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 GLY HCl GTE HOCl 1,3-DCP 2,3-DCP NaOH a-MCH b-MCH GLUA DCP PFD HPS LPS MPS AA CAPEX
Epichlorohydrin Glycerol Hydrogen chloride Glycerol to epichlorohydrin Hypochlorous acid 1,3-dichloro propan-2-ol 2,3-dichloro propan-1-ol Sodium hydroxide a-Monochlohydrin b-Monochlorohydrin Glutamic acid Dichloropropanols Process flow diagram High pressure steam Low pressure steam Medium Pressure steam Adipic acid 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).
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
GTE Process Design
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 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 Process Design
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
Glycerol
C3H8O3
99.9
$/ton 800
Hydrogen
HCl
100
360
NaOH
99.99
120
chloride gas Sodium hydroxide
GTE Process Design
Table of Contents 1.
INTRODUCTION......................................................................................................................................2
2.
Literature Review.......................................................................................................................................5
3.
4.
5.
6.
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
Aspen Simulation of GTE process............................................................................................................17 3.1.
Chlorination of GLY Reactor Selection............................................................................................18
3.2.
Dehydrochlorination of DCP............................................................................................................31
Equipment sizing and cost........................................................................................................................37 4.1.
Reactors and columns.......................................................................................................................37
4.2.
Pumps and Compressors...................................................................................................................38
4.3.
Heat exchangers and decanters.........................................................................................................39
Economic evaluation................................................................................................................................41 5.1.
ASPEN Economics Input..................................................................................................................41
5.2.
Results..............................................................................................................................................42
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
1
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. 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 2
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 chlorination and hypochlorination steps, resulting to the formation of unwanted chlorinated organics 5
(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.
7
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 inappropriate choice for large scale production due to significant losses at the reaction temperature (ca. 120 oC). 8
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 pK a value in relation to the performance of the catalyst was fortified. Specifically catalysts with pK a greater than 4 showed high selectivity to DCP, while those with less than 1.2 demonstrated high selectivity to a-MCH and bMCH. 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] Propionic acid has been discussed as another alternative [6] but with no good perspective because of its slightly lower volatility than AA.
9
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]. 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 2,3DCP for EPH production.
10
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.
11
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.
Figure 2.10 DCP yield as function of stirrer rotation speed [9].
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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]. Rate Constant k (min-1) 2.56 9.07 5.03 11.37
Reaction 1 2 3 4
2.2.
Activation Energy Ea(kJ/mol) 30.7 41.8 29.4 45.9
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.
2.
C3 H 6 Cl2 O+ NaOH yields C 3 H 5 ClO+ NaCl+ H 2 O →
C3 H 5 ClO+ NaOH + H 2 O yields C3 H 8 O 3 + 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]. MA et.al [13] showed that both reactions can be well represented by 2 nd 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. 13
For temperature in the range: 313 - 333 K
Main Reaction −¿ OH ¿ r 1=8.97∗1020 e
−123200 RT
[ DCP ] ¿
(2.3)
Side Reaction −¿ OH ¿ r 2=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
Production of EPH: 14
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 1,3DCP 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
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
16
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
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.
17
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 m 3, 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.
18
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].
SGLY /1,3 DCP
R1,3 DCP Ra MCH
k3Ca MCH k1CGLY
(3.1)
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.
19
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 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 Molar Flow (kmol/hr)
3 285
HCl-1 107.99
HCl-2 107.99
HCl-3 53.9
HCl4 15.02
Table 3.2 Stream 6 (GLY) splitting. Stream Molar Flow (kmol/hr)
6 140
GLY-1 70
GLY-2 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
20
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 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 non21
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 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.
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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.
24
W1
39(IN)
W2
PFRA2IN
B3
PFRA1OUT
B2
40(OUT)
B1 PFRB2OUT
38(IN)
41(OUT)
PFRB1OUT
HCL C-1 HE-1
B5
1HP
PFR1
HCL1
PFR1OUT
1HC
HIERARCHY
B2
HCL2
PRODUCTS HE-2 GLY
B3
P-1
GLY1 PFR2
S32
1HT
GLY2
PFR2OUT
HIERARCHY
V-1
PL
16 DC-2 DC-1
DICHL 17
HE-3
PLTL
19 18
HE-4
21
Figure 3.8 Water extraction from PFR reactors. 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.
25
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. Table 3.4 Injection points of HCl for the first PFR. Reactor 1
26
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 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
27
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 could be distributed evenly and it can even be extrapolated to other reactor size and other production requirements.
