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A Systematic Procedure for Retrofitting Chemical Plants To Operate Utilizing Different Reaction Paths David A. Nelson+and James M. Douglas* Chemical Engineering Department, University of Massachusetts, Amherst, Massachusetts 01003
A systematic procedure has been developed and a software code created t o examine continuous petrochemical plant retrofit problems. The procedure is hierarchical, and it provides stopping points so t h a t little work needs t o be done t o reject a poor project. It is structured in a way t h a t first sets minimum targets for the retrofit project, and it uses a n efficient flow sheet structure filter t o look at only the subset of possible flow sheet structures that have the potential for being profitable. T h e procedure then uncouples the flow sheet evaluation problem into two parts. First, t h e flow sheet is examined by neglecting the energy integration network. Second, the best flow sheet found without considering t h e heat-exchanger network is then examined by using existing energy integration procedures and considering the existing heat-exchanger area. In this paper, we have looked at t h e specific problem of retrofitting an existing plant to produce a new chemical by a different reaction path. As a n example, we examined the conversion of an existing plant t o produce benzene by t h e hydrodealkylation of toluene t o t h e production of mixed xylenes by toluene disproportionation.
Introduction Because of changing economic forces in the chemical industry, there is great interest in the retrofitting of chemical processes. Retrofitting involves modifying some of the process units and/or changing the structure of the flow sheet in an attempt to significantly improve the process economics. Retrofit studies are also useful if we want to increase the capacity of an existing plant; to introduce a new technology, such as a new catalyst, a new tower packing, a new membrane separator, etc.; and/or to switch to a new feedstock which is less expensive. Different Classes of Retrofit Problems. There are many different kinds of retrofit problems demanding different levels of effort and expertise. Four broad classes of retrofit problems which are discussed in the literature and are commonly addressed by industry are shown in Table I. The first three types of problems concentrate on a particular unit operation and either make it operate more efficiently or replace it with something new and different. Integrating new technology into an existing process can also involve operations that were not considered when a plant was constructed. For example, gas-permeable membranes have become useful in separating a stream of light gases. When many existing plants were constructed, light gas separations were considered to be too expensive and were avoided. Now it is possible to put a membrane separator on a vent stream that originally would have been combusted in order to recover a valuable chemical. The last class of retrofit problems in Table I involves the development of procedures to look at an entire process flow sheet and by considering the interactions between unit operations and process streams to find a way to increase profits. That is the kind of problem this research addresses. This kind of problem can be thought of as a large optimization problem. It is an optimization not only over the operating parameters in the flow sheet itself but also over the flow sheet structure itself. To obtain the maximum profit, we should consider changes in the connections of unit operations and replacement of the unit operations that are in the flow sheet. How to handle this optimization is an important issue in solving a retrofit problem. In grassroots design, there * T o whom all correspondence should be addressed. ‘Present address: Chevron Research and Technology Co., P.O. Box 1627, Richmond, CA 94802.
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Table I. Classes of Retrofit Problems (a) improvement studies on the efficiency of a particular unit operation (b) integrating new technology (new unit operations) into an existing process (c) studies to debottleneck the process (d) systematic studies that examine the entire process flow sheet and its interactions to either reduce costs or increase profits
have been basically two approaches to solving this structural optimization problem. One approach has been to use sophisticated optimization theory to rigorously solve the underlying optimization problem. In the past, this approach suffered from the difficulty of the optimization problem, the amount of time required to solve the problem, and the difficulty in selecting a representative superstructure for a specific problem, although improved techniques are being developed. The other approach has used heuristics derived from physics and economics and provided by experienced designers to try to “pick” a flow sheet structure or a t least pick a small number of structures to evaluate. In this research, we build on the works of Kirkwood (1987), Kirkwood et al. (19881, and Douglas (19881, which select flow sheet structures via heuristics for grassroots design. We essentially use their techniques to reduce the number of flow sheet structures we have to evaluate down to a manageable number (one to five) and then evaluate these flow sheets in the context of the existing plant. The above three references demonstrate the utility of using heuristics for selecting flow sheet structures. They also explain why heuristics derived from physics can be very accurate and discuss their limitations. Kocis and Grossmann (1989) solve Douglas’s HDA (hydrodealkylation of toluene to form benzene) process as a MINLP (mixed-integer nonlinear programming) problem. The results they achieve are similar to Douglas’s result (Douglas, 19881, although they have one additional process unit (a gas-permeable-membrane separator) that Douglas did not consider. The similarity of the results demonstrates that the heuristic approach is useful, especially in screening calculations where high accuracy is not expected. Our approach sets out to solve class d problems in Table I. However, classes b and c must be solved in order to completely solve class d problems (which our procedure does). New flow sheet structures are suggested that contain new technology, and our retrofit evaluator can see C 1990 American Chemical Societv
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Table 11. Procedure for Screening Process Retrofit Opportunities (Fisher et al., 1987) _______ - . (1) use an operating cost diagram to identify the incentive for raw materials and energy savings ( 2 ) determine the incentive for completely replacing the entire plant (a) estimate the optimum values of the design variables with current COSIS (b) identify important process alternatives (3) screen the process alternatives, and find the best flow sheet if we replace the plant ( 4 ) modify the existing equipment sizes for the exiPting flow sheet or a structural alternative ( a ) eliminate the existing heat exchangers. but retain the heating and cooling utilities costs cb) identify the dominant operating variables (c) identify the equipment that constraints the dominant operating variables (d) remove the equipment constraints by adding incremental equipment capacity until the incremental investment costs balance the incremental savings in operating costs (e) develop a heat-exchanger network for the process If) modify the new heat-exchanger network in order to use as much existing heat-exchange equipment as possible (g) reoptimize the process flows and heat-exchanger network (5) refine the retrofit calculations
what effect this new equipment has on the existing flow sheet. We also can easily see how to debottleneck the process as we evaluate a flow sheet of interest at a desired production rate.
