Production of Epichlorohydrin

Manufacturing of Epichlorohydrin

Production via GTE Process

Group 1

Alejandro Lopez Perez

Dragana Stojanovic

Kostantinos Papanikolaou

Priyanka Vaiude

Eindhoven 27/3/2015

GTE Process Design

List of Acronyms

EPH Epichlorohydrin

GLY Glycerol

HCl Hydrogen chloride

GTE Glycerol to epichlorohydrin

HOCl Hypochlorous acid

1,3-DCP 1,3-dichloro propan-2-ol

2,3-DCP 2,3-dichloro propan-1-ol

NaOH Sodium hydroxide

a-MCH a-Monochlohydrin

b-MCH b-Monochlorohydrin

GLUA Glutamic acid

DCP Dichloropropanols

PFD Process flow diagram

HPS High pressure steam

LPS Low pressure steam

MPS Medium Pressure steam

AA Adipic acid

CAPEX Capital expenditure

OPEX Operating expenditure

NaCl Sodium chloride

CSTR Continuous Stirred Tank Reactor

*DCP’s are referred as dichlorohydrins in the text in some instances, expressing the same compounds (1,3-DCP, 2,3-DCP).

GTE Process Design

Nomenclature

Units

r1 EPH formation reaction rate mol/m3 s

R Gas constant J/mol K

k1 Reaction rate constant of EPH formation mol s/m3

[OH-] Concentration of hydroxide ions mol/m3

[EPH] Concentration of EPH mol/m3

[DCP] Concentration of DCP mol/m3

T Temperature K

k2 Hydrolysis reaction rate constant mol s/m3

r2 Reaction rate of hydrolysis mol/m3 s

SGLY/1,3-DCP Selectivity of GLY toward 1,3-DCP -

Cp Specific heat capacity kJ/kg K

F Molar flow kmol/s

FCp Flow heat capacity kW/K

Qinterval Heat available in a temperature interval kW

ΔTinterval Temperature difference of interval K

NHE,AP Number of heat exchange units above pinch -

NHE,BP Number of heat exchange units below pinch -

NS,BP Number of streams below pinch -

NS,AP Number of streams above pinch -

ΔTCW Temperature difference of cooling water K

Δhvap Latent heat kJ/kg

Qutility Heat provided by utility kW

Cpo Free on board cost of equipment $

CBM Installed equipment cost $

FBM Design correction factor -

Fm Material correction factor -

FP Pressure correction factor -

M&S Marshall and Swift indices -

Executive Summary

GTE Process Design

Epichlorohydrin is a valuable fine chemical, mainly dedicated to the manufacturing of epoxy resins.

Currently, several routes are available for the production of this chemical with the predominant one

to be the allylic chlorination of propylene, referred also as conventional process route in this text.

The extensive formation of undesired chlorinated organics from this process that are difficult to be

disposed of and the escalating cost of petrochemical raw materials such as propylene postulate the

investigation of alternative routes for epichlorohydrin manufacturing.

The above disadvantages of the conventional process in combination with the growing availability of

glycerol, as consequence of the increase of biodiesel production have played decisive role in the

rapid development of the glycerol to epichlorohydrin route (known as GTE process) which was

historically prevented due to the high cost of glycerol.

GTE route is divided into two steps; chlorination of glycerol is the first and dehydrochlorination of

dichloropropanols the second one. In the present study both steps have been examined thoroughly

and the entire process has been simulated on Aspen Plus V8.6 software.

Two reactor configurations have been proposed for the first step, utilising adipic acid as a liquid

catalyst, whilst a reactive distillation column has been designed for the second step accompanied by

a separation train for production of almost 99% pure final product.

Sizing of process equipment and economic evaluation of the process have been performed, revealing

strong potential of GTE route compared to the conventional technology. China has been chosen as

the location of the plant and the payback period has been estimated to endure 3 years for 1800 $/ton

selling price of epichlorohydrin. Heat integration by Pinch point analysis showed 0.167 M$/year

savings from utilities and finally a preliminary control scheme has been suggested for the process.

GTE Process Design

Design Considerations

The basis of design considered in the current project is specified in this section. The topics

considered in the Basis of Design are the plant capacity, plant location, composition and prices of

materials (i.e. raw material), physical and chemical properties of the substances involved in the

process and storage information about some of these materials.

Plant capacity

The capacity of the GTE process in order to produce epichlorohydrin (EPH) is given to be 100

kton/year.

Plant location

The location of the glycerol (GLY) to EPH plant is assumed to be in China.

Compositions and prices

The raw materials used in this process are GLY, hydrochloride (HCl) gas and sodium hydroxide

(NaOH) solution. The GLY is assumed to be obtained from a nearby biodiesel production plant. The

HCl gas is assumed to be obtained from a plant nearby producing HCl gas as one of the byproducts

of that plant and is assumed to be fed to the plant via pipelines. The raw material composition and

purchase price is listed in Table 1.

The scope of the project claims that the feedstock composition is 99.9% glycerol with 0.1%

impurities that may be present due to processing parameters involved in the previous plant. The

following table contains the classification of the most important substances involved in the process

and their purity and price.

Table 1 Composition and price of the main raw materials present in the GTE process

Name Formula Purity wt.% Purchase price

$/ton

Glycerol C3H8O3 99.9 800

Hydrogen

chloride gas

HCl 100 360

Sodium

hydroxide

NaOH 99.99 120

GTE Process Design

Table of Contents1. INTRODUCTION........................................................................................................................................2

2. Literature Review..........................................................................................................................................5

2.1. First Step –Chlorination of Glycerol.....................................................................................................8

2.2. Second Step –Dehydrochlorination of DCP.......................................................................................13

2.3. General Process Considerations and Operation Conditions...............................................................14

2.4. Literature Review Conclusions...........................................................................................................15

3. Aspen Simulation of GTE process..............................................................................................................17

3.1. Chlorination of GLY Reactor Selection.............................................................................................18

3.2. Dehydrochlorination of DCP..............................................................................................................31

4. Equipment sizing and cost..........................................................................................................................37

4.1. Reactors and columns.........................................................................................................................37

4.2. Pumps and Compressors.....................................................................................................................38

4.3. Heat exchangers and decanters...........................................................................................................39

5. Economic evaluation...................................................................................................................................41

5.1. ASPEN Economics Input....................................................................................................................41

5.2. Results.................................................................................................................................................42

6. Conclusions and Recommendations...........................................................................................................46

Appendix.............................................................................................................................................................46

Heat Integration of GTE Process-Pinch Point Analysis.....................................................................................49

GTE Heat Exchanger Network Design above Pinch Point.............................................................................52

GTE Heat Exchanger Network Design below Pinch Point............................................................................54

Annual Savings Estimation.............................................................................................................................56

Control Scheme of GTE Process........................................................................................................................58

Control of Reactors.........................................................................................................................................58

Control of Distillation columns......................................................................................................................59

Heat Exchangers and Decanters......................................................................................................................59

References...........................................................................................................................................................63

GTE Process Design

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GTE Process Design

Introduction

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1. INTRODUCTION

Several routes are known to manufacture EPH, most is made in a two-step process from allyl

chloride and hypochlorous acid procuring a mixture of two chlorinated alcohols which can be

converted into EPH by treating with a base sodium hydroxide.

In the different reactions occurring in the above process, large amount of undesired organic

compounds which are very expensive to be discarded resulting to a high required selling price. In

combination with the high price of propylene, which is used a raw material, this route of EPH

production becomes less attractive leading to a need for alternative routes.

Owing to the disadvantages mentioned above, other routes have been investigated and GTE is

highlighted as the most promising technology representing an economically and environmentally

advantageous process. The commercial development of the process was obstructed until recently,

because of the high cost of glycerol. The recent advances in the technology related to biofuels have

caused the GLY price to drop dramatically since it can be obtained as a biodiesel by-product. This

new development and subsequent reduction in prices of GLY has indicated a great potential for the

feasibility of GTE process.

In 2007 Solvay, a traditional GLY and EPH manufacturer was the first to start GTE process.

Previously in the early 2000s the company was producing GLY from EPH but after the glycerol

price drop, it reversed its process to produce EPH from GLY. The picture below taken from the

company website represents this reversal trend in the process of production of EPH and glycerol.

Figure 1.1 Trend reversal in the processes concerning Glycerol and EPH from Solvay.

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Solvay has acquired many patents in this regard and the innovative process is registered as

Epicerol®, characterized as the most important patented process for the production EPH. The

company claims that this process is very environment friendly and reduces the carbon footprint by

60% as compared to the conventional route. GTE process of EPH production has been patented by

other well-known chemical industries like DOW chemicals.

The aim of this project is the evaluation of technical as well as economic feasibility of GTE process

and comprehensive comparison with the conventional process. To that end, GTE process has to be

thoroughly designed with subsequent estimation of Capital Expenditure (CAPEX) and Operating

Expenditure (OPEX). ASPEN PLUS (version 8.2) will be used for that purpose and the efficiency of

GTE will be established.

3

Literature Review

4

2. Literature Review

EPH is a liquid epoxide being used as an intermediate chemical for the production of epoxy resins

for coatings paintings and electronic circuits but also for non-epoxy applications such as paper

chemicals, water treatment and health care products [1]. China is highlighted as the leading market in

EPH production. Almost 97 % of the annual production is being consumed in the production of

epoxy resins and ca. 5% increase in annual production rate is predicted until 2018 [2].

Many different routes for EPH manufacture have been suggested in the literature; however the

predominant one in industry is the allylic chlorination of propylene to allyl chloride, starting from

propylene and chlorine as primary raw materials in a multi-step process as indicating in Figure 2.1.

Figure 2.1 The conventional route to EPH production in a multi-step process [3].

