- Email: [email protected]
Hydrogen and energy savings using heuristic allocation of mass exchangers in process synthesis: Technical analysis
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
Available online at www.sciencedirect.com
ScienceDirect journal homepage: www.elsevier.com/locate/he
Hydrogen and energy savings using heuristic allocation of mass exchangers in process synthesis: Technical analysis Carlos Daniel Fischer*, Oscar Alberto Iribarren** Institute for Process Design and Development, INGAR UTN-CONICET, Avellaneda 3657 (3000), Santa Fe, Argentina
article info
abstract
Article history:
For our works, Mass Exchangers (MEs) are membrane equipment that exchange some
Received 15 February 2019
component between two streams in a countercurrent arrangement. In previous works
Received in revised form
these were proposed to use them (with the goal of hydrogen recovery, energy saving and
31 March 2019
waste minimization) at different stages of the hierarchical decision procedure for process
Accepted 1 April 2019
synthesis by Douglas: at an early design stage, when deciding the recycle structure of the
Available online 25 April 2019
process (resulting in considerable changes in the process structure); or at a final design stage, before deciding the mass and energy integrations of process streams (resulting in
Keywords: Hydrogen recovery
minor changes in the process structure). The heuristic allocation of MEs was used independently at both design stages in
Hydrogen exchange
previous works, with the goal of comparing with other design alternatives reported in
Energy saving
the literature that use membranes in the conventional configuration (one feed stream
Process synthesis
and two exit streams; retentate and permeate). In contrast, the present work compares
Process integration
the results of using MEs, when they are used individually at both design stages, and
Heuristic allocation
when they are used jointly at both design stages, in different configurations. For the case study example (the HDA Process), the use of a ME at an early design stage reduces the fresh hydrogen consumed by 3.37%, and the recycle compressor power by 4.09%. The use of a ME at a final design stage reduces the fresh hydrogen consumed by 31.64%, but it does not reduce the recycle compressor power. The joint use of the MEs at both design stages reduces the fresh hydrogen consumed by 14.54%, and the recycle compressor power by 2.91%. The joint use of MEs at both design stages retains the principal benefits of its use at early and final design stages (energy saving and hydrogen recovery, principally), so when both a reduction in the fresh hydrogen consumed and in the recycle compressor power is desired, it is the most appropriate option. © 2019 Hydrogen Energy Publications LLC. Published by Elsevier Ltd. All rights reserved.
Introduction The hierarchical decision procedure of Douglas [1,2] is an approach based principally on the use of heuristics in process synthesis. This procedure can generate processes alternatives that can be assessed implementing them in available computer-based models. There are another
* Corresponding author. ** Corresponding author. E-mail addresses: [email protected] (C.D. Fischer), [email protected] (O.A. Iribarren). https://doi.org/10.1016/j.ijhydene.2019.04.001 0360-3199/© 2019 Hydrogen Energy Publications LLC. Published by Elsevier Ltd. All rights reserved.
13510
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
approaches and novel developments that can be used for process synthesis and process intensification [3e5]. Briefly, in the hierarchical decision procedure of Douglas [1,2], the first decision level deals with the “input-output" structure of the flow sheet. Raw materials, final products, and processing routes are defined at this level, giving the global structure of materials that enter and exit the process. The second decision level refines the process alternative already selected by deciding the recycle structure of the flow sheet and selects the unit operations to perform the separations. This is done following heuristics that recommend criteria for recycling process streams and to select the separation operations. Finally, the heat integration (among the process streams defined in the previous stages) is approached. Over the years, the hierarchical decision procedure added new heuristics, unit operations, and applications. The procedure added the ability to contemplate the conceptual design of vapor-liquid-solid processes [6], the synthesis of separation systems for vapor-liquid and liquid-solid mixtures of both organic and aqueous nature [7], and process synthesis for waste minimization [8]. Other authors included membrane technology and applied the procedure for process design and control [9,10]. In addition, a recent work included membrane distillation crystallization unit [11]. Finally, the original procedure (1985) can be benefited adding the mass integration of process streams (mass integration procedures were at that time not sufficiently developed) [12,13]. Within the mass integration techniques, we may mention the approach of direct recycle, recycle with interception and the mass exchange network synthesis. In previous works [14,15], the heuristic allocation of MEs was proposed to be used as an additional option at different stages of the hierarchical decision procedure [1,2]. The heuristic allocations of MEs were done individually (in either the earlier or later design stages) with the goal of recovering hydrogen in the HDA (Hydrodealkilation) process to transform toluene into benzene [14,15], in the production of cyclohexane from hexane [16,17] and in the production of ammonia [18]. For the HDA process [14,15] different process alternatives were analyzed, using ceramic membranes (ZSM5 [19] and SAPO34 [20]) in a countercurrent configuration modeled as a gaseous ME equipment [14]. These alternatives were compared with the conventional process proposed by Douglas [1,2], and process alternatives proposed by other authors [9,10]) that also used polymeric and ceramic membranes, but in the conventional configuration: one input (the feed), two outputs (permeate and retentate). ZSM5 membranes were allocated in both the early and final design stages, while the SAPO34 membranes were allocated only in the final design stage. For the cyclohexane production process [16,17], different alternatives were analyzed (again in countercurrent configuration) and compared with previously proposed configurations and with using membranes in the traditional configuration. Finally, an ammonia production process [18] was analyzed applying the heuristic at the early stage of design and using ZSM5 membranes. For hydrogen recovery from purge and waste streams, other approach a technologies can be used [21]. In addition to the use of MEs in the hierarchical decision procedure [1,2], we also used MEs for integrate
between different process, resorting to Ion Transport Membranes [22,23]. In the works above cited, the analysis carried out in each of the design stages, and the costs analysis of the alternatives are not strictly comparable among themselves. The goal was to compare with design alternatives proposed by other authors, who use membranes in the conventional configuration. In contrast, the present work compares the use of heuristic allocation of MEs when they are used individually at one design stage, with the case when they are used jointly at both design stages, in different configurations. In this way, the economic performance of each case can be quantified more precisely. Next, there is a section that shows the fundamentals behind the heuristic allocation of MEs. Subsequently, there is a section that presents the general design problem. Following, the use of the heuristic allocations at only one and both design stages is applied to an HDA process. Finally, the last section highlights the conclusions of this work.
