 kt/yr of epichlorohydrin 99.9%, 41.5 kt/yr of glycerol, 29 t/y of HCl and 15.9 kt/yr of NaOH are need. The investment adds up to 63.7 M€ with a production cost of 1.28 €/kg.

Keywords: Glycerol, epichlorohydrin, Techno-economic analysis, Process simulation.

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Corresponding author. Tel.: +34 923294479 Email address: [email protected]

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1.-Introduction.

Along with biodiesel, in a ratio 10:1, glycerol is obtained as byproduct. Therefore, by increasing the amount of biodiesel produced, the availability of glycerol increases too. As a result, it is imperative to find alternative markets and uses for glycerol capable of processing the large amount produced within biorefineries, while providing an added value that helps with the biodiesel economy. There are two possible strategies to solve this problem, either to use glycerol as raw material for obtaining fuel additives or to produce higher value products and commodities. Over the last years different chemicals have been produced out of glycerol such as hydrogen and FT-liquids, 1 ethanol, to enhance up to 5% the production of biofuels from an algae based biorefinery, 2 methanol, to reduce by one half the fossil based raw materials in the production of biodiesel 3 or glycerol ethers, increasing the yield to biofuels by 20%. 4 However, the margin of benefits in biofuels production is tight. Alternatively, glycerol can be used in the production of polymers such as PHB or polyesters, 5 with more favourable economics. Epichlorohydrin is a high volume commodity chemical used largely for epoxy resins production, although smaller quantities have, until recently, been employed for the manufacture of synthetic glycerol.6 By looking at 2012 data, the worldwide production of epichlorohydrin raised to about 2.0 million Mg (near to 9 x 106 kt). 7 In spite of the several routes known for epichlorohydrin manufacture, it is mostly obtained from propylene and chlorine in a multi-step process. This process consists of the allylic chlorination of propylene to allyl chloride followed by hypochlorination, resulting in a 3:1 mixture of 1,3-DCH and 2,3-DCH (dichlorohydrin: DCH). The product mixture is treated with an alkali to obtain epichlorohydrin. 8 Although this process is widely used at large scale, it suffers from some undesirable features, particularly the low yield in terms of chlorine usage. Only one out of the four chlorine atoms employed in the manufacture of epichlorohydrin using this route is retained in the product molecule, while the rest emerges as by-product hydrogen chloride or waste chloride anion. Additionally, inefficiencies in the chlorination and hypochlorination steps lead to the formation of unwanted chlorinated organics that, together with the large amount of waste water produced by the conventional production process, are expensive to dispose of.9 Such factors have prompted the search for


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alternative routes to epichlorohydrin that are more atom-efficient and environment-friendly. The escalating cost of petrochemical raw materials such as propylene has also contributed to the accelerated search for processes that employ less expensive and renewable raw materials.8 Furthermore, the raw materials, propylene and chlorine, are flammable and toxic respectively. However, epichlorohydrin can be produced from glycerol in a two step process where first dichlorhydrin is produced in a gas-liquid contactor and later epichlorohydrin is produced in a packed bed one. In this paper we present the conceptual design of the production process to transform glycerin into epichlorohydrin using HCl and NaOH as reagents instead of other crude based sources of carbon. We use a simulation based approach for rigorous simulation of the process integrating MATLAB and CHEMCAD to design a process and evaluate its economics. The rest of the paper is divided as follows. Section two describes the process and presents the flowsheet. In section three, the models for the different units are presented. We used CHEMCAD for modelling the distillation towers and MATLAB-EXCEL for modelling the reactors. Next, we present the production cost and the investment for such a plant. Finally, a discussion on the process and its perspectives is presented.

2.- Process description

The process consists of four stages: feedstock preparation involving glycerol purification and reactants preparation, glycerol hydrochlorination and product recovery, epichlorohydrin production and product purification up to commercial composition. The crude glycerol is fed to the process at standard conditions, 293 K and 100 kPa. We assume that it does not contain any NaOH as residue from the biodiesel production process. We expect that the use of heterogeneous catalysts is employed in the near future and NaOH residues will no longer be an issue. Otherwise, a prior neutralization step would be required to remove this impurity. Next, glycerol has to be heated up to the boiling point before it is fed to the distillation column as saturated liquid. In this column, T-01, the glycerol is separated from the impurities, namely water and methanol. A packed distillation column is used, 3 ACS Paragon Plus Environment


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with an upper bound for the bottoms temperature of 423 K, to avoid glycerol decomposition.10 Thus, the packed column operates under vacuum, at a pressure up to 10 kPa. Basically, all the water and methanol will be separated as distillate from the glycerol, which is fed to the next stage. The bottom stream containing the main reactant, glycerol, is mixed with the recirculation, comprising the recycled catalyst and the heavy intermediate products from the bottom of the column T-02. The mixture is cooled down to the reaction temperature (383 K) and loaded into the semibatch reactor, R-01. A semibatch reactor is chosen to minimize the side-reaction of glycerol to trichloropropane, which occurs in the presence of a high concentration of hydrogen chloride in the reaction mixture. Furthermore, a semibatch reactor allows the continuous addition of hydrogen chloride maintaining its concentration low.8 The other main reactant introduced into reactor R-01 is the chlorinating agent, hydrogen chloride, which needs to be brought to the reaction conditions. A multistep compression system is used, consisting of two compressors and two interstage coolers. The hydrogen chloride is compressed up to 760 kPa, cooled down to 383 K and fed into the semibatch reactor. Intercooling aims at reducing the energy consumption. Next, the hydrogen chloride is bubbled into the reaction mixture. The reaction that occurs in the semibatch reactor is the hydrochlorination of glycerol, presented in Figure 1.

Figure 1 – Hydrochlorination of glycerol reaction.


