chemical engineering research and design 1 0 4 ( 2 0 1 5 ) 346–356Contents lists available at ScienceDirect Chemical Engineering Research and Design journal homepage: www.elsevier.com/locate/cherd Design and control of an improved acrylic acid process Xiaomeng Suo, Hao Zhang, Qing Ye ∗ , Xin Dai, Hao Yu, Rui Li Jiangsu Key Laboratory of Advanced Catalytic Materials and Technology, Institute of Petrochemical Engineering, Changzhou University, Changzhou, Jiangsu 213164, China a r t i c l e i n f o a b s t r a c t Article history: The acrylic acid (AA) process involves the partial oxidation of a flammable and explosive gas Received 17 May 2015 medium (propylene), so considerable attention is paid to the concentration of the reactants. Received in revised form 16 August Water and air is added to dilute the reactant and enhance the thermal stability of the reac- 2015 tion. Different compositions of water and air can lead to different methods to separate AA Accepted 24 August 2015 product. A modified AA process is developed based on the AA process proposed by Turton, Available online 1 September 2015 and consists of a tubular reactor, an absorber and only two distillation columns, one of which is an azeotropic distillation column. This process is characterized by shortening and simplifying AA refining process so that equipment investment cost can be reduced. A plantwide Keywords: Acrylic acid control structure featured with cascade control provides effective control for the multiunit Improved process process and insures safe operation of the two distillation columns. Dynamic results also Dynamic control reveal that it is useful to apply temperature/temperature cascade control and composition/temperature cascade control to the distillation columns with maximum temperature limitations. © 2015 The Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved. 1. Introduction Acrylic acid (AA) is widely used as an intermediate of chemicals and polymer in textile industry (Xu et al., 2006). There are several alternative processes to produce it, but the most common way nowadays is the partial oxidation of propylene (Lin, 2001). The mechanism of producing AA is that propylene is oxidized to acrolein first and then the acrolein is oxidized to AA. However, several side reactions occur, resulting in the oxidation of reactants and products. The main reaction and the typical side reactions are as follows: Main reaction: [o] [o] C3 H6 −→CH2 CHCHO−→CH2 CHCOOH Side reactions: [o] C3 H6 −→CO2 + H2 O [o] CH2 CHCHO−→CH3 COOH + CO2 ∗ There are two safety concerns about the AA process based on the reactions above. Firstly, the reactions involve partial oxidation of a flammable and explosive gas medium (propylene), so considerable attention must be paid to the compositions of propylene and oxygen in the feed stream to the reactor. The second safety concern is associated with the highly exothermic polymerization of AA (Cutie et al., 1997). AA of high concentration is dimerized at temperatures higher than 110 ◦ C, therefore the separation sequence must be operated under vacuum to keep the bottom temperature in the distillation columns below this temperature. In Turton’s design (2008) (Fig. 1), a fluidized-bed reactor is used, which is operated at isothermal condition and ensures safety and stability of the reaction with the addition of large amounts of water and air. The water inlet concentration of the reactor is quite large. As a result, liquid–liquid extraction is a proper way to separate water from the product stream. AA and acetic acid (ACE) are extracted into the organic phase. Corresponding author. Tel.: +86 519 86330355; fax: +86 519 86330355. E-mail address:
[email protected] (Q. Ye). http://dx.doi.org/10.1016/j.cherd.2015.08.022 0263-8762/© 2015 The Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved. That how to construct a plantwide control system using Aspen Dynamics has been well illuminated in Luyben’s book (2006). The separation of water is achieved using a 14-stage column. Then AA and ACE are separated to obtain the final AA product of high purity. The gas from the top of the absorber is composed of N2 .265 bar. The products of the reaction are cooled to 70 ◦ C immediately via a heat exchanger (HX) to avoided further oxidation reactions. The conditions in the figure are the optimal results which are analyzed in the later section of this paper. It does not need extraction column and the recycled system of extracting agent. Water is separated from the AA and ACE by using only one azeotrope column and the entrainer (toluene) is not carried into product stream. the compositions of the AA and ACE in the feed stream to the two distillation columns change. Process description The modified process shown in Fig. Then the stream leaving HX is fed to a flash drum. which ensures the use of cooling water in the overhead of the condenser. In this work. The bottom temperatures can be tightly controlled by manipulating reboiler duty.52 ◦ C. The purity specifications of the bottoms cannot be held when bottom temperature is set at a constant value. The reaction unit is somewhat simplified because it is assumed that two reactions take place in a single reactor to produce AA and other byproducts. and as a result. When the feed flowrate disturbances are introduced.chemical engineering research and design 1 0 4 ( 2 0 1 5 ) 346–356 2003). 