Fundamentals behind the heuristic allocation of MEs The heuristic allocation aims to take advantage of the partial pressure gradients that could occur between gaseous streams of a process. These gradients can be used for performing a mass exchange, with the goal of resource recovery, energy saving, and waste minimization. From this point of view, the global goal is the process intensification [5]. The mass exchange proposed between gaseous streams can be implemented in a gas membrane equipment in a countercurrent configuration. We developed a detailed model for this ME in a previous work [14]. Contrary to what occurs in heat or mass exchange networks (where the process streams to be heated and cooled, or increase and decrease their concentration, are well defined), here there is no definition about which process streams have to increase its partial pressure or decrease it. In fact, in the early design stages of the hierarchy, many streams are not even defined. There is a definition of the mole fraction of reactants at the entrance of the reactor (although these are also defined in a certain range, or can be a variable to optimize the process). This complicates the definition of an accurate systematic reference framework to choose the process streams between which mass will be exchanged. There are streams that can be considered as rich or lean streams in the component to be recovered, depending on its partial pressure. In addition, for the design of the process, it is usually not a necessary (or determining) condition that these streams vary their partial pressure. Moreover, not all streams that are considered as rich or lean can be candidates for mass exchange, depending on the goal pursued by the process. A clear example of this is that the input stream of a reactant cannot be considered a rich stream in that reactant. We can also add that not all possible exchanges between rich and lean streams are useful for the objective pursued by the process that is designed. In this way, the selection of streams to carry out the exchange needs to be done by a designer with criteria.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
When a process is designed following Douglas procedure [1,2], after defining the reactor system, the recycling structure and the separation system can be approached. At this time, we have the inputs and outputs of the reactor. At this early design stage, a ME allocation heuristic can be used, which will influence the later decision about the recycle structure of the process: “explore the implementation of a mass exchange (through a gaseous exchange membrane equipment) between the input and output streams of the reactor”. Here, one of the output streams of the reactor can be selected as rich stream in the component to be recovered, and an input to the reactor can be selected as a lean stream in this same component. When carrying out the mass exchange at this level, the entire scheme of separation and subsequent recycling can be smaller and even non-existent (if full integration between the input and the output is achieved). After defining the recycle structure and the separation system, but before defining the heat exchange network (because it generates new process streams and may change some process conditions) another ME allocation heuristic can be used. At this final design stage we can: “explore the implementation of a mass exchange (through a gaseous exchange membrane equipment) between the purge stream (a process output stream) and process input streams”. Here, the purge streams of the process can be selected as rich streams in the component to be recovered, and inputs to the process can be selected as lean streams in this same component. When carrying out the mass exchange at this level of the hierarchy, there are no significant changes in the design of the process. However, there can be changes in the flow rate of the process streams. Based in these mass exchange opportunities, the heuristic allocation of MEs was proposed. It is important to note that these heuristic allocations of MEs do not attempt to compete or replace the techniques of process optimization based on mathematical programming. The heuristic allocations of MEs should be seen as a previous step to mathematical programming. The heuristic generates alternatives processes (feasible) that can be evaluated later by resorting to process simulation (as done in this work), or by mathematical programming. For each process alternative generated, the variables usually studied can be optimized by mathematical programming, for example, stream flow rates, compositions, temperatures, pressures, reaction conditions, separation structure, etc.
Definition of the study problem Consider a simple process like the one shown in Fig. 1. In this process, a reaction takes place between two components (let's call them reactant A, and reactant B), both gaseous, which react to form another component (let's call it product C). This process has two inputs to the reactor, and one output (it is also assumed that they are completely in the gaseous state, although this is not necessary). The R-Input 1 is rich in reactant A and can contain other undesired or inert components. The R-Input 2 is rich in reagent B and can also contain other undesired or inert components. In general, one of the
13511
reactants entering into the reactor (usually the cheapest one, let us assume reactant B) is in excess to accelerate the reaction [2]. Therefore, reactant B will be found in abundance in the output stream of the reactor (R-Output 1), together with product C and small amounts of reactant An unreacted if the reaction progress is incomplete. At the output stream of the reactor (R-Output 1), we can also find some by-product, catalyst or inert component fed together with the reactants. If reactant B (it can be hydrogen or oxygen) added in excess is valuable, it will be separated from the output stream of the reactor (by means of the separation system) and recycled to the reactor inlet. However, prior to recycling, a purge should be carried out (to avoid the accumulation of inert components and/or byproducts in the process recycle) and then the stream can be compressed and recycled. In addition to gaseous recycling, liquid recycles (in dotted line) can be necessary, which may contain a reactant, a co-product, an inert or even part of the product (acting as heat carriers). Nevertheless, the liquids recycling is done by pumps (does not need expensive equipment as are compressors) and does not consume much energy, so it is not considered in this analysis. If this process were synthesized following the Douglas design procedures [1,2], the R-Output 1 stream is defined at an early stage of the design, when the recycle structure is decided (with the separation system not yet defined, but represented by a black box with a target separation duty). The R-Output 1 streams is highlighted in Fig. 1 drawing it as a bold line and labeling it. This is the stream for which a ME should be decided at an early stage of the design. The separation system following the reactor can consist of different operations such as condensation, distillation, phase separation, adsorption (e.g. pressure swing adsorption) or absorption (e.g. absorber, scrubber or stripper). Moreover, since there are components in the gaseous state, it is feasible that the separation system includes conventional permeation operations through membranes. These permeation operations, usually work with a considerable pressure difference (e.g. 30 bar or higher) through the membranes [24,25]. The most common approach is to take advantage of the available pressure of the rich stream (from which it is desired to recover a component), resulting in a low-pressure permeate stream. This permeate component to be recycled should be compressed up to the pressure of the reactor (basically equal to the pressure of the rich stream). In this case, the cost of compression equipment is lower than if the entire rich stream need be compressed since a smaller volume of gas is involved. In other cases, if the available pressure is not enough, all the rich stream should be compressed, giving higher compression costs. In these recovery and recycling systems, the main cost is usually given by the compressor (capital cost and energy consumed) and, to a lesser extent, by the required membrane area. In a later stage of the design following Douglas procedure [1,2], the synthesis of the separation system is done, which adds detail to the structure of the earlier design stage. At this stage the purge stream is completely defined. The Purge stream is highlighted in Fig. 1, drawing it as a bold line and labeling it. This is the stream for which a ME should be decided at a final stage of the design.
13512
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
Fig. 1 e Simple process.
Use of the heuristic allocation of a ME at an early design stage
be relatively low, giving a low partial pressure driving force for the exchange, requiring a large membrane exchange surface.
