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Calculate the pressure drop for an ideal cross-flow section. Find the correction factor for the effect of baffle leakage on pressure drop Rl from Fig. Curves shown are not to be extrapolated beyond the points shown. Find the correction factor for bundle bypass Rb from Fig. Calculate the pressure drop for an ideal window section. Calculate the pressure drop across the shell side excluding nozzles.

A fouled exchanger will generally give lower heat-transfer rates, as reflected by the dirt resistances incorporated into Eq. Some estimate of fouling effects on pressure drop may be made by using the methods just given by assuming that the fouling deposit blocks the leakage and possibly the bypass areas. The fouling may also decrease the clearance between tubes and significantly increase the pressure drop in cross-flow.

Single-Component Condensers Mean Temperature Difference In condensing a single component at its saturation temperature, the entire resistance to heat transfer on the condensing side is generally assumed to be in the layer of condensate. A mean condensing coefficient is calculated from the appropriate correlation and combined with the other resistances in Eq. The overall coefficient is then used with the LMTD no FT correction is necessary for isothermal condensation to give the required area, even though the condensing coefficient and hence U are not constant throughout the condenser.

If the vapor is superheated at the inlet, the vapor may first be desuperheated by sensible heat transfer from the vapor. This occurs if the surface temperature is above the saturation temperature, and a single-phase heat-transfer correlation is used. If the surface is below the saturation temperature, condensation will occur directly from the superheated vapor, and the effective coefficient is determined from the appropriate condensation correlation, using the saturation temperature in the LMTD.

To determine whether or not condensation will occur directly from the superheated vapor, calculate the surface temperature by assuming single-phase heat transfer. It is generally conservative to design a pure-component desuperheatercondenser as if the entire heat load were transferred by condensation, using the saturation temperature in the LMTD.

The design of an integral condensate subcooling section is more difficult, especially if close temperature approach is required. The condensate layer on the surface is on the average subcooled by onethird to one-half of the temperature drop across the film, and this is often sufficient if the condensate is not reheated by raining through the vapor.

If the condensing-subcooling process is carried out inside tubes or in the shell of a vertical condenser, the single-phase subcooling section can be treated separately, giving an area that is added onto that needed for condensation. If the subcooling is achieved on the shell side of a horizontal condenser by flooding some of the bottom tubes with a weir or level controller, the rate and heat-balance equations must be solved for each section to obtain the area required. Pressure drop on the condensing side reduces the final condensing temperature and the MTD and should always be checked.

In designs requiring close approach between inlet coolant and exit condensate subcooled or not , underestimation of pressure drop on the condensing side can lead to an exchanger that cannot meet specified terminal temperatures. Since pressure-drop calculations in two-phase flows such as condensation are relatively inaccurate, designers must consider carefully the consequences of a larger-than-calculated pressure drop. Horizontal In-Shell Condensers The mean condensing coefficient for the outside of a bank of horizontal tubes is calculated from Eq.

For undisturbed laminar flow over all the tubes, Eq. For a cylindrical tube bundle, where N varies, it is customary to take N equal to two-thirds of the maximum or centerline value. Baffles in a horizontal in-shell condenser are oriented with the cuts vertical to facilitate drainage and eliminate the possibility of flooding in the upward cross-flow sections. Pressure drop on the vapor side can be estimated by the data and method of Diehl and Unruh [Pet. Refiner, 36 10 , ; 37 10 , ]. High vapor velocities across the tubes enhance the condensing coefficient.

There is no correlation in the open literature to permit designers to take advantage of this. Since the vapor flow rate varies along the length, an incremental calculation procedure would be required in any case. In general, the pressure drops required to gain significant benefit are above those allowed in most process applications. Vertical In-Shell Condensers Condensers are often designed so that condensation occurs on the outside of vertical tubes.

Equation is valid as long as the condensate film is laminar.

Perry's Chemical Engineers' Handbook, Eighth Edition (8th ed.)

When it becomes turbulent, Fig. Some judgment is required in the use of these correlations because of construction features of the condenser. The tubes must be supported by baffles, usually with maximum cut 45 percent of the shell diameter and maximum spacing to minimize pressure drop. The flow of the condensate is interrupted by the baffles, which may draw off or redistribute the liquid and which will also cause some splashing of free-falling drops onto the tubes. For subcooling, a liquid inventory may be maintained in the bottom end of the shell by means of a weir or a liquid-level-controller.

The subcooling heat-transfer coefficient is given by the correlations for natural convection on a vertical surface [Eqs. Pressure drop may be estimated by the shell-side procedure. Horizontal In-Tube Condensers Condensation of a vapor inside horizontal tubes occurs in kettle and horizontal thermosiphon reboilers and in air-cooled condensers. At low flow rates, when gravity dominates the flow pattern, Eq. At high flow rates, the flow and heat transfer are governed by vapor shear on the condensate film, and Eq. A simple and generally conservative procedure is to calculate the coefficient for a given case by both correlations and use the larger one.

Pressure drop during condensation inside horizontal tubes can be computed by using the correlations for two-phase flow given in Sec. Vertical In-Tube Condensation Vertical-tube condensers are generally designed so that vapor and liquid flow cocurrently downward; if pressure drop is not a limiting consideration, this configuration can result in higher heat-transfer coefficients than shell-side condensation and has particular advantages for multicomponent condensation. If gravity controls, the mean heat-transfer coefficient for condensation is given by Figs.

If vapor shear controls, Eq. It is generally conservative to calculate the coefficients by both methods and choose the higher value. The pressure drop can be calculated by using the Lockhart-Martinelli method [Chem. Vertical in-tube condensers are often designed for reflux or knock-back application in reactors or distillation columns. In this case, vapor flow is upward, countercurrent to the liquid flow on the tube wall; the vapor shear acts to thicken and retard the drainage of the condensate film, reducing the coefficient.

