Agitated reactors calculation

Suspension of Catalyst Particles. There were concerns about complete catalyst suspension due to the density of the panicles. Calculations indicated that total suspension of the dense A12C>3 panicles should occur by 600 RPM (ref. 14). Such calculations are normally valid for agitated reactors with 1 1 height to diameter ratios, and a turbine impeller at 1/4 height. Lower clearance in the reactor (the present case) will decrease the impeller speed required for complete suspension of the panicles. Visual inspection of the open reactor confirmed complete suspension of the catalyst at 450-600 RPM. 

The reaction of heterofunctional condensation is carried out in reactor 5, which is a steel enameled apparatus with a water vapour jacket and an anchor agitator. Reactor 5 is loaded at a residual pressure of 730 200 GPa with a,eo-dihydroxydiphenylsiloxane, the product of dihydroxydiphenylsilane condensation a calculated amount of the toluene solution of triace-toxymethylsilane self-flows at agitationffom batch box 6. 

The reactor is filled at agitation with calculated portions of hydrochloric acid and the reaction of the medium is tested. It should be neutral (pH is 5.5-7). The neutral solution of varnish is washed with water at 60 °C from SO4"2 ions until the analysis with barium chloride is neutral. 

While a number of good correlations for the gas holdup in mechanically agitated reactors are available (Joshi et ai, 1982), the best correlations are those by Hughmark (1980) and Sridhar and Potter (1980), and they are recommended. The critical impeller speed, N0, required for effective gas-liquid dispersion, and the impeller speed at which gas above liquid is first entrapped, Nc, can be reasonably well calculated using Eqs. (2.4) and (2.5), respectively. 

The liquid-phase mixing in a multistage mechanically agitated reactor is best correlated by Eq. (2.31) in the absence of gas flow and by Eq. (2.32) in the presence of gas flow. The mixing time can be estimated from the study of Paca et al. (1976). Experimental work is needed to estimate gas-phase back-mixing. The use of Eq. (2.36) for the calculation of the gas-liquid volumetric mass transfer coefficient in a multistage mechanically agitated column is recommended. 

Just as the critical gas velocity is required to suspend the particles in a three-phase fluidization in an agitated reactor, some minimum agitation intensity is required to keep the particles in suspension. Calderbank11 has described the methods of estimating this minimum agitation intensity. For an agitated liquid-solid slurry vessel, the correlation of Zweitering150 for the calculation of the minimum impeller speed which completely suspends the particles, namely... 

Example 10-3 Calculation of Solid-Liquid Mass Transfer Coefficient. (Adapted from Doraiswamy and Sharma, 1984) It is desired to prepare a 25°C aqueous solution of potassium sulfate containing 0.09 g K2S04/g solntion in an agitated 48 in. diameter stainless steel reactor. Calculate ... 

The reactor volume is taken as the volume of the reactor physically occupied by the reacting fluids. It does not include the volume occupied by agitation devices, heat exchange equipment, or head-room above liquids. One may arbitrarily select the temperature, pressure, and even the state of aggregation (gas or liquid) at which the volumetric flow rate to the reactor will be measured. For design calculations it is usually convenient to choose the reference conditions as those that prevail at the the inlet to the reactor. However, it is easy to convert to any other basis if the pressure-volume-temperature behavior of the system is known. Since the reference volumetric flow rate is arbitrary, care must be taken to specify precisely the reference conditions in order to allow for proper interpretation of the resultant space time. Unless an explicit statement is made to the contrary, we will choose our reference state as that prevailing at the reactor inlet and emphasize this choice by the use of the subscript zero. Henceforth,... 

When the reactor is scaled up to 60 cm radius, however, the operating point is between the two curves. This means that the reaction can be safely run at 50°C in a well-agitated process vessel of 60 cm radius with the heat transfer coefficient as stated above becauseerating point is below the Semenov curve. In case the agitation is lost, however, the Frank-Kamenetskii curve becomes the better predictor of runaway temperatures, and because the operating point is above this curve, the estimate is that the reaction will run away. The calculation of the Frank-Kamentskii method is available in ASTME-1231 [166]. 

It is important to calculate U accurately to determine the required heat transfer area for a reactor. Typical expressions to calculate overall heat transfer coefficients for agitated vessels are presented in [174,180] and generally in standard chemical engineering texts and reference books. 

For the semi-batch stirred tank reactor, the model was based on the following assumptions the reactor is well agitated, so no concentration differences appear in the bulk of the liquid gas-liquid and liquid-solid mass transfer resistances can prevail and finally, the liquid phase is in batch, while hydrogen is continuously fed into the reactor. The hydrogen pressure is maintained constant. The liquid and gas volumes inside the reactor vessel can be regarded as constant, since the changes of the fluid properties due to reaction are minor. The total pressure of the gas phase (P) as well as the reactor temperature were continuously monitored and stored on a PC. The partial pressure of hydrogen (pnz) was calculated from the vapour pressure of the solvent (pvp) obtained from Antoine s equation (pvpo) and Raoult s law ... 

The quantities to be calculated are (a) the fraction of the oxygen supplied that reacts, and (b) the corresponding rate of production of o-methylbenzoic acid in kmol per hour. There is also the question of whether an agitated tank is the most suitable reactor for this process. 