The first reaction is the allylic chlorination of propylene to allyl chloride. In parallel hypochlorous

acid (HOCl) being produced via dissolving of chlorine into water and subsequently reacts with the

produced allyl chloride from the first step to yield a mixture of 1,3-dichloropropan-2-ol (1,3-DCP)

and 2,3-dichloropropan-1-ol (2,3-DCP). The last step of the process includes the reaction between

dichloropropanols (DCP) and a base (e.g. NaOH or Ca(OH)2) for the formation of the final product.

The described process can yield EPH of very high purity but suffers from numerous undesirable

features such as very low chlorine atom efficiency (i.e. only one of the four chlorine atoms

participating in the reaction is retained in the product molecule), significant inefficiencies in the

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chlorination and hypochlorination steps, resulting to the formation of unwanted chlorinated organics

(ca. 0.5t/tEPH) and finally the continuous increase in the cost of petrochemical raw material like

propylene. [4]. In order these problems to be addressed different routes were examined in the past

based on less expensive raw materials. One such a route relies on the conversion of GLY through

DCP to EPH, known as glycerine to epichlorohydrin process (GTE). The high cost of GLY

prevented the development and the consideration of this process previously. This situation has been

changed recently since GLY can be obtained as by-product of biodiesel production (ca. 0.1 tn/tn

Biodiesel), and investigation on this process being conducted extensively [5,6]. Figure 2.2 shows the

rise in biodiesel production from 2000 to 2008, implying the potential of GTE process in the near

future.

Figure 2.2 World biodiesel production rate 2000-2008 [2]

The GTE is a two-step process as it can be seen from Figure 2.3. In the first step (or first reaction)

chlorination of GLY is taking place by reaction with HCl in the presence of liquid catalyst providing

a rich stream to DCP. This stream subsequently is driven the next unit to react with a base for EPH

formation.

6

Figure 2.3. EPH production from GTE process [3].

Apart from the reaction steps also separation of the final product is included in the process. Thus a

block-flow diagram, such the one presented in Figure 2.4, can give insight into the different steps of

GTE process presenting the initial idea before the execution of the design part. Literature details and

findings on the different steps are discussed in the following sections.

Figure 2.4 GTE process block flow diagram.

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2.1. First Step –Chlorination of Glycerol

The first step of the process is the chlorination of GLY as mentioned previously, where four

reactions in series and parallel (i.e. complex reactions) are taking place. The reaction of GLY with

HCl yields a mixture of a-Monochlorohydrin (a-MCH) and 2-Monochlorohydrin (b-MCH) as

indicated in Figure 2.5. Subsequently a-MCH reacts with HCl to produce the final product of the

first step, i.e. DCP. All the reactions involved in reaction network are reversible but kinetics studies

have shown negligible kinetic constants of the reverse reactions, thus they can be considered as

irreversible.

Figure 2.5 Chlorination of GLY reaction network.

Before the design of GLY chlorination system several aspects have to be considered such as:

Catalyst Selection

Catalyst Concentration

HCl state

Operating conditions

Reaction kinetics

2.1.1. Catalyst Selection and Concentration

Catalyst selection and concentration are very important aspects as the selectivity to the desired

product and GLY conversion are strongly influence by them. The suitability of the catalyst can be

judged from three factors, namely activity, selectivity and low volatility. Carboxylic acids are

commonly utilised for the reaction, with acetic acid to be the usual choice owing to its very high

activity and selectivity towards 1,3-DCP; however its high volatility (ca. B.P. 117 oC) renders it as

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inappropriate choice for large scale production due to significant losses at the reaction temperature

(ca. 120 oC).

E. Santacesaria [4] tested different carboxylic acids by using an apparatus operating in continuous

mode for HCl and batch for GLY in an effort to define catalysts with similar performance as acetic

acid but significantly lower volatilities. In this study the crucial role of pKa value in relation to the

performance of the catalyst was fortified. Specifically catalysts with pKa greater than 4 showed high

selectivity to DCP, while those with less than 1.2 demonstrated high selectivity to a-MCH and b-

MCH. The experiments were conducted at 100 oC and 5.5 bar HCl pressure and the downstream

composition was measured after 3 h. Adipic acid (AA) found to be very selective to 1,3-DCP, with

molar concentration of 72.23 % mol and complete conversion of GLY. This performance in

combination with high boiling point of AA (337.5 oC) highlights it as a promising catalyst.

R. Vitello et. Al [5] performed experiments with two series of catalysts, specifically glycolic acid

series and amminoacid series. The runs conducted at 100 oC for different pressures of gaseous HCl

and different concentration of catalysts. Glutamic acid (GLUA) found to be the best one in terms of

performance and further investigation followed. Higher concentration of catalyst proved to be very

beneficial regarding the selectivity to the desired product as seen in Figure 2.6, as from 2% catalyst

loading to 8%, a rise of 40% to selectivity towards 1,3-DCP is noted.

Figure 2.6 Effect of catalyst loading in selectivity for GLUA catalyst at P=4.5bar and T=100 oC [5]

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Propionic acid has been discussed as another alternative [6] but with no good perspective because of

its slightly lower volatility than AA.

2.1.2. Hydrogen Chloride State and Reaction Conditions

The chlorination agent can be introduced in the reactor as a solution of hydrochloric acid with water

or as a gaseous hydrogen chloride. The big disadvantage of the first case is the presence of large

amount of water in the reactor and therefore larger reactor volumes because of slower kinetics.

Dimitriev et.al [7] investigated the effect of water concentration in regards to a-MCP formation at a

range of temperature between 80-117 oC. As it can be concluded from Figure 2.7 increase in water

concentration implies significant decline to the reaction rate constant and consequently slow kinetics.

Figure 2.7 Effect of water concentration to the reaction rate constant of a-MCP production [7].

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Apparently, hydration degree strongly affects the reaction rates, resulting to smaller amount of

products formed and consequently to large reaction volumes. This problem can be tackled by feeding

gaseous HCl in the reactor. The feed pressure of HCl is a critical parameter as influences strongly the

conversion of GLY but also the selectivity toward 1,3-DCH which can react 20 times faster than 2,3-

DCP for EPH production.

Bruce M. et. al [3] demonstrated the effect of HCl pressure on product molar composition by

utilising 2 wt% acetic acid catalyst. Three different values of pressure were tested, 20, 30 and 50 psi.

The formation of 1,3-DCP in the case of 20 psi was observed to be negligible, while for 50 psi it has

the highest concentration in the reactor outlet. This behaviour can be explained by the reaction

nature. The reaction seems to be equilibrium limited at low HCl pressures resulting to very low

conversions and after 1 h run the concentrations of a-MCP and GLY reach a limit and no change can

be observed afterwards. Additionally, in the same study proved that the absorption of HCl to GLY is

higher working at higher pressures, indicating that higher conversion of GLY can be achieved.

Figure 2.8 demonstrates product evolution for different HCl pressures [3].

Similar results presented by R. Tesser et. al. [8] where monochloro-acetic acid catalyst tested under

2, 5.5 and 9 bar pressures. At 9 bar pressure complete conversion of GLY achieved after almost 3 hr

of experiment and 30 % rise of selectivity appeared after 225 min as shown in Figure 2.9.

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Figure 2.9 Selectivity to dichlorohydrins (i.e. DCP) as a function of HCl partial pressure [8].

S.H. Lee et al. [9] highlighted the importance of proper mixing of the GLY and HCl for good yields

and selectivity. Experiments were conducted in a batch reactor revealing that at low stirrer speed

chlorination of GLY is mass transfer limited but above 600 rpm HCl is dissolved effectively in the

liquid phase providing higher yields of 1,3-DCP. This a critical observation that should be taken into

account for the reactor design and proper mixing will be required.

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Figure 2.10 DCP yield as function of stirrer rotation speed [9].

2.1.3. Reaction Kinetics

The reaction kinetics were retrieved from the literature [10]. AA referred as promising catalyst for

industrial production and the kinetics by utilising this catalyst are presented in Table 2.1 for

temperature equal to 120 oC. In this a study a model developed for describing kinetics of DCP

production assuming 1st order kinetics for all the reactions and found quiet reliable compared to the

experimental results. These kinetic values were used later from us for simulation of the process.

Table 2.1 Reaction kinetic values of chlorination of GLY at 120 oC [10].

Reaction Rate Constant k (min-1) Activation Energy Ea(kJ/mol)

1 2.56 30.7

2 9.07 41.8

3 5.03 29.4

4 11.37 45.9

2.2. Second Step –Dehydrochlorination of DCP

In the second step of the process dehydrochlorination of DCP occurs for the production of the final

product EPH. The base to be used is defined from the formulation of the problem as sodium

hydroxide (NaOH). The reaction taking place are the main one but also in parallel hydrolysis of DCP

occurs resulting to the consumption of value product as illustrated from Reaction 1 and 2.

1. C3 H 6 Cl2O+NaOH yields→

C3 H 5 ClO+NaCl+H 2 O

2. C3 H 5 ClO+NaOH+H 2 O yields→

C3 H 8O3+NaCl

Apparently, to avoid undesired consumption of the valuable product from hydrolysis reaction it

should be removed promptly. This can be achieved by performing the reaction in a reactive

distillation column with instantaneous removal of EPH. Reactions conditions are relatively mild at

around 60 oC and atmospheric pressure [11] and the optimal ration of base to DCP is defined at

0.89:1 obtaining 97 % selectivity to EPH and 88 % of DCP [12].

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MA et.al [13] showed that both reactions can be well represented by 2nd order kinetics using titration

technique as seen from Equations (2.3) and (2.4). These kinetic parameters were used from us for

the simulation of the reaction unit afterwards.

For temperature in the range: 313 - 333 K

Main Reaction

r1=8.97∗1020e−123200

RT [ DCP ]¿ (2.3)

Side Reaction

r2=5.66∗1010 e−70790

RT [ EPH ]¿ (2.4)

Where [OH-] the concentration of hydroxide ions and [EPH] the one for EPC.