If we follow the design heuristic: explore the implementation of a mass exchange between the input and output streams of the reactor”; we note that the reactant B (added in excess) is copiously present in the reactor output stream (R-Output 1), and can be exchanged from this stream into the reactor input stream (R-Input 1), which is poor in this component (if it has any). In this case, the reactant B in R-Output 1 is at a high partial pressure, while in R-Input 1, it is at a small (if not null) partial pressure. Thus, there exists a considerable partial pressure gradient between these two streams that can be used to carry out a mass exchange of reactant B without the need for a compressor. This can be done resorting to a countercurrent gaseous mass exchange membrane equipment (as proposed previously [14]). This is shown in Fig. 2. Performing the exchange between the input and the output of the reactor has the advantage that a direct recycle of reactant B is carried out, and if the exchange is important, the entire separation system following the reactor is smaller. However, as the reactant B is found together with all the other effluents of the reactor, its partial pressure may in some cases
Use of the heuristic allocation of a ME at a final design stage The purge stream (Purge) of Fig. 1, is seen to be rich in reactant B. To recover reactant B present in the purge stream, several separation operations can be used. If conventional permeation operations are used, the permeate stream with reactant B is obtained at low pressure. To recycle this stream to the process, it should be compressed up to the inlet pressure of the reactor. Other configurations can be used, such as pressurizing the purge stream before the permeation unit. However, the compressor is still necessary. Conventional permeation operations can be used in other streams of the processes (not only in the purge stream) e.g. in an early stage of design over the gaseous stream exiting the separation system, or even on the stream exiting the reactor. However, using them at that stage is not usually found in the industry. If we follow the design heuristic: “explore the implementation of a mass exchange between purge stream and process inputs streams”; we can use the partial pressure gradient to exchange B between the purge stream (Purge) rich in this
Fig. 2 e Use of the heuristic allocation of a ME at an early design stage.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
13513
Fig. 3 e Use of the heuristic allocation of a ME at a final design stage.
component and the input stream to the reactor (R-Input 1) poor in this component. This configuration is shown in Fig. 3. Performing the exchange between the input stream (RInput 1) and the purge stream (Purge) has the benefit that reactant B already passed through the separation system and is already separated from product C. In this way, its partial pressure is relatively greater than in the R-Output 1 stream, giving a greater partial pressure gradient for the exchange. This exchange can be decided in a final stage of design, or in an existing process (without making considerable changes in its structure). A new trade-off can also be made by varying the purge fraction to send more purge stream to the exchange unit. As a disadvantage, it does not reduce the size of the separation system, so the costs of distillation columns and recycle compressors are not reduced. Again, component B can
only be exchanged up to the point when the partial pressure of it is equal at both sides of the membrane.
Joint use of the heuristic allocation of MEs at both design stages Depending on the process, the MEs can be used in both places. Fig. 4 shows an example of this. When applying the heuristic allocation at an early stage, the first exchange of reactant B is between the R-Output 1 stream (rich in B) and the input stream of the reactor (R-Input 1) (free of B). When applying the heuristic allocation at final stage, the second exchange generated is between the purge stream and the input stream to the reactor after performing the first exchange (R-Input 1.1). In this case, the input stream
Fig. 4 e Joint use of the heuristic allocation of MEs at both design stages.
13514
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
to the reactor (R-Input 1) first receives a part of the reactant B coming from the output stream of the reactor (R-Output 1) and then another part from the purge stream (Purge). Consequently, the partial pressure of reactant B in the R-Input 1.1 stream is no longer zero and will depend on the amount of reactant B received in the first exchange. In addition, since part of reactant B is already recovered in the first exchange, the purge stream (Purge) has less of this reactant, and this reactant is at lower partial pressure. The increase of the partial pressure of reactant B in the R-Input 1.1 stream and the decrease of it in the purge stream results in a lower partial pressure gradient to perform the second exchange. Fig. 5 shows an alternative configuration to that of Fig. 4. In this alternative, both heuristic allocations were also applied together, but the difference is that the lean stream (R-Input 1), first receives it from the exchange with the purge stream (Purge), and later from the output stream of the reactor (ROutput 1). With this alternative, the partial pressure of reactant B in the stream R-Input 1, is zero. Also, the amount of reactant B in the purge stream (Purge) remains smaller (as in the previous configuration), than it would be if the exchange between stream R-Output 1 and R-Input 1.1 stream was not performed. This alternative configuration increases the partial pressure gradient for the exchange at final design stage, but decreases it for the exchange at an early design stage. In this way, there are two possible alternatives for jointly using the heuristic allocation, just by changing the order of the first exchange for the lean stream in reactant B. Another alternative configuration for a simple process is shown in Fig. 6. In this alternative configuration, the lean stream in reactant B (R-Input 1) is split into two streams: R-Input 1-1 and RInput 1e2. Therefore, in both exchanges, the lean streams are free of the reactant B. Thus, the initial partial pressure gradient for the exchange is considerable in both cases. However, the new lean streams (R-Input 1-1 and R-Input 1-2) have smaller flow rates than R-Input 1 (from which they were split) and can more quickly be saturated of reactant B; that is, the logarithmic mean partial pressure gradient may be
smaller. In addition, it is necessary to decide in what proportion R-Input 1 is split towards each exchange. The previous description deals with a simple process with two reactants and one product, in which it is only wanted to recover one reactant using either individually or jointly the heuristic allocations. However, alternative designs are not limited to this. In a process in which it is desired to recover more than one component/reactant, both heuristics allocations could be used together (or not) for each component/ reactant, having the options described above for each of the components/reactants. In addition, if there is more than one input lean stream, then it could be chosen with which one to exchange or in what proportion to split each of them for the exchanges. In this way, the possible exchange options are multiplied, making the problem of design be a problem of superstructures of greater complexity, given the nature of the exchanges and how they affect a particular process.
Example: HDA process Brief description of HDA process In this work, the use of the heuristics allocations to an alternative of the classic HDA process [14,15] is technically analyzed. The most extensively studied alternative of the HDA process is that which includes the separation of the produced diphenyl in a secondary reaction (as a co-product). This alternative has a train of three distillation towers to separate it. However, diphenyl does not have a market comparable to that of benzene, so it is an unwanted co-product. In this work, the last separation tower (toluene-diphenyl separation) is eliminated from the conventional process, respecting the same reaction conditions as well as the dimensions of the reactor. The recirculation of diphenyl has the consequence that it decomposes in benzene (in an equilibrium reaction), so that the production of benzene is slightly increased. For this reason, it is necessary to increase the amount of benzene separated from the benzene-toluene separation tower.
Fig. 5 e Alternative joint use of heuristics at both design stages.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
13515
Fig. 6 e Alternative joint use of heuristics at both design stages splitting the lean stream.
