Fixed residence time reactors

Suppose is a function of a alone and that neither dSt Ajda nor d Alda change sign over the range of concentrations encountered in the reactor. Then, for a system having a fixed residence time distribution. Equations (15.48) and (15.49) provide absolute bounds on the conversion of component A, the conversion in a real system necessarily falling within the bounds. If d S A/dc > 0, conversion is maximized by maximum mixedness and minimized by complete segregation. If d 0i A/da < 0, the converse is true. If cf- A/da = 0, micro-mixing has no effect on conversion. 

The conversion of 2 to 3 was optimized at full scale in the VRT reactor. In addition to confirming the productivity, safety, product quality, and economic benefit of the process, the robustness of the process was also demonstrated. Finally, this pilot study provided the basis for a full-scale commercial manufacturing design specification. Having fixed the optimum residence time, the process was then transferred into a plant Fixed Residence Time (FRT) cyanation reactor which employed a fixed length of jacketed static mixer for commercial manufacture. This FRT was capable of producing 300 metric tonnes per year of 3, with the same purified step yield of 80% that was achieved in the laboratory capillary reactor. 

If the conversion at a fixed residence time in the reactor does not change with changing flow rate (and respectively catalyst amount), e.g. cca,i= cza,2, when W /FA1 =W2/F°2 (Figure... 

The ideal plug flow reactor PFR is a simplified picture of the motion of a fluid in a tubular reactor as it is assumed that all fluid elements move with a uniform velocity along parallel streamlines and thus have a fixed residence time r. Strictly speaking, this assumption breaks the hydrodynamic rule that the velocity is zero at the wall (no slip condition. Figure 3.2.22). The steady-state mass balance of a PFR for a constant volume reaction can be deduced from the one-dimensional mass balance for a differential small element with thickness Az in direction of flow ... 

An ideal plug flow reactor has a fixed residence time Any fluid (plug) that enters the reactor at time t will exit the reactor at time t + x, where x is the residence time of the reactor. The residence time distribution function is therefore a dirac delta function at x. A real plug flow reactor has a residence time distribution that is a narrow pulse around the mean residence time distribution. 

Another type of reactor is the plug flow reactor, in which we often can find conversion data for a fixed residence time, but at different inlet concentrations. In this case, we can show that the rate is a function of conversion, initial concentration, and time, according to the simplified mass balance. 

Fixed-bed reactors in the form of gas absorption equipment are used commonly for noncatalytic gas-liquid reactions. Here the packed bed serves only to give good contact between the gas and liquid. Both cocurrent and countercurrent operations are used. Countercurrent operation gives the highest reaction rates. Cocurrent operation is preferred if a short liquid residence time is required. 

A distinc tion is to be drawn between situations in which (1) the flow pattern is known in detail, and (2) only the residence time distribution is known or can be calculated from tracer response data. Different networks of reactor elements can have similar RTDs, but fixing the network also fixes the RTD. Accordingly, reaction conversions in a known network will be unique for any form of rate equation, whereas conversions figured when only the RTD is known proceed uniquely only for hnear kinetics, although they can be bracketed in the general case. 

The three main types of reactors shown in Fig. 27-6 are in aclual commercial use the moving bed, the fluidized bed, and the entrained bed. The moving bed is often referred to as a. fixed bed because the coal bed is kept at a constant height. These differ in size, coal feed, reactant and product flows, residence time, and reaction temperature. 

Tubular reactors have been the main tools to study continuous flow processes for vapor or gas-phase reactions. These are also used for reaction in tv o flowing phases over a solid catalyst. When the catalyst is in a fixed bed, the contact between the liquid on the outside surface of the particulate is uncertain. For slurry-type solid catalyst the residence time of the catalyst or the quantity in the reactor volume can be undefined. 

Various experimental methods to evaluate the kinetics of flow processes existed even in the last centuty. They developed gradually with the expansion of the petrochemical industry. In the 1940s, conversion versus residence time measurement in tubular reactors was the basic tool for rate evaluations. In the 1950s, differential reactor experiments became popular. Only in the 1960s did the use of Continuous-flow Stirred Tank Reactors (CSTRs) start to spread for kinetic studies. A large variety of CSTRs was used to study heterogeneous (contact) catalytic reactions. These included spinning basket CSTRs as well as many kinds of fixed bed reactors with external or internal recycle pumps (Jankowski 1978, Berty 1984.)... 

In the fixed catalyst method, the residence time in the reactor may be easily controlled to generate fibers of selected length and diameter, both dimensions which can vary over several orders of magnitude. Most of the physical properties which have been measured for VGCF have been made on this type of fiber. 

Figure 8-38. Residence time distributions of some commerciai and fixed bed reactors. The variance, equivaient number of CSTR stages, and Peciet number are given for each reactor. (Source Wales, S. M., Chemicai Process Equipment—Seiection and Design, Butterworths, 1990.)...

An arbitrary decision was made to fix the mass of catalyst in the reactor, rather than the feed rate of catalyst. The feed rate is calculated from the loading and the mean residence time ... 

The results of Example 5.2 apply to a reactor with a fixed reaction time, i or thatch- Equation (5.5) shows that the optimal temperature in a CSTR decreases as the mean residence time increases. This is also true for a PFR or a batch reactor. There is no interior optimum with respect to reaction time for a single, reversible reaction. When Ef < Ef, the best yield is obtained in a large reactor operating at low temperature. Obviously, the kinetic model ceases to apply when the reactants freeze. More realistically, capital and operating costs impose constraints on the design. 