Neither the fluid dynamics nor the heat transfer is well understood in this case, but Soliman, Schuster, and Berenson [J. Heat Transfer, 90, — ] discuss the problem and suggest a computational method. The Diehl-Koppany correlation [Chem. If the vapor velocity is great enough, the liquid film will be carried upward; this design has been employed in a few cases in which only part of the stream is to be condensed.

This velocity cannot be accurately computed, and a very conservative high outlet velocity must be used if unstable flow and flooding are to be avoided; 3 times the vapor velocity given by the Diehl-Koppany correlation for incipient flooding has been suggested as the design value for completely stable operation. Multicomponent Condensers Thermodynamic and Mass-Transfer Considerations Multicomponent vapor mixture includes several different cases: all the components may be liquids at the lowest temperature reached in the condensing side, or there may be components which dissolve substantially in the condensate even though their boiling points are below the exit temperature, or one or more components may be both noncondensable and nearly insoluble.

Multicomponent condensation always involves sensible-heat changes in the vapor and liquid along with the latent-heat load. Compositions of both phases in general change through the condenser, and concentration gradients exist in both phases. Temperature and concentration profiles and transport rates at a point in the condenser usually cannot be calculated, but the binary cases have been treated: condensation of one component in the presence of a completely insoluble gas [Colburn and Hougen, Ind.

It is necessary to know or calculate diffusion coefficients for the system, and a reasonable approximate method to avoid this difficulty and the reiterative calculations is desirable. To integrate the point conditions over the total condensation requires the temperature, composition enthalpy, and flow-rate profiles as functions of the heat removed.

These are calculated from component thermodynamic data if the vapor and liquid are assumed to be in equilibrium at the local vapor temperature. This assumption is not exactly true, since the condensate and the liquid-vapor interface where equilibrium does exist are intermediate in temperature between the coolant and the vapor. In calculating the condensing curve, it is generally assumed that the vapor and liquid flow collinearly and in intimate contact so that composition equilibrium is maintained between the total streams at all points.

If, however, the condensate drops out of the vapor as can happen in horizontal shell-side condensation and flows to the exit without further interaction, the remaining vapor becomes excessively enriched in light components with a decrease in condensing temperature and in the temperature difference between vapor and coolant. The result may be not only a small reduction in the amount of heat transferred in the condenser but also an inability to condense totally the light ends even at reduced throughput or with the addition of more surface.

To prevent the liquid from segregating, in-tube condensation is preferred in critical cases. The condensate film coefficient is calculated from the appropriate equation or correlation for pure vapor condensation for the geometry and flow regime involved, using mean liquid properties. The sensible-heat-transfer coefficient for the vapor-gas stream hsv is calculated by using the appropriate correlation or design method for the geometry involved, neglecting the presence of the liquid.

As the vapor condenses, this coefficient decreases and must be calculated at several points in the process. Tv and Tc are temperatures of the vapor and of the coolant respectively. It may be nonconservative for condensing steam and other high-latent-heat substances, in which case it may be necessary to increase the calculated area by 25 to 50 percent.

Pressure drop on the condensing side may be estimated by judicious application of the methods suggested for pure-component condensation, taking into account the generally nonlinear decrease of vapor-gas flow rate with heat removal. In multicomponent systems, the light components are preferentially vaporized at the surface, and the process becomes limited by their rate of diffusion. The net effect is to decrease the effective temperature difference between the hot surface and the bulk of the boiling liquid.

If one attempts to vaporize too high a fraction of the feed liquid to the reboiler, the temperature difference between surface and liquid is reduced to the point that nucleation and vapor generation on the surface are suppressed and heat transfer to the liquid proceeds at the lower rate associated with single-phase natural convection. The only safe procedure in design for wide-boiling-range mixtures is to vaporize such a limited fraction of the feed that the boiling point of the remaining liquid mixture is still at least 5.

Positive flow of the unvaporized liquid through and out of the reboiler should be provided. Kettle Reboilers It has been generally assumed that kettle reboilers operate in the pool boiling mode, but with a lower peak heat flux because of vapor binding and blanketing of the upper tubes in the bundle. There is some evidence that vapor generation in the bundle causes a high circulation rate through the bundle. The result is that, at the lower heat fluxes, the kettle reboiler actually gives higher heattransfer coefficients than a single tube.

Present understanding of the recirculation phenomenon is insufficient to take advantage of this in design. Available nucleate pool boiling correlations are only very approximate, failing to account for differences in the nucleation characteristics of different surfaces. The Mostinski correlation [Eq. Experimental heat-transfer coefficients for pool boiling of a given liquid on a given surface should be used if available.

The bundle peak heat flux is a function of tube-bundle geometry, especially of tube-packing density; in the absence of better information, the Palen-Small modification [Eq. A general method for analyzing kettle reboiler performance is by Fair and Klip, Chem. It is effectively limited to computer application.

Kettle reboilers are generally assumed to require negligible pressure drop. It is important to provide good longitudinal liquid flow paths within the shell so that the liquid is uniformly distributed along the entire length of the tubes and excessive local vaporization and vapor binding are avoided. This method may also be used for the thermal design of horizontal thermosiphon reboilers.

The recirculation rate and pressure profile of the thermosiphon loop can be calculated by the methods of Fair [Pet. Refiner, 39 2 , — ]. Vertical Thermosiphon Reboilers Vertical thermosiphon reboilers operate by natural circulation of the liquid from the still through the downcomer to the reboiler and of the two-phase mixture from the reboiler through the return piping. The flow is induced by the hydrostatic pressure imbalance between the liquid in the downcomer and the two-phase mixture in the reboiler tubes.