Otherwise, it would accumulate in excess in the process recycle, which is undesired. In the HDA processes alternatives, the raw materials (the fresh hydrogen and toluene) streams are mixed with the gaseous and liquid recycle streams, are heated up to the reaction temperature, and then fed to the reactor. The reactions that occur in it are: Toluene þ H2 / Benzene þ CH4 2Benzene 4 Diphenyl þ H2 For this reaction, the reactor is modeled as a plug flow reactor, taken the kinetic rate parameters from Bouton and Luyben [9]. Kinetic rate parameters are given in Eqs. (1) and (2), where R 1 and R 2 represents the reactions rates (expressed in units of kmol/s m3); PH, PT, PB and PD, are the partial pressures of hydrogen, toluene, benzene and diphenyl respectively (given in Pa, pascals); R is the universal gas constant expressed in units of cal/mol K, and T is the temperature (given in Kelvin), while the activation energies are given in cal/mol. 50976 2:4 exp R 1 ¼ PT P0:5 H RT
(1)
50976 50976 R 2 ¼ P2B 0:001 exp PD PH 0:0071 exp RT RT
and produced in the reactor), benzene, toluene, and diphenyl. The largest amount of hydrogen and methane are separated from the aromatics using a partial condenser, and flash separation at a temperature of 48.9 C and pressure of 33 atm. The separated hydrogen and methane are recycled, after purging a part of this stream (around 12%) to avoid the accumulation of methane in the process. This gaseous recycle stream also carries some benzene, toluene and small amounts of diphenyl. On the other hand, a part of the separated liquid is used to quench hot gases coming from the reactor, and thus avoiding hydrocracking. The remaining liquid goes to the distillation train. The first distillation column (a stabilizer) separates the hydrogen and methane that could not be separated in the flash separator. The second distillation column separates the produced benzene from toluene and diphenyl. Finally, toluene and diphenyl are recycled to the reactor. For this alternative (with these small changes), the benzene production goes from around 125 kmol/h to 128.74 kmol/h. Fig. 7 shows a flow sheet of this HDA process alternative (also very studied) [1,2], previous to the heat exchanger network synthesis. To implement the steady-state process alternatives, Aspen Plus V8.6 was used (with the Peng-Robinson equation of state as a physical property model). All physical properties data are taken automatically from Aspen property database.
Use of the heuristics allocations (2)
The homogeneous reaction takes place in the temperature range of 621 Ce667 C (below it the reaction rate is very low and above it triggers hydrocracking reactions) and at a pressure circa 36 atm. An excess of hydrogen of at least 5:1 is necessary to prevent the burning (or charring) in the reactor at this high temperature. The reactor gaseous effluent should be quickly quenched at a temperature of 621 C to prevent burning in the following heat exchanger. The output stream of the reactor contains hydrogen, methane (fed with hydrogen
The MEs are used with the goal of exchanging hydrogen between different streams of the process. Streams will be called rich if they have an appreciable hydrogen molar fraction, and lean streams, those in which the hydrogen molar fraction is low or negligible. Although the driving force it is the partial pressure gradient, as both rich and lean streams have a similar order of total pressure, the molar fraction provides a good reference. To model the hydrogen exchange, a custom mass exchange equipment with membranes in countercurrent configuration was
13516
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
Fig. 7 e Conventional HDA without diphenyl separation.
modeled with Aspen Custom Modeler V7.2 [14] and selected SAPO34 membranes [20]. For a more detailed description of this custom model, and for the membrane performance, refer to Refs. [14,19,20].
For the use of the heuristic allocation of a ME at an early design stage, the output stream of the reactor is taken as the rich stream, after being quenched and cooled only up to 325 C (and at a pressure of 34 atm), so that, in this way, all
Fig. 8 e HDA flow sheet with the use of the heuristic allocation of a ME at an early design stage.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
the components are in gaseous state. After the hydrogen exchange, the stream can be further cooled to the flash separation temperature (48.9 C). As a lean stream, the input stream of toluene to the reactor is taken, heated only up to 325 C. This stream is the summation of the fresh toluene stream, plus the recycled toluene with a small amount of benzene and diphenyl. This ME will be identified as “ME 1". Fig. 8 shows this alternative. For the use of the heuristic allocation of a ME at a final design stage, the purge stream of the process heated up to a temperature of 325 C is taken as the rich stream; and as a poor stream, the same one used for the first heuristic. This ME will be identified as “ME 2". Fig. 9 shows the flow sheet of this alternative. For the joint use of MEs, two configurations are studied. In a configuration (configuration A), the toluene input stream to the reactor (the lean stream), first exchanges hydrogen with the reactor output stream (ME 1), and later with the purge stream (ME 2). In the other configuration (configuration B), the order of the exchanges is inverted. That is to say, that the lean stream, first exchanges (receives hydrogen) with the purge stream, and later with the reactor output stream. These configurations are depicted in Figs. 10 and 11.
Technical analysis After used the heuristics allocations of MEs, finding a new global optimum can require modifying many process parameters, including temperatures, pressures, molar fractions, the extent of reactions, etc. However, as the goal of this work is to compare between the alternatives (with and without used of
13517
MEs) it is convenient not to affect the reaction conditions and other process conditions. For the reaction conditions not to change appreciably, it is necessary that the initial conditions of the reaction are similar. Therefore, the quantity of the components that are present in the reactor inlet will be kept as constant as possible. The alternatives considered for to the HDA process have hydrogen, methane, toluene, benzene and diphenyl at the reactor inlet. When recovering hydrogen by the use of the MEs in any of the design stages, the amount of it that enters from the hydrogen feed stream has to be reduced. But the hydrogen feed stream is not pure, it has 3% of methane. In this way, a reduction of this stream brings with it a reduction of methane entering to the reactor. The methane is also produced in the reactor by the main reaction. In turn, as the selectivity of the membranes is not ideal, they exchange also methane, in addition to hydrogen. In this way, the exchange of hydrogen produces variations of methane at the reactor inlet that should be controlled (if you want to maintain the reaction conditions without sensitive changes). Methane acts as a thermal heat carrier and it is traditionally controlled by varying the purge fraction (which is usually a variable subject to optimization). In this work, the purge fraction will be used as the control variable, to keep the amount of methane constant at the reactor input. The gaseous recycle stream, in addition to hydrogen and methane, also has toluene, a little benzene, and very little amount of diphenyl, so that a variation of the purge fraction, produces a variation of the amount recycled to the reactor thereof. In turn, the liquid recycle stream of toluene also has benzene and diphenyl. The variation in the amount of toluene at the reactor inlet can be controlled by varying the
Fig. 9 e HDA flow sheet with the use of the heuristic allocation of a ME at a final design stage.
13518
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
Fig. 10 e HDA flow sheet with the joint use of MEs at both design stages (configuration A).
amount that is supplied by the toluene feed stream. In contrast, benzene and diphenyl are not supplied by feed streams. These last components enter into the reactor only through the recycles (benzene due to incomplete separations
and diphenyl (mainly) because it is an unwanted co-product to be is reprocessed). The benzene can be controlled by varying very slightly the amount of it that is distilled vs. the amount recycled. The variations of diphenyl are due to the
Fig. 11 e HDA flow sheet with the joint use of MEs at both design stages (configuration B).
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
same causes that produce the variations of benzene. Thus, when controlling the amount of benzene recycled this control it is also acting in a similar way on diphenyl (since both are present in the same stream). The use of MEs, introduce changes in the heat exchanger network design. However, in previous works [14,15] the heat exchanger network was analyzed and no important changes were found, so it is not considered in this work.