28
f(x) = NaN ln(x) R² = NaN
Normalized injection
f(x) = -0.02 ln(x) + 0 R² = 0.98
Reactor 1
Logarithmic (Reactor 1)
Reactor 2
Logarithmic (Reactor 2)
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 βMCH 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 Components Glycerol
Boiling points (°C) 289 29
HCL α -MCH β-MCH 1:3 DCP 2:3 DCP
HCL C-1 HE-1
B5
-85.05 213 220.35 174.3 184 PFR1
HCL1
PFR1OUT
1HP of these undesired 1HC The extraction products from the mix is doneHIERARCHY by employing two distillation HCL2
columns as seen in Figure 3.11. In the first column DC-1, water and traces of HCl are extracted from
B2
PROD
o
the product mix. The P-1 inlet stream B3 of the column stream 15 is at 130 C and at 1 bar. A partial GLY1 HE-2 condenser GLY
PFR2 has been column to flash remaining HCl before the condenser 1HT used on the S32 PFR2OUT (stream-16). GLY2
HIERARCHY
The bottom stream is heated up by HE-4 and is sent to the second column DC-2, the desired product V-1
of the first reaction, DCP, are retrieved on stream 20 with 99.5 % purity. PL
16 DC-2 DC-1
DICHL 17
HE-3
PLTL
19 18
HE-4
21
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
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
30
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
DC2
Bottom: 0,999 Recovery and purity of 1,3-DCP in Distillate: 0,99 Table 3.8 Dimensions of distillation columns
Column Parameters Reflux ratio Number of stages Feed stage
DC1 0.5 10 5
DC2 0.5 25 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 C1 C2
3.2.
Qc (MW) -0.80 -2.61
Qreb (MW) 1.22 0.60
Tc (°C) 98.88 172.52
Tr (°C) 174.52 215.75
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
31
lime milk by NaOH or catholyte. The waste water after the concentration and purification could be recycled to electrolysis [26].
RDISTILL TPROD
DC-2
B1 DICHL
DCHLFEED NAOH
BTM
21
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
Feed stage
5
Distillate rate Reflux ratio
365 kmol/hr 0.1
Reboiler duty
6.11 MW
Condenser duty
-4.7 MW
32
Temperature (feed)
63°C
Temperature (top)
88 °C
Temperature (bottom) Reaction zone (stage numbers) Time Residence
148.8 °C Stages 4-6 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 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 33
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 2 nd 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.
34
Figure 3.14 T-x-y EPH/Water diagram from experimental work [23]. Figure 3.16 shows the PFD of the 2 nd 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. Block DC-3: Composition Profiles
1,00
Liquid mole fraction ALPHA-01
0,95 0,90 0,85
Liquid Mole fraction of EPH
0,80 0,75 0,70 0,65 0,60 0,55 0,50 0,45 0,40 0,35 0,30
1
2
3
4
5
6
7
8
9
10
Stage Number
Figure 3.17 EPH liquid mole fraction per stage in unit DC-3.
35
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
36
Equipment Sizing and Cost
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 37
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
15.2
24.8
15.2
24.8
Cost (k$)
Table 4.2 Costs and sizing of distillation columns and reactive distillation (RDC-1). DCW-1
DC-1
DC-2
RDC-1
DC-3
DC-4
Diameter (m)
1.22
0.91
1.37
1.83
1.37
0.76
Number of trays
33
12
12
12
12
12
Tray spacing (m)
0.61
0.61
0.61
0.61
0.61
0.61
Height (m)
20.13
7.32
7.32
7.32
7.32
7.32
Reboiler duty (MW)
-2.86
-0.80
-2.61
-4.70
-1.72
0.55
Condenser duty MW
1.97
1.22
0.60
6.12
2.10
0.55
Equipment cost (k$)
586
520
557
724
557
442
38
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
Efficiency
0.8
Outlet pressure
5
Outlet temperature
kW
Electricity
1.5
kW
Flow
3
l/s
bar
Outlet pressure
5
bar
250
C
NSPHa
8.5
m
Isentropic outlet
207
C
Head
34
m
temperature 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
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. 39
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
504
128
162
658
-1485
213
1925
Energy / kW
-432
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
40
GTE Aspen Economics
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$/ton) was used to extract the price of the NaOH (aq) 41
30% mol (120$/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.
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. Table 5.2 Raw materials price Utility
Fluid
Electricity
Rate
Units
Cost per
Cost
690.5
kW
Hour 77.1
Units $/ hr
Cooling Water
Water
1065.6
m3/hr
33.7
$/ hr
Steam @690KPA
Steam
19.7
ton/hr
353
$/ hr
Steam @1135KPA
Steam
5.74
ton/hr
123
$/ hr
Steam @2760KPA
Steam
5.20
ton/hr
134
$/ hr
The summary of the economic results is the following:
42
Table 5.3 Raw materials price Parameter
Price
Total Capital Cost [M$]
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
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 $/ton [25], the return over investment (R.O.I) is 33%/year. Results of ASPEN economics for cash flows and detailed revenues and costs are found in appendix. Both EPH and GLY prices vary significantly over the time. In order to study the viability of the project a sensitivity analysis was performed. This analysis is shown in “Figure 5.1”.