2 produces 51. The separation process is much more simplified compared with Turton’s flowsheet. 1 – Turton’s flowsheet. Azeotrope column is operated at 0. The vital and primary safety consideration in this process is keeping the bottom temperatures in the columns below 110 ◦ C to prevent the highly exothermic polymerization of AA (Chen and McAvoy. the reactor performance changes. The organic phase is recycled to azeotrope column while part of the aqueous phase is fed to the absorber to recover AA and ACE. two distillation columns are required for the separation of the extractant from the organic phase and the aqueous phase. O2 . Meanwhile. The liquid streams from the bottoms of the flash tank and the absorber are fed to azeotrope column. Because of the introduction of the extractant. and as a result.094 ton per year of 99. Nomenclature AA ACE TOL HX Ns1 Ns2 Ntubes RR TAC VLL LLE PI 347 acrylic acid acetic acid toluene heat exchanger total stages of the azeotrope column total stages of De-ACE column the number of reactor tubes reflux ratio total annual cost vapor–liquid–liquid equilibrium liquid–liquid equilibrium proportional and integral settings 2. The operation hour is taken to be 8000 h/year. The top stage temperature is 50. Plantwide control should be noticed. the inlet water concentration of the reactor is cut down and air composition is significantly increased so that water composition in the out-let stream of the reactor is significantly reduced and a different way based on the azeotropic principle for dehydration can be considered. capital investment is reduced. The gas steam leaving the flash drum is fed to an absorber column in which AA and ACE are recovered by a steam coming from the decanter of azeotrope column. an effective control structure featured with cascade control is put forward to solve this problem. In this proposed design. It should be noted that the bottom temperature of the column Fig. but the purity specifications of the bottoms cannot be held easily at the same time.5% by mole AA product. The vapor withdrawn from the top is condensed and fed to a decanter in which organic and aqueous phases are formed. water. too. AA and a small amount of extracting agent (diisopropyl ether) get into the aqueous phase. the function of which is to separate water from the AA and ACE by the formation of water–toluene azeotrope using toluene as the entrainer. water and a small amount of propylene. . 2 – Improved flowsheet.a i Ei (kcal/kmol) ko.1 bar. 3(a) and (b). the function of which is to separate AA and ACE apart.36 0.02 14.i exp − RT p(C3 H6 )p(O2 ) Phase equilibrium The NRTL-HOC model phase equilibrium parameters.59 × 105 8.2. too. Kinetics and phase equilibrium 3.3 0. so the maximum temperature limitation is not violated. Aij (J/mol).99 ◦ C.76 0. And the bottom temperature is 108. The separation is achieved by using a 34-stage column and a reflux ratio of 79.265 bar with the pressure drop of 0.3 0. absorber.65 609. Table 1 – Reaction kinetics. The bottom product of azeotrope column is fed to the second distillation column (De-ACE column). The reaction kinetics for the catalyst are as follows. azeotrope column and De-ACE distillation column.759.89 42. Component i Aij (J/mol) Aji (J/mol) ˛ij H2 O + AA H2 O + ACE AA + ACE H2 O + TOL ACE + TOL 919.000 1. n xk kj Gkj k=1 n G x k=1 kj k .6. the bottom temperature will exceed the maximum temperature limitation at the condenser pressure of 0.2 0. The reactions taking place are kinetically controlled under the conditions used in the process. Gij = exp(−˛ij ij ).89 283. The main reason for toluene as the favorable entrainer can be explained Table 2 – The NRTL binary parametersa of the system.006895 bar through the trays.00 −293.57 −27. The preexponential terms and the activation energies for reaction (1)–(3) are given in Table 1.2 ◦ C.83 × 105 1.i 1 2 3 15.000 25. Partial pressures are in kPa. so inexpensive cooling water can be used in the condenser.269.3 a Values are calculated as follows: n ln i = j=1 n ji Gji xj G x k=1 ki k + n x Gij nj j=1 G x j=1 kj k ij − Where Aij −Aji ij = RT . 3.000 20.1.81 × 108 a Overall reaction rates have the units of kmol/m3 reactor h/(kPa)2 . If so many stages are used. 3.46 −723. must be kept below 110 ◦ C for safety concern. The ternary diagrams of AA–water–toluene and ACE–water–toluene is shown in Fig.00 0. The top stage temperature is 54. as is shown in Table 2.3 0.348 chemical engineering research and design 1 0 4 ( 2 0 1 5 ) 346–356 Fig. are used in the simulations of reactor. This column is operated at 0. 