Thermosiphons do not require any pump for recirculation and are generally regarded as less likely to foul in service because of the relatively high two-phase velocities obtained in the tubes. Heavy components are not likely to accumulate in the thermosiphon, but they are more difficult to design satisfactorily than kettle reboilers, especially in vacuum operation. Several shortcut methods have been suggested for thermosiphon design, but they must generally be used with caution. The method due to Fair loc. Fair also suggests a shortcut method that is satisfactory for preliminary design and can be reasonably done by hand.

Forced-Recirculation Reboilers In forced-recirculation reboilers, a pump is used to ensure circulation of the liquid past the heattransfer surface. Force-recirculation reboilers may be designed so that boiling occurs inside vertical tubes, inside horizontal tubes, or on the shell side. Excess pressure required to circulate the two-phase fluid through the tubes and back into the column is supplied by the pump, which must develop a positive pressure increase in the liquid.

In this case the hydrostatic-headpressure effect through the tubes is zero but must be considered in the two-phase return lines to the column. The same procedure may be applied in principle to design of forced-recirculation reboilers with shell-side vapor generation. Little is known about two-phase flow on the shell side, but a reasonable estimate of the friction pressure drop can be made from the data of Diehl and Unruh [Pet. No void-fraction data are available to permit accurate estimation of the hydrostatic or acceleration terms.

These may be roughly estimated by assuming homogeneous flow. Heating surface areas are normally, but not always taken as those in contact with the material being evaporated. Such losses include pressure drop through entrainment separators, friction in vapor piping, and acceleration losses into and out of the piping. The latter loss has often been overlooked, even though it can be many times greater than the friction loss.

Boiling-point rise, the difference between the boiling point of the solution and the condensing point of the solvent at the same pressure, is another loss. A further loss may occur when the heater effluent flashes as it enters the vapor-liquid separator. Simply basing the liquid temperature on the measured vapor head pressure may ignore both—or only the latter if temperature rise through the heater is estimated separately from known heat input and circulation rate.

Forced-Circulation Evaporators In evaporators of this type in which hydrostatic head prevents boiling at the heating surface, heattransfer coefficients can be predicted from the usual correlations for condensing steam Fig. The liquid film coefficient is improved if boiling is not completely suppressed. When only the film next to the wall is above the boiling point, Boarts, Badger, and Meisenberg [Ind. In such cases, the course of the liquid temperature can still be calculated from known circulation rate and heat input. When the bulk of the liquid is boiling in part of the tube length, the film coefficient is even higher.

However, the liquid temperature starts dropping as soon as full boiling develops, and it is difficult to estimate the course of the temperature curve. It is certainly safe to estimate heat transfer on the basis that no bulk boiling occurs. Fragen and Badger [Ind.

This equation is based primarily on experiments with copper tubes of 0. Long-Tube Vertical Evaporators In the rising-film version of this type of evaporator, there is usually a nonboiling zone in the bottom section and a boiling zone in the top section. The length of the nonboiling zone depends on heat-transfer characteristics in the two zones and on pressure drop during two-phase flow in the boiling zone. The work of Martinelli and coworkers [Lockhart and Martinelli, Chem. In estimating pressure drop, integrated curves similar to those presented by Martinelli and Nelson are the easiest to use.

The curves for pure water are shown in Figs. Similar curves can be prepared if one or both flows are laminar or if the properties of the liquid differ appreciably from the properties of pure water. Film coefficients for the boiling of liquids other than water have been investigated. Coulson and McNelly [Trans. The Reynolds numbers are calculated on the basis of each fluid flowing by itself in the tube. The frictional pressure drop is derived from Fig.

Pressure drop due to hydrostatic head can be calculated from liquid holdup R1. Liquid holdup, which represents the ratio of liquid-only velocity to actual liquid velocity, also appears to be the principal determinant of the convective coefficient in the boiling zone Dengler, Sc. In other words, the convective coefficient is that calculated from Eq. Process Des. McAdams, Woods, and Bryan Trans. Schweppe and Foust [Chem. The simplified method of calculation outlined includes no allowance for the effect of surface tension. Stroebe, Baker, and Badger loc.

Coulson and Mehta [Trans. The higher coefficients at low surface tension are offset to some extent by a higher pressure drop, probably because the more intimate mixture existing at low surface tension causes the liquid fraction to be accelerated to a velocity closer to that of the vapor. Figure gives the general range of overall long-tube vertical- LTV evaporator heat-transfer coefficients usually encountered in commercial practice. The higher coefficients are encountered when evaporating dilute solutions and the lower range when evaporating viscous liquids.

The dashed curve represents the approximate lower limit, for liquids with viscosities of about 0. The LTV evaporator does not work well at low temperature differences, as indicated by the results shown in Fig. Badger Associates, Inc. PB The feed was at its boiling point at the vapor-head pressure, and feed rates varied from 0. Falling film evaporators find their widest use at low temperature differences—also at low temperatures. Under most operating conditions encountered, heat transfer is almost all by pure convection, with a negligible contribution from nucleate boiling.

The same Dukler correlation presents curves covering falling film heat transfer to nonboiling liquids that are equally applicable to the falling film evaporator [Sinek and Young, Chem. Kunz and Yerazunis [ J. Heat Transfer 8, ] have since extended the range of physical properties covered, as shown in Fig. The boiling point in the tubes of such an evaporator is higher than in the vapor head because of both frictional-pressure drop and the head needed to accelerate the vapor to the tube-exit velocity.

These factors, which can easily be predicted, make the overall apparent coefficients somewhat lower than those for nonboiling conditions. Figure shows overall apparent heat-transfer coefficients determined in a falling-film seawater evaporator using the same tubes and flow rates as for the rising-film tests W. Performance is primarily a function of temperature level, temperature difference, and viscosity.