2.3. General Process Considerations and Operation Conditions

Production of Dichlorohydrins :

In this process the catalyst, glycerol and HCL are fed to a reactor where the chlorohydrination

reaction takes place. The products are subjected to distillation to remove impurities like unreacted

raw materials and undesired products. The top product from this operation is subjected to decantation

where the DCP’s are separated from water [15].

Operating conditions of DCP production

Reaction temperature: 120 ⁰C

Reaction pressure: 5 bar

Separation (distillation) temperature: 130 - 195⁰C

Separation (distillation) pressure: 1 bar

Condenser temperature: 25 ⁰C

Residence time: 7-11 hours

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Production of EPH:

The EPH is manufactured in two steps: the first step is reaction and extraction followed by

distillation separation train. The distillation operation can be in combination with adsorption, but this

is not the most preferred option. The reactor has to be equipped with stream stripping to remove the

EPH from the reaction environment as soon as it is formed to avoid the side reaction.

Operating conditions of EPH production [16]:

REACTORS

Excess of DCP as compared to NaOH: 0.89 effective equivalent (in order to reduce EPH

degradation reactions especially hydrolysis)

Reaction temperature: 60-90 ⁰C

Pressure range: 1– 1.5 bar

Reactor type: Reactive distillation column

Residence time: 7 – 10 min

2.4. Literature Review Conclusions

Regarding to the first step of the process (chlorination of GLY), the catalyst can strongly influence

the efficiency of the reaction. A catalyst is characterised as successful when is active, selective to

1,3-DCP and non-volatile. According to the above criteria AA seems to be the optimal choice at a

concentration around 8 % mol. HCl should be fed as a gas inside the reactor to avoid accumulation

of water and therefore slow kinetics. Water is also produced as a product of chlorination of GLY,

thus a combination of reaction and intermediate separation might be a promising solution for water

removal and lower reactor volume. The pressure of HCl and mixing intensity are highlighted as very

important parameters in relation to yield and selectivity toward DCP. HCl at 5 bar pressure can

provide very good results, whilst good mixing should be provided to overcome mass transfer

limitations. Upon the dehydrochlorination of DCP, EPH produced should be removed instantly from

the reaction system, and to that end a reactive distillation column has to be utilised. The optimum

ratio between the base and DCP was found to be 0.89:1 giving high selectivity on EPH and

conversion of DCP. Finally the kinetics of both reactions were found in the literature and will be

used for the simulation of the process on Aspen Plus.

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GTE Aspen Simulation

16

3. Aspen Simulation of GTE process

In the present chapter the simulation and conceptual design of GTE process by Aspen Plus software

are presented. As it can be seen from Figure 3.1 GTE process is composed of four different steps.

The first step is chlorination of GLY, where HCl, AA and GLY are fed to the first reactor producing

mainly DCP’s and water, but also a-MCH and b-MCH. In the next step water with

monochlorohydrins are removed giving a stream rich in DCP’s which react with NaOH in the

dehydrochlorination unit to produce EPH and water. At the same time a solution of sodium chloride,

one of the products of the reaction, is removed from the bottom of the unit. Finally the separation of

EPH from water is performed in the purification unit. For the simulation Electrochemical NRTL has

chosen as base method due to the presence of electrolytes in the process, such as HCl and NaOH.

Figure 3.1 Block diagram of the GTE process

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3.1. Chlorination of GLY Reactor Selection

The reaction is taking place in liquid phase and HCl decided to be fed in gas phase. As pointed out in

the previous chapter effective mixing is required in order to increase the yield and selectivity to

DCP. The most commonly used reactors on bench scale apparatus are either batch or semi-batch;

however these types of reactors are inappropriate for large scale production and particularly when the

annual capacity of the product is 100 kt. In industrial scale the previous types can be replaced by

continuous stirred tank reactors (CSTR’s), retaining the capability of intense mixing in a continuous

mode. Initially, the reaction system was simulated on Aspen only with one CSTR but an extremely

large reactor volume was needed , ca. 360 m3, in order to obtain the desired conversion of GLY and

at the same time the necessary mole flow of DCP for 100 kt/yr production of EPH. This result most

likely is attributed to the fact that CSTR operates at the same operating conditions (i.e. temperature

and concentration) as the exit stream. To reduce the reactor volume, CSTR’s in series were simulated

in order to approximate plug flow behaviour but at the same time retaining perfect mixing. The

following configurations are based on the same principle, but lower volumes for the same conversion

can be achieved in the second case making them the preferred option.

3.1.1. First Reactor Configuration Proposal-CSTR Model

The process flow diagram of the first proposed configuration is seen in Figure 3.2. 285 kmol/hr of

gaseous HCl at Stream 1 are driven to the compressor C-1, where they compressed at 5 bar and

afterwards the stream is moving towards HE-1 where it is cooled down to 90 oC. The stream is

splitted by S-1 to four streams, HCl-1, HCl-2, HCl-3 and HCl-4 for distribution into the first four

reactor units. Stream 2 contains pure GLY in a mole flow rate of 140 kmol/hr. After heating at 90 oC

by HE-2 the stream is pumped toward S-2. After splitting into two streams, GLY-1 and GLY-2,

GLY is fed to the first two reactor units follows CSTR-1 and CSTR-2.

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Figure 3.2 PFD of GLY chlorination step.

The first reactor block is focused on the production of monochlorohydrins (a-MCH, β-MCH), whilst

the second block is dedicated to the production of DCP (1,3-DCP and 2,3-DCP). This can be better

understood by the definition of selectivity as given by Equation 3.1 and the kinetics of the reaction

[17].

The production of a-MCH and 1,3-DCH are 1st order reactions as mentioned before and they occur in

series. Thus by formulating the selectivity of GLY toward 1,3-DCH becomes clear that for high

concentration of GLY high selectivity to a-MCH can be achieved and this occurs in the first reaction

block where the entire stream of GLY is distributed. After the first reaction unit high concentration

of a-MCH has been obtained and therefore the process is primarily selective to 1,3-DCH in the

second reaction unit. Feed of GLY is interupted after CSTR-2 so as to speed up the kinetics of 1,3-

DCH formation and at the last unit the feed of HCl is stopped where the remaining a-MCH reacts

with the unreacted HCl. The distribution of HCl is targeting to complete consumption of GLY in the

first block, preventing 1,3-DCH formation.

Water removal has been pointed out as very beneficial for the kinetics of GLY chlorination and to

that end DC-1 unit is present between the two reaction blocks where 177 kmol/hr out 186 kmol/hr of

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water are removed from the top stream, while the bottom stream composed of MCP, DCP and GLY

moves to the next reaction unit (CSTR-3). This amount of water removal can be achieved with a

column of 25 bubble cap trays and reflux ratio around 1.3. The downstream of the reactor is

composed of large amount of DCP, around 126 kmol/hr and therefore the yield of GLY to DCP is

estimated around 90 %. Also water (ca. 77 kmol/hr) and small amount of monochlorohydrins (ca.12

kmol/hr) are present, with negligible amounts of GLY and HCl. Table 3.1 and Table 3.2 show the

moleflow of HCl and GLY after S-1 and S-2 respectively.

Table 3.1 Stream 3 (HCl) splitting.

Stream 3 HCl-1 HCl-2 HCl-3 HCl4

Molar Flow (kmol/hr) 285 107.99 107.99 53.9 15.02

Table 3.2 Stream 6 (GLY) splitting.

Stream 6 GLY-1 GLY-2

Molar Flow (kmol/hr) 140 70 70

The volumes of the reaction units with the corresponding dimensions by assuming cylindrical vessel

geometries are tabulated in Table 3.3.

Table 3.3 Reaction units diamensions.

Unit Volume (m3) Diameter (m) Height (m)

CSTR-1 31.5 3 4.4

CSTR-2 36.3 3 5.1

CSTR-3 47 3 6.6

CSTR-4 40 3 5.6

CSTR-5 24.2 3 3.4

Total Volume 179

3.1.2. Second Reactor Configuration Proposal-PFR Model

As discussed in the previous chapter, using a CSTR in series improves the yields of DCP and reduces

the overall volume of reaction respect a conventional CSTR. The next logical step is to improve this

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configuration by developing a custom model of a real PFR reactor. The main idea behind this model

is to overcome some of the difficulties still present in the previous configurations:

-Elevated residence times.

-High water production.

-Two mol of Hydrogen chloride in gas phase is needed for each mol of GLY.

As it is shown in Figure 3.3, in the reaction between HCl and GLY, the production of 1,3-DCP is

increased significantly for high residence times (8-9 h). This is exactly the same relation obtained in

the simulation with ASPEN plus.

Figure 3.3 Kinetics Comparison between experimental results and the developed model for

chlorination of GLY [10].

As a consequence, the reactors needed for the DCP production have very high volumes. In each

CSTR, water is being generated in the system reducing the concentration of GLY, there is non-

reacted HCl and very high pressure is needed to reduce the overall volume of the gas phase for the

non-reacted HCl.

The first solution to deal with these problems is working at very high pressures (>20 bar) therefore

the volumetric flow of HCl is reduced. However, the residence times are still high and the volume

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cannot be further reduced. The idea to decrease the residence time and achieve high yields was to

increase the number of CSTR in series until the configuration assembled a more realistic PFR.

Figure 3.4 Absorption of HCl as a function of time for different pressures [3].

The model of the real PFR is developed by using many CSTR in series. As described in the literature

review chapter, the HCl is not participating on the kinetics, the flow of HCl introduced is optimized

to be the exact amount that can react in each reactor. Introducing more than required increases the

overall volume significantly and an additional separation is needed. Splitters are used to distribute

the flows of HCl. The split ratio is optimized so the outlet of each reactor does not contain unreacted

HCl, this ensures that the correct amount is used. The custom model of this reactor is shown in

Figure 3.5.

Introducing the exact amount of HCl in each reactor decreased the overall volume of reaction and

allowed to decrease the pressure of operation from 20 bar to 5 bar. All the amount of HCl has reacted

and therefore there is no gas phase in the reactor. Increasing the number of CSTRs used has shown to

improve the distribution of HCl among the reactor and reduces the overall volume and residence

times required to achieve the required yield of DCP.