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The hydrochlorination is performed isothermally at 383 K and under hydrogen chloride pressure of 760 kPa, in the presence of hexanoic acid (5 % mol) as catalyst. Working at high pressure the liquid-vapor contact is improved, and the existing equilibrium limit shown for the absorption of the HCl in glycerol at atmospheric pressure is overcome, 8 resulting in a better absorption of the gas in the liquid and consequently a more efficient production of DCH. Operating isothermally allows controlling the side reactions. 8 The reaction product mixture comprises dichlorohydrins, in a mole ratio of 30-50 to 1 of 1,3-DCH and 2,3-DCH respectively, monochlorohydrins (MCH), water and catalyst. This ratio represents a very important advantage over the current industrial process, showing a ratio only of 3 to 1. The improved 1,3-regioselectivity undergoes cyclization with a base to form epichlorohydrin about 300 times faster than from the 2,3-DCH. This translates into smaller equipment in the second reaction stage. The chlorinating agent is bubbled to the reaction mixture gradually, in an excess of 10 % by weight. Keeping its concentration in the reaction mixture at a low value, together with the choice of the optimal working catalyst, inhibits the side reaction of trichloropropane formation. 8 The hydrogen chloride not dissolved in the reaction mixture, together with the steam formed by the evaporation of the water in the mixture, is evacuated from the isothermal semibatch reactor when the internal pressure exceeds the upper bound. After 3 h of reaction, the mixture containing the dichlorohydrins is cooled and expanded to the pressure conditions of the product recovery column. At this stage, dichlorohydrins, hydrochloric acid and water formed in the reaction step are separated from the heavy components, namely monochlorohydrins, catalyst and unreacted glycerol. The recirculation comprises the heavy components, obtained at the bottom of the column. This stream is returned to the reaction step. To maintain low pressure drop, a packed column is usedAlmost sharp separation of the substances can be achieved by operating under vacuum, 5 kPa.11 The distillate is sent to a liquid-liquid separator where the two phases, organic and aqueous, are separated. The purpose of this embodiment is to make use of the amount of water formed in the first reaction stage to prepare the sodium hydroxide solution, a reactant of the second reaction stage. The organic phase, comprising mostly of dichlorohydrins and a small amount of water dissolved, is heated up to the reaction temperature and sent to reactor R-02. On the other hand, the aqueous phase with 5 ACS Paragon Plus Environment


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most of the water, and saturated with dichlorohydrins and other species such as hydrochloric acid and traces of catalyst, is mixed with a concentrated solution of sodium hydroxide in mixer M-03 to be first neutralized and later diluted to the appropriate concentration for the reaction: 20 w% aqueous solution. The original solution of NaOH is prepared in a stirred tank, M-02. The sodium hydroxide fed to the process represents an excess of 5 mol % of the dichlorohydrins fed to the reactor R-02, plus the necessary amount to neutralize the hydrochloric acid. Both stirred tanks operate adiabatically, reducing the energy consumption in the subsequent heating of the solution to reaction temperature. The reactant streams are only mixed in the reactor, to prevent the formation of epichlorohydrin and its decomposition before the reaction stage. Reactants are fed to the top of the reactor tower R-02, a packed column where dehydrochlorination of dichlorohydrins occurs. It operates isothermally at 363 K to avoid an increase in the yield of the hydrolysis of epichlorohydrin and under vacuum conditions, 30 kPa, to make easier the azeotropic distillation12. Therefore, external refrigeration is needed. The main reaction is presented in Figure 2:

Figure 2 – Dehydrochlorination of dichlorohydrins reaction.

The advantage of a reactive distillation column is the short contact time between the water and the epichlorohydrin, minimizing the hydrolysis of the produced epichlorohydrin to glycerol, which represents the major yield losses. Steam is injected from the bottom of the column, stripping the undissolved epichlorohydrin produced by the reaction. The amount of steam used for the stripping of epichlorohydrin is such that the composition at the top has a water/epichlorohydrin ratio by weight of from about 1 to 2.5.12


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The bottoms of the column are mainly waste water, with all the substances dissolved in it, such as epichlorohydrin, sodium chloride, unreacted dichlorohydrins and sodium hydroxide excess. This stream is cooled to make use of its energy and disposed of as waste product. The top stream comprising the epichlorohydrin-water azeotrope, a mixture 24 w% of water at atmospheric pressure,

13

is cooled down and condensed. Next, it is sent to a decanter to separate the two

phases, according to the identification of this azeotrope as heterogeneous. 14 The organic phase, with a water concentration of 1.48 w%, set by the solubility limit of water in epichlorohydrin, goes to an evaporator to follow the purification process, while the aqueous phase is treated as waste.15 Further studies of this process may found how to treat the wastewater so as to reuse it joining the stripping steam to the R-02. A previous treatment of this stream would be needed to remove the epichlorohydrin that accompany in a 6.58 w%,15 according to the solubility of epichlorohydrin in water, which would affect the equilibra in the reactor. The next stage in the purification of the epichlorohydrin is a pressure swing absorption (PSA) system. The evaporated organic phase obtained in the decanter is compressed to 4.5 bar to improve the water absorption in the molecular sieve. The PSA unit consists of two parallel 4A zeolite beds, with an absorption capacity of at least 0.215 mg/g.16 When one of the molecular sieves is working, the other one is being regenerated using hot air. Thus, continuous operation is ensured and the purity of the epichlorohydrin is raised to 99.9%. Finally, the product is expanded to atmospheric pressure, recovering energy, and then it is condensed and stored. A schematic flow chart of the studied process is represented in Figure 3 as follows:



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Figure 3 – Process flow chart.

3. - Modelling issues. 3.1. - Distillation columns. Rigorous simulation using CHEMCAD. In the process there are two distillation columns that are modelled in CHEMCAD. The first one, T-01, is a packed distillation column that purifies the glycerol, from the composition obtained in a biodiesel plant, to the appropriate one for reaction. This column has as a constraint that the bottoms, consisting mainly on glycerol, must exit the column below 423 K to avoid decomposition. Therefore, it operates under vacuum, at 10 kPa. 15 stages and total condenser are used to reach a 99.99 % recovery of glycerol. The thermodynamic model for this column is UNIFAC. Water and methanol exit as distillate. In the second tower, the packed distillation column T-02, heavy components comprising monochlorohydrins, catalyst and unreacted glycerol are recovered and recycled while the main product is sent to further reactions. CHEMCAD database has only one of the dichlorohydrins and one of the monochlorohydrins isomer defined, 1,3-DCH and 3-MCH respectively. The similar boiling points and characteristics of each pair of isomers, together with the fact that the defined ones are found in a much larger proportion, made us assume that the two monochlorohydrins will behave as 3-chloropropane-1,2-diol, while 8 ACS Paragon Plus Environment

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both dichlorohydrins behave as 1,3-dichloropropan-2-ol. The thermodynamic model selected for this column is NRTL. The column is expected to recover 99.99 % of the dichlorohydrins in the distillate, accompanied by the water in the feed to the column, with a package height of 30.2 m and operating at 5 kPa. Total condenser is used. Nearly 100 % of the monochlorohydrins, catalyst and glycerol are obtained in the bottoms.