While liquid level can also have an important influence, this is usually encountered only at levels lower than considered safe in commercial operation. Overall heat-transfer coefficients are shown in Fig. Liquid level was maintained at the top tube sheet. Foust, Baker, and Badger [Ind. In the normal range of liquid levels, their results can be expressed as 0. This work was done with water. No detailed tests have been reported for the performance of propeller calandrias. Not enough is known regarding the performance of the propellers themselves under the cavitating conditions usually encountered to permit predicting circulation rates.

In many cases, it appears that the propeller does no good in accelerating heat transfer over the transfer for natural circulation Fig. Miscellaneous Evaporator Types Horizontal-tube evaporators operating with partially or fully submerged heating surfaces behave in much the same way as short-tube verticals, and heattransfer coefficients are of the same order of magnitude. Some test results for water were published by Badger [Trans. When operating unsubmerged, their heat transfer performance is roughly comparable to the falling-film vertical tube evaporator.

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Heat-transfer coefficients in clean coiled-tube evaporators for seawater are shown in Fig. London , 1B 7 , ]. The tubes were of copper. Heat-transfer coefficients in agitated-film evaporators depend primarily on liquid viscosity. This type is usually justifiable only for very viscous materials. Figure shows general ranges of overall coefficients [Hauschild, Chem. When used with nonviscous fluids, a wipedfilm evaporator having fluted external surfaces can exhibit very high coefficients Lustenader et al. However, the high degree of cleanliness needed for high coefficients was difficult to achieve, and the tube layout and liquid level were changed during the course of the tests so as to make direct comparison of results difficult.

Other workers have found little or no effect of conditions of surface or tube material on boiling-film FIG. Nauk, no. Work in connection with desalination of seawater has shown that specially modified surfaces can have a profound effect on heattransfer coefficients in evaporators. Figure Alexander and Hoffman, Oak Ridge National Laboratory TM compares overall coefficients for some of these surfaces when boiling fresh water in 0.

The area basis used was the nominal outside area. Tube 20 was a smooth 0. Tube 23 was a clean aluminum tube with 20 spiral corrugations of 0. Tube 48 was a clean copper tube that had 50 longitudinal flutes pressed into the wall General Electric double-flute profile, Diedrich, U. Patent 3,,, Apr. Tubes 47 and 39 had a specially patterned porous sintered-metal deposit on the boiling side to promote nucleate boiling Minton, U. Patent 3,,, May 21, Both of these tubes also had steamside coatings to promote dropwise condensation—parylene for tube 47 and gold plating for tube Of these special surfaces, only the double-fluted tube has seen extended services.

Most of the gain in heat-transfer coefficient is due to the condensing side; the flutes tend to collect the condensate and leave the lands bare [Carnavos, Proc. First Int. Water Desalination, 2, ]. There is as yet no accepted explanation for the superior performance in salt water.

This phenomenon is also seen in evaporation from smooth tubes. Effect of Fluid Properties on Heat Transfer Most of the heattransfer data reported in the preceding paragraphs were obtained with water or with dilute solutions having properties close to those of water.

Heat transfer with other materials will depend on the type of evaporator used. For forced-circulation evaporators, methods have been presented to calculate the effect of changes in fluid properties. For natural-circulation evaporators, viscosity is the most important variable as far as aqueous solutions are concerned. When handling molasses solutions in a forced-circulation evaporator in which boiling was allowed to occur in the tubes, Coates and Badger [Trans. Little work has been published on the effect of viscosity on heat transfer in the long-tube vertical evaporator.

Cessna, Leintz, and Badger [Trans. Kerr Louisiana Agr. In Fig. These curves show apparent coefficients, but sugar solutions have boiling-point rises low enough not to affect the results noticeably. Effect of Noncondensables on Heat Transfer Most of the heat transfer in evaporators does not occur from pure steam but from vapor evolved in a preceding effect. This vapor usually contains inert gases—from air leakage if the preceding effect was under vacuum, from air entrained or dissolved in the feed, or from gases liberated by decomposition reactions. To prevent these inerts from seriously impeding heat transfer, the gases must be channeled past the heating surface and vented from the system while the gas concentration is still quite low.

The primary effect, however, results from the formation at the heating surface of an insulating blanket of gas through which the steam must diffuse before it can condense.

The latter effect can be treated as an added resistance or fouling factor equal to 6. Inert-gas concentrations may vary by a factor of or more between vapor inlet and vent outlet, so these relationships should be integrated through the tube bundle. Nomenclature Use consistent units. Applications One typical application in heat transfer with batch operations is the heating of a reactor mix, maintaining temperature during a reaction period, and then cooling the products after the reaction is complete.

This subsection is concerned with the heating and cooling of such systems in either unknown or specified periods. The technique for deriving expressions relating time for heating or cooling agitated batches to coil or jacket area, heat-transfer coefficients, and the heat capacity of the vessel contents was developed by Bowman, Mueller, and Nagle [Trans.

The assumptions made were that 1 U is constant for the process and over the entire surface, 2 liquid flow rates are constant, 3 specific heats are constant for the process, 4 the heating or cooling medium has a constant inlet temperature, 5 agitation produces a uniform batch fluid temperature, 6 no partial phase changes occur, and 7 heat losses are negligible. The developed equations are as follows.

If any of the assumptions do not apply to a system being designed, new equations should be developed or appropriate corrections made. Heat exchangers are counterflow except for the exchangers, which are one-shell-pass, two-tube-pass, parallel-flow counterflow. The subscript 0 refers to the makeup. An alternative method for all multipass-exchanger gases, including those presented as well as cases with two or more shells in series, is as follows: 1. Determine UA for using the applicable equations for counterflow heat exchangers. Use the initial batch temperature T1 or t1.

Calculate the outlet temperature from the exchanger of each fluid. This will require trial-and-error methods. Note the FT correction factor for the corrected mean temperature difference. See Fig. Repeat steps 2, 3, and 4 by using the final batch temperature T2 and t2. FT can be raised by increasing the flow rate of either or both of the flow streams.

