The final stage is the process of product rectification. Enriched phase of ethyl acetate is fed to the rectification column, where, by distilling the triple azeotrope, a pure ethyl acetate is obtained as the bottom product [ 8 ].
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After azeotropic distillation, there is an extraction process. To obtain enriched phase of ethyl acetate from the distillation product, it is necessary to wash it with water and extract ethanol. The ethanol content of the organic mixture affects the solubility of water in this phase. The higher is the content of alcohol, the greater the water content in the organic phase. By extracting ethanol, water is also removed from the organic phase [ 8 ].
First is the process of rectifying the reaction mixture, i.e., stripping acetic acid by a rectification column and distilling the ethyl acetate&#;ethanol&#;water (EA&#;ET&#;W) ternary azeotrope. In order to generate an azeotrope by the distillation of reaction products, and to avoid accumulating water in the reactor, an azeotropic agent should be fed to the column. By esterification of the acetic acid with ethanol, some excess water is released (approximately 17%) [ 29 ] which is more than contained in the azeotrope mixture (approximately 7.8%). Therefore, the azeotropic agent is the organic phase in the extraction process, recycled to the column. This fraction is richer in ethyl acetate than the azeotrope (>93%) [ 8 ], so it binds the excess water.
Both membrane methods and extractive distillation to produce a high purity product require a combination with another method. This multiplies the disadvantages of the selected methods. Therefore, the classic method of separating the reaction mixture by distillation was chosen.
Complete dehydration with membranes does not require a lot of energy, but leaves a waste stream of 17 wt% ethyl acetate on the permeate side [ 28 ]. This waste stream requires additional treatment. Therefore, the pervaporation process is combined with the distillation process.
Pervaporation membranes can be used to selectively separate water and ethyl acetate to break down any azeotrope present. The energy consumption in this case is lower than in the case of classical extraction with water [ 27 ].
Membrane processes belong to a large group of techniques for separating the components of liquid and gas mixtures. The membrane technique used to separate azeotropic mixtures is called pervaporation [ 26 ].
Extractive distillation is a partial vaporization process in the presence of a non-volatile separating agent with a high boiling point, which is generally called solvent or entrainer, and which is added to the azeotropic mixture to alter the relative volatility of the key component with no additional formation of azeotropes [ 23 ]. The principle driving extractive distillation is based on the introduction of a selective solvent that interacts differently with each of the components of the original mixture and which generally shows a strong affinity with one of the key components [ 23 ]. The extraction agent after extractive distillation is purified in the recovery column and recycled to the process. Thanks to this it can work in a closed circuit, which reduces the amount of generated waste [ 24 ]. Typical examples of effective agents used in extractive distillation in the purification of ethyl acetate are: dimethyl sulfoxide, glycerin and diethyl glycol, 1-naphthol, hydroquinone and-dimethylformamide [ 25 ]. Due to the complex nature of the reaction mixture, in order to obtain a pure product, the extraction distillation has to be combined with another method, e.g., a membrane method [ 22 ].
The classic method of purifying the reaction mixture is a three-stage distillation in combination with a water extraction process. There are also other methods used to purify the reaction mixture such as extractive distillation and membrane methods.
Distillation is one of the most common methods of separating components of liquid mixtures. However, it is not always possible to separate a liquid mixture by simple distillation. Separation of the EA&#;ET&#;W mixture is one of such cases [ 21 ]. Ethyl acetate cannot be separated from water and ethanol by simple distillation, because a ternary azeotrope consisting of: 83.1% ethyl acetate, 8.7% ethanol and 8.2% water, is formed [ 22 ]. Therefore, the separation of reaction products is far more complex, in this particular case.