The overall rate of reaction calculated for the three-phase fluidised-bed reactor above is approximately one tenth of the rate calculated for the agitated tank slurry reactor in Example 4.6. The main reasons are the very poor effectiveness factor and the relatively smaller external surface area for mass transfer caused by using the larger particles. Even the gas-liquid transfer resistance is greater for the three-phase fluidised-bed, in spite of the larger particles being able to produce relatively small bubbles these bubbles are not however as small as can be produced... 

Subscripts s and b stand for small and big reactor, d is typical diameter, n is revolutions of agitator per unit time, N is power input of agitator, 17 is viscosity, p is density, and g is gravity constant. Using this set of equations, it is impossible to calculate the power input (Nb) and the... 

This calculation can be found in a paper by Rushton (2).] This result means that in the big reactor the viscosity must be higher than it is in the small one to achieve geometrically similar agitation patterns. From this, another conclusion may be drawn if one does not vary the... 

Thus for a given reaction mass, the heat transfer coefficient of the internal film can be influenced by the stirrer speed and its diameter. The value of the equipment constant (z) can be calculated using the geometric characteristics of the reactor. The value of material constant for heat transfer (y) can either be calculated from the physical properties of the reactor contents-as far as they are known-or measured by the method of the Wilson plot in a reaction calorimeter [4, 5]. This parameter is independent of the geometry or size of the reactor. Thus, it can be determined at laboratory scale and used at industrial scale. The Wilson plot determines the overall heat transfer coefficient as a function of the agitator revolution speed in a reaction calorimeter ... 

The value of z, characterizing the internal part of the equipment factor, can be calculated using the geometric characteristics of the reactor. Some typical values of the agitator constant are given in Table 9.4 [2],... 

The synthesis of chloromethylsilatrane is carried out in steel enameled reactor 5 with an agitator and a water vapour jacket. First the reactor is loaded with ethyl alcohol and freshly distilled triethanolamine from batch boxes a calculated amount of potassium hydroxide is loaded through a hatch. The mixture is intensively agitated, the temperature raised till KOH dissolves completely, as shown by the appearance of reflux in the rundown box situated after cooler 4. Then chloromethyltriethoxysilane is fed at such speed that the reactive mixture boils uniformly. After the reaction is finished, the mixture is cooled to 15°C, sent to nutsch filter 6 and filtered through coarse calico. The technical chloromethylsilatrane in the nutsch filter is washed with ethyl alcohol twice, thoroughly pressed and dried in draft 7 at a temperature below 100 °C till its weight is constant. 

After the drying, the contents of reactor 10 are cooled to 40 °C by sending water into the jacket. From weight batch box 9, a calculated amount of silicone oligomer is loaded through the hatch the reactor receives trifluoroacetic acid (0.12% of the amount of the reactive mixture) and potassium hydroxide (0.04%). The reactive mixture is agitated and heated to 120-140 °C at this temperature re-etherification takes place. The released ethyl alcohol is withdrawn out of the system in the form of azeotropic mixture with toluene. The vapours of the azeotropic mixture rise up tower 13 and condense in refluxer 14. From there, part of the condensate in the form of reflux is returned into the tower, and the rest is collected in receptacle 16. 

When the medium becomes acid, at 45-48 °C reactor 1 receives a calculated amount of poisonous methanol from batch box 6 at such speed that the temperature in the reactor does not exceed 50°C. Then, at 67-75 °C the released methylacetate is sent from reactor 1 through cooler 7 into neutraliser 8 with an agitator. 

The reactor is an enameled apparatus with an agitator and a water vapour jacket. The production of sodium dihydroxyphenylsilanolate is carried out in butanol and toluene or ethanol and toluene medium at 35-50 °C. The consumption of other components is calculated by the amount of the loaded condensation product. After loading the product of condensation, the reactor is filled with toluene and butanol (or ethanol and toluene) from batch boxes 10 and 11. The ratio of the solvents should be 1 1.4 to obtain 10% silanol solution. The calculation takes into account toluene contained in the product of hydrolytic condensation. The loaded mixture is agitated in the reactor for 30 minutes after that it receives 20% alkali solution from batch box 12 at agitation. The reaction forms sodium dihydroxydiphenylsi-lanolate and water. 

In the first set of runs In Table I, the H2/CO feed ratio exceeds the consumption ratio, 7/12 m 0.58 In the second It Is less than the consumption ratio. In each case, the H2 partial pressure In the reactor Increased with decreased agitation, as conversion dropped. In the absence of mass transfer resistance this would be expected to Increase the P/0 ratio. The fact that the P/0 ratio In both cases Instead decreased is consistent with the postulate that the H- concentration In the liquid has decreased. The corresponding mass transfer resistance K, back-calculated from equation (8), is given at each stirring speed together with the hydrogen and carbon monoxide liquid-phase concentrations that are estimated by equations (6) and (7). 

The power in aerated slurry reactors in regime c can be calculated using Equation (3.3). In general, relations summarized by Baldi (1986) can be used for the calculation of power consumption. The most widely used correlation for the minimum rotational speed of agitation required for complete suspension of solids is that of Zweitering (Equation (3.6)). The most versatile... 

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