Increasing flow rates to give values well above 0. If FT varies widely from one end of the range to the other, FT should be determined for one or more intermediate points. The average should then be determined for each step which has been established and the average of these taken for use in step 6. Effect of External Heat Loss or Gain If heat loss or gain through the vessel walls cannot be neglected, equations which include this heat transfer can be developed by using energy balances similar to those used for the derivations of equations given previously.

Basically, these equations must be modified by adding a heat-loss or heat-gain term. A similar energy balance would apply to a vessel being cooled. Equivalent-Area Concept The preceding equations for batch operations, particularly Eq. However, different surfaces of a tank, such as the top which would not be in contact with the tank contents and the bottom, may have coefficients of heat transfer which are different from those of the vertical tank walls. Tanks on legs, indoors, not insulated Tanks on legs, indoors, insulated 1 in. Flat-bottom tanks, indoors, not insulated Flat-bottom tanks, indoors, insulated 1 in.

U is usually taken as Us. Table lists typical values for Us and expressions for Ae for various tank configurations. Nonagitated Batches Cases in which vessel contents are vertically stratified, rather than uniform in temperature, have been treated by Kern op. These are of little practical importance except for tall, slender vessels heated or cooled with external exchangers. The result is that a smaller exchanger is required than for an equivalent agitated batch system that is uniform. Storage Tanks The equations for batch operations with agitation may be applied to storage tanks even though the tanks are not agitated.

This approach gives conservative results. The important cases nonsteady state are: 1. Tanks cool; contents remain liquid. This case is relatively simple and can easily be handled by the equations given earlier. Tanks cool, contents partially freeze, and solids drop to bottom or rise to top. This case requires a two-step calculation. The first step is handled as in case 1. The second step is calculated by assuming an isothermal system at the freezing point.

It is possible, given time and a sufficiently low ambient temperature, for tank contents to freeze solid. Tanks cool and partially freeze; solids form a layer of selfinsulation. This complex case, which has been known to occur with heavy hydrocarbons and mixtures of hydrocarbons, has been discussed by Stuhlbarg [Pet. Refiner, 38, Apr.

The contents in the center of such tanks have been known to remain warm and liquid even after several years of cooling. It is very important that a melt-out riser be installed whenever tank contents are expected to freeze on prolonged shutdown. The purpose is to provide a molten chimney through the crust for relief of thermal expansion or cavitation if fluids are to be pumped out or recirculated through an external exchanger.

An external heat tracer, properly located, will serve the same purpose but may require more remelt time before pumping can be started. Typical coil coefficients are listed in Table More exact values can be calculated by using the methods for natural convection or forced convection given elsewhere in this section.

One method of allowing for shutdown is to add a safety factor to Eq. In the case of a tank maintained at temperature with internal coils, the coils are usually designed to cover only a portion of the tank. Steam — lb. Agitated r. Because of the many factors affecting heat transfer, such as viscosity, temperature difference, and coil size, the values in this table should be used primarily for preliminary design estimates and checking calculated coefficients.

The safety factor used in the calculations is a matter of judgment based on confidence in the design. A value of 1. Typical design parameters are shown in Tables and Helical and spiral coils are most commonly shop-fabricated, while the hairpin pattern is generally field-fabricated. The helical coils are used principally in process tanks and pressure vessels when large areas a b for rapid heating or cooling are required.

In general, heating coils are placed low in the tank, and cooling coils are placed high or distributed uniformly through the vertical height. Stocks which tend to solidify on cooling require uniform coverage of the bottom or agitation. A maximum spacing of 0. For smaller pipe or for low-temperature heating media, closer spacing should be used. In the case of the common hairpin coils in vertical cylindrical tanks, this means adding an encircling ring within mm 6 in of the tank wall see Fig.

The coils should be set directly on the bottom or raised not more than The coil inlet should be above the liquid level or an internal melt-out riser installed to provide a molten path for liquid expansion or venting of vapors. Coils may be sloped to facilitate drainage. When it is impossible to do so and remain close enough to the bottom to get proper remelting, the coils should be blown out after usage in cold weather to avoid damage by freezing. Most coils are firmly clamped but not welded to supports. Supports should allow expansion but be rigid enough to prevent uncontrolled motion see Fig.

Nuts and bolts should be securely fastened. Reinforcement of the inlet and outlet connections through the tank wall is recommended, since bending stresses due to thermal expansion are usually high at such points. The tube-side heat-transfer coefficient, high-pressure, or layout problems may lead to the use of smaller-size pipe. The wall thickness selected varies with the service and material. Carbon steel coils are often made from schedule 80 or heavier pipe to allow for corrosion. When stainless-steel or other high-alloy coils are not subject to corrosion or excessive pressure, they may be of schedule 5 or 10 pipe to keep costs at a minimum, although high-quality welding is required for these thin walls to assure trouble-free service.

Methods for calculating heat loss from tanks and the sizing of tank coils have been published by Stuhlbarg [Pet. Refiner, 38, April ]. Fin-tube coils are used for fluids which have poor heat-transfer characteristics to provide more surface for the same configuration at reduced cost or when temperature-driven fouling is to be minimized. Fin tubing is not generally used when bottom coverage is important.

ISBN 13: 9780071422949

Fin-tube tank heaters are compact prefabricated bundles which can be brought into tanks through manholes. These are normally installed vertically with longitudinal fins to produce good convection currents. To keep the heaters low in the tank, they can be installed horizontally with helical fins or with perforated longitudinal fins to prevent entrapment. Fin tubing is often used for heat-sensitive material because of the lower surface temperature for the same heating medium, resulting in a lesser tendency to foul.

Plate or panel coils made from two metal sheets with one or both embossed to form passages for a heating or cooling medium can be used in lieu of pipe coils. Panel coils are relatively light in weight, easy to install, and easily removed for cleaning. They are available in a range of standard sizes and in both flat and curved patterns.