A Systematic Procedure for Retrofitting Chemical Plants The retrofit screening strategy described by Fisher et al. (1987) provides a framework for screening retrofit projects. The procedure is easy to describe but is very difficult to implement with conventional process simulators. The work described in this paper refined the procedure, and implemented the result in a way that demonstrates its utility in a wide variety of industrial problems. The obvious way to implement and demonstrate the procedure was to develop an interactive computer code suitable for use by a process engineer and use the code to retrofit as many industrial plants as possible. This follows the work by Kirkwood et al. (1988),who developed PIP (process invention procedure) to perform grassroots plant design. The code described here extends the existing PIP code, capitalizing on its large base of models, methods, and subroutines but adding a set of performance models and the logic for switching between performance and design calculations. The resulting new PIP code can examine both grassroots design and retrofit design problems. Differences from Fisher's Procedure. Fisher's procedure (shown in Table I1 (adapted from Fisher et al. (1987))) contains steps that are obviously useful and necessary but are not implemented in the software. A more complete discussion of Fisher's procedure and a detailed discussion of what additions and modifications were made for this work can be found in Nelson (1989). The following paragraphs briefly discuss the additions and modification. Preparing an operating cost diagram is a necessary step done before an engineer approaches the computer. It is the fundamental step of gathering data about the process and placing it on an appropriate diagram. On the basis of this diagram, a decision is made on whether additional retrofit work on a specific project is merited. Screening process alternatives using Douglas and Woodcock's (1985) quick screening procedure (step three in Fisher's procedure) are an appropriate tool when suitable computer software is not available, but when software like
Table 111. Procedure for Screening Process Retrofits (1) design a new plant, and evaluate the significant structural alternatives ( 2 ) enter the existing structure of the flow sheet ( 3 ) enter the existing equipment sizes, and tune the equipment in od e1s ! 4 retrofit ~ the existing proce.ss flow sheet (neglecting process heat exchangers) ( 5 ) retrofit any significant structural alternatives (neglecting process heat exchangers) 16) energy integrate the best remaining few flow sheets
PIP is available, it is simpler and faster (for the user) to simply use the software to design in more detail all the significant structural alternatives. With PIP, all of the significant alternatives can be evaluated in less than one-half day. Additional Steps. Several features of the retrofit problem that Fisher et al. (1987) did not anticipate became ohvious in our work. Existing plants often have extensive amounts of operational data available. Since differences between how our simple equipment models predict, equipment operation and actual operation (which we know about from the plant data) might lead us into incorrect decisions, it was determined that a model tuning step would be appropriate. Tuning our models is not thought of as a way around physically incorrect models. Tuning a simplified model will never give it the ability to predict phenomena it wasn't derived for. Tuning simply gives simple models more accuracy in domains where they are reasonably appropriate. Having more accuracy is important for the correct prediction of equipment constraints. Another feature that had bo be added was the details on how to remove process constraints. There are always several ways to remove an equipment constraint, and a variety of these techniques were implemented. Description of the Procedure. Table I11 details the procedure we implemented in our software development effort. In this section, we describe t,he function of each of these steps. The basic approach is to first get a target by designing the best new plant in order to focus our attention on a smaller group of flow sheets. We then input the existing, plant and resolve any modeling differences. The fourth and fifth steps examine the best flow sheets found in step 1, neglecting the energy integration but in the context of our existing equipment sizes. The final step performs energy integration targeting on the best remaining flow sheet (or flow sheets). Design a New Plant. A key step in the procedure is to tic, a preliminary design of the best new plant producing the same product. This step is critical since it provides retrofit targets and acts as a very efficient filter, screening out poor flow sheet alternatives for the retrofit problem. Douglas's hierarchical procedure (Douglas. 1985, 1988) is used to perform the new plant design. This procedure is summarized in Table IV. It starts with a minimum amount of information about the products, feed streams, reactions. and thermodynamics. In a series of steps, it makes decisions to determine a flow sheet structure. At e x h slep, a partial solution is constructed and the econoniics evaluated of the candidate flow sheet structure (see Douglas (1988) for detailed description). Only processes that meet a defined economic crit,eria are carried forward t o the next step. All other possible flow sheet alternatives are tabulated at each level of detail. Designing the best new plant using Douglas's procedure f o r o u r retrofit makes several assumptions that the retrofit procedure depends upon. It assumes that Douglas's procedure does a good joh of designing the best new plant.
Ind. Eng. Chem. Res., Vol. 29, No. 5 , 1990 821 Table IV. Hierarchical Procedure for Process Design. Including Level by Level Design Decisions (1) batch versus continuous (2) input-output structure of the flow sheet (a) feedstream purification (b) excess reactants (c) reversible by-product destination (d) light component destination (3) recycle structure of the flow sheet (a) number of reactors (b) reactor type (c) heating/cooling policy (d) external heating/cooling specifications (e) modifications of equilibrium limited reactions (4) general structure of the separation system (a) type of vapor recovery system (b) location of vapor recovery system (c) distillation column sequencing ( 5 ) heat-exchanger networks
While this is an open question, it is reasonable to assume that for screening calculations will be near the best new plant conditions. In addition, the most important design aspect provided by this step is the flow sheet structure, not the exact equipment sizes and operating parameters. Douglas’s procedure does a very good job of identifying the best flow sheet structures (best 1-5). For the case where we are simply improving an existing process, their is an assumption that something has changed like market conditions and utility or feedstock costs or that the original process design was poorly optimized. If none of these effects are present, the best new plant design will be similar to the current plant, and this step will have concluded that no retrofit is in order. We also assume that partial solutions evaluated in the new plant design stage act as a useful filter to eliminate flow sheet alternatives from further consideration (they act as a filter in the sense that only “good designs” pass through this step; “bad designs” are rejected and not considered further). This should not be confused with a mathematical filter to eliminate white noise in data (like a Karmon filter). In the initial steps of Douglas’s hierarchical procedure, very little equipment is considered (they primarily focus on raw materials costs). Since there is little difference between the new plant and retrofit economics at this level of detail, it is reasonable to reject retrofit alternatives in these steps. Filtering potential structures via a new plant design procedure has one drawback. When the complete new plant design (final step in Douglas’s procedure) is used, the presence of existing equipment significantly affects the economics and what the optimum flow sheet structure and flows might be. While it is reasonable and correct to reject poor flow sheets in the early steps of Douglas’s procedure, it is not in the later steps. The message of using Douglas’s new plant design procedure for eliminating possible flow sheets is to reject flow sheets early in the procedure but not to reject marginal flow sheets in the later steps to avoid missing optimal solutions. Entering the Existing Structure of the Flow Sheet. This is just the simple step of entering the flow sheet structure of the users plant. This consists of making the structural decisions required by PIP, which correspond to the users plant. By answering the specific questions that fill out the information required in Table IV about the users plant, we force consistency of flow sheet structures. This eliminates, to some degree, any differences in plant structure that are due to semantics or conventions used when drawing flow sheet structure. As the software gets answers to questions, it places equipment in the flow sheet