22

Figure 3.5 Custom model of PFR with multi-injection of HCl.

In the real PFR, each CSTR reactor of the model is considered as one injection of HCl in the bottom.

The gas flow must consist of very thin bubbles to ensure optimum dispersion of HCl and increase the

mass transfer to be similar to each CSTR of the simulation and a depiction of the bubbling HCl into

the reaction mixture is illustrated in Figure 3.6. At the beginning higher amount of HCl is needed

because the reaction occurs faster and each CSTR was modelled with identical residence time.

Figure 3.6 Representation of HCl bubbling into the reaction mixture.

The model of the reaction part in Aspen is as shown in Figure 3.8 and 3.9. The model is equivalent

to use a total of four PFR. The hierarchies A and B are equivalent. These blocks are composed of

two PFR in series. They are interconnected with a distillation column in the middle that extracts all

the water formed during the reaction in the first reactor and sends back the products and unreacted

species to the second reactor.

23

Figure 3.7 PFD of DCP production.

Figure 3.8 Water extraction from PFR reactors.

24

B1B2B3

W1 W2

PFRB1OUT38(IN)

PFRA1OUT39(IN)PFRA2IN 40(OUT)

PFRB2OUT 41(OUT)

C-1

HE-1

HE-2 P-1

B5

HIERARCHY

PFR2

B3

HIERARCHY

PFR1

B2

HE-3

V-1

DC-1DC-2

HE-4

HCL

1HP

GLY 1HT S32

1HC

PFR2OUTGLY2

HCL2

GLY1

HCL1PFR1OUT

PRODUCTS

PL

PLTL

17

16

19

DICHL

2118

The water extraction step reduced the overall volume of the reactors by 40%. Both identical reactors

A1 and B1 are connected to the column. Water is removed from the tops, and the products and

unreacted species are sent back to reactors A2 and B2 respectively.

Figure 3.9 PFR Reactor 1 and 2 in series (blocks A and B equivalents) with water extraction.

The injection point was arithmetically determined according to the volume of each CSTR respect the

total volume of the PFR. Table 3.4 and 3.5 show the injection point of HCl, the exact flows of HCl

needed to completely react, the total length of the reactors, and the total volume of each reactor.

The diameter of the PFRs is fixed at 2.5m. As can be checked in bold letters in the following tables,

the first reactor has 24 m3 and 4.89 m of length, and the second reactor has 44 m3 and 9 m of length.

25

Table 3.4 Injection points of HCl for the first PFR.

Reactor 1

VCSTR t / h VPFR(L) / cc L / m Injection / m HCl / kg/h HCl / L/s

1567 0.25 1567 0.02 0.01 799 36.23

1701 0.50 3268 0.05 0.04 608 27.53

1806 0.75 5074 0.07 0.06 480 21.74

1891 1.00 6965 0.10 0.09 384 17.39

1971 1.25 8936 0.13 0.12 288 13.04

2036 1.50 10972 0.16 0.14 352 15.94

2085 1.75 13057 0.19 0.17 224 10.14

2127 2.00 15184 0.22 0.21 192 8.70

2162 2.25 17346 0.25 0.24 160 7.25

2193 2.50 19539 0.28 0.27 144 6.52

2221 2.75 21760 0.32 0.30 128 5.80

2245 3.00 24005 0.35 0.33 96 4.35

Table 3.5 Injection points of HCl for the second PFR

Reactor 2

26

t / h V(Length) / cc PFR2Length / m Injection / m HCl / kg/h HCl / L/s

3.25 1762 0.37 0.19 13.5 0.61

3.50 3545 0.40 0.39 9.6 0.44

3.75 5347 0.43 0.41 12.5 0.57

4.00 7167 0.45 0.44 10.6 0.48

4.25 9002 0.48 0.47 8.3 0.38

4.50 10853 0.51 0.49 8.2 0.37

4.75 12716 0.53 0.52 57.6 2.61

5.00 14591 0.56 0.55 54.4 2.46

5.25 16477 0.59 0.58 48.0 2.17

5.50 18372 0.62 0.60 41.6 1.88

5.75 20275 0.64 0.63 36.8 1.67

6.00 22186 0.67 0.66 38.4 1.74

6.25 24105 0.70 0.69 35.2 1.59

6.50 26031 0.73 0.71 32.0 1.45

6.75 27963 0.76 0.74 28.8 1.30

7.00 29901 0.78 0.77 25.6 1.16

7.25 31843 0.81 0.80 22.4 1.01

7.50 33790 0.84 0.83 19.2 0.87

7.75 35741 0.87 0.86 22.4 1.01

8.00 37697 0.90 0.88 16.0 0.72

8.25 39656 0.93 0.91 19.2 0.87

8.50 41618 0.95 0.94 16.0 0.72

8.75 43583 0.98 0.97 12.8 0.58

9.00 44369 0.99 0.99 0.5 0.02

Plotting the injection ratio of HCl/GLY mole ratio against the normalized length of reaction (PFR1

and PFR2 length), it is obtained a logarithmic expression that relates the optimized HCl injection

divided by the GLY mole flow (140 kmol/h of GLY at the inlet of the reactors) and the inject zones

across the reactor. This would be an advantage for designing the real reactor, where the nozzles

27

could be distributed evenly and it can even be extrapolated to other reactor size and other production

requirements.

0.30 0.40 0.50 0.60 0.70 0.80 0.90 1.000.00

0.01

0.02

f(x) = − 0.0155410550078427 ln(x) + 0.00144716881721868R² = 0.983599034486075

f(x) = NaN ln(x) NaNR² = NaN Normalized injection

Reactor 1 Logarithmic (Reactor 1)Reactor 2 Logarithmic (Reactor 2)

L/Lreactor (m)

HCl/G

Ly ra

tio

Figure 3.10 Plotting of HCl injection points across PFR1 and PFR2.

3.1.3. DCP Separation-1st Train

DCP are obtained on removal of water, unreacted GLY, HCl and intermediate compounds α and β-

MCH from the product stream exiting the reactors. As can be checked on Table 3.6, the difference in

volatilities allows to separate all the components easily.

Table 3.6 Boiling points of components in the system

28

Components Boiling points (°C)

Glycerol 289

HCL -85.05

α -MCH 213

β-MCH 220.35

1:3 DCP 174.3

2:3 DCP 184

The extraction of these undesired products from the mix is done by employing two distillation

columns as seen in Figure 3.11. In the first column DC-1, water and traces of HCl are extracted from

the product mix. The inlet stream of the column stream 15 is at 130oC and at 1 bar. A partial

condenser has been used on the column to flash remaining HCl before the condenser (stream-16).

The bottom stream is heated up by HE-4 and is sent to the second column DC-2, the desired product

of the first reaction, DCP, are retrieved on stream 20 with 99.5 % purity.

C-1

HE-1

HE-2 P-1

B5

HIERARCHY

PFR2

B3

HIERARCHY

PFR1

B2

HE-3

V-1

DC-1DC-2

HE-4

HCL

1HP

GLY 1HT S32

1HC

PFR2OUTGLY2

HCL2

GLY1

HCL1PFR1OUT

PRODUCTS

PL

PLTL

17

16

19

DICHL

2118

Figure 3.11 DCP separation-1ST separation train.

Optimization first separation train

The important criteria considered for obtaining the optimal dimensions of the column are:

Number of stages

Feed stage

Mole recovery

29

Sensitivity analysis was done to find the optimal parameters considering the above criteria.

Optimization was performed by varying number of stages, feed stage and reflux ratio, the minimum

duties are found (iterative process). The composition profile of the components of interest must

reveal a changing profile across the column, without pinches (feed stage not optimal) or stages at

constant compositions (excess of stages). For the recovery, the bottoms/distillate rate was varied to

obtain maximum recovery of desired product.

Sieve trays are used for the columns due to their low cost, low maintenance requirements, low

fouling tendency. The design specs were applied to mole recovery and mole purity of 1,3-DCP in the

first column and second column respectively, since that is most desired product. The Table 3.7 and

3.8 give information about the design specs and dimensions of the columns.

Table 3.7 Design spec of the columns

Column Design Spec

DC1 Recovery and purity of 1,3-DCP in the

Bottom: 0,999

DC2 Recovery and purity of 1,3-DCP in

Distillate: 0,99

Table 3.8 Dimensions of distillation columns

Column Parameters DC1 DC2

Reflux ratio 0.5 0.5

Number of stages 10 25

Feed stage 5 18

The energy consumption of the columns were minimized is to be noted since it is an important

factor considering the overall economics of the process. The Table 3.9 shows the condenser and

reboiler duties of the columns.

Table 3.9 Energy Requirements of the columns

Column Qc Qreb Tc Tr

(MW) (MW) (°C) (°C)

C1 -0.80 1.22 98.88 174.52

C2 -2.61 0.60 172.52 215.75

30

3.2. Dehydrochlorination of DCP

Dehydrochlorination of GLY dichlorohydrin is the process in which the DCP’s react with a base to

form EPH, salt and water. In industry this process is carried out with lime milk. It creates many

ecological problems caused by waste water containing calcium chloride. It is advisable to replace the

lime milk by NaOH or catholyte. The waste water after the concentration and purification could be

recycled to electrolysis [26].

DC-2B1

RDISTILL

DICHL

21

DCHLFEED

NAOH

TPROD

BTM

Figure 3.12 Process flow diagram of reactive distillation column

The reactive distillation column is fed with the DCP from column DC-2 (cooled down to the reaction

temperature of 60°C) and NaOH solution (30 mol% concentrated). The column specifications are as

follows:

Table 3.20 Column Parameters

Column Parameters

Pressure 1 bar

Number of stages 10

31

Feed stage 5

Distillate rate 365 kmol/hr

Reflux ratio 0.1

Reboiler duty 6.11 MW

Condenser duty -4.7 MW

Temperature (feed) 63°C

Temperature (top) 88 °C

Temperature (bottom) 148.8 °C

Reaction zone (stage

numbers)

Stages 4-6

Residence Time 10 mins.