Process tanks have been built by using panel coils for the sides or bottom. A serpentine construction is generally utilized when liquid flows through the unit. Header-type construction is used with steam or other condensing media. Standard glass coils with 0.

Also available are plate-type units made of impervious graphite. Teflon Immersion Coils Immersion coils made of Teflon fluorocarbon resin are available with 2. The flexible bundles are available with , , , , and tubes with standard lengths varying in 0. These coils are most commonly used in metal-finishing baths and are adaptable to service in reaction vessels, crystallizers, and tanks where corrosive fluids are used. Bayonet Heaters A bayonet-tube element consists of an outer and an inner tube.

These elements are inserted into tanks and process vessels for heating and cooling purposes. Often the outer tube is of expensive alloy or nonmetallic e. In glass construction, elements with External Coils and Tracers Tanks, vessels, and pipe lines can be equipped for heating or cooling purposes with external coils. These are generally 9. External coils spaced away from the tank wall exhibit a coefficient of around 5. Direct contact with the tank wall produces higher coefficients, but these are difficult to predict since they are strongly dependent upon the degree of contact.

The use of heat-transfer cements does improve performance. These puttylike materials of high thermal conductivity are troweled or caulked into the space between the coil and the tank or pipe surface. Costs of the cements in varied from 37 to 63 cents per pound, with requirements running from about 0. A rule of thumb for preliminary estimating is that the per-foot installed cost of tracer with cement is about double that of the tracer alone. Jacketed Vessels Jacketing is often used for vessels needing frequent cleaning and for glass-lined vessels which are difficult to equip with internal coils.

The jacket eliminates the need for the coil yet gives a better overall coefficient than external coils. However, only a limited heat-transfer area is available. The conventional jacket is of simple construction and is frequently used. It is most effective with a condensing vapor. A liquid heat-transfer fluid does not maintain uniform flow characteristics in such a jacket. Nozzles, which set up a swirling motion in the jacket, are effective in improving heat transfer.


Wall thicknesses are often high unless reinforcement rings are installed. Spiral baffles, which are sometimes installed for liquid services to improve heat transfer and prevent channeling, can be designed to serve as reinforcements. A spiral-wound channel welded to the vessel wall is an alternative to the spiral baffle which is more predictable in performance, since cross-baffle leakage is eliminated, and is reportedly lower in cost [Feichtinger, Chem.

The half-pipe jacket is used when high jacket pressures are required. The flow pattern of a liquid heat-transfer fluid can be controlled and designed for effective heat transfer. The dimple jacket offers structural advantages and is the most economical for high jacket pressures. The low volumetric capacity produces a fast response to temperature changes. A few typical fin configurations are shown in Fig. Longitudinal fins are used in double-pipe exchangers. Transverse fins are used in cross-flow and shell-and-tube configurations.

High transverse fins are used mainly with low-pressure gases; low fins are used for boiling and condensation of nonaqueous streams as well as for sensible-heat transfer. Finned surfaces have been proven to be a successful means of controlling temperature driven fouling such as coking and scaling.

Fin spacing should be great enough to avoid entrapment of particulate matter in the fluid stream 5 mm minimum spacing. The area added by the fin is not as efficient for heat transfer as bare tube surface owing to resistance to conduction through the fin. Pressure drop is particularly sensitive to geometrical parameters, and available correlations should be extrapolated to geometries different from those on which the correlation is based only with great caution and conservatism.

The best correlation is that of Robinson and Briggs [Chem. Low Fins Low-finned tubing is generally used in shell-and-tube configurations. For sensible-heat transfer, only minor modifications are needed to permit the shell-side method given earlier to be used for both heat transfer and pressure [see Briggs, Katz, and Young, Chem.

The efficiency curves for some common fin configurations are given in Figs. High Fins To calculate heat-transfer coefficients for cross-flow to a transversely finned surface, it is best to use a correlation based on experimental data for that surface. Such data are not often available, and a more general correlation must be used, making allowance for the possible error. Fouling refers to any change in the solid boundary separating two heat transfer fluids, whether by dirt accumulation or other means, which results in a decrease in the rate of heat transfer occurring across that boundary.

Fouling may be classified by mechanism into six basic categories: 1. Corrosion fouling. The heat transfer surface reacts chemically with elements of the fluid stream producing a less conductive, corrosion layer on all or part of the surface. Organisms present in the fluid stream are attracted to the warm heat-transfer surface where they attach, grow, and reproduce.

The two subgroups are microbiofoulants such as slime and algae and macrobiofoulants such as snails and barnacles. Particulate fouling. Particles held in suspension in the flow stream will deposit out on the heat-transfer surface in areas of sufficiently lower velocity. Chemical reaction fouling ex. Chemical reaction of the fluid takes place on the heat-transfer surface producing an adhering solid product of reaction. Precipitation fouling ex. A fluid containing some dissolved material becomes supersaturated with respect to this material at the temperatures seen at the heat-transfer surface.

Freezing fouling. Control of Fouling Once the combination of mechanisms contributing to a particular fouling problem are recognized, methods to substantially reduce the fouling rate may be implemented. For the case of corrosion fouling, the common solution is to choose a less corrosive material of construction balancing material cost with equipment life. In the case of particulate fouling, one of the more common types, insuring a sufficient flow velocity and minimizing areas of lower velocities and stagnant flows to help keep particles in suspension is the most common means of dealing with the problem.

For water, the recommended tubeside minimum velocity is about 0. This may not always be possible for moderate to high-viscosity fluids where the resulting pressure drop can be prohibitive. In general, shellsideparticulate fouling will be greatest for segmentally baffled bundles in the regions of low velocity and the TEMA-fouling factors which are based upon the use of this bundle type should be used.

