RECYCLE STRUCTURE OF THE FLOWSHEET CHAPTER 6 RECYCLE STRUCTURE OF THE FLOWSHEET

6.1 DECISIONS THAT DETERMINE THE RECYCLE STRUCTURE The decisions that fix the recycle structure of the flowsheet are listed in Table 6.1-1.

TABLE 6.1-1 Decision for recycle structure How many reactor system are required? Is there any separation between the reactor system? How many recycle stream are required? Do we want to used an excess of one reactant at the reactor inlet? Is a gas compressor required? What are the cost? Should the reactor operated adiabatically, with direct heating or cooling, or is a diluent or heat carries required?

TABLE 6.1-1 Decision for recycle structure (continuous) Do we want to shift the equilibrium conversion? How do the reactor costs affect the equilibrium potential?

Number of reactor Systems For example HDA process. } Toluene+H2  Benzene+CH4 2Benzene  Diphenyl+H2 1150-1300F, 500 psia (6.1-1) Both take place at the same T. and P. without catalyst. one reactor requied.

Number of reactor Systems The anhydride process } Acetone  Ketene+CH4 Ketene  CO+1/2C2H4 700C, 1 atm (6.1-2) Ketene + Acetic Acid  Acetic Anhydride 80C, 1 atm  two reactor requied.

Number of recycle stream From the anhydride process Acid feed Reactor R1 Reactor R2 Acetone feed Acid recycle Acetone recycle FIGURE 6.1-1 Acetic anhydride

-List of all components leaving the reactor that has been ordered normal boiling points. e.g., Table 5.1-4 -List the reactor number as the destination code for each recycle stream. Next group recycle component having neighboring boiling points if they have the same reactor destination. Do not separate two components and then remix them at a reactor inlet. (6.1-3)

Example 6. 1-1 Number of recycle stream Example 6.1-1 Number of recycle stream. Consider the components and the destinations given below in order of their normal boiling point: Primary product Reactant-recycle to R2 Reactant-recycle to R1 Valuable by-product Waste by-product Reactant-recycle to R1 Fuel by-product 4 product stream A+B, D+E, F and J 3 recycle stream C and I (go to R1), G+H (go to R2)

TABLE 6.1-2 HDA process. The component and their destination for the HDA process are as follows: NBP, C Destination H2 -253 Recycle + purge-gas CH4 -161 Benzene 80 Primary product Toluene 111 Recycle-liquid Diphenyl 255 Fuel by-product Product stream3 Purge Benzene 3) Diphenyl Recycle stream2 H2 and CH4 Toluene

A recycle flowsheet is given in Fig. 6.1-2 Purge Compressor H2 feed Benzene Reactor Separator Toluene feed Diphenyl Toluene recycle FIGURE 6.1-2 HDA recycle structure

TABLE 6. 1-3 Anhydride process TABLE 6.1-3 Anhydride process. The component and their destination for the anhydride process are given: Component NBP, F Destination CO -312.6 Fuel by-product CH4 -258.6 C2H4 -154.8 Ketene -42.1 Unstable reactant-completely converted Acetone 133.2 Reactant – Recycle to R1-liquid Acetic acid 244.3 Reactant – Recycle to R2-liquid Acetic anhydride 281.9 Primary product Liquid-recycle stream2 Acetone Acetic Acid Product stream2 CH4 + CO + C2H4 anhydride

FIGURE 6.1-3 Acetic anhydride recycle Acid feed CO, CH4, C2H4 Anhyd. Reactor R1 Reactor R1 Reactor R2 Acetone feed Acetic Acid recycle Acetone recycle FIGURE 6.1-3 Acetic anhydride recycle

Excess reactants In some cases the use of an excess reactant can shift the product distribution. For example, the production of isooctane via butane alkylation as: Butene + Isobutane  Isooctane Butene + Isooctane  C12 6.1-4 If the kinetics match the stoichiometry, the use of excess of isobutane leads to an impoved selectivity to produce isooctane. The larger the excess, greater the improvement in the selectivity, but the larger cost to recover and recycle isobutane. Thus, an optimum amount of excess must be determined from economic analysis.

The use of an excess component can also be used to force another component to be close to complete conversion. For example, the production of phosgene CO +Cl2  COCl2 6.1-5 Which is an intermediate in the production of di-isocyanate, the product must be free of Cl2. Thus, an excess of CO is used to force Cl2 conversion to be very high.

Similarly, an excess can be used to shift equilibrium conversion For example, the production of cyclohexane Benzene + 3H2  cyclohexane 6.1-6 Molar ratio of reactant inlet is often a design variable

6.2 RECYCLE AND MATERIAL BALANCE Our goal is to obtain a quick estimate of the recycle flow. We have not specified any detail of separation system as yet, and therefore we assume that greater than 99% recoveries of reactants are equivalent 100% recoveries.