Comparison between the heuristics allocations at an early and a final design stage The variations of the main variables of the HDA process are explored in the range of 0e1000 m2 of membrane exchange area, for the cases where the heuristics are applied independently (see Figs. 8 and 9). This range was selected because at the maximum exchange area, the amount of hydrogen exchanged tends to reach a maximum value. The amount of hydrogen at the maximum area is 65.58 and 77.91 kmol/h for the ME 1 and ME 2 respectively. On the contrary, the quantity of methane exchanged increases almost linearly with the exchange area and does not tend to reach a maximum in the range analyzed. This characteristic indicates that beyond a certain exchange area, it stops being convenient to continue increasing it (increases the methane exchange, but not the hydrogen exchange). The amount of methane in the maximum area is 7.67 and 12.07 kmol/h for the ME 1 and ME 2 respectively. For the use of the ME 1, the amount of hydrogen fed to the process is reduced by increasing the exchange area by 3.37% (due to the hydrogen exchanged), while it is reduced 31.64% for the ME 2. The variations of toluene fed and benzene distilled are minimal (0.15% the larger variation), and can be explained in terms of the controls on the purge fraction and the amount of benzene distilled (with the goal of keeping constant the reaction conditions). The variation of the purge fraction necessary to avoid the methane accumulation is very slight for the use of the ME 1 (0.31%), while for the use of the ME 2 it is appreciable (7.54%). When the ME is used at an early design stage, the molar flow of the gaseous recycle stream and the hydrogen molar fraction of it are reduced considerably (3.96% and 4.77% respectively). When the ME is used at an early design stage, the molar flow of the recycle stream decreases only 1.02%, while the molar fractions of its components remain constant (since the exchange takes place in the purge stream). The recycle compressor power exhibit a reduction of 4.09% for the exchange at early design stage, while that for a final design stage the reduction is not appreciable. After these appreciations, a question arises: Why does the exchange with the purge stream (ME 2) reduces more the consumption of fresh hydrogen than the exchange with the outlet stream of the reactor (ME 1)? This can be explained mainly for two reasons: First, the amount of hydrogen exchanged with the purge stream is greater, since hydrogen is found at a higher concentration in the purge stream than in the reactor outlet stream, therefore, there is a greater partial pressure for the exchange (the driving force). Second, the
13519
hydrogen exchanged with the purge stream is a hydrogen that was purged from the process, thereof the amount of it that is recycled by the gaseous recycle stream is not directly affected by this exchange. While the hydrogen exchanged with the outlet stream of the reactor reduces the amount of hydrogen in the gaseous recycle stream. It is a change in the path of hydrogen recycling (it is as a direct recycle). It is for this reason that the recycle compressor power is considerably reduced. For the exchange with the purge stream, and when the purge is taken after the recycle compressor, there is no reduction in the recycle compressor power.
Joint use of heuristics allocations at both design stages In this subsection, the variations of the main variables of the HDA process are explored, also in the range of 0e1000 m2 of membrane exchange area, when using MEs at both design stages jointly, for the two configurations proposed, called above configuration A (see Fig. 10) and configuration B (see Fig. 11). The exploration of the variables is performed by varying the membrane area of ME 1, for different membrane areas of ME 2. For the configuration A, hydrogen exchange (in ME 1) increases as the area of ME 1 increases, but is not affected by variations in the area of ME 2. For a certain area of ME 2, the amount of hydrogen exchanged in ME 2 decreases when increasing the area of ME 1, and this decrease is larger at higher membrane areas of ME 2. This is mainly due to the fact that the lean stream receives a higher amount of hydrogen from ME 1 (when increasing its area), and therefore the partial pressure gradient for exchange in ME 2, is lower. Another effect, although of much less weight, is that when exchanging hydrogen in ME 1, there is less hydrogen left for the second exchange, whereby the rich stream for the second exchange (with the purge stream) has a slightly lower partial pressure of hydrogen (and consequently, a lower partial pressure gradient). The total flow rate of hydrogen exchanged (hydrogen exchange in ME 1 plus hydrogen exchange in ME 2) tends asymptotically to a value that is the maximum that the lean stream can receive (where the exchange area would be infinite and the partial pressure gradient would be zero). In the case in when the area of ME 2 is zero, the amount of hydrogen feed can be reduced by increasing the area of ME 1. For larger areas of ME 2, the hydrogen feed increases when the area of ME 1 increases, although it is always smaller than in the case when the areas of both MEs are zero. This behavior is due to the fact that the ME 2 reduces more than ME 1 the quantity of the hydrogen feed, and that when increasing the area of ME 1, the hydrogen exchange in ME 2 is reduced. Finally, the recycle compressor power is only affected by the variation of the area of ME 1. This is consistent with the fact that the exchange with the purge stream does not reduce the recycle compressor power, and in this case, the increase of the area of ME 2 does not produce variations in the quantity of hydrogen exchanged in ME 1. On the other hand, for the configuration B, the amount of hydrogen exchanged in ME 1 decreases considerably when with increasing the area of ME 2. A slight reduction of the
13520
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
hydrogen exchanged (in ME 2) is noted when increasing the area of ME 1. This slight reduction is due to a slight reduction of the hydrogen partial pressure in the purge stream, since part of the hydrogen was exchanged in ME 1. The total flow rate of hydrogen exchanged tends asymptotically to a value that is the maximum amount that the lean stream can receive, but their upper bound is lower than for the configuration A. This establishes that the configuration B has less potential to exchange hydrogen than the configuration A. The amount of hydrogen feed decreases more when the area of ME 2 increases, than when that area of ME 1 increases. Finally, when increasing the area of ME 2, there is a smaller reduction of the recycle compressor power (because of the ME 1 exchanges less hydrogen). From the actual analysis, it can be noted that each ME and each configuration (A and B) has its own characteristics. The most economical alternative will depend on (not exclusively) the particular costs of energy, hydrogen (which is reducing its consumption), the cost of conventional equipment and to some extent on the cost of the membranes. For the HDA process, the cost of the compressor represents a large proportion of the investment cost, almost one-third of the total capital cost for some process alternatives [10]. Therefore, a small reduction of this equipment (and of the consumed energy by this equipment) may be more important (in the total annualized cost) than a reduction in the amount of hydrogen fed to (or consumed by) the process. Table 1 shows a comparison of the effects of used the heuristic allocations of MEs independently and jointly (configurations A and B). It assumes that when the
MEs are used independently, 1000 m2 of membrane area was selected for each ME; while that when the MEs are used jointly (configurations A and B), 500 m2 of membrane area was selected for each ME. It can be noted that the ME is used at an early design stage is which that most reduces the recycle compressor power (and its cost), while the ME at final design stage reduces more the amount of hydrogen fed. On the other hand, configuration A presents a significant reduction in hydrogen consumption and retains much of the reduction in the recycle compressor power. Configuration B has similar characteristics than configuration A, but less pronounced, from which it deduces that configuration A has more potential. Given the characteristics of the ME, it is possible that the economic optimum is in configuration A, with some area distribution between the MEs to be determined by optimization.