Some examples are the plate and frame exchanger, the spiral plate exchanger, and the twisted tube exchanger, all of which have dispensed with baffles altogether and use the heat-transfer surface itself for bundle support. The general rule for these designs is to provide between 25 and 30 percent excess surface to compensate for potential fouling, although this can vary in special applications. For the remaining classifications—polymerization, precipitation, and freezing—fouling is the direct result of temperature extremes at the heat-transfer surface and is reduced by reducing the temperature difference between the heat-transfer surface and the bulk-fluid stream.

Conventional wisdom says to increase velocity, thus increasing the local heat-transfer coefficient to bring the heat-transfer surface temperature closer to the bulk-fluid temperature. However, due to a practical limit on the amount of heat-transfer coefficient increase available by increasing velocity, this approach, although better than nothing, is often not satisfactory by itself. A more effective means of reducing the temperature difference is by using, in concert with adequate velocities, some form of extended surface.

In cases where unfinned tubing in a triangular tube layout would not be acceptable because fouling buildup and eventual mechanical cleaning are inevitable, extended surface should be used only when the exchanger construction allows access for cleaning. Fouling Transients and Operating Periods Three common behaviors are noted in the development of a fouling film over a period of time.

One is the so-called asymptotic fouling in which the speed of fouling resistance increase decreases over time as it approaches some asymptotic value beyond which no further fouling can occur. This is commonly found in temperature-driven fouling. A second is linear fouling in which the increase in fouling resistance follows a straight line over the time of operation. This could be experienced in a case of severe particulate fouling where the accumulation of dirt during the time of operation did not appreciably increase velocities to mitigate the problem.

The third, falling rate fouling, is neither linear nor asymptotic but instead lies somewhere between these two extremes. The rate of fouling decreases with time but does not appear to approach an asymptotic maximum during the time of operation. This is the most common type of fouling in the process industry and is usually the result of a combination of different fouling mechanisms occurring together. The optimum operating period between cleanings depends upon the rate and type of fouling, the heat exchanger used i. As noted above, care must be taken in the use of fouling factors for exchanger design, especially if the exchanger configuration has been selected specifically to minimize fouling accumulation.

An oversurfaced heat exchanger which will not foul enough to operate properly can be almost as much a problem as an undersized exchanger. Removal of Fouling Deposits Chemical removal of fouling can be achieved in some cases by weak acid, special solvents, and so on. Other deposits adhere weakly and can be washed off by periodic operation at very high velocities or by flushing with a high-velocity steam or water jet or using a sand-water slurry. These methods may be applied to both the shell side and tube side without pulling the bundle. Many fouling deposits, however, must be removed by positive mechanical action such as rodding, turbining, or scraping the surface.

These techniques may be applied inside of tubes without pulling the bundle but can be applied on the shellside only after bundle removal. Even then there is limited access because of the tube pitch and rotated square or large triangular layouts are recommended. In many cases, it has been found that designs developed to minimize fouling often develop a fouling layer which is more easily removed.

Fouling Resistances There are no published methods for predicting fouling resistances a priori. The accumulated experience of exchanger designers and users was assembled more than 40 years ago based primarily upon segmental-baffled exchanger bundles and may be found in the Standards of Tubular Exchanger Manufacturers Association TEMA.

In the absence of other information, the fouling resistances contained therein may be used.

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Values from these tables may be used for preliminary estimating purposes. They should not be used in place of the design methods described elsewhere in this section, although they may serve as a useful check on the results obtained by those design methods. For process and mechanical-design considerations, see the referenced sections. Thermal design concerns itself with sizing the equipment to effect the heat transfer necessary to carry on the process. There are two general methods for determining numerical values for Uco, Ucv , Uct, and Ura. One is by analysis of actual operating data.

Values so obtained are used on geometrically similar systems of a size not too different from the equipment from which the data were obtained. The second method is predictive and is based on the material properties and certain operating parameters. Conductive Heat Transfer Heat-transfer equipment in which heat is transferred by conduction is so constructed that the solids load burden is separated from the heating medium by a wall. Values of Uco are reported in Secs. To convert British thermal units per hour-square foot-degrees Fahrenheit to joules per square meter-second-kelvins, multiply by 5.

For details of terminology, equation development, numerical values of terms in typical equipment and use, see Holt [Chem. Equation is applicable to burdens in the solid, liquid, or gaseous phase, either static or in laminar motion; it is applicable to solidification equipment and to divided-solids equipment such as metal belts, moving trays, stationary vertical tubes, and stationaryshell fluidizers. Fixed or packed bed operation occurs when the fluid velocity is low or the particle size is large so that fluidization does not occur.

For such operation, Jakob Heat Transfer, vol. Other correlations are those of Leva [Ind. NOTE: To convert British thermal units per hour-square foot-degrees Fahrenheit to joules per square meter-second-kelvins, multiply by 5. Fluidization occurs when the fluid flow rate is great enough so that the pressure drop across the bed equals the weight of the bed. Wen and Fau [Chem. The low end of the range is recommended. Volume specific heat, B.

Thermal diffusivity, sq. For dilute fluidized beds on the shell side of an unbaffled tubular bundle Genetti and Knudsen [Inst. London Symp.

Gmf, the minimum fluidizing velocity, is defined by 0. Wender and Cooper [Am. Solidification involves heavy heat loads transferred essentially at a steady temperature difference. It also involves the varying values of liquid- and solid-phase thickness and thermal diffusivity. Later work by Hashem and Sliepcevich [Chem. The heat-transfer rate is found to be substantially higher under conditions of agitation. The heat transfer is usually said to occur by combined conductive and convective modes.

A discussion and explanation are given by Holt [Chem. Prediction of Uco by Eq. To date so little work has been performed in evaluating the effect of mixing parameters that few predictions can be made. However, for agitated liquid-phase devices Eq. Holt loc. This is applicable for such devices as agitated pans, agitated kettles, spiral conveyors, and rotating shells.