and uses its built-in knowledge of process design to connect the equipment. The user can also accomplish this task as the user designs the best new plant. When the user designs the plant structure corresponding to the existing structure, they are essentially completing this step. When they actually perform this step, the software will use their earlier results from the new plant design step to set default answers for all of the questions. The user is just confirming the answers. The user is not asked to enter information about equipment we neglect. While a typical plant will require many pumps and valves for operation, we neglect them since their costs normally are negligible. The user is also asked to specify the distillation sequence in the existing plant. This is a new query, since in PIP distillation sequences are selected by exhaustive enumeration. The default answer will be the least expensive grassroots design sequence. A more straightforward way, from the user’s point of view, to enter his existing flow sheet structure might be to provide a point-and-draw interface simill: to what many of the current process simulators are providing (see Simulation Sciences Inc., 1987). This would allow the user to draw the flow sheet in a form he is familiar with. The next step would involve the software resolving the user’s information into a form it could readily use, while checking for consistency of the user’s input as well as to see if the data are consistent with the limitations of the program. We decided not to use this method because of the large amount of programming involved and the research nature of this work. It is also probably faster to enter the data in the way we have designed, especially since the exact same task is accomplished as the user designs the best new plant. Entering the Existing Equipment Sizes and Tuning the Equipment Models. Here the user enters the sizes of the equipment in his plant. The equipment present was indicated in step 2 of this procedure. The user can then tune the equipment models. The software will indicate to the user the size of each piece of equipment. The software predicts these sizes based on a new plant design a t the base case operating variable values. The user then indicates the actual size of the equipment. Based on the differences between the predicted and actual size, the software will calculate a value for a tuning parameter that will force the model to absolutely agree a t the base case. The user can modify this tuning parameter based on the degree of excess capacity the particular piece of equipment has. Basically, the user is asked to convey two facts about a piece of equipment: (1)the size and ( 2 ) how much extra capacity it has. Retrofitting the Existing Process Flow Sheet. In step 1 of the retrofit procedure, we use PIP to design a new plant and examine the process alternatives. In this step, we want to optimize the significant operating variables for the existing flow sheet, but in the context of the limitations of the existing equipment. This involves calculating the material flows and then determining if the equipment can accomplish its task. If a piece of equipment is constrained, we must remove the constraint. The program has a portfolio of methods to break the constraints the user can select from. The available constraint removal polices available currently are shown in Table V (see Nelson (1989) for a detailed description). Generally, when a piece of equipment is found to be constrained, there are several methods available to break the constraint. We define the “retrofit policy” as a method
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Table V. Constraint Removal Policies (1) gas compressors parallel unit(s) new unit(s) (2) chemical reactors parallel unitsk) new unit(s) change reactor conditions (temperature, molar ratio) ( 3 ) distillation columns parallel unit(s) new unitis) retray column (smaller tray spacing) minimum allowable reflux rate (to remove a weeping constraint)
of breaking an equipment constraint. The best choice is usually a matter of economics. However, many times the economic differences between various retrofit policies are not large. and it is not an important economic decision to select the best policy. In addition, the economics of the retrofit policy chosen may be very plant-specific. For example, there may not be room for an additional compressor in an existing building, so the user would know that adding new compressors would be excessively expensive. A new catalyst and reactor have been developed, so we will always either have to repack the existing reactor or have to put in a new reactor. Maybe a distillation column will exhibit serious corrosion problems, and we would prefer to replace it rather than add capacity in parallel. The point here is that there are many different methods of breaking constraints on a particular piece of equipment. Rather than treat this as an economic optimization problem, we will provide the user with several choices on how to break constraints. They will indicate before any analysis is done how constraints should be handled for each particular piece of equipment in the plant being examined. During the analysis, if a piece of equipment becomes constrained, the program will take appropriate action to break the constraint, only notifying the user that a constraint was encountered and the cost of breaking the constraint. During this step and the following step (retrofit significant structural alternatives), we neglect the process heat exchangers present in the flow sheet. At this stage, we are concentrating on finding significant raw material savings. After we find the best flow sheet structure without energy integration, we perform energy integration on that structure After removing the equipment constraint, we continue to change the operating variable, and we continue to change until the incremental savings in operating costs is j u s t balanced by the incremental annualized investment for the new equipment. We can stop when the incremental return on investment satisfied some corporate criteria for return on investment. If another equipment constraint is encountered before we reach the new optimum, we break this constraint also. Actually, instead of making these incremental changes, we simply run a set of case studies over the range of the operating variables where we might expect to obtain the new optimum and have the logic in the code break constraints as needed (see Nelson (1989) for a description of how the code breaks the constraints). If we are looking a t not only the changes in operating variables but also the changes in the plant production rate, we must also include increased operating profits in the economic analysis. In this case, we will look at the process economic potential (Douglas, 1985). The economic potential is simply the value of all the value products minus
the value of the feed streams, other operating costs, and incremental annual capital costs. Picking the production rate is not really an optimization problem like selecting values for the operating variables. The production rate is going to be set by market forces. Increasing the production rate will almost always result in an increased economic potential. The return on investment will decrease as we hit constraints on various pieces of equipment, but for a valuable product, the profit will just increase. If the product is marginally valuable, this may not be true, but then management is usually not interested in increasing the production rate for marginally profitable processes. For looking a t changes in the production rate, we recommend a new desired production rate be set by market forecasting. When the process is evaluated, the software will identify and remove any equipment constraints corresponding to this new production rate. We also have the facility to look at plots of production rate versus costs and economic potential. The changes in slope on plots indicate equipment constraints, and the user could determine a production rate to operate based on how much capital is available for plant expansion. Retrofitting Significant Structural Alternatives. In the previous step, we analyzed the existing flow sheet structure and optimized the operating variables. In this step, we want to perform the same task but on new flow sheet structures we identified in the new plant design step of this procedure. &’e first must make the choices corresponding to the new flow sheet structure. T o make these choices, we go to the appropriate menu in the program (this is very simple to do since menus are grouped together by hierarchical level and function) and make a new choice. For example, if we wanted to recycle a reversible byproduct to extinction, we would