The catalytic reaction was carried out at 270°C and 101.3 kPa in a stainless steel tubular fixed-bed reactor. The premixed reaction solution, with a molar ratio catechol. methanol water of 1 1 6, was fed into the reactor using a micro-feed pump. To change the residence time in the reactor, the catechol molar inlet flow (Fio) and the catalyst mass (met) were varied in the range 10 < Fio <10 mol-h and 2-10 < met < 310 kg. The products were condensed at the reactor outlet and collected for analysis. The products distribution was determined quantitatively by HPLC (column Nucleosil 5Ci8, flow rate, 1 ml-min, operating pressure, 18 MPa, mobile phase, CH3CN H2O =1 9 molar ratio). 

When a number of competing reactions are involved in a process, and/or when the desired product is obtained at an intermediate stage of a reaction, it is important to keep the residence-time distribution in a reactor as narrow as possible. Usually, a broadening of the residence-time distribution results in a decrease in selectivity for the desired product. Hence, in addition to the pressure drop, the width of the residence-time distribution is an important figure characterizing the performance of a reactor. In order to estimate the axial dispersion in the fixed-bed reactor, the model of Doraiswamy and Sharma was used [117]. This model proposes a relationship between the dispersive Peclet number ... 

The figure shows the ratio of the widths of initially delta-like concentration tracers at the reactor exits as a function of the micro-channel Peclet number for different values of the porosity. Taking a value of = 0.4 as standard, it becomes apparent that the dispersion in the micro-channel reactor is smaller than that in the fixed-bed reactor in a Peclet number range from 3 to 100. Minimum dispersion is achieved at a Peclet number of about 14, where the tracer width in the micro-channel reactor is reduced by about 40% compared with its fixed-bed counterpart. Hence the conclusion may be drawn that micro-channel reactors bear the potential of a narrower residence time than fixed-bed reactors, where again it should be stressed that reactors with equivalent chemical conversion were chosen for the comparison. 

In cases where the operation time-scale is independent of the channel diameter, as for a homogeneous reaction, it is necessary to keep the residence time fixed when downscaling a reactor in order keep the efficiency constant. When the flow-rate Qtot of the process gas is given, this means that a reduction in the channel diameter has to be accompanied by an increase in the channel length L or the number of channels N, according to... 

GP 3] [R 3b] The maximum selectivity of about 33% was the same for the best micro reactor and the fixed bed (Figure 3.34) at the same conversions from 73 to 85% with a V205/P205/Ti02 catalyst [103]. At stiU higher conversion, the fixed bed has a better performance. However, the residence times needed for comparable conversion are one order of magnitude shorter than in the fixed-bed reactor. 

GP 3] [R 3b] The space-time yield of chemical micro processing was a factor of five larger than that of a conventional fixed-bed reactor (Figure 3.35) (0.4% 1-butene in air 0.1 MPa 400 °C) [103]. This is due to the shorter residence time needed in the micro reactors for the same conversion as in the fixed bed. Differences from the fixed bed become smaller when operating at very high conversion, up to 95%. 

GL 18] ]R 6a]]P 17/Using the same experimental conditions and catalysts with the same geometric surface area, the performance of micro-channel processing was compared with that of a fixed-bed reactor composed of short wires [17]. The conversion was 89% in the case of the fixed bed the micro channels gave a 58% yield. One possible explanation for this is phase separation, i.e. that some micro channels were filled with liquids only, and some with gas. This is unlikely to occur in a fixed bed. Another explanation is the difference in residence time between the two types of reactors, as the fixed bed had voids three times larger than the micro channel volume. It could not definitively be decided which of these explanations is correct. 

A fixed-bed reactor system was employed (Figure 32.2). Each of the two reactors was charged with 38 cc of Amberlyst BD20 catalyst. Sample ports located at the exit of each reactor enabled increased acquisition of residence time data. Pressure was maintained by a back pressure control valve to maintain methanol in the liquid phase. After charging, the 1st and then 2nd reactors were connected to the pumps and filled with the reaction mixture while vapor was released from each through the top vent valve. Once each reactor was filled with liquid and emptied of vapor, the pressure regulator was connected to the output and both reactors were immersed into the water bath. 

Another example of performance enhancement using a zeolite/TUD-1 catalyst is shown in n-hexane cracking using a series of zeolite-Beta-embedded TUD-1 catalysts (29) 20, 40 and 60 wt% zeolite Beta in Al-Si-TUD-1 (Si/Al = 150). These are compared to pure zeolite Beta, and to a physical mixture of 40% zeolite Beta and 60% Al-Si-TUD-1. These catalysts were tested in a fixed bed reactor, at atmospheric pressure, with constant residence time at 538°C. The pseudo-first-order rate constants are shown in Figure 41.8. Note that the zeohte-loaded catalysts were clearly superior to both the pure zeolite Beta catalyst and the zeohte-TUD-1 physical mixture. Again, this is evidence that catalyst performance benefits from a hierarchical pore stracture such as zeolite embedded in TUD-1. 

In this work it has been demonstrated that parallel fixed bed reactor system facilitates experimentation of heterogeneous catalysts in three-phase systems. The residence times and Reynolds numbers were determined. The Reynolds numbers were veiy... 

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