Other configurations and analyzes performed The alternatives were also analyzed taking the purge stream before the recycle compressor. The advantage is that the recycle compressor has lower power (446 kWe393.08 kW in the conventional design without the heuristics). These alternatives have as a disadvantage a smaller pressure in the purge stream (33.0 atm instead of 41.25 atm), resulting in a lower partial pressure gradient for hydrogen and a lower amount of hydrogen exchanged thereof, in ME 2. Thus, the amount of hydrogen exchanged when used the ME at a final design stage is 64.05 kmol/h for an area of 1000 m2 (previously it was 77.91 kmol/h). This alternative does not cause any
Table 1 e Comparison between alternatives (purge stream taken after recycle compressor). At an early design stage Exchange area (m2) Hydrogen exchanged (kmol/hr) Methane exchanged (kmol/hr) Hydrogen fed (kmol/hr) -Reduction (%) Recycle compressor power (kW) -Reduction (%) Purge fraction -Increase (%)
At a final design stage
Configuration A
Configuration B
ME 1
ME 2
ME 1-ME 2
ME 1-ME 2
1000 65.78 7.69 223.48 3.37 427.75 4.09 0.1191 0.31
1000 77.91 12.05 158.10 31.64 446.00 0.00 0.1276 7.54
500e500 78.18 9.47 197.70 14.52 433.04 2.91 0.1248 3.68
500e500 69.75 9.60 174.45 24.57 441.92 0.91 0.1246 3.52
Table 2 e Comparison between alternatives (purge stream taken before recycle compressor).
2
Exchange area (m ) Hydrogen exchanged (kmol/hr) Methane exchanged (kmol/hr) Hydrogen fed (kmol/hr) -Reduction Recycle compressor power (kW) -Reduction Purge fraction -Increase
At an early design stage
At a final design stage
Configuration A
Configuration B
ME 1
ME 2
ME 1-ME 2
ME 1-ME 2
1000 65.78 7.67 223.48 3.37 376.83 4.14 0.1191 2.98
1000 64.05 9.19 170.58 26.25 390.10 0.76 0.1254 5.63
500e500 66.20 8.18 209.42 9.45 380.11 3.30 0.1222 2.98
500e500 65.12 8.33 183.75 20.55 386.75 1.61 0.1219 2.72
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
13521
Table 3 e TAC comparison between alternatives (purge stream taken after recycle compressor). Cost of membrane area (U.S.$/m2)
3000 1500 750
Alternative
At an early design stage
At a final design stage
Configuration A
Configuration B
Optimum Exchange area (m2) Minimum TAC (U.S.$) Optimum Exchange area (m2) Minimum TAC (U.S.$) Optimum Exchange area (m2) Minimum TAC (U.S.$)
ME 1
ME 2
ME 1-ME 2
ME 1-ME 2
0 75,680,369 0 75,680,369 100 75,674,827
250 75,534,170 500 75,344,919 900 75,164,777
0e250 75,534,170 0e500 75,344,919 0-9}00 75,164,777
0e250 75,534,170 0e500 75,344,919 0-9}00 75,164,777
Table 4 e TAC comparison between alternatives (purge stream taken before recycle compressor). Cost of membrane area (U.S.$/m2)
3000 1500 750
Alternative
At an early design stage
At a final design stage
Configuration A
Configuration B
Optimum Exchange area (m2) Minimum TAC (U.S.$) Optimum Exchange area (m2) Minimum TAC (U.S.$) Optimum Exchange area (m2) Minimum TAC (U.S.$)
ME 1
ME 2
ME 1-ME 2
ME 1-ME 2
0 75,538,163 0 75,538,163 100 75,533,466
200 75,459,601 450 75,303,442 850 75,142,831
200 75,459,601 450 75,303,442 850 75,142,831
200 75,459,601 450 75,303,442 850 75,142,831
disadvantage for the mass exchange of ME 1 (its performance retained same). Table 2 shows a comparison of these alternatives.
Cost analysis To provide a brief assessment of Total Annual Cost (TAC) of the alternatives, we considered the raw material cost, utilities cost and installed cost of large equipment. We take special attention on the varying cost of raw material, MEs and recycle compressor. To annualize the installed costs, we considered a capital charge factor of 0.351. For raw material, we considered a price of 0.5058 U.S.$./kg for hydrogen and 0.6399 U.S.$./kg for toluene [26]. The cost of energy consumed by the compressors was determined considering an electrical-mechanical efficiency of 0.9 and a cost of 0.18 U.S. $/kW-h. For MEs, we refer to Babita et al. [27], which report a cost of 3000 U.S.$./m2 of membrane area. To consider the uncertainty of the membrane area cost after their industrialization, we consider a cost reduction of one half and one quarter. Tables 3 and 4 show the TAC for the alternatives, when the purge stream is taken after and before of the recycle compressor, respectively. From these tables, it can be seen that the economic benefit of the incorporation of the MEs with SAPO34 membranes are marginal. That is because, although the SAPO34 have good selectivity, also have a low permeance, and the reduction of consumption of raw materials and energy of the recycle compressor, is offset by the cost of the MEs units added. Under the condition analyzed (maintaining constant the process reaction, costs and prices, etc.), the alternative with MEs
allocated at the final design stage has a lower TAC. Moreover, when the purge stream is taken after the compressor, a lower TAC it is reached. Note than the alternatives in configuration A and B, have zero area for ME 1, even at the lower cost of membrane area. This is because, under the conditions analyzed, it is more efficient their use at the final design stage.
Conclusions This work presents a general discussion of the incorporation of the heuristic allocations of mass exchangers (ME) within the hierarchical design procedure of Douglas [1,2]. The additional contribution of this work, with respect to the papers cited [14e18] is a more detailed description of the general problem of the incorporation of the MEs, using them independently and jointly in the early and final stages of the design procedure, including its order of application. The heuristics allocations were applied to an alternative of the HDA process, keeping constant most of the process variables (especially the reaction conditions), to independence the impact (hydrogen recovery and energy saving) of their used. When analyzing results obtained by using the heuristic allocation at an early stage of the design procedure, it is noted that a significant reduction in the power of the recycle compressor (of the order of 4.09%). In addition, a reduction of the amount of hydrogen fed to the process (of the order of 3.37%) is obtained. When applying the heuristic allocation at a final stage of the design procedure (or to an already existing process), the benefit was a
13522
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
significant reduction in the amount of hydrogen fed into the process (of the order of 31.64%), but there is no appreciable decrease in the power of the recycle compressor. Using the heuristics allocations jointly in both design stages, both benefits are obtained, although less pronounced (2.91% for the recycle compressor power and 14.52% for the hydrogen fed to the process). This suggests that, if the goal is to maximize the profits, the joint use of the MEs in configuration A would be the most appropriate. However, it is not always like that. The most profitable alternative will depend on the process conditions, the optimization criteria, and especially, the permeance (and selectivity) of the membranes, and of the costs of raw materials and energy. It is important to note that, the implementation of these MEs does not require any compressor to provide the driving force necessary for the exchange of hydrogen. These results suggest the convenience of studying new applications of ceramic membranes in a countercurrent configuration, both in the recycling structures and separation systems, and in the design of hydrogen recovery systems from the process purge stream.