go the reversible byproduct menu in the input-output structure of the flow sheet level (see Table 111) and change the default choice. As we make the choices, the software automatically identifies how these choices affect other flow sheet decisions and flags them for the user’s attention as the alternative structure is created. After all the flowsheet decisions are made, we use the same techniques as described in the previous section to optimize the operating variables and analyze the effects of a changed production rate. One new structural selection available during this step is the choice of distillation sequence. For a new plant design, we selected the sequence by calculating the total annual cost for each sequence and selecting the least expensive. In retrofit design, this task is somewhat harder because we need guidance on how to use our existing distillation columns. Nikolaides (1987) proposes a method of resequencing a existing set of distillation columns. This method suggests the new plant sequencing be examined by using Glinos’s (1984) sequencing algorithm but concentrating only on the operating costs. If any of the sequences found with this algorithm offer significant savings in operating costs over the existing sequence, an attempt should be made to modify the existing sequence. Nikolaides does not offer any suggestions on how to integrate the existing columns into the target selected above. We suggest using the user’s insight into the problem to place the columns. Our method for examining the distillation train is shown in Table VI. This method does not map out the entire space of possible sequences. Its only intention is to target a new sequence of interest and to reasonably place the existing
Ind. Eng. Chem. Res., Vol. 29, No. 5, 1990 823 Table VI. Column Placement in a Target Sequence (1) choose a sequence to investigate via Glinos (1984) (2) substitute the existing columns into the new sequence by comparing the existing columns to the required columns; this step utilizes the users physical insight rather than enumerating all the possible substitutions (3) add parallel or new capacity to each existing column if required so each separation can be performed (4) calculate the incremental capital costs from step 3 and calculate a total annual cost
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columns in the sequence. Grossmann et al. (1987) shows the retrofit sequencing problem is much larger than the new plant design sequencing problem. For example, separating a 7-component mixture into 7 pure products would involve evaluating 132 sequences for the new plant design but potentially 4.5 million in a retrofit analysis. Clearly for screening calculation, where we are simply trying to get an idea of the potential savings, a quick, reasonable, heuristic method as outlined above is required. Step 2, placing the existing columns, is the key step in this procedure. The user is given a table of the current distillation column sizes and what sizes would be required if new columns were inserted. Figure 1is such a table from the P I P code. The user can directly see the similarities and differences between what is available and what is needed. The existing columns should be placed in a location with similar sizes. If they are placed in a dissimilar location, they will be used as much as possible and additional capacity added to make up the difference. Energy Integrating the Best Flow Sheet. The final step in the retrofit code is to energy integrate the best resulting flow sheet. Our strategy here is to try and find a reasonable solution very rapidly in order to show the potential for further study. The following strategy targets potential additional heat-exchanger capital and operating costs and includes the existing exchangers. The retrofit procedure uses the same targeting procedure as the grassroots design option in PIP, see Kirkwood (1987), with a few additions to account for the existing-heat exchanger area. The basic idea is to estimate the required heat-exchanger area and capital costs along with the utility costs using surface area targeting (Townsend and Linnhoff, 1983a,b). We then make a projection of the actual area required-based on the performance of our existing plant (Tjoe and Linnhoff, 1986) and estimate the required new area we have to purchase. All of the above is performed on the background process only. In the previous steps in our retrofit analysis, we considered the operating and incremental capital cost of column reboilers and condensers. Distillation columns act as energy degradation devices. They input high-temper-
ature energy in the reboiler and reject low-temperature energy out the condenser and thus do not energy integrate directly with the background process. We can, however, look a t a grand composite curve with the column energy boxes; see Androcovich and Westerberg (1985) along with Hindmarsh and Townsend (1984). By looking a t such a figure, we can see the potential for integrating a distillation column with the background process or another distillation column. If a column will fit directly above or below the background process grand composite curve or another distillation column, we can input energy into the distillation column and reject it to the other device (or vice versa). If the column does not fit above or below the grand composite curve, energy integration is not possible without changing the column conditions. The grand composite curve gives insight on how to change column conditions. If the column only needs to be moved up a few degrees (in temperature), this indicates that pressure shifting the column is probably a good possibility. Hindmarsh and Townsend (1984) give recommendations on integrating columns. Figure 2 is a grand composite curve plot for the mixed-xylene process. The grand composite curve is a new feature put into PIP for the retrofit procedure and also works for grassroots design. It is used only as a illustrative device to indicate to the user the potential for integrating the distillation sequence and the background process. The economics PIP calculates doing energy integration do not account for these integrations. They only include a problem table analysis with a target on the minimum number of exchangers and the minimum area. Tjoe and Linnhoff s (1986) retrofit targeting method is used to estimate how much more heat-exchanger area over the targeted amount will be required. This approach makes an estimate based on how much heat-exchange area is in the existing plant and how much is targeted. It calculates a factor to relate the actual amount of heatexchanger area in a process to the amount targeted:
This factor is then used to estimate the actual area required under new retrofit conditions to the targeted amount under the new conditions. Tjoe and Linnhoff used their Q factor to estimate the heat-exchange area when the flow sheet structure, flows, and temperatures remain constant. They only considered changes in the minimum approach temperature a t the pinch. This work requires a heat-exchanger retrofit after structures and flows have been changed. A procedure we
824 Ind. Eng. Chem. Res., Vol. 29, No. 5 , 1990 Table VII. Heat-Exchanger Network Retrofit Targeting (1)find Tjoe and Linnhoffs (Y factor to relate the amount of additional area the existing plant has over the area-energy targeting (2) for the flow sheet structure, mass flows and temperature under consideration perform the area-energy targeting (reference) (3) use c i found in step 1 to predict the new expected area for the plant, A, = @A,,, (4) find the incremental heat-exchnger area required, A,,, = A, A,
(5) cost of the new area, including some kind of charge for possible repiping of the existing exchangers
developed to extend Tjoe and Linnhoff's work is shown in Table VII. This procedure assumes that the value of a found in step 1 can be used to estimate the amount of additional heatexchanger area required above the standard targeting after any variables have been changed. Step 4 implicitly assumes all of the existing heat-exchange area can be utilized. This may be problematic for plants with large exchangers that might not be fully utilizable in a new flow sheet location. However, it should be reasonable even in these cases, since only small incremental errors will be obtained when the heat-exchange equipment is underestimated. The cost of repiping in step 5 is neglected. As an approximation, the user can input large installation factors that the code will use to cost the new exchanger capacity, which will approximate higher piping costs. A t this level of detail in the energy integration, very little is known about the resulting network layout and how extensively the existing exchangers will have to be repiped. Since the primary goal is to see if significant savings can be obtained by retrofitting and to see if more detailed retrofit studies can be justified, it is not expected that piping costs will change our conclusions.