3.2.-Synthesis reactor R-01. This reactor is responsible for the first reaction stage of the process. In this section, the mechanism and kinetic model for the production of monochlorohydrins and dichlorohydrins from glycerol are described. The hydrochlorination of glycerol, catalyzed by a carboxylic acid, can be considered to take place following a two step mechanism, see Figure 4. It starts with glycerol chlorination to monochlorohydrins, mainly to 3-MCH, which suffers a second chlorination to dichlorohydrin. For each of the five chlorinations, a hydroxyl group of the glycerol is substituted by a chlorine atom and, consequently, a water molecule is formed as byproduct. The 1,3-dichlorohydrin is the main product of this reaction and represents a great advantage over the current industrial process.

Figure 4 – Dehydrochlorination of glycerol scheme. At this point, it is important to perform a detailed analysis of the first reaction step of the process, describing the assumptions within the mechanism of dehydrochlorination of glycerol. Based on experimental



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observations, 17 it was found that the reaction that converts 2-MCH into 2,3-DCH is not possible because of the absence of vicinal OH groups, that inhibits the second hydroxyl mechanism substitution, and thus, it can be neglected. Furthermore, reactions 2 and 4 can be considered irreversible because 2-MCH is produced during the reaction and 2,3-DCH is always obtained in small quantities. Therefore, the scheme can be simplified into the one shown in Figure 5.

Figure 5 – Dehydrochlorination of glycerol simplified scheme.

The four reactions comprising the catalyzed glycerol hydrochlorination, which best explains the experimental data, can be represented by the following equations: 1)

C3H8O3 + HCl + Catalyst

K1

3-C3H7ClO2 + H2O + Catalyst

K-1

2)

C3H8O3 + HCl + Catalyst

3)

3-C3H7ClO2 + HCl + Catalyst

K2

K3

2-C3H7ClO2 + H2O + Catalyst

1,3-C3H6ClO + H2O + Catalyst

K-3

4)

3-C3H7ClO2 + HCl + Catalyst

K4

2,3-C3H6ClO + H2O + Catalyst

Each of the reactions presented above are based on the substitution of a hydroxyl group with a chlorine one. The elementary steps are beyond the scope of this paper. There are two hypothesis, considering all the steps of the mechanism elementary (more complex) or assuming the ester formation as the rate determining step, which agrees with experimental evidence. To simplify the kinetic model for each substitution, where the ester formation is considered as the rate-determining step, a pseudo constant is introduced, that 10 ACS Paragon Plus Environment

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lumps all the elementary steps of the complex hydrochlorination mechanism. Thus, the kinetic model becomes as follows: 17

= . . − . .

(1)

= . . .

(2)

= . . − . , .

(3)

= . . .

(4)

Values of the kinetic constants are listed in Table 1.17 They are obtained under atmospheric pressure and using malonic acid as carboxylic acid catalyst. Table 1. – Kinetic constant and Arrhenius parameters for catalyst glycerol hydrochlorination.

K 1*

T (ºC) 80 90 100 110 120

K 2*

K 3*

K 4*

7 667 ± 940 11 704 ± 1 272 13 274 ± 1 692 19 433 ± 2 216 27 411 ± 2 861

450 ± 41 764 ± 60 1 089 ± 87 1 465 ± 123 2 215 ± 170

714 ± 227 1 109 ± 307 1 784 ± 407 2 383 ± 532 2 179 ± 685

8±3 13 ± 5 26 ± 7 32 ± 9 31 ± 13

Reaction 1

Reaction 2

Reaction 3

Reaction 4

35.2 ± 0.3 20.9 ± 9

44.3 ± 0.2 21.3 ± 0.7

34.9 ± 0.8 18.6 ± 2.2

42.1 ± 1.0 16.5 ± 2.8

-1

Ea (kJ mol ) ln A

KE1*

KE3*

3 846 3 064 2 470 2 015 1 660

194 167 146 128 113

*Kinetic constants are expressed in cm6/(mol2 min) Even under these pressure conditions, it is admitted that the equilibria represented by equations 1 and 3 are almost completely shifted towards products.17 Further studies show that by increasing the HCl pressure, the equilibrium limit is overcome.8, 18 Moreover, according to the results of Santacesaria, et al., 18 by increasing the HCl pressure, the yield to the main product increases by a factor by almost 3. Therefore, the kinetics of the production of 1,3-DCH under pressure is speeded up. Bell, et al.8 found that the optimal HCl pressure is 760 kPa. Finally, the use of a more active catalyst, selective to the main product and easy to recycle, is expected for industrial scale. Therefore, the carboxylic acid of choice is hexanoic acid. The reason to select it over acetic acid, the catalyst with the best performance studied so far, is the fact that hexanoic acid is as 11 ACS Paragon Plus Environment


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effective as acetic acid at the end of the reaction time but it is also less volatile, even heavier than dichlorohydrins.8,19 Therefore, the recovery of the catalyst is easier, allowing its recycle. The kinetic rates listed in Table 1 are obtained experimentally at atmospheric pressure using malonic acid for catalyzing the glycerol hydrochlorination. Therefore, the kinetic data available in the literature are not obtained under the favorable conditions described above. To provide simulation results under these conditions, we have used the mechanisms described in the literature to come up with a modified kinetic model. For this model to match the experimental results at high pressure and using a different catalyst, we use a parameter estimation problem based on an adjustable parameter (θP). Typically the absorption of a gas in a liquid is directly proportional to the pressure and thus, this parameter multiplies the kinetic rates of the literature. The value of this parameter is calculated using an error function in Matlab, so that the model allows fitting the experimental results at the optimal pressure at 3 hours of reaction time, Bell, et al.,8 Santacesaria, et al.18 and Herliati, et al.20

Thus, after the elimination of the equilibrium limitation by increasing the HCl pressure, and adjusting the reaction rate, the following model becomes as given by equations (5) – (8):

= . "# . . .

(5)

= . "# . . .

(6)

= . "# . . .

(7)

= . "# . . .

(8)

The optimal conditions to achieve a high yield to 1,3-DCH include an operating temperature of 383 K, gradual feed of HCl to the reactor in 10 w% excess with respect to the stoichiometric mass and a pressure of 760 kPa. The catalyst ratio needed is 0.08 mol of hexanoic acid per mol of glycerol fed to reactor. 8 The most appropriate reactor to carry out the catalyzed glycerol hydrochlorination is a stirred isothermal semibatch unit. The glycerol and the hexanoic acid are charged into the reactor at reaction 12 ACS Paragon Plus Environment

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conditions. Then, the gaseous chlorinating agent is bubbled gradually as the reaction proceeds. Avoiding a great excess of hydrogen chloride in the reaction mixture makes negligible a side reaction to trichloropropene, while the operating conditions minimizes the glycerol and monochlorohydrins esterification. Mass balances of components in the semicontinuous reactor, according to the reactions seen before, are shown in equations (9) – (15): .- Glycerol: $%&'(

= −) + + . ,)-+

(9)

= 0 1 − 0 2 − ) + + + + . ,)-+

(10)

$

.- Hydrogen chloride: $%./' $

.- 3-chloropropane-1,2-diol: $%34/. $

= ) − − + . ,)-+

(11)

.- 2-chloropropane-1,3-diol: $%34/. $

= ) + . ,)-+

(12)

.- 1,3-dichloropropan-2-ol: $%,35/.