There are also other methods of synthesis, such as ethanol dehydrogenation and synthesis from ethylene and acetic acid. These methods require difficult process conditions (T > 200 °C,> 20 bar), which significantly increases the investment costs of the plant. Selectivity of the ethyl acetate synthesis from ethylene and acetic acid is almost 100% [ 17 ], but the efficiency of this process is not so high. Depending on the catalyst used, the maximum achieved conversion degree of this process ranges between 40% [ 18 ] and 50% [ 5 ]. On the other hand, many by-products are formed in the ethanol dehydrogenation reaction such as other esters, alcohols, aldehydes and ketones. The reaction mixture contains components with boiling points similar to ethyl acetate, including components which can form azeotropes [ 19 ]. This is a particular problem when high purity ethyl acetate is desired [ 20 ].
An alternative method is the Tishchenko reaction, in which acetaldehyde disproportionates in the presence of base to the alcohol and the acid that then esterify in situ [ 14 ]. The most common catalyst of this reaction are aluminum alkoxides. In this reaction, the obtainable yield of ethyl acetate by adding aluminum ethoxide to acetaldehyde at &#;20 °C is 61% [ 15 ]. However, this method is less popular than the classical Fischer esterification due to the availability and price of the raw material [ 16 ].
The process can also take place in the steam phase. The catalysts are then oxides of metals such as ZrO, TiO, Al, Fe, and the reaction is carried out at a temperature that allows the evaporation of the reactants [ 12 13 ]. In the case of ethyl acetate, the conversion rate of the reactants can be as high as 100%; however, depending on the catalyst used, by-products may be formed [ 12 13 ].
Ionic liquids can also be used as a catalyst for the reaction. In the esterification process, it is possible to carry out the reaction already at room temperature, while obtaining high yields. Another advantage of using this type of catalyst is the fact that the reaction product is eliminated from the reaction medium as a separate phase [ 9 ], thus shifting the reaction equilibrium towards the product. The homogeneous catalyst, which is an ionic liquid, can therefore be reused by returning it to the process after purification [ 10 11 ]. Due to their corrosiveness, ionic liquids are rarely used on an industrial scale to produce ethyl acetate.
The reaction of liquid phase esterification is reversible and insignificantly exothermic. The value of the equilibrium constant of the reaction depends primarily on the structure of alcohol and acid. In the reaction of acetic acid with ethyl alcohol, the equilibrium constant may vary in the range of 4&#;4.5, which corresponds to a conversion of 66&#;68% at a stoichiometric ratio of substrates [ 8 ].
Esters, including ethyl acetate, can also be synthesized in a number of other reactions using compounds such as acid anhydrides, acid chlorides, amides, nitriles, unsaturated hydrocarbons, ethers, aldehydes, ketones, alcohols and other esters (transesterification reaction) [ 6 ]. However, due to the relatively high price of raw materials and the possibility of undesired by-products forming by these pathways of alternative synthesis, the classic Fischer esterification is found to be the most commonly used reaction for the EA synthesis. The Fischer reaction of esterification of alcohols with carboxylic acids is carried out at elevated temperatures, in the presence of catalysts. Both homogenous and heterogeneous catalysts are employed [ 7 ]. The former are usually inorganic acids such as sulfuric and phosphoric acid, and hydrochloric acid, while the latter include acidic ion exchange resins. In the presence of the mentioned catalysts, the ethyl acetate synthesis is carried out under the temperature ranging between 70&#;90 °C. The temperature of the reaction mixture is controlled at a level assuring efficient removal of the vapors of the lighter key product in the esterification reaction [ 8 ].
Ethyl acetate (EA) in the chemical industry is obtained mainly by the classic Fischer esterification reaction, where the substrates are ethanol (ET) and acetic acid (AA), and water (W) is a by-product of the reaction [ 5 ].
In order to minimize the losses of raw materials and the consumption of auxiliary media, it is necessary to optimize the technology. This will both improve the economy of the process and minimize the environmental impact by reducing pollutant emissions. Our improved variant of the classic ethyl acetate plant includes additional subcooling of azeotrope, which improves extraction efficiency, making the organic phase enriched with ethyl acetate. This increases the efficiency of the other installation units, and reduces the heat duty of the plant and the emission of pollutants. A closed circulation of the extractant was also applied, which contributes to reducing the amount of wastewater.