Table VIII. Toluene Disproportionation Process Input Information primary product: mixed xylenes (equilibrium mixture) production rate: unknown value: $44.99"/kmol ($20.45/lb-mol) primary reaction: toluene = xylene + benzene temperature: 510 O C (950 "F) pressure: 21.8 atm (320 psia) phase: vapor secondary reactions: 2(toluene) + H, = benzene + CH4 toluene + 5.3H, = Z.d(propane) toluene + 4.4H2 = l.4pentane feed stream: 80% H,, 15% methane, 5% propane phase: vapor pressure: 1 atm (15 psia) cost: $2.2/kmol ($l/lb-mol) feed stream: 100% toluene phase: liquid pressure: 1 atm (14.7 psia) cost: $33.68/kmol ($15.31/lb-mol) byproduct values: benzene: $40.96/ kmol ($18.62/lb-mol) constraints: Hz/toluene molar ratio: 16.0 design variables conversion: unknown purge composition: unknown Chemical prices from the Chemical Marketing Reporter. Jan 30, 1989.
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Example: Retrofitting an HDA Plant To Produce Mixed Xylenes As an example of using our retrofit procedure, we examine retrofitting an existing plant to produce benzene via the hydrodealkylation of toluene (HDA). Rather than having our plant continue to make benzene, we will examine changing the reaction chemistry of the plant to make an equilibrium mixture of xylenes via toluene disproportionation. This retrofit is of interest because xylenes are somewhat more valuable than benzene, and we could retain some of our benzene production due to the stoichiometry of the primary reaction
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Current plant structures for the two plants are similar, so we expect a retrofit to be possible. The information about a toluene disproportionation plant is shown in Table VIII. This type of retrofit poses several problems. A reasonable production rate will have to be found. If we set the xylene production rate equal to the original benzene production rate, this will a t least double the plant flow. The primary reaction for production of xyelene is equilibrium limited a t about 45% toluene conversion compared to HDA plant operation of 75% conversion. This will result in substantially higher toluene recycle flow. A new reactor will be used since the existing reactor uses a different catalyst (we could also try to pack the new catalyst in the old reactor).
Figure 3. Toluene disproportionation optimal grassroots design flow sheets. The top flow sheet is for the optimal distillation sequence (sequence A). The bottom flow sheet is for the distillation sequence in the existing plant (sequence B).
Step-by-stepRetrofit to Mixed-Xylene Production. Step 1. Designing a New Plant. If we design the best new plant and examine the flow sheet alternatives, we find two profitable flow sheets as shown in Figure 3. The only difference between these two flow sheets is the distillation sequence. The most profitable sequence (sequence A) separates xylene from toluene and benzene in the second column and toluene from benzene in the final column. However, the sequence (sequence B) that separates
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Table IX. Existing Plant Sizes for the HDA Plant equipment size feed comuressor (HP) 2070.0 133.0 recycle compressor (HP) 245 reactor volume (FT3) distillation column 1 1.0 diameter (FT) 36 trays (number) distillation column 2 5.0 diameter (FT) 54 trays (number) distillation column 3 3 diameter (FT) 46 trays (number)
Figure 4. Toluene disproportionation grassroots design optimization.
benzene from toluene and xylene in the second column and toluene from xylene in the final column is only slightly less profitable. Since sequence B is the one in the existing HDA plant, it might be less expensive in our retrofit study. Both need to be studied further. Other flow sheet alternatives for this process are clearly not economical and not considered further. These include feed stream purification (the H, feed stream is very expensive to purify), light component destination (either venting the gas out of the flash, which is not economical, or using a membrane separator, which is beyond the scope of the program), reactor configuration (fixed by the cost of heat removal), or other liquid separation technology. This highlights the benefit of using heuristics for selecting flow sheet structures. When good heuristics are available, they limit the scope of the problem and eliminate timeconsuming (human as well as computer) analysis of poor designs. The economic potential versus toluene conversion and hydrogen purge composition plot is shown in Figure 4 for the flow sheet structure with distillation sequence A. This figure illustrates one important aspect of the xylene chemistry. The reaction is equilibrium limited a t about 45% conversion, and this bounds our search space in conversion with a physical limit. Figure 4 also shows this acts as an economic constraint. since selectivity losses are not great in this plant, there is no economic incentive to recycle excess toluene (low conversion). It only costs additional money in separation expenses. We would thus expect all plots of conversion versus economic potential to display a maximum a t close to the equilibrium conversion. The equilibrium conversion limit places another restraint on this retrofit problem. Since this value of conversion is significantly smaller than the existing HDA plant, the recycle flows will be high relative to the existing plant. Equipment constraints will come into play sooner than a standard retrofit problem, and we might not even be able to operate the plant a t all, without some new capital. This problem can be attacked in two ways. We can set a production rate and break constraints by standard methods to achieve operation, or we can look a t plots of production rate versus capital cost. By looking a t capital cost plots, we gain an understanding of where constraints become important and can intelligently set production rates. Both methods will be examined latter in this section. In summary, this step identified two flow sheet structures we want to examine further in our retrofit study. These flow sheets have economic potentials of $5000 000 per year and $4800000 million per year respectively.
Type t h e h . i r . d Option and llERRH mTROrIT P R..ctor
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Figure 5. Input forms for the toluene disproportionation. The example forms are for the reactor and distillation column two.