= ) + . ,)-+

(13)

$%,35/.

= ) + . ,)-+

(14)

$

.- 2,3-diclhoropropan-1-ol: $

.- Water: $%.6 $

= ) + + + + . ,)-+ − 0 2

(15)

The gas phase exiting the semibatch has been computed using an offline simulation in CHEMCAD to account for the gas-liquid equilibrium determining the amount of HCl and water vapor not dissolved in the liquid phase. The main reaction takes place in the liquid phase so, although an amount of gas is added, the



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density of the reaction mixture has a negligible variation. Thus, the reaction mixture volume has a linear relationship with time: ,)-+ = ,7 + 87 . -

(16)

Finally, it is necessary to mention that we consider that the 2-MCH fed to the semibatch reactor by the recirculation stream reacts to dichlorohydrin following the equations (7) and (8), just like the 3-MCH formed inside the reactor does. This consideration is reasonable due to the small amount that is recycled, the heavy key is the catalyst, and the fact that the time available for 2-MCH to react is 3 hours, which appears to be sufficient for the substitution of the hydroxyl group occurs.8 Equations (5) – (16) comprise the synthesis reactor model, an ordinary differential equations system that is solved using MATLAB R2013. The initial conditions for the Matlab model are the charged kilomoles of stream 6 *, in one hour of reactor charging time, and the following parameters: T= 383 K; Vo= 4272 dm3; vo= 0.288 cm3/s; FHCle=9.2510-3 kmol/ s; FHCls=6.9310-4 kmol/ s; FH2Os=3.2910-4 kmol/ s.

3.3.-Reactive distillation column. The second reaction stage is carried out in a reactive distillation column, where the dehydrochlorination of dichlorohydrins takes place. The kinetics of the reaction and the modeling of the reactor are depicted below. The dehydrochlorination of dichlorohydrins is a one step reaction where a dichlorohydrin molecule reacts with a sodium hydroxide molecule in aqueous medium, as follows: C3H6Cl2O + NaOH

C3H5ClO + NaCl + H2O

Both dichlorohydrins react to epichlorohydrin following the above expression, but the rate of the transformation of 1,3-DCH is about 300 times faster than of 2,3-DCH.8 The presence of epichlorohydrin in an aqueous basic medium causes epichlorohydrin hydrolysis as a side reaction, which represents an important yield loss: 21 C3H5ClO + NaOH + H2O

C3H8O3 + NaCl 14

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However, the use of a reactive distillation column makes possible a short contact time between water and epichlorohydrin that, together with the optimal conditions, enhance the main reaction and minimize the epichlorohydrin hydrolysis. The optimal reaction conditions are 363 K and 30 kPa.9,21 The basic aqueous solution reactant must have a concentration of 20 w%, in an amount of sodium hydroxide that represents an excess of 5 mol% related to the dichlorohydrins fed to the reactor.12 Furthermore, the column operates isothermally at the optimal temperature, removing the heat of reaction that would enhance the main reaction, but also the hydrolysis. The kinetics of the dehydrochlorination and hydrolysis reactions has been studied experimentally. It turns out that second order kinetic model provides the best fit in both cases.21 Thus, the kinetic equation for the dehydrochlorination of dichlorohydrins can be expressed as given by eq. (17): = 9 . [DCH] . [NaOH]

(17)

The reaction rate follows the Arrhenius model. In the temperature range 313 K – 333 K, close to the reaction temperature, the kinetic constant 9 is given as follows: 21 9 = 8.97 x 107 . e The reaction temperature, 363 K,

9, 21

77K I .J

(l/mol s)

(18)

is above the upper limit, but it is assumed that the Arrhenius

model holds. Equation (17) was obtained for the 1,3-DCH. 2,3-DCH reacts to epichlorohydrin 300 times slower than the isomer 1,3-DCH.8 Therefore, it is assumed that the kinetic constant of the 2,3-DCH (9 ∗ ) is 300 times smaller than 9 . The reaction rate for 2,3-DCH is named ∗ . On the other hand, the experimental results for the hydrolysis of the epichlorohydrin can be represented by assuming a second order kinetic model, such as given by eq. (19): M = 9M . [NOP] . [QP ]

(19)

The reaction rate of the hydrolysis of epichlorohydrin follows an Arrhenius type of model too. The kinetic constant 9M is computed using eq. (20): 21 9M = 5.66 x 107 . e

T7TU7K I .J

(l/mol min)

(20) 15



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The main reaction rate is more than ten orders of magnitude larger than the side reaction. However, the hydrolysis rate depends on the OH- anions, including those of the water, so that the yield of the side reaction could be important because of the large amount of water available within the reaction mixture. To minimize the hydrolysis loss, it is necessary to reduce the contact time between water and epichlorohydrin. The reactor of choice is an isothermal reactive distillation column, as it was discussed in the process description. The column packing consists of metal Pall rings of 50 mm diameter and disposed randomly, according to the rules of thumb and the design specifications. At the bottom, steam is injected to strip the undissolved epichlorohydrin produced by the reaction. The countercurrent vapor stream is not considered for the global mass balance given by the set of differential equations, but the performance of its flow across the package is studied and plotted later. External refrigeration is needed to remove a large part of the energy produced during the reaction, which exceeds the needs to evaporate the epichlorohydrin generated, in order to maintain an isothermal operation. The plug flow model is assumed to hold for the operation of the column reactor. The countercurrent vapor stream is not considered in the reactor model, but the performance of its flow across the package is studied and plotted later. Thus, a mass balance is performed to a differential volume of the column (see Figure 6), where the reaction rate is considered invariable. The balance can be expressed as given by eq. (21): V X∆VW

)0 +VW − )0 +VWX∆VW − ZV W W

)[# + \, = 0

(21)

Figure 6 – Mass balance in a differential volume of reactor R-02. 16 ACS Paragon Plus Environment

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Since the cross section is constant along the column, the differential of volume can be expressed as the product of transversal section (]), the porosity of the package (^), the liquid hold up, (εL) , where the reaction actually takes place, and a differential of height (∆_). Applying the First Mean Value Theorem for Integrals and the Mean Value Theorem for Derivatives, the eq. (22) is obtained: $à $b

= ^c · ^ . ] . ) +

(22)