Limiting Reactant For HDA process.(Fig. 6.2-1) Purge H2, feed Benzene,PB Separation system Reactor R1 FT(1-x) FT Diphenyl FFT FT(1-x) FT(1-x) Toluene feed Acetone recycle FIGURE 6.2-1 HDA, liquid recycle

Thus,the feed to the reactor is Limiting Reactant FFT + FT(1-x) = FT (6.2-1) Thus,the feed to the reactor is (6.2-2)

In some case, some of limiting reactant might leave the process in a gas recycle and purge stream (ammonia synthesis), or it may leave with the product (ethanol synthesis). If we consider a simplified version of the ethanol process, the reaction are (6.2-3)

FIGURE 6.2-2 Ethanol synthesis If we want to produce 783 mol/hr of an EtOH-H2O azeotrope that contains 85.4 mol% EtOH, from an ethylene feed stream containing 4%CH4 and pure water C2H4, CH4 C2H4, CH4 EtOHH2O Aceotrope Separation system Reactor H2O Ether H2O FIGURE 6.2-2 Ethanol synthesis

Overall material balances start with the production rate of the azeotrope (6.2-4) This contains (6.2-5) The amount of water in the product stream is (6.2-6)

Thus, from Eq. 6.2-3 and the result above ,the required feed rate of water, which is the limiting reactant, is (6.2-4)

Suppose that we let the water leaving with the product be Fw,p = 114 and the fresh feed water required for the reaction be Fw,R. Now refering to the schematic in Fig. 6.2-3 for water FP+FR Reactor Separator FP F F(1-x) F(1-x)-FP (6.2-8) (6.2-9) We let the amount entering the reactor be Fw, the amount leaving the reactor be Fw(1-x), the amount leaving with the product be Fw,P, and amount recycled be Fw(1-x)-Fw,P Then a balance at the mixing point before the reactor gives

Other Reactants :For example , HDA process (6.2-10) (6.2-11) H2, feed 95%H2, 5% CH4 FG Purge, H2 CH4 RG ,yPH Benzene,PB MR Separation system Reactor R1 FT Diphenyl Toluene feed FIGURE 6.2-4 gas recycle

Design Heuristics For the case single reactions, choose x= 0.96 or x=0.98xeq as a first guess. This rule of thumb is discussed in Exercise 3.5-8 (6.2-12)

Reversible By-products If we recycle a by-product formed by a reversible reactions and let the component build up to its equilibrium level. Such as the diphenyl in the HDA process. 2Benzene  Diphenyl+H2 Or the diethylether in ethanol synthesis (Eq.6.2-3) We find the recycle flow by using the equilibrium relationship at the reactor exit. (6.2-13)

6.3 REACTOR HEAT EFFECTS Reactor Heat Load For the single reaction (6.3-1)

Example 6.3-1 HDA process. (6.3-1) If we want to estimate the reactor heat load for a case where x=0.75, PB=265, and FFT=273 mol/hr, we might neglect the second reaction and write (6.3-1) Where HR is the heat of reaction at 1200 F and 500 psia and heat is liberated by the reaction.

Example 6.3-2 Acetone can be produced by the dehydrogenation of isopropanol (CH3)2CHOH(CH3)2CO+H2 (6.3-2) If we desire to produce 51.3 mol/hr of acetone and 51.3 mol/hr of IPA is required. The heat of reaction at 570 F and 1 atm is 25,800 Btu/mol, So that the reactor heat load is (6.3-3) Heat is consumed by the endothermic reaction.

Adiabatic Temperature change Estimate the adiabatic Temp. change from the expression: (6.3-4)

Example 6.3-3 HDA process.The flow and heat capacities of the reactor feed stream for case where x=0.75 and yPH=0.4 are given below. Stream Flow, mol/hr Cp,Btu/(molF) Makeup gas 496 0.95(7)+0.05(10.1)=7.16 Recycle gas 3371 0.4(7)+0.6(10.1)=8.86 Toluene feed 273 48.7 Toluene recycle 91 From, from Ex.6.3-1 and Eq. 6.3-4 with TR,in=1,150 F (6.3-5)

Example 6.3-4 IPA process. If the feed stream to acetone process described Eq. 6.3-2 is an IPA-H2O azeotrope 70% IPA) and if we recycle and azeotropic mixture, then it is to show that 22.0 mol/hr of water enters with the feed. Also, for a conversion of x=0.96, the recycle flow will be 2.1 mol/hr of IPA and 0.9 mol/hr of water. If the reactor inlet Temp. is 572 F, then from Eqs.6.3-1 and 6.3-4 the adiabatic Temp. change is (6.3-6) This is unreasonable result. Thus, instead of using an adiabatic reactor, we attempt to achieve isothermal operation by supplying heat of the reaction to the process. In particular, we might attempt to pack the tubes of a heat exchanger with a catalyst.