In the reactive distillation model, either the residence time or the liquid hold up can be specified. The

reaction stages are from 4 to 6, with 10 minutes of residence time [16]. The thermodynamic property

method used to carry out this reactive distillation is Electro-NRTL. The separation of EPH and water

is an important aspect of this process since EPH forms an azeotrope with water at about 88 °C and

101.3kPa. Aspen database’s does not include the binary interaction for this mixture. In order to

achieve a proper simulation the parameters for the azeotropic distillation have to be introduced in the

properties environment.

Table 3.11 Binary parameters for NRTL method for EPH-water azeotropic distillation [23]

The top of this column i.e. the top product consist mainly of water and EPH. Sodium chloride

solution was obtained at the bottom. Design specifications were employed to achieve maximum

recovery of EPH at the top by varying the distillate rate. The yield of EPH was found to be 99%.

Almost no GLY was formed in the system.

3.2.1. EPH Purification-2nd Separation Train

The last part of GTE process is the separation of water from the final product, EPH. As mentioned

previously these two components form an azeotrope at 88 oC, atmospheric pressure. The traditional

techniques to tackle with azeotrope via using distillation are pressure swing, when the equilibrium is

32

sensitive to pressure changes, entrainers or to resort to modern types of separation such as

membranes [18]. After assigning the binary NRTL parameters to Aspen plus, T-x-y diagram

retrieved as illustrated in Figure 3.13, whilst Figure 3.14 shows the corresponding graph derived

from experimental results. Apparently, the two graphs are almost identical indicating the azeotrope

composition to be at 0.34 mole fraction of water and 0.66 for EPH. This observation leads to the

conclusion that a promising separation can be simulated on Aspen Plus, but besides this fact , it is

evident that the forming azeotrope is heterogeneous with the formation of two liquid phases, one

organic and one aqueous.

Figure 3.13 T-x-y EPH/Water diagram obtained from Aspen Plus simulator after the assignments of

binary parameters.

For the simulation of the 2nd separation train NRTL base method employed since is capable of

predicting vapour-liquid-liquid equilibrium. After the above observation the separation becomes

simpler due to ability to cross distillation boundaries in the case of heterogeneous azeotropic

distillation. The proper piece of equipment in order to achieve this is a simple decanter while phase

splitting is not constrained by distillation boundaries.

33

Figure 3.14 T-x-y EPH/Water diagram from experimental work [23].

Figure 3.16 shows the PFD of the 2nd separation train. The total mole flow rate of stream 28 is 216

kmol/hr with 90 kmol/hr water and 126 kmol/hr EPH and it is fed at the fifth stage of DC-3 column.

At the top of the column the azeotropic composition of the mixture is obtained while the bottom

stream contains pure EPH (ca. 83 kmol/hr).

Figure 3.16 PFD of 2nd separation train.

Pure EPH can be retrieved at the bottom of the column because of the significant difference of its

boiling point (118 oC) and the azetrope temperature (88 oC). From the liquid mole fraction of EPH

throughout the column in Figure 3.17, it is seen that after stage six pure EPH is present in the liquid

phase, whilst at the top of the column (Stage 1) EPH exists in its azeotropic composisiton.

34

Block DC-3: Composition Profiles

Stage Number

Liqu

id M

ole

fra

ctio

n o

f EPH

1 2 3 4 5 6 7 8 9 100,30

0,35

0,40

0,45

0,50

0,55

0,60

0,65

0,70

0,75

0,80

0,85

0,90

0,95

1,00

Liquid mole fraction ALPHA-01

Figure 3.17 EPH liquid mole fraction per stage in unit DC-3.

The distillate of DC-1 is driven to the first decanter where significant amount of water (73.5 kmol/hr)

is removed due to the immiscibility of the two liquid phases. With utilisation of one more column

(DC-2) and a second decanter complete separation of EPH, retrieving 123.5 kmol/hr out of 125

kmol/hr of product with 99.1% purity. In such a way the principle of crossing the distillation

boundaries with decanters in the case of azeotrope is verified and the separation becomes simple

without any need of introduction of new materials (e.g. entrainers) to the process or pressure change.

Finally the relatively high difference in the boiling point of the azeotrope and EPH allows the usage

of small columns and therefore low capital cost.

Table 3.12 Column Input Specifications

Parameter DC-1 DC-2

Number of stages 10 10

Distillate rate(kmol/hr) 133 25

Reflux ratio 0.1 0.5

Feed stage 5 5

Feed Temperature (°C) 65 65

Pressure (bar) 1 1

35

Equipment Sizing and Cost

36

4. Equipment sizing and cost

In order to evaluate the total capital costs, it is first necessary to size the equipment. After

introducing the desired parameters, ASPEN Plus can estimate the sizing of most of the equipment

used in the simulation.

4.1. Reactors and columns

The reaction unit for the chlorination of GLY was simulated by many CSTR’s in series as presented

on Chapter 3; however in reality four PFR’s will be utilised and thus the installed equipment cost

should be estimated accordingly. For that purpose Guthrie method has been employed by using

Marshall and Swift (M&S) indices and costing was carried out by considering every PFR as

horizontal pressure vessel. M&S indices could not be retrieved from the literature for 2015 and a

sufficiently high value (1800) was assumed compared to M&S indices in 2007 (ca. 1363) [19] in

order to avoid underestimation of the cost. Moreover, 30 % of the calculated cost was added to the

final price, taking into account the coating (i.e. glass line coating) as well as additional charges for

piping and construction. The total installed cost of the reactor found to be 1.04 M$ and the installed

cost of each individual unit is shown in Table 4.1, whilst the sizing and the installed cost of the

columns is demonstrated in Table 4.2.

Table 4.1 Dimensions PFR reactors for Chlorination of glycerol.

PFR1A PFR2A PFR1B PFR2B

Volume (m3) 24 44 24 44

Residence Time (hr) 3 5.7 3 5.7

Diameter (m) 2.5 2.5 2.5 2.5

Length (m) 4.8 9.0 4.8 9.0

Installed Equipment

Cost (k$)

15.2 24.8 15.2 24.8

Table 4.2 Costs and sizing of distillation columns and reactive distillation (RDC-1).

DCW-1 DC-1 DC-2 RDC-1 DC-3 DC-4

Diameter (m) 1.22 0.91 1.37 1.83 1.37 0.76

37

Number of trays 33 12 12 12 12 12

Tray spacing (m) 0.61 0.61 0.61 0.61 0.61 0.61

Height (m) 20.13 7.32 7.32 7.32 7.32 7.32

Reboiler duty (MW) -2.86 -0.80 -2.61 -4.70 -1.72 0.55

Condenser duty MW 1.97 1.22 0.60 6.12 2.10 0.55

Equipment cost (k$) 586 520 557 724 557 442

4.2. Pumps and Compressors

It is assumed that the raw materials enter the plant at 30ºC and 1 bar. They need to be compressed to

the working pressure of the reactor (5 bar). The details for the equipment needed is shown in Table

4.3 and 4.4.

Table 4.3 Costs and sizing of pressure changers for raw materials

Isoentropic compressor HCl Centrifugal pump GLY

Net work required 478 kW Electricity 1.5 kW

Efficiency 0.8 Flow 3 l/s

Outlet pressure 5 bar Outlet pressure 5 bar

Outlet temperature 250 C NSPHa 8.5 m

Isentropic outlet

temperature

207 C Head 34 m

Equipment cost 1.1 M$ Equipment cost 80 k$

Installed cost 5.1 M$ Installed cost 140 k$

Table 4.4 Costs and flows of other pumps

Flow rate (m3/s) Installed cost / $ Capital cost / $

SW.DW-reflux pump 2.85 32500 5300

38

SW.B1 pump 4.62 90300 66200

DC-1-reflux pump 0.83 28200 5100

DC-2-reflux pump 13.21 43900 6123

DC-3-reflux pump 3.49 33900 5900

DC-4-reflux pump 0.89 31700 5100

RDC-1-reflux pump 10.00 40100 8100

4.3. Heat exchangers and decanters

The results for required heat exchangers and decanters are shown in the tables below.

Table 4.5 Costs and flows of additional pumps.

HE-1 HE-2 HE-3 HE-4 HE-5 HE-6 HE-7 SW.HE

Heat exchange area / m2 13.9 6.2 3.8 53.2 12.4 36.8 4.7 56.5

Capital cost / $ 10900 10800 8500 8400 10800 15200 9600 24800

Installed cost / $ 61600 63200 61300 61200 60700 74300 58800 103200

Energy / kW -432 504 128 162 658 -1485 213 1925

Table 4.6 Costs and flows of additional pumps

D-1 D-3 D-2

Liquid volume m3 170 170 170

Vessel diameter / m 9.8 9.8 9.8

Design gauge pressure / bar 3 3 3

Heat duty / kW 148 278 1.2

Design temperature / C 122 121 157

Operating temperature / C 88 87 130

Capital cost / $ 15400 15400 15400

Installed cost / $ 119900 116200 117600

39

GTE Aspen Economics

40

5. Economic evaluation

The selling price for conventional EPH ranges from 1700 to 2200 $/ton in the Asian market. In this

section, this prices will be compared to the required selling price and to recommended selling price.

5.1. ASPEN Economics Input

As the required amount of hydrogen chloride is very high (85848 kton/year) for the production of

100kton/year of EPH, ideally a neighbour plant which produces HCl may be required. However, in

order to make an economic evaluation it is necessary make assumptions to establish a price for each

stream. This price was extrapolated from the price of 36% hydrochloric acid (100$/ton), dividing by

0.36 (concentration), the price per ton HCl is extracted (277$/ton). Assuming a factor of 30% for the

separation of hydrogen chloride from water, the value of the HCl stream was obtained (360$/ton).