Solvents, such as ethyl acetate, are the substances used in many industrial processes, including the chemical industry. Due to their wide range of application, world demand for solvents is constantly growing [ 1 3 ]. Increasingly stringent regulations on emissions of harmful pollutants from production processes make the importance of so called &#;green chemistry&#; (non-toxic chemicals to the environment and living organisms) well recognized [ 4 ]. There is, therefore, an urgent need for commonly used solvents to have a lower health and environmental impact. Ethyl acetate is characterized by low toxicity and, importantly, it is biodegradable. These advantageous features caused a significant increase in the market demand for this product of &#;green chemistry&#;. On the other hand, a sustainable development and good engineering practice are the drivers towards a lowered consumption of energy and utilities, as well as a reduced wastes production&#;an economy of the closed cycle. A technology improvement presented herein responds to all these environmental and economy requirements.
According to Atalay, the best fit of the experimental data is offered by the reaction rate described with the following equation:where:
The kinetics of the esterification reaction is also a subject of numerous studies [ 36 39 ]. Atalay [ 36 ] determined the activation energies and coefficients of the classic Arrhenius equation in the reaction of acetic acid esterification with ethanol, where sulfuric acid (VI) was used as a homogenous catalyst. He also studied the effect of catalyst concentration, temperature and the ratio of substrate concentrations on the reaction kinetics.
Numerous data sets regarding liquid&#;liquid and liquid&#;vapor equilibria for the ethyl acetate&#;ethanol&#;acetic acid&#;water systems, are reported in the literature [ 30 35 ]. This quaternary system exhibits non-ideal behavior with formation of azeotropes and with the possibility of formation of two liquid phases. Calvar studied the liquid&#;vapor equilibria for a system at atmospheric pressure [ 30 ]. He received a very good approximation of the experimental data, using the UNIQUAC thermodynamic model. Arce, in his work [ 31 ] studied the liquid&#;liquid equilibria. He also acheived the best fit using the UNIQUAC model. Therefore, for the modeling of either liquid&#;vapor or liquid&#;liquid equilibria, the UNIQUAC model is used in the modelling exercise.
The classic approach with the product and wastewater quality, resulting from the heat duties by rectification columns (SCDS 11, 15, Figure 1 ) equivalent to case 1, and the increased recycle of the azeotrope (Stream 27, Figure 1 ). The heat duty of the reactor (KREA 4, Figure 1 ) is increased to maintain the conversion rate.
The classic approach with product (Stream 16, Figure 1 ) and wastewater quality (Stream 25, Figure 1 ) maintained at a level comparable to case 1, resulting in the increased heat duties applied by the reactor (KREA 4, Figure 1 ) as well as rectification columns (SCDS 11, 15, Figure 1 ).
cooling the azeotrope before extraction (Streams 8 and 14, Figure 1 ) from 70 °C down to 30 °C (HTXR 6, 12, Figure 1 ), shifting a phase equilibrium towards higher concentrations of EA in the organic phase, resulting in a reduced reflux on the azeotropic column as well as the reduced flowrate of the water phase directed to the wastewater treatment plant.
a closed circuit of the extractant, demineralized water, reducing the consumption of the fresh extractant, but also a deep recovery of raw materials, i.e., ethyl acetate and ethanol from wastewater, significantly reducing the TOL;
The impact of key parameters on the performance of individual installation nodes was examined by sensitivity analysis. On the basis of the analysis, the optimal conditions for the process were selected.
The aqueous phase after extraction (Stream 20, Figure 1 ), containing ethanol and ethyl acetate is directed to the wastewater column (SCDS 15, Figure 1 ), where ethyl acetate and ethanol are recovered. The distillate (Stream 21, Figure 1 ) is recycled to the reactor and the bottom product (Stream 22, Figure 1 ) is recycled as a washing water to the extraction process, partially refreshed with a fresh portion of water (Stream 17, Figure 1 ).