3 C3 H6 + O2 → CH2 CHCOOH + 3H2 O 2 (1) C3 H6 + 5 O2 → CH3 COOH + 3H2 O 2 (2) C3 H6 + 9 O2 → 3CO2 + 3H2 O 2 (3) E i −ri = ko. Kinetics The kinetic equations given by Turton are easily implemented in Aspen Plus because they are in the convenient power-law form. The bottom stream from the column contains 0. As a result. So toluene is a good choice as the entrainer for this system.98 MW.265 bar.8. which means that toluene. Effect of important design variables 4.1 bar generated by Aspen plus. however all these distillation columns exhibit high reflux ratios. which gives a heterogeneous azeotrope temperature of 38. A small amount of impurities (ACE. 5 gives the T–xy diagram of the AA/ACE system at a pressure of 0. in this work) can be removed by distillation. 4. The composition of the vapor from the top of the column is close to the composition of the toluene/water azeotrope.265 bar. . and the solid density is 2000 kg/m3 . a high reflux ratio is required.chemical engineering research and design 1 0 4 ( 2 0 1 5 ) 346–356 349 Fig. as the entrainer. Fig. Reactor The feed vapor stream is feed to a tubular reactor containing 16.14 wt% water. The pressure is 0. less entrainer of toluene is needed to be inside of the column. The Fig. The phase equilibrium in the AA process is quite nonlinear with a heterogeneous azeotrope formed in the binary system of toluene and water. The cooling duty is 24.8 ◦ C closed to the decanter temperature. The exothermic heat is removed by heat conduction oil cooling at 290 ◦ C. by the residue curve maps and the LLE of this system.. 0. is more capable of carrying water to the top of the column. Thus. Because of a small amount of light component (ACE) entering in the feed. 3 for the points between organic phase composition and azeotropic composition can be better and further apart. The heat-transfer coefficient in the tubular reactor is 1 kW m−2 K−1 (Guan et al.026 m in diameter and 10 m in length. The distance in Fig. the distillate rate of the De-ACE column is small. 6 shows the effect of the tube numbers (Ntubes) and coolant temperature on the production of AA and ACE. 4. The separation of ACE and AA in the De-ACE column is not easy. 3 – The ternary diagram for (a) AA–water–toluene and (b) ACE–water–toluene at 0. The valid phases in azeotrope column are VLL. 4 – T–xy diagram for toluene/water at 0.265 bar.000 tubes. as is shown in the T–xy diagram given in Fig. Fig. 2002). The void fraction of the catalyst in the tubular reactor is 0.1. The azeotropic compositions of toluene are better because this mixture contains much more water. 16. is withdrawn from the overhead.000 tubes are used to realize the same convention in Turton’s process. but it has maximum temperature limitation in its bottom too. The coolant temperature is set at 290 ◦ C. The production of ACE increases as coolant temperature increases. . So. The composition of water in the bottom stream is specified below 0. However.2.14 wt%. 4. 6 – Effect of Ntubes and coolant temperature of AA and ACE production. 5 – T–xy diagram for AA and ACE at 0.7 ◦ C and located at about 0. 1998). 4.350 chemical engineering research and design 1 0 4 ( 2 0 1 5 ) 346–356 Fig. TAC decreases as more trays are added while the bottom temperature increases. Fig. 8 shows the effects of Ns1 bottom temperature and the total annual cost (TAC) of this column (Douglas. which should be controlled when the reactor is in operation.1 bar. and the cost of azeotrope column is minimized at the same time. production of AA and ACE increase as more tubes are used but the increase is at an increasing smaller rate. The maximum temperature is 300. which can be achieved by increasing the number of theoretical plates (Ns1). the bottom temperature increases at the same time. It is a conventional distillation column. 7 gives the temperature profile of the tubular reactor. Here. So Ns1 is set at 14 so that the bottom temperature doesn’t exceed 110 ◦ C. De-ACE column Azeotrope column The function of this column is to separate water from the AA and ACE by the formation of water–toluene azeotrope which The function of column is to separate AA and ACE. where there exists a peak value of the production of AA. 9 shows the effects Fig. Fig.33 m down the 10-m reactor. Fig. the reaction temperature should be well controlled to ensure the largest production of AA.3. the set point of the bottom temperature controller is required. the 11th tray of azeotrope column is selected as temperature control stage when using reboiler duty as the manipulated variable. the bottom temperature increases with more trays being used. reflux drums and column bases of the two distillation columns and the absorber base are sized to provide 5 min of hold-up when at half full. The decanter is sized to provide 10 min of holdup when at the 50% level. A small change is made in the reboiler duty of the two distillation columns and reflux ratio of De-ACE column. 