The price of solid pearls of pure NaOH (300$/ton) was used to extract the price of the NaOH (aq)

30% mol (120$/ton). The refined vegetal glycerol price ranges from 700-800$/ton [24] in the Asian

market. The summary of the prices introduced is shown in the following table:

Table 5.1 Raw materials price

Raw materials price

HCl 360 $/ton

GLY 800 $/ton

NaOH 30% vol 120 $/ton

Another additional parameters where needed to be introduce:

Rate return of the total costs is 20%/year.

The life expectancy of the plant is assumed equal to 30 years.

41

5.2. Results

After introducing the raw materials price, extracting the price of the CSTR in series and adding the

price for the PFR reactors it is possible to use ASPEN economics to obtain all the desired economic

information.The utility costs are found in Table 5.2.


Utility Fluid Rate Units Cost per

Hour

Cost

Units

Electricity 690.5 kW 77.1 $/ hr

Cooling Water Water 1065.6 m3/hr 33.7 $/ hr

Steam @690KPA Steam 19.7 ton/hr 353 $/ hr



The summary of the economic results is the following:


Parameter Price

Total Capital Cost [M$] 30.0

Total Operating Cost [M$/Year] 163.3

Total Raw Materials Cost [M$/Year] 140.8

Total Product Sales [M$/Year] 199.6

Total Utilities Cost [M$/Year] 6.3

Desired Rate of Return [Percent/'Year] 20

Equipment Cost [M$] 4.3

Total Installed Cost [M$] 10.7

Required selling price [$/ton] 1633

Asia market price [$/ton] 1800-2,500

42

The required selling price (1633 $/ton) was calculated by dividing the total operating cost of the

plant (163 M$/year) by the production capacity 100,000 ton/year.

The percentage of investment returned each year is calculated in aspen assuming the total cash flow

for each year. Taking into account the total expenses and the total income, the percentage of

investment returned each year is defined as return over investment. Assuming a final recommended

selling price for the EPH in Asia of 1800 $/ton [25], the return over investment (R.O.I) is

33%/year.

Results of ASPEN economics for cash flows and detailed revenues and costs are found in appendix.

Both EPH and GLY prices vary significantly over the time. In order to study the viability of the

project a sensitivity analysis was performed. This analysis is shown in “Figure 5.1”.

600 650 700 750 800 850 9001000

1200

1400

1600

1800

2000

f(x) = 1.34551428571429 x + 556.918571428572R² = 0.999999999549093

Sensitivity analysis: GLY price

Market price for GLY in ($/ton)Min

imum

req

uire

d se

lling

pri

ce f

or E

PH

($

/ton

)

Figure 5.1 Sensitivity analysis - GLY price with RSP of EPH

Using the relation found on “Figure 5.1”, it is possible to predict the minimum required selling price

for EPH. This process is economically viable when the price of GLY in the market is lower than

900$/ton. The price market trend for glycerol is shown in “Figure 5.2”.

43

Figure 5.1 Sensitivity analysis - GLY price with RSP of EPH

Figure 5.1 shows that glycerol price has decreased significantly over the years, making the

alternative GTE process viable and very profitable. It is worth remarking that the crude glycerine

price is eight times lower than refined glycerine.

A further study of the economics for glycerine refining process is recommended in order to

determine if the GTE process can be even more profitable by refining in-situ the crude glycerine.

Crude glycerine price has sharply decreased over the last years as mass production of biodiesel is

continuously increasing over time.

44

Conclusions

45

6. Conclusions and Recommendations

Market price of GLY has sharply decreased over the years due to increased biodiesel production

making GTE process feasible. In the present report the economic feasibility of a plant with a capacity

100 kton/year for EPH manufacturing via GTE process has been investigated. The payback period

has been estimated 3 years for 1800 $/ton EPH selling price, highlighting GTE process as

economically viable and a promising alternative.

For the first step of the process (i.e. chlorination of GLY) two reactor configurations have been

proposed and intermediate water removal is strongly recommended in order to achieve smaller

reaction volume. Custom PFR model is suggested for further reduction of residence times and

therefore reaction volume. For the second step, a reactive distillation column is used for fast removal

of EPH and improved yields. The purification of EPH has been done with consecutive decanters and

distillation column in order to tackle the heterogeneous azeotrope, giving 99% pure product (EPH).

The overall yield of GLY to EPH found to be 89%.

Pinch point analysis including only the heat exchanger units of the process showed annual savings

ca. 168 k$/year in the utility costs. More thorough heat integration including the column of the

process is recommended for further reduction.

Appendix

The following figures (Figure 1, Figure 2 and Figure 3) represent the 3D design regarding the first

proposal of the reaction system for chlorination of GLY from two different perspectives. The first

reaction block is composed of two CSTR’s in series, focusing on the production of a-MCH. GLY is

fed only to the first reaction unit after Splitter-1, whilst HCl is fed to the first five reactors for the

reasons explained on Chapter 3. In order to reduce the number of units, reaction towers with stages

will be utilised. In each stage a downcomer will be designed for the flow of the reaction mixture

from on stage to the other so as to achieve the required residence time for every reactor. GLY lines

are painted in red colour and HCl lines are in blue. In each reaction column a shaft is employed

having had multi-impellers incorporated.

46

Figure 1 Chlorination of GLY 3D reactor design.

47



PFR cost estimation – Guthrie method

The free on board cost (f.o.b.) for each reactor is estimated by Equation (1):

Where H is the length of the reactor, D the diameter and Cp0 the f.o.b. cost in 1968.

Then the installed cost can be calculated from Equation (2),

Where FBM is a correction factor equal to 4.23, Fm the material correction factor which is equal to 1

for Carbon Steel and FP the pressure correction factor equal to 1.05 for pressure less or equal to 6.7

bar. The calculated cost is for 1968 and by using Equation (3) the corresponding for 2015 can be

found.

48

Heat Integration of GTE Process-Pinch Point Analysis

Pinch point analysis was performed on GTE process targeting to the reduction of hot (QHmin) and cold

(QCmin) utilities as well as for defining the minimum number of the heat exchanger units. The

minimum temperature difference was chosen equal to 20 oC in order to achieve balance between the

capital and utility costs. In the process there are five cold stream and the same number of cold

streams which can be combined for minimising energy losses. Table 1 demonstrates the process

streams and their properties such as heat capacities, flows and heat capacities flowrates (FC p) as

obtained from the Aspen simulation. The number of each stream refers to the number of the stream

of the PFD while the letter indicates whether the stream is hot (i.e. able to transfer heat) or not.

Table 1 Hot and cold process streams properties

Stream Tin (K) Tout (K) Cp (kJ/kmol· K) F (kmol/sec) FCp (kW/K)

H2 523 363 28.8 268 2.1

H21 446 333 167.2 126 5.8

H28 472 338 141 217 8.5

H33 403 338 236.1 58 3.8

H35 361 353 183.9 25 1.2

C4 303 363 217.3 140 8.4

C14 393 403 137.9 198 7.6

C18 448 468 172.4 140 6.7

C25 361 472 54.3 365 5.5

C31 361 403 54.3 133 2

From the inlet and outlet temperatures of hot streams is evident that the requirements of cold utilities

are significantly higher than the requirements for hot utilities. Figure 4 represents the process

streams as vectors and the temperature scale for cold streams is shifted by ΔΤmin as pinch point

analysis dictates. The problem, as it can be observed is divided into fifteen temperature intervals (I1,

I2 etc.) where hot streams and cold streams are able to exchange heat and heat surplus can be

transferred from one interval to the other owing to the driving force or i.e. temperature gradient.

49

523 K 503 K

492 K 472 K

488 K

468 K

472 K 452 K

468 K 448 K

446 K 426 K

433 K 403 K

413 K 393 K

403 K 383 K

383 K 363 K

381 K 361 K

363 K 343 K

361 K 341 K

353 K 333 K

338 K 318 K

333 K 303 K

Figure 4 Stream population of GTE process.

50

I-1

I-3

I-2

I-4

I-7

I-5

I-6

I-8

I-9

I-10

I-11

I-13

I-12

I-15

I-14

H2

H35

C18

H28

H21

H33

C42

C25

C14

C31

The amount of heat of each interval is found by multiplying the total flow heat capacity of the

interval by the corresponding temperature difference as indicated from Equation 4 [18]:

Where FCH,i is the heat capacity of the hot stream i in the specific interval, FCC,i the heat flow

capacity of cold stream j and ΔΤinterval the temperature difference between the limits of the interval.

The heat duty of each interval was calculated and afterwards cascade calculations were performed so

as to define the minimum hot and cold utilities of the process as demonstrated in Table 2.

Table 2 Minimum hot and cold utilities estimation.

Interval FCp,interval ΔΤinterval (K) Qinterval (kW) Cascade

ΣQj (kW)

0 +113.8

1 2.1 31 +66.7 +66.7 180.6

2 -3.3 4 -13.4 +53.3 167.2

3 -10 16 -160.9 -107.6 6.2

4 -1.5 4 -6.2 -113.8 0

5 5.1 22 +113.3 -0.56 113.3

6 11 13 +143.2 +142.6 256.5

7 3.4 20 +67.9 +210.6 324.5

8 9 10 +90 +300.7 414.5

9 12.8 20 +256.5 +557.2 671.0

10 4.3 2 +8.7 +565.9 679.8

11 11.8 18 +214 +779.9 893.8

12 9.7 2 +19.4 +799.4 913.3

13 11 8 +88.1 +887.5 1001.4

14 9.7 15 +146 +1033.5 1147.4

15 -2.5 5 -12.9 +1020.6 +1134.5

As it can be observed the heat duty of each interval is calculated by Equation 4, and fourth column

is derived by cascade calculation of column 3. Negative values on column 4 imply heat transfer from

lower temperatures to higher temperatures, violating in such a way the second law of

51

thermodynamics. The most negative value (-113.8 kW) is observed in the fourth row and thus in

order to obstruct further violation of the second thermodynamic law this amount of heat should be

provided as an absolute value to the first interval (I-1) as being done in the sixth column. This

amount of heat simultaneously is the minimum requirements in hot utilities. The minimum amount

of cold utility is found by proceeding again with cascade calculations and it can be seen at the bottom

cell of the sixth column (+1134.5 kW). Finally by definition the point where the flux of heat equals

zero is the pinch point and it is located at interval 4 or at 472 K for hot streams and 452 K for cold

streams; because of the zero heat flux in the pinch point, the problem now is separated into two

different problems, namely above pinch point and below pinch point.