The organic phase from the extraction process is partly recycled (Stream 26, Figure 1 ) to the azeotropic column (SCDS 5, Figure 1 ) as reflux, while the remainder (Stream 12, Figure 1 ) is directed to the product rectification column (SCDS 11, Figure 1 )&#;final product distillation. In this column, pure ethyl acetate is obtained as the bottom product (Stream 16, Figure 1 ), and the subcooled distillate (triple azeotrope) is recycled to the extraction process (EXTR 8, Figure 1 ).
Acetic acid and ethanol are mixed (MIX 1, Figure 1 ), heated up (HTXR 2, Figure 1 ) and directed to the reactor. The esterification reaction takes place in the reactor&#;evaporator (KREA 4, Figure 1 ) at a temperature of about 90 °C, under atmospheric pressure. The vapors from the reactor (Stream 6, Figure 1 ) are directed to the column (SCDS 5, Figure 1 ), in which unreacted acetic acid is separated from the product by azeotropic distillation&#;the first stage of EA purification. The bottom product (Stream 7, Figure 1 ) is recycled to the reactor&#;evaporator. The distillate (Stream 9, Figure 1 ) is a triple EA&#;ET&#; W azeotrope. The azeotrope is washed with water (EXTR 8, Figure 1 ) to extract ethanol&#;the second stage of EA purification. In case of the improved approach, the extractor feed is cooled down to enhance extraction&#;advantageous shift of extraction equilibrium (HTXR 6, Figure 1 ).
The main goal of this paper is to optimize the process conditions of synthesis and purification of ethyl acetate. A study was carried out based on a model built with the use of flowsheeting software Chemcad 7. The shame of the modeled installation is shown in Figure 1
3. Results
3.1. Optimization of the Reaction System
The continuous stirred tank reactor CSTR was used to model the reaction system (KREA 4, Figure 1 ). The key parameters affecting the degree of conversion are the temperature and the residence time of the reaction mixture in the reactor. The influence of these parameters on the composition of the post-reaction mixture was examined using the sensitivity analysis tool. The analysis assumed a constant composition of the reactor input stream. The results of the analysis are shown in Figure 2
The results obtained by simulation prove that the temperature increase in the reactor positively affects the EA content in the output stream, but only up to a certain point. The highest degree of reagent conversion was obtained at 85 °C. Above 85 °C, the EA content drops significantly. This is due to the increased and excessive evaporation of ethanol from the reaction system, which shifts the equilibrium of the reaction towards the substrates. Apparently, the optimum residence time at elevated temperatures should be kept in the range of 3&#;4 h.
3.2. Optimization of Azeotropic Distillation
The azeotropic column (SCDS 5, Figure 1 ) is used to strip acetic acid from the reactor vapors and for distillation of the EA&#;ET&#;W triple azeotrope. The reactor vapors are fed to the last stage of the column, while the reflux (Stream 27, Figure 1 )&#;a distillate washed with water&#;is fed onto the first stage. The obtained distillate is an azeotrope&#;with a mass composition of 83.1% ethyl acetate, 8.7% ethanol, 8.2% water&#;which is directed to the water extraction process. The bottom product from the column (Stream 7, Figure 1 ) containing the stripped acetic acid is recycled to the reaction node.
Azeotropic distillation is limited by the amount of acetic acid in the distillate. The product of appropriate quality should not contain more than 0.005 wt%. To assure the required contents of AA in the distillate, both reflux ratio (R/V) and the number of stages were adjusted. The results are shown in Figure 3
The acetic acid content strongly depends on the amount of the reflux applied to the column. The required acidity (<0.005 wt% of acetic acid) can be obtained with an R/V ratio greater than 0.5 and with a minimum number of 14 theoretical stages.