7 – Reactor temperature profile. Reboiler duty has the fastest and the most direct effect on bottom temperatures. Several methods about how to select the best control trays are well summarized in Luyben’s book and only steady-state information is needed. The reactor is simulated as a tubular reactor with constant coolant temperature. 8 – Effect of Ns1 on bottom temperature and TAC. the absorber. a sensitivity criterion is used to find the tray on which there is the largest change in temperature for a change in the manipulated variable. It is a 30-lump model and runs without no difficulty in Aspen Dynamics. 2. and two distillation columns. This dilemma is resolved by using 11th tray temperature/bottom temperature cascade control structure. so bottom temperature is controlled by manipulating reboiler duty. For the controls of the two distillation columns. optimal control trays are selected to hold the composition profile by adjusting manipulated variables. A temperature controller is added to control the temperature of the 11th tray by manipulating the set point of bottom temperature controller. 5. The tray with the largest temperature difference is considered to be the most “sensitive”. However. TAC of the De-ACE column decreases with the increasing of Ns2. Here. Fig. The size of the tubular reactor has been determined above. Fig. 10 shows the temperature profiles of azeotrope column and temperature differences of every stage by changing its reboiler duty. This structure is to change the set point of bottom temperature of azeotrope column and De-ACE column to control water concentration in its bottom. the flash tank. and the resulting change in temperature of all trays is obtained. Plantwide control In this part. It is necessary to determine the volumes of all vessels before it is converted to a pressure-driven dynamic simulation. The output signals generated by the tray temperature controller are controlled within a range from 106 ◦ C to 110 ◦ C so that the maximum . Because bottom temperature is not held constant. 9 – Effect of Ns2 on bottom temperature and TAC. The absorber and two distillation columns are simulated with radfrac model in Aspen. Fig. The flash tank. because reboiler duty cannot be manipulated in two independent temperature control loops at the same time. it precludes the direct use of the control stage when bottom temperature is used to manipulating reboiler duty. It also should be noticed that there exists maximum temperature limitation in the bottom of both distillation columns. However. bottom temperatures of azeotrope column must be tightly controlled. Therefore the total stages of De-ACE column is configured at 34 so that the bottom temperature is not violated and TAC of the De-ACE column is minimized at the same time. Therefore. dynamic control of this process is studied based on the steady-state design which has been determined above and shown in Fig. As a result. Fig.chemical engineering research and design 1 0 4 ( 2 0 1 5 ) 346–356 351 of Ns2 on TAC and bottom temperature. and it is often possible to achieve effective control. The units with dynamics include the tubular reactor. The flash tank is simulated with flash2 model in Aspen. The reason why dual-end control structure is necessary can be explained as follows. Another advantage of this control structure we proposed is that the reboiler duty is manipulated to maximum extent under the condition that maximum temperature limitation is not violated. and the design value of reflux ratio is indicated by yellow line.352 chemical engineering research and design 1 0 4 ( 2 0 1 5 ) 346–356 Fig. . temperature limitation is not violated. 11 shows temperature profiles of De-AC column and temperature differences of every stage generated by changing reboiler duty and reflux ratio by 0. it is not good to use the 3rd tray temperature-bottom temperature cascade loop like azeotrope column. Here. 11 – Temperature profile of De-ACE column and temperature differences of every stage. Because reflux ratio is quite large.1%. a dual control structure (two compositions. 12 – The effect of feed composition of ACE to reflux ratio of De-ACE column. The 3rd tray is the most sensitive to both reboiler duty and reflux ratio. However. Conventional distillation control wisdom suggests that a tray near the bottom of the column should be controlled to achieve effective inferential composition control of the bottom product by using a tray temperature. The 3rd tray temperature is controlled by manipulating reflux ratio. 