Step 2. Entering an Existing Flow Sheet Structure. In this step, we make the choices corresponding to the existing flow sheet structure. Since the best new plant design structure is identical with the existing structure except for the distillation sequence, we need only indicate the correct distillation sequence. Step 3. Entering Existing Equipment Sizes. The existing plant sizes for the HDA plant are summarized in Table IX. These sizes were determined by a design calculation a t the published design variable values. These data are entered in the forms for the toluene disproportionation as shown in Figure 5 for the reactor and distillation column 2. For the distillation column, we input values of 1.0 for the tuning parameters in addition to the sizes, which indicate that we assume the models are correct. When the software generated default sizes for the equipment, it designed a new toluene disproportionation and calculated equipment sizes based on this. We actually have equipment from an HDA plant that has no relationship to the these default sizes. By indicating that the tuning parameters have values of 1.0, we show that the models are correct, and any differences between the actual sizes and model predicted sizes represent a difference in equipment capacity. In Figure 5, we see that the predicted number of trays is 48 and the actual number of trays is
826 Ind. Eng. Chem. Res., Vol. 29, No. 5 , 1990
h - 1
23.
Figure 6. HDA plant retrofitted to produce mixed xylenes. Production rate set to 135 Ib-mol/h of mixed xylenes. The "P" next to the columns indicates new columns in parallel were added. M - X Y L E N E PLOWSHEET Retrofit halysit
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Figure 7. Plot of capital costs versus production rate. This plot is for an HDA plant retrofitted to produce mixed xylenes
56. This indicates to the code that their are eight more trays in the column than required. These extra trays may result in a lower operating reflux ratio. In actual operation, the column might actually be providing higher purity separations then required. For the reactor, we specify a retrofit policy of NEW, since the reaction chemistry requires a new catalyst, and we choose to completely replace the reactor. This will require capital expenditure for any potential retrofit to produce xylenes. Step 4. Retrofitting the Existing Process Flow Sheet. Figure 6 shows the optimized flow sheet for the production of xylene with a production rate of 135 lbmol/h (80 x lo6 lb/year). This production rate was arbitrarily chosen at one-half the HDA plant production rate due to the xylene reaction stoichiometry. At this production rate, we still encounter flooding constraints in two distillation columns. The total capital cost for additional column capacity and a new reactor is $220 000 per year, and an economic potential of $5900 000 per year is achieved. This is significantly better than the $2600000 per year economic potential for the best HDA retrofit. It is instructive to look a t a plot of capital costs versus production rate for the toluene disproportionation (Figure 7) when trying to set the production rate. By looking a t such a plot, we can see graphically where different equipment becomes constrained. The method we used to construct this plot of capital costs versus production rate needs to be mentioned before we interpret Figure 7. T o construct the figure, we do our retrofit calculations a t six equally spaced values of production rate. This curve is not continuous but just plots the six discrete points, with the points connect to form a curve. At each point, if equipment is constrained, the
Figure 8. HDA plant retrofitted to produce mixed xylenes. Production rate set to 85 (top flow sheet) and 110 ib-mol/h (bottom flow sheet) of mixed xylenes. The "N"next to the column indicates a new column was added.
constraint is removed, and the total amount of new capital required is summed up. If this curve was continuous rather than discrete points, we would see sharp discontinuities whenever a equipment constraint becomes active. For the discrete curve, we only see a change in the slope between points where a constraint becomes active. We can continue to reduce the production rate scale to obtain a more accurate value of where a constraint becomes active. In Figure 7 there are four distinct regions shown. From 50 to 90 lb-mol/ h, there are active constraints in the second and third distillation columns. Figure 8 shows, for a production rate of 85 lb-mol/h, the second column is weeping and had to be replaced by a new, smaller column (we chose this retrofit policy), and the third column is flooding and an additional parallel column was added. In the region 90-130 lb-mol/h in Figure 7 , capital costs dropped (there would be a discontinuity not a negative slope section of the curve). The bottom flow sheet in Figure 8 for a production rate of 110 lb-mol/h shows the weeping constraint in column two is not active. Since this constraint is no longer active, no new capital was required for the separation, and the capital cost for the retrofit dropped. At about the same production rate, column one began to flood, and some additional capacity was required. This section of the curve illustrates one method of selecting a production rate. We could choose to operate in the region of minimum additional capital costs. This is especially interesting if we have only a limited amount of capital for this project. This minimum in the capital cost versus production rate curve is purely a function of the way we are handling minimum flow constraints in distillation columns. If we chose another method of removing weeping constraints, such as increasing the column reflux to prevent weeping, we would not observe the minimum. We would, however,
Ind. Eng. Chem. Res., Vol. 29, No. 5, 1990 827
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Figure 10. HDA plant retrofitted to produce mixed xylenes. Production rate set to 135 lb-mol/h. This flow sheet is for an alternative distillation sequence.
1 CQitIl cost- 0.69 ms/m Operating Costa30. .Si!?. Cmnmic Potential. 11.
msm
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Figure 9. HDA plant retrofitted to produce mixed xylenes. Production rate set to 210 (top flow sheet) and 250 lb-mol/h (bottom flow sheet) of mixed xylenes.