3.3. Optimization of the Extraction System
The extraction node (EXTR 8, Figure 1 ) is found to be one of the most important installation nodes with regard to the overall process performance. Analysis was focused on the temperature of the azeotrope stream as well as the mass ratio of water:azeotrope (W/A), both affecting the composition of the organic phase (richer in ethyl acetate) and losses of the EA with wastewater. The results are shown in Figure 4
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It was observed that lowering the temperature has a positive effect on the content of ethyl acetate in the organic phase. A temperature drop of every 10 °C increases the weight fraction of ethyl acetate in the organic phase by about 0.15%. Lowering the temperature from 50 °C to 20 °C increases the EA content by approximately 0.45%. This relatively small change has a significant impact on the amount of reflux recycled to the azeotropic column as well as on the amount of azeotrope recycled from the product distillation column. It is therefore profitable to deeply cool down the azeotrope before extraction. Yet, due to the fact that the process is limited by the cooling water temperature, the azeotrope can be cooled down to 20&#;30 °C.
The content of ethyl acetate in the organic phase increases with an increased amount of water during rinsing. However, by increasing the water flowrate, the ethyl acetate losses are bigger. The losses of EA also increase with the temperature. Furthermore, the increased flowrate of the rinsing water causes an excessive amount of wastewater to be produced, and by this, the cost of raw material recovery by wastewater distillation increases. Analyzing the modelling results, it can be concluded that the optimum mass ratio of water to azeotrope is in the range between 1 and 2.
3.4. Optimization of Product Distillation
Organic-phase, enriched ethyl acetate is fed to the rectification column (SCDS 11, Figure 1 ). The purification of ethyl acetate involves distilling off the triple azeotrope consisting of water, ethanol and ethyl acetate. The distillate is recycled to the extraction process. The bottom product is ethyl acetate with required purity greater than 99.9%.
First, the effect of the reflux ratio (R/D) on the composition of the bottom product as well as reboiler duty were examined at a set, rigid product rate. The results are presented in Figure 5 . Next, at a set reflux ratio, the influence of the feed stage location and the number of stages on the product quality as well as evaporator/reboiler heat duty were analyzed ( Figure 6 ).
The amount of reflux has an obvious key impact on the product quality. By increasing the reflux, the purity of the product increases significantly. The required EA content in the product (99.9 wt%) can be obtained for minimum R/D equal to 1.5. The limit parameter in this case is also the water content, which should be less than 0.03 wt% and an ethanol content that should be less than 0.04 wt%. These parameters are met for R/D greater than 2.25. The optimum R/D ratio is in the range 3&#;3.5. A higher reflux does provide an insignificant increase in the product yield, while heat duties on both the evaporator and the condenser increase proportionally.
The product quality is also influenced by both the EA feed location and the number of column stages. As presented in Figure 6 , the number of stages has a significant impact on the product purity. However, the feed location is also crucial. The best results were obtained when the feed was located between the 6th and 4th column stage. The introduction to the higher or lower stages of the column resulted in a deterioration of the product quality.
3.5. Optimization of Wastewater Distillation
From an ecological and economic point of view, the quality and the quantity of the sewage produced is also of great interest. The aqueous phase after extraction contains a large share of ethanol and ethyl acetate (over 15 wt%). Therefore, it is directed to a sewage column, where the organics are separated from the water. Sewage processing is meant to reduce the loss of raw materials, fresh water consumption and improve wastewater quality.
A mixture of ethyl acetate, ethanol and a small amount of water is collected as a distillate in the column. This stream is recycled to the reactor. Water is obtained as the bottom product, which is partly recycled to the extraction process. Only a small amount of cleaned wastewater is released as sewage.
By modeling the wastewater distillation column (SCDS 15, Figure 1 ), the effect of reflux ratio on the distillate composition and the bottom product, wastewater was examined. It was assumed that the bottom product must contain less than 0.01% of its weight in organic compounds. The influence of the reflux ratio on heat duty of the reboiler was also analyzed. The results are shown in Figure 7 a,b.
The distillate from the column is recycled to the reactor (Stream 21, Figure 1 ), so its composition affects the reaction equilibrium. The more water we returned with the stream, the more reflux we need on the azeotropic column. It is therefore important to keep the water concentration as low as possible. Due to the fact that the three-component EA&#;ET&#;W system creates the azeotrope with the lowest boiling point, we are not able to remove water completely.