Acknowledgment The authors thank the financial support of the National Council of Scientific and Technical Research of Argentina, through Project PIP 0352.
references
[1] Douglas JM. A hierarchical decision procedure for process synthesis. Am Inst Chem Eng J 1985;31:353e61. [2] Douglas JM. Conceptual design of chemical processes. New York: McGraw-Hill; 1988. [3] Biegler LT, Grossmann IE, Westerberg AW. Systematic methods for chemical process design. Prentice Hall PTR; 1997. [4] Re-engineering the chemical processing Plant: process intensification edited by Andrzej Stankiewicz and Jacob A. Moulijn. Marcel Dekker: New York. 2004. $165. 529 pp. ISBN 0-8247-4302-4. Org Process Res Dev 2004;vol. 8:299. [5] Tian Y, Demirel SE, Hasan MMF, Pistikopoulos EN. An overview of process systems engineering approaches for process intensification: state of the art. Chem Eng Process Process Intensification 2018;133:160e210. [6] Rajagopal S, Ng KM, Douglas JM. A hierarchical procedure for the conceptual design of solids processes. Comput Chem Eng 1992;16:675e89. [7] Douglas JM. Synthesis of separation system flowsheels. AIChE J 1995;41:2522e36. [8] Douglas JM. Process synthesis for waste minimization. Ind Eng Chem Res 1992;31:238e43. [9] Bouton GR, Luyben WL. Optimum economic design and control of a gas permeation membrane coupled with the
[10]
[11]
[12]
[13] [14]
[15]
[16]
[17]
[18]
[19]
[20]
[21]
[22]
[23]
[24] [25] [26] [27]
hydrodealkylation (HDA) process. Ind Eng Chem Res 2008;47:1221e37. Murthy Konda NVSN, Rangaiah GP, Lim DKH. Optimal process design and effective plantwide control of industrial processes by a simulation-based heuristic approach. Ind Eng Chem Res 2006;45:5955e70. Pantoja CE, Nariyoshi YN, Seckler MM. Membrane distillation crystallization applied to brine desalination: a hierarchical design procedure. Ind Eng Chem Res 2015;54:2776e93. El-Halwagi MM, Manousiouthakis V. Synthesis of mass exchange networks. Am Inst Chem Eng Journal 1989;35:1233e44. El-Halwagi MM. Process integration. New York: Elsevier Science Inc.; 2006. Fischer CD, Iribarren OA. Mass integration as a design heuristic: improvements in the HDA process. Ind Eng Chem Res 2011;50:12664e77. Fischer CD, Iribarren OA. Hydrogen recovery from the purge stream of an HDA process using the concept of mass exchange networks. Int J Hydrogen Energy 2013;38:6381e90. Fischer CD, Iribarren OA. Improvements in the design of the cyclohexane production process implementing mass exchange with a counter current gas permeation membrane. Ind Eng Chem Res 2013;52:1698e704. Fischer CD, Iribarren OA. Hydrogen recovery from the purge stream of a cyclohexane production process using a mass exchange heuristic. Int J Hydrogen Energy 2014;39:20094e104. Fischer CD, Iribarren OA. Improvements in the design of the ammonia synthesis process implementing counter current gas permeation modules. Ind Eng Chem Res 2012;51:16410e8. Welk ME, Nenoff TM, Bonhomme F. Defect-free zeolite thin film membranes for H2 purification and CO2 separation. In: van Steen IMC E, Callanan LH, editors. Studies in surface science and catalysis. Elsevier; 2004. p. 690e4. Li S, Falconer JL, Noble RD. SAPO-34 membranes for CO2/CH4 separations: effect of Si/Al ratio. Microporous Mesoporous Mater 2008;110:310e7. Hamad TA, Agll AA, Hamad YM, Bapat S, Thomas M, Martin KB, et al. Hydrogen recovery, cleaning, compression, storage, dispensing, distribution system and End-Uses on the university campus from combined heat, hydrogen and power system. Int J Hydrogen Energy 2014;39:647e53. Fischer CD, Iribarren OA. Oxygen integration between a gasification process and oxygen production using a mass exchange heuristic. Int J Hydrogen Energy 2016;41:2399e410. Fischer CD, Iribarren OA. Oxygen integration of autothermal reforming of ethanol with oxygen production, through ion transport membranes in countercurrent configuration. Comput Chem Eng 2017;99:245e54. Yampolskii Y, Freeman B. Membrane gas separation. The Atrium: John Wiley & Sons Ltd; 2010. Pabby AK, Rizvi SSH, Sastre AM. Handbook of membrane separations. Boca Raton, FL: CRC Press; 2009. ICIS. Chemical insdustry news & chemical market intelligence. Babita K, Sridhar S, Raghavan KV. Membrane reactors for fuel cell quality hydrogen through WGSR e review of their status, challenges and opportunities. Int J Hydrogen Energy 2011;36:6671e88.
Available online at www.sciencedirect.com
ScienceDirect journal homepage: www.elsevier.com/locate/he
Hydrogen and energy savings using heuristic allocation of mass exchangers in process synthesis: Technical analysis Carlos Daniel Fischer*, Oscar Alberto Iribarren** Institute for Process Design and Development, INGAR UTN-CONICET, Avellaneda 3657 (3000), Santa Fe, Argentina
article info
abstract
Article history:
For our works, Mass Exchangers (MEs) are membrane equipment that exchange some
Received 15 February 2019
component between two streams in a countercurrent arrangement. In previous works
Received in revised form
these were proposed to use them (with the goal of hydrogen recovery, energy saving and
31 March 2019
waste minimization) at different stages of the hierarchical decision procedure for process
Accepted 1 April 2019
synthesis by Douglas: at an early design stage, when deciding the recycle structure of the
Available online 25 April 2019
process (resulting in considerable changes in the process structure); or at a final design stage, before deciding the mass and energy integrations of process streams (resulting in
Keywords: Hydrogen recovery
minor changes in the process structure). The heuristic allocation of MEs was used independently at both design stages in
Hydrogen exchange
previous works, with the goal of comparing with other design alternatives reported in
Energy saving
the literature that use membranes in the conventional configuration (one feed stream
Process synthesis
and two exit streams; retentate and permeate). In contrast, the present work compares
Process integration
the results of using MEs, when they are used individually at both design stages, and
Heuristic allocation
when they are used jointly at both design stages, in different configurations. For the case study example (the HDA Process), the use of a ME at an early design stage reduces the fresh hydrogen consumed by 3.37%, and the recycle compressor power by 4.09%. The use of a ME at a final design stage reduces the fresh hydrogen consumed by 31.64%, but it does not reduce the recycle compressor power. The joint use of the MEs at both design stages reduces the fresh hydrogen consumed by 14.54%, and the recycle compressor power by 2.91%. The joint use of MEs at both design stages retains the principal benefits of its use at early and final design stages (energy saving and hydrogen recovery, principally), so when both a reduction in the fresh hydrogen consumed and in the recycle compressor power is desired, it is the most appropriate option. © 2019 Hydrogen Energy Publications LLC. Published by Elsevier Ltd. All rights reserved.