Heuristic for Heat Loads If adiabatic operation is not feasible, such as in the isopropanol example, then we attempt to use direct heating or cooling. However, in many cases there is limit to the amount of heat-transfer surface area that we can fit into a reactor. Consider the of high T. gas-phase reaction Let U=20 Btu/(hr ft2 F) and T= 5 F ,Then for the heat load 1x106 Btu/hr (6.3-7)

6.4 EQUILIBRIUM LIMITATIONS Equilibrium Conversion Example 6.4-1 Cyclohexane production. Cyclohexane can be produced by the reaction (6.4-1) We consider a case where we desire to produce 100 mol/hr of C6H12 with a 99.9%purity. A pure benzene feed stream is available, and the hydrogen makeup stream contains 97.5 %H2 , 2%CH4, 0.5 %N2. A flowsheet for recycle structure is shown in Fig. 6.4-1 for a case where we recycle some of the benzene(which is not necessarily the best flowsheet).

Solution Overall material balancees. Assume no losses. Then Production of C6H12: Pc=100 (6.4-2) (6.4-3) Benzene fresh feed: FFB=Pc=100 Assume we use a gas recycle and a purge stream.Let FE=Excess H2 Fed to process (6.4-4) Total H2 Feed =3Pc+FE=0.975FG (6.4-5)

(6.4-6) (6.4-7) Recycle balances (6.4-8) Let molar ratio of H2 to Benzene be MR. THen (6.4-9)

Reactor exit flows (6.4-10) (6.4-11) (6.4-12) (6.4-13) (6.4-14)

Equilibrium relationship (6.4-15) (6.4-16) Then (6.4-17)

Separator Reactors If one of the product can be removed while the reaction is taking place, then an apparently equilibrium-limited reaction can be forced to go to complete conversion.

IsopropanolAcetone+H2 Example 6.4-2 Acetone production. Acetone can be produced by dehydrogenation of isopropanol IsopropanolAcetone+H2 (6.4-19) In the liquid phase as well as gas phase. At 300 F the equilibrium conversion for the liquid-phase process is about xeq=0.32. However, by suspending the catalyst in a high-boiling solvent and operating the reactor at a Temp. above the boiling point of Acetone, both H2 and Acetone can be removed as a vapor from reactor. Thus equilibrium conversion is shift to right. A series of three continuous stirred tank reactor with a pump around loop containing a heating system that supplies the endothermic heat to reaction, can be used for process.

Example 6. 4-3 Production of ethyl acrylate Example 6.4-3 Production of ethyl acrylate. Ethyl acrylate can be produced by the reaction Acrylic Acid + Ethanol Ethyl Acrylate + H2O (6.4-19) Acrylic Acid , Ethanol are monomers, which tend to polymerize in the reboilers of distillation columns. We can eliminate a column required to purify and recycle acrylic Acid from the process if we can force the equilibrium-limited reaction to completion, by removing the water. Hence we use an excess of ethanol to shift the equilibrium to the right, and we carry out the reaction in the reboiler of retifying column. With this approach, the ethanol, water, and ethyl acrylate are taken overhead, and acrylic acid conversion approaches unity.

For Example x ↓ as T  Reversible Exothermic Reactions (6.4-21) Sulfuric acid process : SO2+1/2H2O SO3 In ammonia synthesis Water-gas shift : CO +H2O CO2+H2 (6.4-22) ammonia synthesis: N2+3H2O 2NH3+H2 (6.4-22) x ↓ as T 

Diluents In some case an extraneous component (a diluent) is added which also causes a shift in the equilibrium conversion. For example, styrene can be produced by the reactions Ethylbenzene styrene +H2 (6.4-24) (6.4-25) Ethylbenzene Benzene+ethylene Ethylbenzene Toluene+Methane (6.4-26) Where the reactions take place at about 1100 F and 20 psia. The addition of steam so decrease the reverse reaction rate in Eq. 6.4-24. The stream serves in part as a heat carrier to supply endothermic heat of reaction.

6.5 COMPRESSOR DESIGN AND COST Whenever a gas-recycle stream is present, we will need a gas recycle compressor. The design equation for the theoretical horsepower(hp) for a centrifugal gas compressor is (6.5-1) Where Pin=lbf/ft2 , Qin=ft3/min and  =(Cp/Cv-1)/Cp/Cv The exit Temp. from the compression stage is (6.5-2)

More complex gases(CO2,CH4) 0.23 Other gases R/Cp TABLE 6.51 Values of  Monotonic gases 0.40 Diatomic gases 0.29 More complex gases(CO2,CH4) 0.23 Other gases R/Cp

Efficiency For the first designs, we assume a compressor efficiency of 90% to account for the fluid friction in suction and discharge values, ports, friction of moving metal surface fluid turbulence, etc. also we assume a driver efficiency of of 90% to account for the conversion of the input energy to shaft work.

Multistage Compressors For a three-stage compressor with intercooling, the work required is (6.5-3) The intermediate pressures that minimize the work are determined from (6.5-4)

Which lead to the results (6.5-5) Design heuristic: The compression ratios for each stage in a gas compressor should be equal.

Annualized Install Cost The brake horsepower bph is obtained by introducing the compressor efficiency in to Eq. 6.5-1: (6.5-7) (6.5-7) Then, Guthrie’s correlation(page.573) (6.5-8)

Operating Cost By dividing the brake horsepower by the driver efficiency. We can calculate the utility requirement. Then from utility cast and using 8150 hr/yr, we can calculate the operating cost.