The solids passage time through rotary devices is given by Saemann [Chem. From these equations a predictive equation is developed for rotaryshell devices, which is analogous to Eq. Contactive Direct Heat Transfer Contactive heat-transfer equipment is so constructed that the particulate burden in solid phase is directly exposed to and permeated by the heating or cooling medium Sec.

The carrier may either heat or cool the solids. A large amount of the industrial heat processing of solids is effected by this mechanism. Physically, these can be classified into packed beds and various degrees of agitated beds from dilute to dense fluidized beds. Heat Transfer Conf. This temperature difference is also applicable for well-fluidized beds of small particles in cross-flow as in various vibratory carriers.

The packed-bed-to-fluid heat-transfer coefficient has been investigated by Baumeister and Bennett [Am. Kunii and Suzuki [Int. Heat Mass Transfer, 10, ] discuss heat and mass transfer in packed beds of fine particles. Particle-to-fluid heat-transfer coefficients in gas fluidized beds are predicted by the relation Zenz and Othmer, op. A more general equation is given by Frantz [Chem. Bed-to-wall coefficients in dilute-phase transport generally can be predicted by an equation of the form of Eq. Farber and Morley [Ind. Physical properties used were those of the transporting gas.

See Zenz and Othmer op. This term indirectly includes an area factor so that thermal performance is governed by a cross-sectional area rather than by a heated area. Another such equation, for stationary vertical-shell and some horizontal rotary-shell and pneumatic-transport devices in which the gas flow is parallel with and directionally concurrent with the fluidized bed, is the same as Eq. If the operation involves drying or chemical reaction, the heat load Q is much greater than for sensibleheat transfer only. Also, the gas flow rate to provide moisture carry-off and stoichiometric requirements must be considered and simultaneously provided.

A good treatise on the latter is given by Pinkey and Plint Miner. Evaporative cooling is a special patented technique that often can be advantageously employed in cooling solids by contactive heat transfer.

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The drying operation is terminated before the desired final moisture content is reached, and solids temperature is at a moderate value. The cooling operation involves contacting the burden preferably fluidized with air at normal temperature and pressure.

Perrys Chemical Engineers Handbook Eighth Edition

The air adiabatically absorbs and carries off a large part of the moisture and, in doing so, picks up heat from the warm or hot solids particles to supply the latent heat demand of evaporation. Using m3 ft3 of airflow at normal temperature and pressure at 40 percent relative humidity to carry off 0. Using the lowered solids temperature as t3 and calculating the remainder of the heat to be removed in the regular manner by Eq. The required air quantity for 2 must be equal to or greater than that for 1.

When the solids heat capacity is higher as is the case for most organic materials , the temperature reduction is inversely proportional to the heat capacity. A nominal result of this technique is that the required airflow rate and equipment size is about two-thirds of that when evaporative cooling is not used.

See Sec. Convective Heat Transfer Equipment using the true convective mechanism when the heated particles are mixed with and remain with the cold particles is used so infrequently that performance and sizing equations are not available. Such a device is the pebble heater as described by Norton Chem. For operation data, see Sec. Convective heat transfer is often used as an adjunct to other modes, particularly to the conductive mode.

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A pseudo-convective heat-transfer operation is one in which the heating gas generally air is passed over a bed of solids. Its use is almost exclusively limited to drying operations see Sec. The operation, sometimes termed direct, is more akin to the conductive mechanism. For this operation, Tsao and Wheelock [Chem. Note here that the temperature-difference determination of the operation is a simple linear one and of a steady-state nature. Also note that the operation is a function of the airflow rate. Further, the solids are granular with a fairly uniform size, have reasonable capillary voids, are of a firm texture, and have the particle surface wetted.

The reader should refer to the full reference article by Tsao and Wheelock loc. Presence of air is not necessary see Sec. In fact, if air in the intervening space has a high humidity or CO2 content, it acts as an energy absorber, thereby depressing the performance. For the emissivity values, particularly of the heat source es, an important consideration is the wavelength at which the radiant source emits as well as the flux density of the emission. Data for these values are available from Polentz [Chem. Both give radiated flux density versus wavelength at varying temperatures.

Often, the seemingly cooler but longer wavelength source is the better selection. For some emissivity values see Table Important considerations in the application of the foregoing equations are: 1. Since the temperature of the emitter is generally known preselected or readily determined in an actual operation , the absorptivity value er is the unknown.

This absorptivity is partly a measure of the ability of radiant heat to penetrate the body of a solid particle or a moisture film instantly, as compared with diffusional heat transfer by conduction. Such instant penetration greatly reduces processing time and case-hardening effects. Moisture release and other mass transfer, however, still progress by diffusional means.

Therefore, the absorptivity, color, and nature of the solids are of little importance. For drying, it is important to provide a small amount of venting air to carry away the water vapor. This is needed for two reasons. Second, water-vapor accumulation depresses further vapor release by the solids. Now updated to reflect the latest technology and processes of the new millennium, the Eighth Edition of this classic guide provides unsurpassed coverage of every aspect of chemical engineering-from fundamental principles to chemical processes and equipment to new computer applications.

Filled with over detailed illustrations, the Eighth Edition of Perry's Chemcial Engineering Handbook features: Comprehensive tables and charts for unit conversion A greatly expanded section on physical and chemical data New to this edition: the latest advances in distillation, liquid-liquid extraction, reactor modeling, biological processes, biochemical and membrane separation processes, and chemical plant safety practices with accident case histories Inside This Updated Chemical Engineering Guide.

Donald W. Green is chair and the Deanne E. Ackers distinguished professor of chemical and petroleum engineering at the University of Kansas. See All Customer Reviews. Shop Textbooks. Add to Wishlist. USD Ship This Item — This item is available online through Marketplace sellers.