10 – Temperature profile of azeotrope column and temperature differences of every stage. Reflux rate cannot be used to control the tray temperature directly. the level of reflux drum must be controlled by manipulating reflux rate. Plot of reflux ratio to ACE feed composition is presented in Fig. Changing reflux ratio influences the reflux rate by holding reflux level constant. There is a methodology to decide whether dual-end control structure is required. a tray near the bottom or using a Fig. and the 3rd tray temperature is controlled indirectly. The 3rd tray is far from the bottom and the strip section of the column. Reflux ratio changes significantly with ACE feed composition. Fig. or one of each) is required to handle feed disturbances. A steady-state simulation is run and AA and ACE purities are held constant with feed composition changing around the design value. Fig. 12. The purity of the bottom product is controlled by using a tray temperature. so dual-end control structure is necessary. two temperatures. If reflux ratio shows significant change to ACE feed composition. a dual-end control scheme is applied to hold the specifications of AA and ACE. and large deadtime or lag is introduced and poor dynamic performance cannot be prevented when the control loop is used. 8 MW 0.28 66 ◦ Fig. Temperature control is fast but it may not hold purity specifications. (2) Peak temperature of the tubular reactor is controlled by manipulating the temperature of heat conduction oil (reverse acting). (3) The pressure of the reactor is controlled by manipulating the discharge valve of the reactor (direct acting).00253–0 MW/mol 0. The set point ranges from 106 ◦ C to 110 ◦ C. The base levels of the absorber and two distillation columns are controlled by manipulating the flow rate of the stream leaving the bottom (direct acting). AA composition cannot be held when the temperature of the trays is held near the bottom or the bottom temperature is held constant by manipulating reboiler duty. The pressure in the azeotrope column is controlled by manipulating the flowrate of the vapor stream from the top (direct acting). The bottom temperatures of the azeotrope column and De-ACE column are controlled by manipulating the reboiler heat inputs of the two columns. The temperature of the decanter is controlled by manipulating the condenser duty (reverse acting). and composition–temperature cascade control structure is used in the bottom of De-ACE column. composition controller. The pressure in the absorber is controlled by manipulating the flow rate of the off-gas (direct acting).353 chemical engineering research and design 1 0 4 ( 2 0 1 5 ) 346–356 Table 3 – Parameters of temperature controller. The other various control loops are listed as follows: (5) (6) (7) (8) (9) (10) (1) Fresh feed of the mixture of propylene.7 250–300 ◦ C 0. The flow rate of the wash stream to the absorber is rationed to the gas feed coming from the flash tank. Organic reflux from the decanter is rationed to the azeotrope column feed.96 5. The liquid level of the flash tank is controlled by manipulating the stream leaving the bottom (direct acting).08 105–110 ◦ C 8. and it performs better than the direct temperature control. so maximum temperature limitation is not violated.1722 89.2 I (min) 3.3185 70 −0. and the ratio (condenser duty/mass feed flow rate of HX) is manipulated by the temperature of the (11) (12) (13) stream leaving HX (direct acting).16 0–0. The pressure in the flash tank is controlled by manipulating the gas leaving the top of the vessel (direct acting). However. respectively (reverse acting).8 MW 0. Controllers TC1 TC2 TC3 TC4 TC5 TC6 SP ( C) OP range Kc (gain) 300.796 108. according to the simulation we have run.28 5.1401 60.2 0–6. (4) The condenser duty of HX is ratioed to the feed flow rate of HX. 13 – Plantwide control structure. So composition control is considered. This feed forward control structure works well to control temperature changes of the flow leaving HX caused by fresh feed flow rate disturbances.02186 D/R ratio 212.7 0–7.2 5.28 13. . water and air is flow controlled (reverse acting). Composition control is slow but it can drive AA purity to the set point. The composition controller (reserve acting) is used to determine AA composition and generates set point signals for bottom temperature controller (reserve acting). This control structure combines the advantages of composition control strategy with temperature control strategy and avoids the defects of both strategies.0601 108. which is the Aspen default turning. (20) A deadtime of 1 min is inserted to all temperature control loops to fit measurement lags while a deadtime of 3 min is necessary for composition control loops. The level controls are only proportional with gain (Kc ) = 2. (17) Liquid level in the reflux drum of De-ACE column is controlled by manipulating reflux flowrate because the reflux ratio is quite large (direct acting) (Luyben. The .354 chemical engineering research and design 1 0 4 ( 2 0 1 5 ) 346–356 Fig. (18) Distillate flowrate of De-ACE column is rationed to reflux flowrate. The pressure controls of flash tank. 2000). 