According to the developed model and considering the reactions that take place in the column reactor, the mass balances for each component are shown in eqs. (23) – (29): .- 1,3-dichloropropan-2-ol: $`,35/. $b

= − ^c · ^ . ] . ) +

(23)

= − ^c · ^ . ] . ) ∗ +

(24)

.- 2,3-dichloropropan-1-ol: $`,35/. $b

.- Sodium hydroxide: $èf6. $b

= − ^c · ^ . ] . ) + ∗ + M +

(25)

= ^c · ^ . ] . ) + ∗ − M +

(26)

.- Epichlorohydrin: $`. $b

.- Sodium chloride: $èf/' $b

= ^c · ^ . ] . ) + ∗ + M +

(27)

$`.6

= ^c · ^ . ] . ) + ∗ − M +

(28)

.- Water: $b

. - Glycerol: $`&'( $b

= ^c · ^ . ] . )M +

(29)

The water stream dissolves epichlorohydrin until it is saturated and takes on other species such as NaOH, NaCl, glycerol and DCH’s that are below the solubility limit so that the stream exits the column through 17 ACS Paragon Plus Environment


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bottoms as a waste. The ordinary differential equations system consists of eq. 17 – 20 and eq. 23 – 29 and it is solved using MATLAB R2013, being the initial conditions the shown kinetics constants, the molar inflow of the different components, the porosity of de packing (^ = 0.94),22 and the section of the column (] = 0.636 i ).23 The reaction takes place in the liquid phase and therefore it only considers a fraction of the void volume of the packed column. However, for simplicity due to the unknown hydrodynamics of the column, we assumed that the kinetic rates and the porosity already account for that effect in the parameters. Thus, the liquid hold up is assumed to be one in the computations (^c = 1.)Further validation of the model is however required, but the lack of data in the literature prevented this fact. Therefore, we oversize the column height by 10% to be on the safe side. Furthermore, we will compute the epichlorohydrin in vapor and dissolved to determine the one that exits on top and from the bottoms, in which case takes the solubility limit of epichlorohydrin in water with a value of 0.0658 kg/kg. 15 4.- Results In this section, results on the simulation of an epichlorohydrin production plant using glycerol as raw material are presented. The economics are performed assuming that the plant is allocated in Spain. Based on the market for epichlorohydrin and the availability of glycerol, economic and financial viability studies are carried out. It turns out that an economic plant size is 26,000 t/yr. 23

4.1.- Mass and energy balances The simulation of the process provides the mass and energy balances, the needs of raw materials and utilities. The raw materials fed to the production process comprise 41,500 t/yr of crude glycerol, 29,300 t/yr of anhydrous hydrogen chloride and 15,200 t/yr of sodium hydroxide. Their prices, updated to 2014, are given as Delivery Duty Paid (DDP), which represents the final price of the substance at the plant’s door. In this regard, crude glycerol costs 20.69 €/t DDP, 24 with an estimated mean concentration of 80 w% of glycerol, 15 w% of methanol and 5 w% of water as impurities.25 On the other hand, anhydrous hydrogen chloride with a 99 w% of purity has a price of 261.34 €/t DDP, 26 while 99 w% pure sodium hydroxide costs 252.74 €/t.27


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Furthermore, 56,400 t/yr of 4x106 Pa steam are needed, together with 15,900 t/yr of 6x106 Pa steam and 5.3x106 t/yr of cooling water. These utilities are needed across the plant including heat exchangers, condensers and reboilers of distillation columns and reactors cooling. The results of the reactors simulation are shown below. The semibatch reactor for the production of mono and dichlorohydrins is charged initially with glycerol, catalyst and other accompanying substances as described before; loading the reactor takes one hour. On the other hand, the anhydrous hydrogen chloride is added gradually, with a flow rate of 20 kg/min as required by the reaction conditions. As the reaction time last three hours, four semibatch reactors are necessary to assure a continuous flow of products and good temperature control. According to the kinetic mechanism developed in section 3, the adjustable parameter ("# ) has a value of (8.5 10-4)-1. Thus, the profile of the products in the semibatch during the three hours is shown in Figure 7. It can be seen the formation of 3-MCH peaking near 5500 s. Next, the second hydroxyl substitution begins to be important and the 1,3-DCH is formed. The HCl concentration in the reaction mixture is kept at low values, in order to avoid the formation of trichloropropane. It can be seen that a large amount of water is also obtained as byproduct, while glycerol is totally converted and the other species, namely 2,3-DCH, 2-MCH and dissolved HCl, are obtained at a low concentration. It is important to mention that there is no 2-MCH entering the reactor, as we can see in figure 7. As it was explained in the reactor modelling section, 2-MCH inflow is included as 3MCH inlet.8



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Figure 7 – Evolution in time of the components moles in semibatch reactor. The second reaction step of the process occurs in the reactive distillation column, where the dichlorohydrin stream and the aqueous sodium hydroxide solution stream are fed separately at the top of the column. Steam is injected at the bottoms as stripping agent to recover the undissolved epichlorohydrin, forming an organic liquid phase, which is evaporated by a fraction of the heat of reaction. Optimal results were obtained at a packing height of 14.8 m. The solution of the reactive distillation column model, including water is shown in Figure 8. The presence of so high flow of water is due to the NaOH solution that is needed for the reaction, accompanying as solvent. It can be seen that there is not any distinction between liquid epichlorohydrin or countercurrent vapor epichlorohydrin. Thus, the molar flow of the epichlorohydrin inside the reactor grows due to the reaction and decrease of the reactants NaOH and both dichlorohydrins.


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Figure 8 – Evolution of the reactive distillation column model.

The evolution of the concentration of the aqueous liquid phase and the countercurrent vapor phase, inside the column packing, is shown in Figure 9.



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Figure 9 – Evolution of the components molar flows along the reactive distillation column.

In Figure 9, it can be seen the formation of epichlorohydrin and its dissolution into the aqueous phase until the solubility limit is reached. The dissolved epichlorohydrin achieves the solubility limit quickly, due to the high reaction rate, and then it remains almost constant, only increasing proportionally to the water production by the reaction. At this moment, the organic phase is formed and starts to evaporate, becoming part of the countercurrent stripping gas stream. The vapor phase is free of epichlorohydrin at the moment of its injection at the end of the package. The concentration of the epichlorohydrin in the countercurrent stream increases along the packing, while the epichlorohydrin in the organic phase is evaporated and stripped, getting the saturation at the top of the column. Thus, the epichlorohydrin-water azeotrope is obtained at the top of the column that at the vacuum operation conditions achieves a 22.8 w% of water. The heat generated in the main reaction, according to the energy balance, takes a value of 55.5 MJ/kg of epichlorohydrin produced. Out of it 410 kJ are consumed per kg of epichlorohydrin evaporated. Thus the energy to be removed from the system is 2.17 x 105 MJ/h, so the 2% of the reaction heat will evaporate the undissolved epichlorohydrin and join the


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countercurrent vapor stream. The excess will be indirectly refrigerated. The other substances, such as NaOH, NaCl, glycerol and DCHs do not reach their solubility limit along the reactor, so they remain in the aqueous phase and constitute the waste stream at the bottoms. The high amount of liquid water that flows inside the reactive column is not plotted in Figure 9, for clarity.