The amount of reflux is crucial to the composition of the distillate. By increasing the amount of reflux, the content of EA and ethanol in the distillate increases, and the water content decreases. With an increase in reflux (R/D) in the range from 0.1&#;1.5, you can see a very large decrease in the water content in the distillate. With an R/D ratio greater than 1.5, no major changes can be seen.
The effect of the feed location and the number of stages on the composition of distillate and the bottom product as well as reboiler heat duty were examined at a set reflux ratio 1.5 ( Figure 7 c,d).
The number of stages and the feed location affect the composition of the distillate and the heat duty of the evaporator. Analyzing the graphs in Figure 7 c,d, it was found that the column must have reached at least seven stages in the stripping section and seven in the reinforcing section.
A multicomponent model system is presented in this work. The basic model system contains ethyl acetate (EtAc), ethanol (EtOH), water (HO), and acetic acid (AA); one case study is extended by ethylene oxide (EO) and monoethylene glycol (MEG). This system is strongly non-ideal. Two homogeneous binary azeotropes (EtOH&#;HO, EtOH&#;EtAc), one heterogeneous binary azeotrope (EtAc&#;HO), and one homogeneous ternary azeotrope (EtAc&#;EtOH&#;HO) are reported in the literature [ 35 36 ] and databases [ 26 ]. Despite acetic acid forming no azeotropic mixture with other participating compounds, it is known for its strong association in the vapor phase and the formation of dimers [ 37 ]. Monoethylene glycol does not form an azeotropic mixture with other mentioned components. Unlike other components, ethylene oxide is a gas at room temperature [ 38 ]. For such a system, the NRTL-HOC thermodynamic model is highly recommended [ 15 28 ] as it is capable of calculating the VLLE (vapor&#;liquid&#;liquid phase equilibria) correctly including two liquid phases, azeotropic mixtures composition, and boing points, dimerization in the vapor phase. All the above-mentioned papers have shown simulation results to be in good agreement with experiment data. Moreover, reliable parameters for the NRTL-HOC model can be obtained from available databases (Aspen Plus [ 26 ], DECHEMA, NIST).
As it was mentioned, an auxiliary chemical reaction was included to enhance the reactive distillation process in the last case study. This auxiliary reaction was chosen to ensure the removal of the main esterification reaction by-product&#;water. Ethylene oxide (EO) hydration was used, Equation (7). Monoethylene glycol (MEG) is produced as the main product of the auxiliary reaction. Further reactions towards higher glycols (diethylene glycol, triethylene glycol) were omitted because of lower reaction rate and negligible change in the composition of product streams [ 23 41 ]. The auxiliary reaction rate is expressed by Equation (8) [ 23 41 ].
In this work, homogeneous catalysis using sulfuric acid was assumed, Equation (5). The reaction rate ([kmol m]) is expressed by Equation (6), which has been used in works [ 23 40 ]. The reaction occurs in the liquid phase; liquid phase molar concentrations ([kmol m]) are used. The concentration of sulfuric acid catalyst was low [ 15 25 ], so its presence in the phase equilibria calculation was neglected.
Mechanism and reaction kinetics of the esterification reaction of ethanol and acetic acid in the presence of acid catalysts has been studied in many works [ 1 40 ]. Three types of catalysis have been reported: autocatalysis, homogeneous catalysis, and heterogeneous catalysis. The reaction rate is low and achievable conversion is up to 20% at high residence times in the case of autocatalysis reaction [ 39 ]. Strong mineral acids, such as sulfuric acid and hydrochloric acid, are traditionally used as homogeneous catalysts [ 15 ]. High conversion, of up to 65.5%, is achieved in industrial applications using sulfuric acid [ 1 ] in the range from 0.2 to 1.0 volume percent of the reactive mixture [ 15 23 ]. Acidic ion exchange resins in various forms have been used as heterogeneous catalysts to increase conversion (slightly below 70%); however, no solid heterogeneous catalyst has been found to increase the reaction rate in favor of ethyl acetate production better than sulfuric acid [ 15 ]. High reaction rate, high conversion, and a smaller amount of catalyst are preferred from the industrial point of view. In addition, process set-up and equipment design are much easier in the case of homogeneous catalysis. On the other hand, equipment corrosion and catalyst recycling are major drawbacks of homogeneous catalysis with sulfuric acid.