43
Sensitivity analysis: GLY price
f(x) = 1.35x + 556.92 R² = 1
Figure 5.1 Sensitivity analysis - GLY price with RSP of EPH
Using the relation found on “Figure 5.1”, it is possible to predict the minimum required selling price for EPH. This process is economically viable when the price of GLY in the market is lower than 900$/ton. The price market trend for glycerol is shown in “Figure 5.2”.
44
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.
45
Conclusions
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 46
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.
47
Figure 1 Chlorination of GLY 3D reactor design.
48
Figure 2 Chlorination of GLY 3D reactor design.
Figure 3 Chlorination of GLY 3D reactor design.
PFR cost estimation – Guthrie method The free on board cost (f.o.b.) for each reactor is estimated by Equation (1):
C po ($@1968) 645.4 H 0.78 D 0.98
(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), CBM [( FBM 1) Fm FP ]Cop
(2)
Where FBM is a correction factor equal to 4.23, F m 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
49
bar. The calculated cost is for 1968 and by using Equation (3) the corresponding for 2015 can be found.
( M & S ) 2015 CBM , ( M & S )1968 CBM ,
(3)
Heat Integration of GTE Process-Pinch Point Analysis Pinch point analysis was performed on GTE process targeting to the reduction of hot (Q Hmin) 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, 50
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. H2
523 K
I-1
492 K
503 K 472 K
I-2 468 K 488 K I-3
472 K
452 K
I-4
468 K
448 K
I-5
446 K
426 K
I-6
433 K
403 K
I-7
413 K
393 K
H28
C18 H21
C14 I-8
403 K
383 K
I-9
383 K
363 K
I-
381 K
361 K
H33
C25
10 I-11 I-12
I-13
I-14
C31 363 K
343 K
H35 361 K
341 K
353 K
333 K
338 K
318 K
51
I-15
333 K
303 K
C4
Figure 4 Stream population of GTE process. 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]:
Qint erval
FCH ,i
hot streams ,i
FC int erval C,j hot streams , j int erval
(4)
Where FCH,i is the heat capacity of the hot stream i in the specific interval, FC C,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
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15
FCp,interval
2.1 -3.3 -10 -1.5 5.1 11 3.4 9 12.8 4.3 11.8 9.7 11 9.7 -2.5
ΔΤinterval (K)
31 4 16 4 22 13 20 10 20 2 18 2 8 15 5
Qinterval (kW)
Cascade
+66.7 -13.4 -160.9 -6.2 +113.3 +143.2 +67.9 +90 +256.5 +8.7 +214 +19.4 +88.1 +146 -12.9
ΣQj (kW) 0 +113.8 +66.7 180.6 +53.3 167.2 -107.6 6.2 -113.8 0 -0.56 113.3 +142.6 256.5 +210.6 324.5 +300.7 414.5 +557.2 671.0 +565.9 679.8 +779.9 893.8 +799.4 913.3 +887.5 1001.4 +1033.5 1147.4 +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 thermodynamics. The most negative value (-113.8 kW) is observed in the fourth row and thus in 52
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],
N HE ,AP NS,AP 1 4 1 3 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.
53
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.
QHE 1 FC p,C 25 g(Tin,C 25 Tout ,C 25 )
Tout ,C 25 469 K
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,
54
5.5(472 469) x(472 468) x 4.125 kW / K 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,
N HE ,BP N S,BP 1 9 1 8
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.
55
Figure 6 GTE heat integration below pinch point.
56
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 Cooling water (CW)
Price ($/tn) 0.06 Cost ($/tn) 3.5
Cp (kJ/kg K) 4.2 Δhvap (kJ/kg) 3.8
C) Medium Pressure steam (MPS,11 bar-
7
2.54
184 oC) Low Pressure steam (LPS,6 bar-160
18
2.15
High Pressure steam (HPS,42 bar-254 o
o
C)
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.
Qstream FC p stream g
Qstream mCW C p ,CW CW g
Qstream m steam hvap
(5) (6) (7)
Where FCp the flow heat capacity of the stream, m cw the mass flow of cooling water, m steam the mass flow of steam and Δhvap the latent heat of steam as given in Table 3.
57
Table 4 shows the calculation of the annual utility cost without heat integration of the process and it is estimated around 287 k$/year. By performing pinch point and combining hot and cold streams this cost can be reduced by ca. 56 % as demonstrating in Table 5.