Following this process, with the quantity of the components fed indicated at the beginning of this section, we produce 26,540 t/yr of epichlorohydrin with at least 99.9 w% of purity, the commercial concentration of this product. The process yield reaches 80 %, related to the mass of pure glycerol fed to the plant, higher than the yield of the current industrial process set at 73 % related to the pure propylene fed, 28 and using a sustainable raw material. See the stream compositions in the supplementary material.

4.2.- Economic evaluation In this section, we estimate the investment costs of the designed plant, the payout time, the production costs, the total invested capital and the sales revenue. The epichlorohydrin production cost comprises the manufacturing costs and the management costs: -- The manufacturing costs are the expenses involved in the stages that process the raw materials into the main product. They include the raw materials cost, direct and indirect labor, utilities, supplies, maintenance, laboratory, equipment depreciation, insurance and taxes. Silla’s method

29

is used to compute

the cost of utilities, supplies, maintenance, laboratory, insurance and taxes cost items. In the case of labor costs, the salaries of the various employees are taken from the XVII Spanish Labor of Chemical Engineering Agreement (2013) and updated to 2014.30 The equipment depreciation is computed assuming straight-line depreciation 31 and the raw material cost is the sum of the products of each substance quantity and its price (DDP). The manufacturing costs add up to 24 M€/yr. --On the other hand, management costs are also estimated by Silla’s

29

method and add up to 10

M€/yr.



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Therefore, the epichlorohydrin production cost is estimated to be 36 M€/yr, and its breakdown is shown in Figure 10. This cost, according to the capacity of production of the plant, results in an epichlorohydrin production unit cost of 1,282 €/t. The current epichlorohydrin selling price in 2014 prices is 1,976.22 €/t, 32 at the exit of the plant (ExW).

Figure 10 – Production cost breakdown The total invested capital is also computed comprising the investment necessary to the construction of the plant and its start up, including equipment purchase and installation, ground acquisition and its adjustment and all related to the project development. The total invested capital includes tied-up capital and working capital: --The tied-up capital comprises the direct capital, building expenses, research, start up and other expenses items. As part of the direct capital, the equipment purchase, see Table 2 for the complete list, and installation cost are shown in Table 3. The design of the various units across the flowsheet such as tanks or heat exchangers is based on typical design rules in Walas 33 and Sinnot 34. The compressors are designed based on the power and the reactors based on the kinetics and/or residence times in the literature or a combination of both. For further details of the equipment units see Almena. 23 The equipment purchase is mainly obtained from Matche, 35 or else in Peters, Timmerhaus and West’s tool, 36 in which case the price is updated to 2014


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values. The installation factor used for each unit is fixed by Madrigal et al.37 In addition, the correlations used for estimating the cost of most of the units are presented in the supplementary material. Table 2. – Equipment design variables.

Unit Crude Glycerol Tank Epichlorohydrin Tank Wastes Tank M-01 M-02 M-03 M-04 R-01 Unit T-01 T-02 R-02 Unit Dec-01 Dec-02

Type Large Vertical API Large Vertical API Large Vertical API Static inline mixer Stirred Tank Stirred Tank Static inline mixer Semibatch (x4)

Type Packed column Packed column Packed column Type Horizontal cylinder Vertical tank

3

Material Carbon steel &API Carbon steel & API Carbon steel & API Stainless steel AISI 316 Stainless steel AISI 304 Stainless steel AISI 304 Stainless steel AISI 316 Stainless steel AISI 304

Material Carbon steel Stainless steel Stainless steel

Volume (m ) 4367 2655 8896 4.35 4.35 12.25 (x4)

Diameter (m) 0.90 2.70 0.90

Material Stainless steel AISI 304 Carbon steel

Height (m) 10.0 33.2 16.0

Diameter (m) 0.762 0.90

Length (m) 13.1 2.70 2

Unit I-01 I-02 I-03 I-04 I-05 I-06 I-07 I-08 I-09 I-10 I-11 CD-01 EV-01 CD-02 EV-02

Type Shell and tubes Shell and tubes Shell and tubes Double pipe Shell and tubes Double pipe Shell and tubes Shell and tubes Shell and tubes Shell and tubes Shell and tubes Shell and tubes Shell and tubes Shell and tubes Shell and tubes

Material Carbon steel Carbon steel Stainless steel AISI 316 (only tubes) Stainless steel AISI 316 Stainless steel AISI 304 (only tubes) Carbon steel Stainless steel AISI 304 (only tubes) Carbon steel Stainless steel AISI 304 (only tubes) Carbon steel Carbon steel Carbon steel Carbon steel Stainless steel AISI 316 (only tubes) Stainless steel AISI 316

Heat exchange area (m ) 12.7 13.7 25.8 3.9 62.7 5.3 17.1 72.7 370.4 27.9 94.3 25.4 46.0 264.1 108.0

Unit C-01 C-02 C-03 Exp-01

Type Centrifugal Centrifugal Centrifugal Radial expander

Material Stainless steel AISI 316 Stainless steel AISI 316 Carbon steel Carbon steel

Power (kW) 141.56 97.52 70.15 52.61 (Produced)

Unit S-01

Bed Type 4 A Zeolite

Package weight (kg) 220 (x2)

Material Carbon steel

Diameter (m) 0.60

Height (m) 2.40



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Meanwhile, piping (including valves and pumps) represents 45 % of equipment cost, instrumentation accounts for 20 %, isolation 7%, electric installation adds up to 15 % and utilities installation represent 40 % of equipment cost based on typical values.29, 33 In addition, the cost of the ground needed for the allocation of the project ascends to 7.25 M€. The sum of it corresponds to fixed capital. Besides, the honoraries must be considered, which represents 20 % of fixed capital. Direct capital is computed as the sum between the fixed capital and the honoraries. Other expenses and building expenses represent 20 % and 7 % of the direct capital respectively. Furthermore, fixed investments also include research, representing 12 % of them, and start up expenses, which suppose the 8 % of the fixed investment. As a result, the tied-up capital adds up to 55.6 M€. Table 3. – Equipment purchase and installation cost.