2.3. Equipment Model
Aspen Plus V10 simulation environment provides several options to compile a process model. In this work, three main types of equipment models were used: chemical reactor, heat exchanger, and distillation column/reactive distillation column.
A chemical reactor is simulated by a model of continuous stirred tank reactor (CSTR) which assumes ideal mixing along with rate-controlled chemical reaction based on known kinetics. The reactor can be operated as an isothermal as well as an adiabatic one. The residence time parameter was used to achieve the desired conversion. Valid phases (liquid, vapor, vapor&#;liquid) for the chemical reaction were specified; in case of esterification (5), the chemical reaction rate is expressed by Equation (6) and takes place only in the liquid phase.
Heat exchangers were simulated by the Heater and HeatX models, respectively. A shortcut set-up was applied to reach the desired stream temperature. Heat integration was applied to improve the optimal process design. The minimum stream temperature difference was set to 10 °C.
N
), reactive zone (
NR
), feed stage position (
f
), reflux ratio (
R
). These parameters can be found in the literature [
A Rigorous RadFrac column model was used for RD modeling as well as for conventional distillation. This model allows both EQ and NEQ approaches. Building an NEQ model of reactive separation or separation is not as straightforward as it is in the EQ model. The NEQ model requires much more reliable parameters compared to the EQ model. Consequently, the NEQ model is more difficult to calculate and convergence problems often occur. To improve NEQ model convergence, a good initial guess of stage temperature, liquid phase composition, and vapor phase composition have to be used. For this purpose, the EQ model of each column was made in the first step of the simulation using initial column parameters such as the number of theoretical stages (), reactive zone (), feed stage position (), reflux ratio (). These parameters can be found in the literature [ 5 15 ]. Results of the EQ model simulation provide a very good starting point for building the NEQ model [ 19 27 ]. Another advantage of first building the EQ model is the possibility of faster testing of individual case studies [ 28 ]. When the suitable case study concept is selected, the NEQ model is built based on the EQ model results.
d
), packing height (
H
), packing dimensions) were set during the calculation procedure with regard to reasonable column hydraulics, pressure drop, and approach to flood.
Rate-based set-up must be enabled in the Aspen Plus [ 26 ] in the NEQ model. Therefore, detailed column internal configuration is required next. A packed column is selected similarly to simulation-experimental works [ 22 28 ]. Mass and heat transfer correlation methods were selected according to the recommendation for the packing type (Rashing Ralu-Ring). Column hydraulics was simulated by Aspen Plus built-in hydraulic function assuming correlation for Rashing Ralu-Ring packing type. Column internal configuration (internal diameter (), packing height (), packing dimensions) were set during the calculation procedure with regard to reasonable column hydraulics, pressure drop, and approach to flood.
NR
) only, where acetic acid is presented. On the other hand, the hydration reaction (7) is enabled in the whole RD column because ethylene oxide reacts with water whenever they meet in the liquid phase. All the above-mentioned column parameters (
N
,
NR
,
f
,
R
,
d
,
H
&#;) were optimized to meet the design criteria of individual columns as well as of the whole process.
In the case of an RD column, the chemical reaction rate is expressed by Equations (6) and (8). A homogeneous catalyst (sulfuric acid) is fed to the column together with acetic acid [ 17 23 ]. The esterification reaction (5) is enabled in the reactive zone () only, where acetic acid is presented. On the other hand, the hydration reaction (7) is enabled in the whole RD column because ethylene oxide reacts with water whenever they meet in the liquid phase. All the above-mentioned column parameters (&#;) were optimized to meet the design criteria of individual columns as well as of the whole process.
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