be paying more in operating costs. Two more constraints are shown in Figure 7 . In the region of production rate from 130 to 215 lb-mol/h, the first and third columns are constrained along with the feed compressor (not indicated in the figure) as shown in Figure 9. For production rates above 210 lb-mol/h, the recycle compressor becomes constrained (shown in the bottom flow sheet of Figure 9). As each new constraint becomes active (for production rates above 125 lb-mol/h), the slope of the capital cost curve becomes steeper. This is another piece of useful information for setting the production rate. By knowing at what value of production rate a constraint becomes active, we can set our production rate to a value slightly below that. This gives us the best policy for adding incremental capital cost. If we go above the constraint, the amount of capital required to increase the production rate increases more steeply than below the constraint. Of course, the production rate a plant can maintain depends on several factors outside of the retrofit design problem. If you are selling your chemical to an outside client or to another section of your chemical complex, they may only be interested in a fixed amount. If you can sell more, the competitive pressure to decrease the selling price may make the new capital expenditure unwarranted. Plots of capital cost versus production rate only assist in setting the production rate within these business limitations. Step 5. Retrofitting Structural Changes in the Flow Sheet. Another flow sheet of interest, found in step 1,was to use a different distillation sequence. Instead of taking the benzene product out of the top of the second distillation column, we might want to take pure xylenes out of the bottom. The first step in looking at this new sequence is to compare the operating costs between the two sequences. If there are no savings in operating costs, there is no incentive to change sequences. The new sequence shows a savings of $160000 per year in operating
costs, and we will look a t it further. The second step is to place the existing columns into the new sequence. To do this, the program prints out the sizes of the existing columns and the sizes that would be required to perform each separation if new columns were used. The user is then asked to direct which existing column will be used for which separation. Figure 1 show the input forms used to place the columns for the toluene disproportionation. In this case, we simply switch the order of the last two distillation columns since they perform essentially the same task as in the original sequence but in a different order. This will involve some expense for repiping the distillation sequence, but a t this stage of retrofitting, we will neglect this cost. After placing the columns, we can look a t the entire flow sheet and repeat our retrofit optimization. Figure 10 shows the new flow sheet a t the optimal values of the design variables. We have set the production rate to 135 lbmol/h, which is comparable to the existing HDA plant. Comparison of the economic potentials of the two flow sheets a t this production rate (Figure 6 and Figure 10) shows that the alternative sequence offers a $200 000 per year improvement in economic potential over the flow sheet with the existing sequence. Step 6. Retrofitting the Heat-Exchanger Network. The temperature-enthalpy diagram and grand composite curve for the mixed-xylene process operating a t a production rate of 135 lb-mol/h is shown in Figure 11. The energy requirements are substantially higher than the HDA process since the recycle flows are substantially higher. The grand composite curve shows that column four (1,4 on the figure) can be partly integrated directly with the background process, and if column three was pressurized so that the condenser operates about 50 O F higher, it could be fully integrated with column four. Alternatively, if the pressure of column four was decreased, additional integration could be achieved. After considering the energy integration, the new optimal values for the design variables are toluene conversion = 0.45 (close to the equilibrium limit), hydrogen pure composition = 0.60, and pinch AT = 48.0 O F . The purge composition was raised to reduce the recycle flows now that the energy integration costs are fully accounted for. The additional energy integration costs were $2550000 per year, with $1610 000 (annualized cost) in new heat-exchanger area required. This results in a final economic potential for the best xylene alternative of $3360 000 per year producing 135 lb-mol/h of mixed xylenes. Summary for Mixed-Xylene Production. For the toluene disproportionation, we identified two promising
828 Ind. Eng. Chem. Res., Vol. 29, No. 5 , 1990 M-XYLENE
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Figure 11. Mixed-xylene process energy integration. Production rate set to 135 lb-mol/h. The top curve is the temperature-enthalpy diagram (composite curve), and the bottom figure is the grand composite curve with distillation columns.
by reoptimizing the design variables and removing some equipment constraints. If we choose to use our existing plant to make mixed xylenes (and a benzene byproduct), we can realize even greater profits. By replacing the reactor and adding some additional distillation capacity, we can achieve an economic potential of $6100000 per year compared to the best HDA plant operation of $2600000 per year. This would be operating a t a production rate of 135 lb-mol/h of mixed xylenes, which roughly corresponds to the HDA plant production rate of 269 lb-mol/h of benzene. Both plants a t this production rate use roughly the same amount of toluene feed. Another interesting alternative would be to increase the production rate in the resulting toluene disproportionation to about 250 lb-mol/h of xylene. In this case, we could still make about the same amount of benzene as in our original plant, if it is required somewhere for further downstream conversion. We would then have the xylene to sell for increased profit. Figure 9 shows that operation a t this production rate would have an annual economic potential of $11 000 000 for a annualized capital investiment of $690 000.
Conclusions A basic strategy for screening the retrofit profitability of an existing plant is simple to device but very tedious to implement by using existing simulators. However, with a special purpose code designed to implement this strategy, it is a simple matter to examine several retrofit alternatives very rapidly. With this approach, it is possible to restrict the number of alternatives that need to be considered for more detailed study. Acknowledgment
T a b l e X. Results of a n HDA Retrofit T o P r o d u c e (before Energy Integration)" design capital operating design case conv purge costs costs HDAl 0.68 0.53 $290000 $38000000 HDA2 0.96 0.15 $250000 $37000000 XYLl 0.45 0.49 $220000 $16100000 XYL2 0.45 0.47 $210000 $16000000
Xylenes
economic potential $1500000 $2600000 $5900000 $6100000
"HDA1: HDA plant with diphenyl byproduct. HDAZ: HDA plant with diphenyl recycled to extinction. XYLl: existing HDA plant flow sheet (135 lb-mol/h of mixed xylenes). XYL2: new distillation sequence (135 lb-mol/h of mixed xylenes).
flow sheets and examined them with the retrofit procedure. The economics for these two flow sheets are summarized in Table X. We also looked a t the problem of how to set the production rate. We found where the minimum capital expenditure for the toluene disproportionation was and were able to identify where different equipment constraints become important. By knowing when equipment constraints become active, we can set our production rate a t a slightly smaller production rate to avoid capital expenditure on a particular unit. Recommendations for the HDA Plant. Table X summarized our results for retrofitting the HDA plant to make mixed xylenes. We can profitably continue to make benzene by the dealkylation route if we choose (see Nelson (1989)). The best retrofit alternative would be to retrofit to a new flow sheet structure and recycle the reversible byproduct (diphenyl) to extinction. We can do this for a limited capital expenditure and increase our profits by $4200000 per year. Even by staying with our existing flow sheet structure, we can improve our profits by $3200 000
This work was supported by the Department of Energy under Grants DE-AC02-81ER10938 and DE-FGO287ER13676. We also acknowledge the sponsors of the Process Design and Control Center at the University of Massachusetts for their support.
Literature Cited Andrecovich, M. J.; Westerberg, A. W. A Simple Synthesis Method Based on Utility Bounding for Heat-Integrated Distillation Sequences. AIChE J. 1985, 31, 363. Chemical Marketing Reporter. Toluene Squeeze to Tighten. Chem. Market. Rep. 1989,235 (Jan 30), 2. Douglas, J . M. A Hierarchical Decision Procedure for Process Synthesis. AIChE J . 1985, 32, 353. Douglas, J. M. Conceptual Design of Chemical Processes; McGrawHill: New York, 1988. Douglas, J . M.; Woodcock, D. C. Cost Diagrams and the Quick Screening of Process Alternatives. Ind. Eng. Chem. Process. Des. Deu. 1985, 24, 970. Fisher, W. R.; Doherty, M. F.; Douglas, J. M. Screening of Process Retrofit Alternatives. Znd. Eng. Chem. Res. 1987, 26, 2195. Glinos, K. A Global Approach to the Preliminary Design and Synthesis of Distillation Trains. Ph.D Dissertation, Department of Chemical Engineering, University of Massachusetts, Amherst, 1984. Grossmann, I. E.; Westerberg, A. W.; Biegler, L. T. Retrofit Design of Processes. In Computer Aided Process Operations; Reklaitis, G. V., Spriggs, H. D., Eds.; CACHE Elsevier: New York, 1987. Hindmarsh, E.; Townsend, D. W. Heat Integration of Distillation Systems into Total Flowsheets-A Complete Approach. 1984 Annual AIChE Meeting, San Francisco, 1984. Kirkwood, R. K. PIP-Process Invention Procedure, A Prototype Expert System for Synthesizing Chemical Process Flowsheets. Ph.D. Dissertation, Department of Chemical Engineering, University of Massachusetts, Amherst, 1987. Kirkwood, R. K.; Locke, M. H.; Douglas, J. M. A prototype Expert System for Synthesizing Chemical Process Flowsheets. Comp. rhem. Eng. 1988, 22, 329.