The minimum hot and cold utilities have been found; however further analysis is needed in order to

define the minimum number of heat-exchange units and also the way at which the different streams

should be combined together for the achievement of minimum utilities. To that end the problem is

separated to two sub-problems, namely above and below the pinch, and both of them are analysed on

the following sections.

GTE Heat Exchanger Network Design above Pinch Point

The minimum number of heat exchange units above pinch is given by Euler’s theorem as following

[19],

Where NHE,AP the number of heat exchange units and NS,AP the number of streams above pinch point.

Taking into account one more unit for the hot utility above pinch, the predicted number of units is

four. Figure 5 represents the problem above pinch. Stream H28 has 8.5 kW/K flow heat capacity, a

higher number than the two cold streams. Therefore, this stream cannot be utilized for heating up the

cold streams C25 and C18 which have lower flow heat capacities [20]. Splitting of stream H28 into

two streams with smaller flow heat capacities (8.5-x and x) is the necessary action to enable heat

transfer.

52

Figure 5 GTE heat integration above pinch point.

Only hot utility should be used above pinch point, thus the main goal is the full energy satisfaction of

the hot streams. HE-1 connects C25 and H2 fulfilling the energy requirements of H2. The necessary

heat duty can be calculated by multiplying the temperature difference of H2 and the flow heat

capacity of the stream (115 kW). Apparently C25 has not reached the target temperature (i.e. 472K),

and therefore the outlet temperature of the stream has to be found by performing an energy balance.

Stream H28 is divided into two streams with flow heat capacities 8.5-x and x. Presuming that we

want to satisfy stream C25 with the stream of x flow heat capacity. The value of x can be found from

the energy balance between the two streams as follows,

53

A heat exchanger with 16.5 kW heat duty is employed for that purpose and the other sub-stream

with flow heat capacity of 8-x (4.375 kW/K) can be combined with C18.

The outlet temperature is calculated in the same manner at 450 K and apparently 116.5 kW of hot

utility (High Pressure Steam) is needed for heating up the stream to the final target (i.e. 468K). This

value is slightly higher than the one calculated in Table 2 most likely due to propagation errors on

excel. Finally the number of heat exchange units verifies Euler’s theorem, since three units are

needed plus one more unit for the utility.

GTE Heat Exchanger Network Design below Pinch Point

Below the pinch point hot streams are more than cold streams (see Figure 4), rendering feasible heat

integration. The same procedure as the previous section was followed with the only difference that

here hot streams are able to exchange heat only with cold streams of lower flow heat capacity and

also the main goal is the full energy satisfaction of cold streams as only cold utilities can be utilised

below pinch [20]. Accordingly heat integration was performed as presented in Figure 6.

The minimum number of heat exchange units is estimated again by Euler’s theorem [19] as follows,

Therefore eight heat exchange units plus one unit for the utility are expected and this result can be

verified by Figure 6 where there are 9 units in total. The minimum cold utility requirements can be

found by summing up all the cold utility duties from Figure 6, and the total duty is found equal to

1134.1 kW, indicating optimum stream combination. Having performed pinch point analysis, the last

step is to provide some rough estimations of annual savings ($/year) achieved by applying the

technique in practice.

54

Figure 6 GTE heat integration below pinch point.

55

Annual Savings Estimation

Initially the total annual utility cost should be estimated for the case where there is no heat

integration and subsequently the corresponding cost with heat integration applied in order to

calculate the annual savings. To that end prices for different hot and cold utilities are provided in

Table 3.

Table 3 Utility properties and prices [20,21].

Utility Price ($/tn) Cp (kJ/kg K)

Cooling water (CW) 0.06 4.2

Cost ($/tn) Δhvap (kJ/kg)

High Pressure steam (HPS,42 bar-254 oC)

3.5 3.8

Medium Pressure steam (MPS,11 bar-

184 oC)

7 2.54

Low Pressure steam (LPS,6 bar-160 oC)

18 2.15

The heat content of each process stream can be estimated by Equation 5, depicting the amount of

heat that should be added or removed from the stream so as the target temperature to be reached. The

annual mass flow rate of water is calculated by Equation 6 and the one for steam from Equation 7

and subsequently the annual cost of utilities.

Where FCp the flow heat capacity of the stream, mcw the mass flow of cooling water, msteam the mass

flow of steam and Δhvap the latent heat of steam as given in Table 3.

Table 4 shows the calculation of the annual utility cost without heat integration of the process and it

is estimated around 287 k$/year. By performing pinch point and combining hot and cold streams this

cost can be reduced by ca. 56 % as demonstrating in Table 5.

56

Table 4 Annual utility cost before heat integration.

Stream Utility Qstream (kW) Utility Mass

Flowrates

(tn/year)

Cost ($/year)

H2 CW 344 258962 15537

H21 CW 662 497763 29865

H28 CW 1139 855827 51349

H33 CW 248 186247 11174

H35 CW 10 7672 460

C4 LPS 507 5898 20644

C14 LPS 76 885 3100

C18 HPS 134 1510 27197

C25 HPS 611 6890 124024

C31 LPS 84 980 3432

Total annual utility cost M$ 0.287

Table 5 Annual utility cost after heat integration.

Utility Qutility Utility Mass Flowrates

(tn/year)

Cost ($/year)

CW 1134.6 1703845 102230

HPS 116.5 1312.6 23626

Total annual utility cost M$ 0.12

57

Control Scheme of GTE Process

Any process flow diagram is incomplete without at least a basic control scheme to explain the

controls so the output of a specific process is maintained within a desired range. The critical factors

which can cause disturbances in the operating conditions of the process are [20]:

Changes in the feed flow rate – for example changes in the temperature or feed composition.

Changes in the conditions of utilities like steam or cooling water temperature fluctuations.

Ambient conditions in the environment like temperature fluctuations or moisture content.

The basic parameters under consideration for designing a control scheme for a process are:

Temperature

Pressure

Level

Flow

The process scheme under consideration consists of various equipment supposed to be working at

specified conditions. This equipment requires a control mechanism so that desired quality of product

is obtained. The main equipment under consideration for this case are the reactors, distillation

columns, the heat exchangers and the decanters.

Control of Reactors

The reactor system consists of the PFRs and the reactive distillation column (which is a hybrid vessel

consisting of a reactor and a distillation column in one equipment). As mentioned previously, the

main parameters considered for the control of the PFRs is the pressure in the vessel, and the level in

the vessel. The temperature of the vessel is controlled by using a jacketed vessel. The flow of the

coolant in the jacket is the manipulated variable used for the control of this parameter. The pressure

of the vessel is controlled by regulating the feed flow of the HCl gas in both the vessel, at the valve

after the splitter. The level of the vessels is controlled by the product outflow from the vessel. The

flow to these vessels is kept in a ratio which is already specified in the splitter equipment.

The reactive distillation column is slightly different equipment since it combines two unit operations

in a single vessel. The temperature of the reactive zone is regulated by the flow of dichlorohydrins to

the column. The composition of the product at the top is regulated by the flow of the NaOH to the

58

column. The control of this equipment is otherwise very similar to a normal distillation column

which is explained in the next section.

Control of Distillation columns

There are in totality six distillation columns in the EPCH production process scheme including the

reactive distillation column. The critical parameters for smooth operation of the columns are

controlled in the following ways:

Pressure:

The pressure of the column can be either controlled by manipulating the coolant flow of the

condenser in case of a total condenser or the vapor outflow of the reflux drum in case of a partial

condenser. In the EPH production process, the pressure in the columns C-2, RDistill, C-4 and C-5 is

controlled by regulating the coolant temperature. The column C-0 and C-1 has a partial condenser

and therefore the flow of vapor of the reflux drum is used to control the pressure in these columns.

Temperature:

The selection of the manipulated variable for the temperature control in a column depends on the

requirement of product quality of the top or bottom product. If the top product quality is required to

be high then the reflux ratio is to be regulated and if the bottoms product quality is the essential, the

temperature is regulated by varying the hot utility flow in the reboiler. The temperature and in turn

the quality of the top product of the columns RDistill and C-4 is controlled by manipulating the

reflux ratio. In the remaining columns the bottom product is important and thus the temperature is

controlled by varying the flow rate of the hot utility in the reboiler.

Level:

The level in all the columns at the top and bottom is controlled by varying the distillate rate and the

bottoms rate respectively. In our case, the level of all the columns at the bottom is controlled by the

bottoms flow rate. The level of the reflux drum is controlled by the distillate flow rate. The flow rate

of the reflux is controlled in the columns where the temperature is controlled at the bottom.

Heat Exchangers and Decanters

The temperature control of the heat exchangers is usually carried out by manipulating the utility flow

i.e the heating or the cooling fluid flow in the exchanger. The same has been done for all the heat

59

exchangers in this scheme. The level in the decanters is controlled by controlling the flow of one of

the liquids flowing out of the decanter, usually the liquid that is sent to the next unit.