Table 4 Annual utility cost before heat integration. Stream
Utility
Qstream (kW)
Utility Mass
Cost ($/year)
Flowrates H2 H21 H28 H33 H35 C4 C14 C18 C25 C31
(tn/year) 344 258962 662 497763 1139 855827 248 186247 10 7672 507 5898 76 885 134 1510 611 6890 84 980 Total annual utility cost M$
CW CW CW CW CW LPS LPS HPS HPS LPS
15537 29865 51349 11174 460 20644 3100 27197 124024 3432 0.287
Table 5 Annual utility cost after heat integration. Utility CW HPS
Qutility 1134.6 116.5
Utility Mass Flowrates (tn/year) 1703845 1312.6 Total annual utility cost M$
Cost ($/year) 102230 23626 0.12
58
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 column. The control of this equipment is otherwise very similar to a normal distillation column which is explained in the next section.
59
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 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
Figure 7 Proposed control scheme for GTE process. 61
DETAILED COSTS I (ASPEN DATABASE) ITEM TW (Number of Weeks per Period) T (Number of Periods for Analysis) DTEPC (Duration of EPC Phase) DT (Duration of EPC Phase and Startup) WORKP (Working Capital Percentage) OPCHG (Operating Charges) PLANTOVH (Plant Overhead) CAPT (Total Project Cost) RAWT (Total Raw Material Cost) PRODT (Total Product Sales) OPMT (Total Operating Labor and Maintenance Cost) UTILT (Total Utilities Cost) ROR (Desired Rate of Return/Interest Rate) AF (ROR Annuity Factor) TAXR (Tax Rate) IF (ROR Interest Factor) ECONLIFE (Economic Life of Project) SALVAL (Salvage Value (Percent of Initial Capital Cost)) DEPMETH (Depreciation Method) DEPMETHN (Depreciation Method Id) ESCAP (Project Capital Escalation) ESPROD (Products Escalation) ESRAW (Raw Material Escalation) ESLAB (Operating and Maintenance Labor Escalation) ESUT (Utilities Escalation) START (Start Period for Plant Startup) PODE (Desired Payout Period (excluding EPC and Startup Phases)) POD (Desired Payout Period) DESRET (Desired Return on Project for Sales Forecasting) END (End Period for Economic Life of Project) GA (G and A Expenses) DTEP (Duration of EP Phase before Start of Construction)
UNITS Weeks/period Period Period Period Percent/period Percent/period Percent/period Cost Cost/period Cost/period Cost/period Cost/period Percent/period Percent/period Period Percent
Percent/period Percent/period Percent/period Percent/period Percent/period Period Period Period Percent/Period Period Percent/Period Period
52 20 0.846154 0.929487 5 25 50 3.06E+07 1.41E+08 2.00E+08 2.37E+06 6.33E+06 20 5 40 1.2 30 20 Straight Line 1 5 5 3.5 3 3 1
10.5 30 8 0.326923
62
OP (Total Operating Labor Cost) MT (Total Maintenance Cost)
Cost/period Cost/period
2.06E+06 305714
CASH FLOW
1st year
2nd year
3rd year
4th year
5th year
6th year
7th year
8th 10
10th year
Period R (Revenue) DEP (Depreciation Expense) E (Earnings Before Taxes) TAX (Taxes) NE (Net Earnings) TED (Total Earnings) TEX (Total Expenses (Excludes
-4.50E+07 817172 -4.58E+07 0 -4.58E+07 -4.50E+07 5.98E+07
4.52E+07 817172 4.44E+07 1.78E+07 2.66E+07 2.75E+07 1.75E+08
5.02E+07 817172 4.94E+07 1.97E+07 2.96E+07 3.04E+07 1.81E+08
5.55E+07 817172 5.46E+07 2.19E+07 3.28E+07 3.36E+07 1.87E+08
6.11E+07 817172 6.03E+07 2.41E+07 3.62E+07 3.70E+07 1.94E+08
6.71E+07 817172 6.63E+07 2.65E+07 3.98E+07 4.06E+07 2.00E+08
7.36E+07 817172 7.27E+07 2.91E+07 4.36E+07 4.45E+07 2.07E+08
8.04E+07 817172 7.96E+07 3.18E+07 4.78E+07 4.86E+07 2.15E+08
8.77E+07 817172 8.69E+07 3.48E+07 5.21E+07 5.30E+07 2.22E+08
Taxes and Depreciation)) CF (CashFlow for Project) FVI (Future Value of Cumulative
-4.50E+07 1.48E+07
2.75E+07 2.38E+08
3.04E+07 5.16E+08
3.36E+07 8.62E+08
3.70E+07 1.29E+09
4.06E+07 1.82E+09
4.45E+07 2.46E+09
4.86E+07 3.25E+09
5.30E+07 4.20E+09
Cash Inflows) PVI (Present Value of Cumulative
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
Cash Inflows) PVOP (Present Value of
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
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
Cash Outfows) 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
Cumulative Cash Outfows Prod.)
ROI (Return over investment)
33%/year
63
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