972,000 € 822,000 € 1,337,000 €

Installation Factor 1.7 1.8 1.8

Installation Cost 680,000 € 657,000 € 1,069,000 €

1,652,000 € 1,479,000 € 2,406,000 €

2,352,000 €

2.1

2,587,000 €

4,939,000 €

43,000 € 909,000 € 578,000 € 18,000 7,031,000 €

1.7 1.9 1.2 2.1 -

30,000 € 801,000 € 116,000 € 20,000 5,977,000 €

73,000 € 1,727,000 € 694,000 € 38.000 13,008,000 €

Equipment set

Purchase Cost

Storage Tanks Mixers Semibatch reactors Columns (distillation and reaction) Liquid-liquid separators Heat exchangers Compressors Molecular Sieves FINAL COST

Total Cost

--The working capital is the second kind of investment and makes it possible to operate the plant. It is calculated for an only one cycle of production, fixing one month as time base, and it includes the following items: raw materials, manufacturing materials, stored products, sales outstanding one month ahead and liquidity. The sum of all this items reaches 7.2 M€, as working capital needed. The total invested capital adds up to 63.7 M€. The breakdown for the invested capital is shown in Figure 11. 26 ACS Paragon Plus Environment

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Figure 11 – Total invested capital breakdown Next, we compute the sales revenue obtained by the operation of the plant. According to the plant production capacity of 26,540 t/yr of epichlorohydrin and the product selling price fixed in 1,976.22 €/t ExW, 32 the sales revenue ascends to 52.9 M€. In order to check the profitability of the epichlorohydrin plant, the economic evaluation results in a gross profit of nearly 18.5 M€, which represents a gross profit percentage of 29 %. Once deducted the corporate income tax, set at 30 % for the chemist industry in Spain, the net profit takes a value of 12.9 M€ (20 %). It means that the plant, expected to be working for an average life or 20 years, is profitable, with an invested capital recovery period of 5 years. The price of raw materials is volatile, above all of them that of glycerol. Therefore, it is interesting to see the effect of the variation of the prices of glycerol, HCl and NaOH in the production cost of epichlorohydrin. Figure 12 shows a parametric figure of the sensitivity analysis of the prices of the main raw materials on the production cost of epichlorohydrin. We cover a rage of 0.75-1.25 times the cost of each one. As expected, there is a linear relationship between the product costs and that of the raw materials from 1.16€/kg to 1.4€/kg. The unit cost of the product is more sensitive to the price of NaOH and HCl than to the price of crude glycerol.



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Figure 12.-Sensitivity analysis of the effect of the raw material prices on the production cost.

5.- Discussion and Conclusions

In this work a novel process for the production of epichlorohydrin has been proposed using glycerol as a sustainable raw material. The aim is to provide further added value to the main byproduct of the biodiesel industry by providing an alternative market for it. The process consists of four stages, glycerol purification, its reaction to mono and dichlorohydrins and their separation, epichlorohydrin production and final product purification. We use a hybrid modularequation based modeling approach to simulate the operation of such process integrating MATLAB and CHEMCAD.


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It was found that the proposed process has yield 10% higher than the current industrial process (80.5% vs. 73 % related to the mass of pure glycerol fed to the plant) using a sustainable raw material. The process consumes 41.5 kt/yr of glycerol, 29 t/yr of HCl and 15.9 kt/yr of NaOH to produce 26.5 kt/y of epichlorohydrin at 99.9 w%. The investment cost adds up to 63.7 M€ with a production cost of 1.28 €/kg. As a result the payout time is 5 years, while the plant is projected to operate continuously during an average life of 20 years.

Nomenclature

A: Cross sectional area (m2). C: Concentration (mol/cm3). Ea: Activation energy (kJ/mol). F: Molar flow (kmol/s). Fe: Molar inflow (kmol/s) Fs: Molar outflow (kmol/s) k: Arrhenius constant. K: Velocity constant. Units dependant on the reaction N: Number of moles in reactor (kmol). r: reaction rate. V: Liquid volume (dm3). ε: Porosity. εL: Liquid hold up parameter. θP: Parameter (Dimensionless). Subscripts: MCH: Monochlorohydrin. DCH: Dichlorohydrin. Gly: Glycerol. EH: Epichlorohydrin. e: Inlet M€: Millions of Euros. s: Outlet o: Initial *S Supporting Information Detailed correlations for units cost estimation and the stream tables for the entire flowsheet are also provided. This material is available free of charge via the Internet at http://pubs.acs.org.



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References

(1) Martín, M.; Grossmann, I.E. Optimal Simultaneous Production of Hydrogen and Liquid Fuels from Glycerol: Integrating the Use of Biodiesel Byproducts. Ind. Eng. Chem. Res. 2014, 53 (18), 7730–774. (2) Martín, M.; Grossmann, I.E. Design of an optimal process for enhanced production of bioethanol and biodiesel from algae oil via glycerol fermentation. Applied Energy. 2014, 135, 108–114. (3) Martín, M.; Grossmann, I.E. ASI: Toward the Optimal Integrated Production of Biodiesel with Internal Recycling of Methanol Produced from Glycerol. Environmental Progress & Sustainable Energy. 2013, 32 (4), 791–801. (4) Martín, M.; Grossmann, I.E. Simultaneous dynamic optimization and heat integration for the coproduction of diesel substitutes: Biodiesel (FAME & FAEE) and glycerol ethers from algae oil. Ind. Eng. Chem. Res. 2014, 53 (28), 11371–11383. (5) Bueno, L.; Toro, C.A.; Martín, M. Techno-economic evaluation of the production of added value polymers from glycerol. Chem Eng. Res. Des. 2014, 93, 432-440. (6) Weissermel, K.; Arpe, H.-J. Industrial Organic Chemistry. 3rd Edition. Weinheim, Germany: Wiley-VCH. 1997; ISBN: 978-3-527-61459-2. 294-299. (7) Krzyzanowska, A; Milchert, E. Apparatus and methods of dehydrochlorination of 1,3-dichloropropane-2ol, reaction times in the pre-reactor and reactor. Chemik. 2013, 67 (9), 779-786. (8) Bell, B. M.; Briggs, J.R.; Campbell, R.M.; Chambers, S.M.; Gaarenstroom, P.D.; Hippler, J.G.; Hook, B.D.; Kearns, K.; Kenney, J.M.; Kruper, W.J.; Schreck, D.J.; Theriault, C.N.; Wolfe, C.P. Glycerin as a Renewable Feedstock for Epichlorohydrin Production: The GTE Process. Clean. 2008, 36 (8), 657–661. (9) Bijsterbosch, J.W.; Das, A.; Kerkhof, F.P.J.M. Clean technology in the production of epichlorohydrin. J. Cleaner Prod. 1994, 2 (3-4), 181-184. (10) Zhang, Y.; Dube, M. A.; McLean, D. D.; Kates, M. Biodiesel production from waste cooking oil: 2. Economic assessment and sensitivity analysis. Bioresour. Technol. 2003, 90, 229. (11) Dow Global Technologies, 2013. Process and apparatus for efficient recovery of dichlorohydrins. United States, Patent Application Publication. Inventor: Mehta, A. US 2013/0259766 A1. 03/10/2013. (12) Dow Global Technologies, 2012. Process and apparatus for producing and purifying epichlorohydrins. France, European Patent Application. Inventors: Kneupper, C. D.; Basile, P.S.; Fan, W.W.; Noorman, S. EP 2 537 837 A1. 26/12/2012. (13) Blasher, H.U.; Schmidt, E. Asymmetric Catalysis on Industrial Scale: Challenges, Approaches and Solutions. Weinheim, Germany: Wiley-VCH. 2004; ISBN: 3-527-30631-5. p.197 (14) McKetta, J.J. Encyclopedia of Chemical Processing and Design. Volume 19: Energy, Costing Thermal Electric Power Plants to Ethanol. New York-Basel: Marcel Dekker Inc. 1983 ISBN:0-8247-2469-0 (15) JP Dyechem. Importers and distributors: Chemicals, Solvents and Indl. Raw Materials. Epichlorohydrin Specification. 2015. Last accessed 16/10/2015. http://www.jpdyechem.com/pdf/epichlorohydrin.pdf 30 ACS Paragon Plus Environment