I n d . Eng. Chem. Res. 1990,29, 829-841
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Tjoe, T. N.; Linnholf, B. Using Pinch Technology for Process Retrofits. Chem. Eng. 1986, April 28,47. Townsend, D. W.; Linnholf, B. Heat and Power Networks in Process. AIChE J . 1983a,29, 742. Townsend, D. W.; Linnholf, B. Part 11: Design Procedure in Equipment Selection and Process Matching. AIChE J . 1983b,29, 748.
Kocis, G. R.; Grossmann, I. E. A Modeling and Decomposition Strategy for the MINLP Optimization of Process Flowsheets. Comp. Chem. Eng. 1989, 13, 797. Nelson, D. A. Preliminary Retrofit Design of Chemical Process Plants. Ph.D. Dissertation, Department of Chemical Engineering, University of Massachusetts, Amherst, 1989. Nikolaides, I. A Retrofit Procedure for Distillation Systems. Ph.D. Dissertation, Department of Chemical Engineering, University of Massachusetts, Amherst, 1987. Simulation Sciences Inc. Process Simulation Program, I / O Graphics Manual; Simulation Sciences Inc.: 1987.
Received f o r review May 22, 1989 Revised manuscript received December 4, 1989 Accepted December 15, 1989
Heat-Transfer Characteristics of a Rotary Disk Processort Pradip S. Mehta* Hoechst Celanese Corporation, Corpus Christi, Texas 78469-9077
Gary S. Donoian Farrel Corporation, Ansonia, Connecticut 06401
Poly(oxymethy1ene) melt was cooled with oil in a novel scraped surface rotary disk heat exchanger. T h e polymer is repeatedly spread as thin films over t h e cooled surfaces of the rotating disks and intermittently collected for surface renewal, as i t moves from one chamber t o the next within t h e processor. T h e total heat removal could be accurately estimated by computing t h e unsteady-state heat transfer during each revolution from a composite flow of the incoming melt and internally recycled films, frozen over t h e moving disks, to the cooling oil. Viscous heat dissipation, especially a t the close clearance areas within the processor, was found t o significantly offset the cooling capacity. Even so, surprisingly, the oil-side transport was found to be rate limiting. In a polymer processing line, it is often necessary to externally heat or cool the polymer to either facilitate a downstream operation or compensate for harsher conditions prevailing during an upstream processing step. For example, polymer melt exiting a polymer reactor may need to be cooled rapidly if the reaction is highly exothermic. Similarly, in many processes, dilute polymer solutions are superheated in heat exchangers before being devolatilized in a flash chamber or a falling strand devolatilizer. Precooling of polymer prior to a final shaping operation, which requires quick solidification,is also not uncommon. In fact, because of high-temperature sensitivity of polymers, it is desirable to have a good temperature control in most postreactor processing steps. In spite of an overwhelming need, inherent properties of polymers make it difficult to effectively transfer heat to or from polymers by conventional means. High viscosities of polymers limit transport via forced convection severely. On the other hand, heat conduction is constrained by the low thermal conductivity of polymers, the attainable temperature gradients, and, in many cases, the available contact area. A variety of heat exchangers have been developed for heating and cooling of viscous fluids, of which scrapedsurface or “closed-clearance” devices have been more popular. Uhl (1970) and Penny and Bell (1967) described some of these equipment which range from agitated vessels to spring-loaded scraper arrangements. Uhl found that addition of scraping augments heat transfer several folds and correlated their performance in terms of penetration and boundary layer theories. In addition to heat-transfer enhancement, continual scraping and renewal of stagnant layers is also desirable for polymers that are sensitive to extended heat histories. On the other end of the spectrum, in-line motionless static mixing elements have been successful in improving transverse temperature distribution ‘This paper was presented at the Annual AIChE Meeting, New York, Nov 1987.
through multiple flow divisions (Collins, 1979). Postreactor polymer processing equipment such as single-screw or multiscrew extruders have been traditionally used to perform a wide range of tasks such as melting of polymers, mixing and compounding of additives, alloying and blending of polymers, polymerization, devolatilization, and pressurization for the final shaping operation. The polymer is conveyed within the equipment through relative movement of some parts which are in close clearance with the stationary parts. For example, flights of a single-screw extruder closely wipe the barrel as the screw rotates relative to the barrel. This feature makes the equipment, in principle, equivalent to a scraped-surface heat exchanger if either surface or both surfaces in relative motion are temperature controlled externally. Indeed, polymer processors today take advantage of this feature by circulating heat-transfer medium through both the barrel and the screw of the extruder in many applications where a close control of the internal polymer temperature is required. Davis (1986) and Hold et al. (1982) have described procedures for calculating heat-transfer coefficients in single-screw extruders. Despite an apparent gain in heat transfer due to surface renewal in the wiped processing equipment, however, Japson (1953) also pointed out the problem of heat generation within the clearances of the wiping elements. Because of the high viscosities of the molten polymers, significant power is dissipated and appears as viscous heat. In processing equipment, therefore, cooling of a viscous polymer is much more difficult than heating it. The rotary disk polymer processor is a novel single-shaft machine (Tadmor et al., 1979), which like extruders has the inherent capability of performing most postreactor processing functions. It differs from screw devices in that it utilizes a twin-drag mechanism for polymer movement, in contrast to drag induction through a single moving surface in screw extruders. It enjoys a distinctly high surface-to-volume ratio even when compared with conventional scraped-surface heat exchangers (Tadmor, 1985). 0 1990 American Chemical Society