1996. (16) Organic level is controlled by manipulating the flow rate of toluene makeup (reverse acting). 2009). (15) Aqueous level is controlled by manipulating the flow rate of the aqueous stream (direct acting). Conventional proportional and integral (PI) settings are used in all control loops (Luyben. 14 – Dynamic response of the control structure to 20% feed flowrate disturbances. (14) The decanter pressure is controlled by manipulating the flow rate of the vent stream (direct acting). (19) The top pressure of De-ACE column is controlled by manipulating the condenser duty (reverse acting). decanter and two distillation columns are PI with Kc = 20 and I = 12 min. 15 – Dynamic response of the control structure to 10% feed composition disturbances. 13. The peak temperature of the reactor is . which consist of deadtime and PI parameters. Fig. 15 shows the response of the control structure to 10% feed composition disturbances with varying of air/propylene feed ratio. 14 gives the dynamic results for 20% step changes in set point of the fresh feed flow controller with orange lines for the increase and green lines for the decrease. dynamic performance of this control structure is evaluated by introducing feed flowrate disturbances and feed composition disturbances. Both of disturbances are introduced at 0. The plantwide control structure we proposed is shown in Fig.3 min. are obtained by running relay-feedback tests and using Tyreus-Luyben tuning rules. Now.16. flow control is PI with Kc = 0.chemical engineering research and design 1 0 4 ( 2 0 1 5 ) 346–356 355 Fig. in which orange lines for the increase and green lines for the decrease. PI parameters of temperature controllers are listed in Table 3. Fig.52 and I = 50. The composition controller is PI with Kc = 1.5 h. and the whole process reaches a new steady-state within 8 h.5 and 1 = 0. The temperature and composition control loops. . and a method for separating water from the product based on azeotropic principle is introduced. R. Catal. Luyben. Boiler Mag. References Chen. 14. the modified process is much more simplified so that equipment investment cost can be reduced.M. McAvoy.Y.L. C. Turton. J. New York. Large feed disturbances are effectively handled with bottom temperatures of two distillation columns and product purities holding close to the set values. W. Powell. New York. Xu. 1997. Acrylic acid polymerization kinetics.. Cutie. A: Gen.. Sci. X. Inc.E. 39. Tuning proportional-integral-derivative controllers for integrator/deadtime processes. Res. M. Tuning proportional-integral controllers for processes with both inverse response and deadtime. 207. New York. B 35.. 997–1008. Chem. Pearson Education. 4753–4771. Douglas.. T.. Res.M. P.. pp. Lin. 9–10. Shaeiwitz. X. Pen. Compared with that of Turton’s design.P. Pol. Chem. Whiting.S. Advances in the research and development of acrylic acid production from biomass.. 1996. Chem. Z.. McGraw Hill.B. 42. 48. . Unusual control structure for high reflux ratio distillation columns. 11048–11059. T.. Bottom temperatures of the two distillation columns are controlled below 110 ◦ C by using cascade control structure.. Henton.G. S.. Bailie. Eng. (411024) and assistance from the staff at the Institute of Petrochemical Technology (Changzhou University). Distillation Design and Control using Aspen Simulation. Chem. Res. Polym. Luyben.B. Smith. Synthesis and Design of Chemical Processes. Cen. Guan. first ed.L. Eng. J. Close control of the high-purity AA product is insured by using temperature/temperature and composition/temperature cascade controls. Luyben. Lin. and the purities of AA and ACE are held close to the set points. D. first ed. Appl. 2008. Conflict of interest The authors declare no financial interest. 2001. R. Chin.. Selective oxidation of propane to acrylic acid with molecular oxygen.. Ind.. Temperature transient deviations of the flow leaving HX are controlled within 1 ◦ C.. Luyben. Staples. 2002. Liu. 973–976.C. Ind. Eng. 1998. Plantwide control system design: methodology and application to a vinyl acetate process. J.. Acknowledgments We are thankful for support from the project fund of the China Petroleum & Chemical Corp.A.L. 6. Ind.L.L... 6. W. The effective dynamic control of the multiunit process has been achieved by adding various PI controllers. Analysis. R. Eng. Conceptual Design of Chemical Processes. Conclusion A modified AA process is developed based on Turton’s flowsheet.. Ind.356 chemical engineering research and design 1 0 4 ( 2 0 1 5 ) 346–356 well controlled.. Chem. Chem. 3480–3483. 2009...B. W. 419–427. 2003. 2006. J. 2000.. P. W. 1–16. W. T. Calculation for convective heat transfer coefficient of organic heat transfer material forced flowing in tube. Eng. 35. The component of water in the feed for reaction is significantly reduced. 2029–2047. J. 2006. Ind.L. third ed. Wiley.