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(16) Alibaba.com.Global trades. Zeolite 4A molecular sieve beads supplier. 2014. Last accesed 16/10/2014. http://www.alibaba.com/product-detail/zeolite-molecular-sieve-4A-beadssupplier_1826881433.html?spm=a2700.7724838.30.9.xhd6m2&s=p (17) Tesser, R.; Santacesaria, E.; Di Serio, M.; Di Nuzzi, G.; Fiandra, V. Kinetics of Glycerol Chlorination with Hydrochloric Acid: A New Route to α,γ-Dichlorohydrin. Ind. Eng. Chem. Res. 2007, 46, 6456–6465. (18) Santacesaria, E.; Tesser, R.; Di Serio, M.; Casale, L.; Verde, D. New Process for Producing Epichlorohydrin via Glycerol Chlorination. Ind. Eng. Chem. Res. 2010, 49, 964–970. (19) Santacesaria, E.; Vitiello, R.; Tesser, R.; Russo, V.; Turco, R.; Di Serio M.. 2014 Chemical and Technical Aspects of the Synthesis of Chlorohydrins from Glycerol Ind. Eng. Chem. Res. 2014, 53, 8939– 8962. (20) Herliati, R.Y.; Intan, A.S.; Abidin, Z.Z. Preliminary Design of Semi-Batch Reactor for Synthesis 1,3Dichloro-2-Propanol Using Aspen Plus. International Journal of Chemistry. 2011, 3 (1), 196–201. (21) Ma, L.; Zhu, J. W.; Yuan, X. Q.; Yue, Q. Synthesis of Epichlorohydrin from Dichloropropanols: Kinetic Aspects of the Process. Chemical Engineering Research and Design. 2007, 85 (A12), 1580–1585. (22) RASCHIG. Ring Division, Porosity of a random packing of 3.5 inches metal pall rings. 2014. Last accessed 09/12/2014. http://raschig.de/Pall-Ring-Metal. (23) Almena, A. Planta de producción de epiclorhidrina a partir de glicerol. MEng Thesis. University of Salamanca (Spain). 2015. (24) Yang, F.; Hanna, M. A.; Sun, R. Value-added uses for crude glycerol, a byproduct of biodiesel production. Biotechnology for Biofuels. 2012, 5:13. (25) Thompson, J.C.; He B.B. Characterization of Crude Glycerol from Biodiesel Production from Multiple Feedstocks. Applied Engineering in Agriculture. 2006, 22 (2), 261–265. (26) Quiminet. Products and specialized services information. Anhydrous hydrogen chloride market price. 2010. Last accessed 04/04/2014. http://www.quiminet.com/productos/cloruro-de-hidrogeno-anhidro-27860706000/precios.htm (27) Alibaba.com. Global trades. Sodium hydroxide market price. 2014. Last accesed 26/03/2014. http://www.alibaba.com/product-detail/2014-Hot-sale-caustic-soda-flakes_1768788307.html?s=p (28) Spolek Pro Chemickou A Hutni Vyrobu . Method of preparing dichloropropanols from glycerin. Czech Republic, International Patent. Inventors: Kubicek, P.; Sladek, P.; Buricova, I. WO 2005/021476. 30/01/2012. (29) Silla, H. Chemical process engineering. Design and economics. New York, United States: Marcel Dekker Inc. 2003; ISBN: 0-8247-4274-5. (30) State official Bulleting (2014). http://www.boe.es/boe/dias/2014/03/20/pdfs/BOE-A-2014-2971.pdf (31) Peters, M.S.; Timmerhaus, K.D. Plant design and economics for chemical engineers. 5th Edition. Singapore: Mc Graw-Hill. 2003; ISBN: 0-07-119872-5



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Figure 1.- Hydrochlorination of glycerol reaction. 90x35mm (300 x 300 DPI)



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Figure 2.- Dehydrochlorination of dichlorohydrins reaction. 90x10mm (300 x 300 DPI)


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Figure 3 – Process flow chart. 199x111mm (300 x 300 DPI)



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Figure 4 – Dehydrochlorination of glycerol scheme. 90x46mm (300 x 300 DPI)


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Figure 5 – Dehydrochlorination of glycerol simplified scheme. 90x23mm (300 x 300 DPI)



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Figure 6 – Mass balance in a differential volume of reactor R-02. 90x76mm (300 x 300 DPI)


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Figure 7 – Evolution in time of the composition in semibatch reactor. 149x123mm (300 x 300 DPI)



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Figure 8 – Evolution of the reactive distillation column model. 192x145mm (300 x 300 DPI)


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Figure 9 – Evolution of the components molar flows along the reactive distillation column. 149x97mm (300 x 300 DPI)



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Figure 10 – Production cost breakdown 90x74mm (300 x 300 DPI)


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Figure 11 – Total invested capital breakdown 90x78mm (300 x 300 DPI)



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Figure 12.-Sensitivity analysis of the effect of the raw material prices on the production cost. 199x159mm (300 x 300 DPI)


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