60

61

DETAILED COSTS I (ASPEN DATABASE)

ITEM UNITS

TW (Number of Weeks per Period) Weeks/period 52

T (Number of Periods for Analysis) Period 20

DTEPC (Duration of EPC Phase) Period 0.846154

DT (Duration of EPC Phase and Startup) Period 0.929487

WORKP (Working Capital Percentage) Percent/period 5

OPCHG (Operating Charges) Percent/period 25

PLANTOVH (Plant Overhead) Percent/period 50

CAPT (Total Project Cost) Cost 3.06E+07

RAWT (Total Raw Material Cost) Cost/period 1.41E+08

PRODT (Total Product Sales) Cost/period 2.00E+08

OPMT (Total Operating Labor and Maintenance Cost) Cost/period 2.37E+06

UTILT (Total Utilities Cost) Cost/period 6.33E+06

ROR (Desired Rate of Return/Interest Rate) Percent/period 20

AF (ROR Annuity Factor) 5

TAXR (Tax Rate) Percent/period 40

IF (ROR Interest Factor) 1.2

ECONLIFE (Economic Life of Project) Period 30

SALVAL (Salvage Value (Percent of Initial Capital Cost)) Percent 20

DEPMETH (Depreciation Method) Straight Line

DEPMETHN (Depreciation Method Id) 1

ESCAP (Project Capital Escalation) Percent/period 5

ESPROD (Products Escalation) Percent/period 5

ESRAW (Raw Material Escalation) Percent/period 3.5

ESLAB (Operating and Maintenance Labor Escalation) Percent/period 3

ESUT (Utilities Escalation) Percent/period 3

START (Start Period for Plant Startup) Period 1

PODE (Desired Payout Period (excluding EPC and Startup Phases)) Period

POD (Desired Payout Period) Period

DESRET (Desired Return on Project for Sales Forecasting) Percent/Period 10.5

END (End Period for Economic Life of Project) Period 30

GA (G and A Expenses) Percent/Period 8

DTEP (Duration of EP Phase before Start of Construction) Period 0.326923

OP (Total Operating Labor Cost) Cost/period 2.06E+06

MT (Total Maintenance Cost) Cost/period 305714

62

CASH FLOW

Period

1st year 2nd year 3rd year 4th year 5th year 6th year 7th year 8th 10 10th year

R (Revenue) -4.50E+07 4.52E+07 5.02E+07 5.55E+07 6.11E+07 6.71E+07 7.36E+07 8.04E+07 8.77E+07

DEP (Depreciation Expense) 817172 817172 817172 817172 817172 817172 817172 817172 817172

E (Earnings Before Taxes) -4.58E+07 4.44E+07 4.94E+07 5.46E+07 6.03E+07 6.63E+07 7.27E+07 7.96E+07 8.69E+07

TAX (Taxes) 0 1.78E+07 1.97E+07 2.19E+07 2.41E+07 2.65E+07 2.91E+07 3.18E+07 3.48E+07

NE (Net Earnings) -4.58E+07 2.66E+07 2.96E+07 3.28E+07 3.62E+07 3.98E+07 4.36E+07 4.78E+07 5.21E+07

TED (Total Earnings) -4.50E+07 2.75E+07 3.04E+07 3.36E+07 3.70E+07 4.06E+07 4.45E+07 4.86E+07 5.30E+07

TEX (Total Expenses (Excludes

Taxes and Depreciation))

5.98E+07 1.75E+08 1.81E+08 1.87E+08 1.94E+08 2.00E+08 2.07E+08 2.15E+08 2.22E+08

CF (CashFlow for Project) -4.50E+07 2.75E+07 3.04E+07 3.36E+07 3.70E+07 4.06E+07 4.45E+07 4.86E+07 5.30E+07

FVI (Future Value of Cumulative

Cash Inflows)

1.48E+07 2.38E+08 5.16E+08 8.62E+08 1.29E+09 1.82E+09 2.46E+09 3.25E+09 4.20E+09

PVI (Present Value of Cumulative

Cash Inflows)

1.23E+07 1.65E+08 2.99E+08 4.16E+08 5.18E+08 6.08E+08 6.86E+08 7.55E+08 8.15E+08

PVOP (Present Value of

Cumulative Cash Outfows Prod.)

4.98E+07 1.84E+08 3.00E+08 4.01E+08 4.88E+08 5.64E+08 6.30E+08 6.87E+08 7.37E+08

PVO (Present Value of Cumulative

Cash Outfows)

4.98E+07 1.84E+08 3.00E+08 4.01E+08 4.88E+08 5.64E+08 6.30E+08 6.87E+08 7.37E+08

NPV (Net Present Value) -3.75E+07 -1.84E+07 -820491 1.54E+07 3.02E+07 4.38E+07 5.63E+07 6.76E+07 7.78E+07

ROI (Return over investment) 33%/year

63

References

[1] Solvay.http://www.solvaychemicals.com/EN/products/chlorinated/Allylicproducts/

Epichlorohydrin.aspx. http://www.solvaychemicals.com/. [Online] May 2013.

[2] Guan, Eric Linak with Akihiro Kishi and Maoshi. July 2014. CEH Report Epichlorohydrin:

IHS Chemicals, July 2014.

[3]

BruceM. Bell, John R. Briggs,RobertM. Campbell,Susanne M. Chambers,Phil D.

Gaarenstroom,Jeffrey G. Hippler,Bruce D. Hook,Kenneth Kearns,John M. Kenney,William J.

Kruper,D. James Schreck,Curt N. Theriault,Charles P. Wolfe. 2008. Glycerin as a Renewable

Feedstock forEpichlorohydrin Production. The GTE Process. Midland USA : Clean Journal , 2008.

657-661.

[4] R. Tesser, *,† M. Di Serio,† R. Vitiello,† V. Russo,† E. Ranieri,† E. Speranza,† and E.

Santacesaria†. 2012. Naples, Italy : American Chemical Society, 2012.

[5] R. Vitiello, V. Russo, R. Turco, R. Tesser *, M. Di Serio, E. Santacesaria. 2014. Glycerol

chlorination in a gas‐liquid semibatch reactor: New catalysts for chlorohydrin production.

Naples, Italy : Chinese Journal of Catalysis, 2014. 663–669.

[6] R. Tesser, † E. Santacesaria,*,† M. Di Serio,† G. Di Nuzzi,† and V. Fiandra‡. 2007. Kinetics

of Glycerol Chlorination with Hydrochloric Acid:A New Route to a,γ-Dichlorohydrin.

Milano, Italy : American Chemical Society, 2007.

[7] Georgy Dmitriev S.*, Leonid Zanaveskin N. 2011. Synthesis of Epichlorohydrin from

Glycerol.Hydrochlorination of Glycero. Moscow : Karpov Institute of Physical Chemistry,

2011.

[8] R. Tesser, *,† M. Di Serio,† R. Vitiello,† V. Russo,† E. Ranieri,† E. Speranza,† and E.

Santacesaria†. Glycerol Chlorination in Gas-Liquid Semibatch Reactor:An alternative Route

for Chlorination Production Naples, Italy : American Chemical Society, 2012.

64

[9] Sang Hee Lee a, Sun Ho Song a, Dong Ryul Park a, Ji Chul Jung a, Ji Hwan Song b, Sung

Yul Wooc,Won Seob Song c, Myong Suk Kwon c, In Kyu Song a,. 2008. Seoul : Catalysis

Communications, 2008.

[10] XIUQUAN LING1, DINGQIANG LU1,2*, JUN WANG1, MINGXIN LIANG1,SHUMIN

ZHANG1, WEI REN1, JIANHUI CHEN1 and PINGKAI OUYANG1. 2010. Investigation of

the kinetics and mechanism of the glycerol chlorination reaction using gas chromatography–

mass spectrometry. Nanjing : Journal of the Serbian Chemical Society , 2010.

[11] S. Carra, E. Santacesaria and M. Morbidelli. 1979. Synthesis of Epichlorohydrin by

Elimination of Hydrogen Chloride from Chlorohydrins simulation of the reaction unit.

Milano : Ind Eng. Chemical Process Design, 1979.

[12] Anna Krzyżanowska, Eugeniusz Milchert, Marcin Bartkowiak. 2014. Dehydrochlorination of

1,3-dichloropropan-2-ol by calcium and sodium hydroxide solutions. Szczecin : Polish

Journal of Chemical Technology, 2014.

[13] L. Ma, J. W. Zhu, X.Q. Yuan and Q. Yue. SYNTHESIS OF EPICHLOROHYDRIN FROM

DICHLOROPROPANOLS Kinetic Aspects of the Process. Shangai, China : Institution of

Chemical Engineers, 2007.

[14] WO2006/100320 method for converting polyhydroxylated aliphatic hydrocarbons into

chlorohydrins

[15] US20130032755 A1 Continuous process for preparing chlorohydrins

[16] WO2008/152045 Epichlorohydrin manufacturing process and use

[17] Fogler, Scott. Elements of Chemical Reaction Engineering. Michigan : Prentice Hall, 1989.

[18] Smith, Robin. 2005. Chemical Process Design and Integration. Manchester : Wiley , 2005.

[19] Koukos, Ioannis. 2009. Design of Chemical Plants. Athens,2009.

[20] Gavin Towler, Ray Sinnot and Gavin. 2010. Chemical Engineering Design. Oxford : BH

Elsevier, 2010.

[21] J.M. Smith, H.C.Van Ness, M.M. Abbott. 2005. Introduction to Chemical Engineering

Thermodynamics. New York : McGraw Hill, 2005.

[22]

65

Thielemans, Geert L.B. 2014. Lecture Notes Process Design. Eindhoven, 2014.

[23]

Qun Yue, Jia-wen Zhu∗, Yan-yangWu, Xiang-qian Yuan, Liang Ma. Liquid–liquid equilibria and

vapor–liquid equilibria for the binary system of epichlorohydrin and water. Shangai : Fluid Phase

Equilibria, 2009.

[24] ICIS Pricing. Glycerol ASIA Price. 2014

[25] Tecnom OrbiChem. CHEM-NET FACTS. Chemical market insight and foresight.

EPICHLOROHYDRIN. 1 November 2013

[26] Milchert E., Goc W.; Lewandowski G.; Myszkowski J. (1995). Dehydrochlorination of

glycerol dichlorohydrin to epichlorohydrin. Chem. Papers 49(3), 133–136.

66

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