Introduction The hierarchical decision procedure of Douglas [1,2] is an approach based principally on the use of heuristics in process synthesis. This procedure can generate processes alternatives that can be assessed implementing them in available computer-based models. There are another
* Corresponding author. ** Corresponding author. E-mail addresses: [email protected] (C.D. Fischer), [email protected] (O.A. Iribarren). https://doi.org/10.1016/j.ijhydene.2019.04.001 0360-3199/© 2019 Hydrogen Energy Publications LLC. Published by Elsevier Ltd. All rights reserved.
13510
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
approaches and novel developments that can be used for process synthesis and process intensification [3e5]. Briefly, in the hierarchical decision procedure of Douglas [1,2], the first decision level deals with the “input-output" structure of the flow sheet. Raw materials, final products, and processing routes are defined at this level, giving the global structure of materials that enter and exit the process. The second decision level refines the process alternative already selected by deciding the recycle structure of the flow sheet and selects the unit operations to perform the separations. This is done following heuristics that recommend criteria for recycling process streams and to select the separation operations. Finally, the heat integration (among the process streams defined in the previous stages) is approached. Over the years, the hierarchical decision procedure added new heuristics, unit operations, and applications. The procedure added the ability to contemplate the conceptual design of vapor-liquid-solid processes [6], the synthesis of separation systems for vapor-liquid and liquid-solid mixtures of both organic and aqueous nature [7], and process synthesis for waste minimization [8]. Other authors included membrane technology and applied the procedure for process design and control [9,10]. In addition, a recent work included membrane distillation crystallization unit [11]. Finally, the original procedure (1985) can be benefited adding the mass integration of process streams (mass integration procedures were at that time not sufficiently developed) [12,13]. Within the mass integration techniques, we may mention the approach of direct recycle, recycle with interception and the mass exchange network synthesis. In previous works [14,15], the heuristic allocation of MEs was proposed to be used as an additional option at different stages of the hierarchical decision procedure [1,2]. The heuristic allocations of MEs were done individually (in either the earlier or later design stages) with the goal of recovering hydrogen in the HDA (Hydrodealkilation) process to transform toluene into benzene [14,15], in the production of cyclohexane from hexane [16,17] and in the production of ammonia [18]. For the HDA process [14,15] different process alternatives were analyzed, using ceramic membranes (ZSM5 [19] and SAPO34 [20]) in a countercurrent configuration modeled as a gaseous ME equipment [14]. These alternatives were compared with the conventional process proposed by Douglas [1,2], and process alternatives proposed by other authors [9,10]) that also used polymeric and ceramic membranes, but in the conventional configuration: one input (the feed), two outputs (permeate and retentate). ZSM5 membranes were allocated in both the early and final design stages, while the SAPO34 membranes were allocated only in the final design stage. For the cyclohexane production process [16,17], different alternatives were analyzed (again in countercurrent configuration) and compared with previously proposed configurations and with using membranes in the traditional configuration. Finally, an ammonia production process [18] was analyzed applying the heuristic at the early stage of design and using ZSM5 membranes. For hydrogen recovery from purge and waste streams, other approach a technologies can be used [21]. In addition to the use of MEs in the hierarchical decision procedure [1,2], we also used MEs for integrate
between different process, resorting to Ion Transport Membranes [22,23]. In the works above cited, the analysis carried out in each of the design stages, and the costs analysis of the alternatives are not strictly comparable among themselves. The goal was to compare with design alternatives proposed by other authors, who use membranes in the conventional configuration. In contrast, the present work compares the use of heuristic allocation of MEs when they are used individually at one design stage, with the case when they are used jointly at both design stages, in different configurations. In this way, the economic performance of each case can be quantified more precisely. Next, there is a section that shows the fundamentals behind the heuristic allocation of MEs. Subsequently, there is a section that presents the general design problem. Following, the use of the heuristic allocations at only one and both design stages is applied to an HDA process. Finally, the last section highlights the conclusions of this work.
Fundamentals behind the heuristic allocation of MEs The heuristic allocation aims to take advantage of the partial pressure gradients that could occur between gaseous streams of a process. These gradients can be used for performing a mass exchange, with the goal of resource recovery, energy saving, and waste minimization. From this point of view, the global goal is the process intensification [5]. The mass exchange proposed between gaseous streams can be implemented in a gas membrane equipment in a countercurrent configuration. We developed a detailed model for this ME in a previous work [14]. Contrary to what occurs in heat or mass exchange networks (where the process streams to be heated and cooled, or increase and decrease their concentration, are well defined), here there is no definition about which process streams have to increase its partial pressure or decrease it. In fact, in the early design stages of the hierarchy, many streams are not even defined. There is a definition of the mole fraction of reactants at the entrance of the reactor (although these are also defined in a certain range, or can be a variable to optimize the process). This complicates the definition of an accurate systematic reference framework to choose the process streams between which mass will be exchanged. There are streams that can be considered as rich or lean streams in the component to be recovered, depending on its partial pressure. In addition, for the design of the process, it is usually not a necessary (or determining) condition that these streams vary their partial pressure. Moreover, not all streams that are considered as rich or lean can be candidates for mass exchange, depending on the goal pursued by the process. A clear example of this is that the input stream of a reactant cannot be considered a rich stream in that reactant. We can also add that not all possible exchanges between rich and lean streams are useful for the objective pursued by the process that is designed. In this way, the selection of streams to carry out the exchange needs to be done by a designer with criteria.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
When a process is designed following Douglas procedure [1,2], after defining the reactor system, the recycling structure and the separation system can be approached. At this time, we have the inputs and outputs of the reactor. At this early design stage, a ME allocation heuristic can be used, which will influence the later decision about the recycle structure of the process: “explore the implementation of a mass exchange (through a gaseous exchange membrane equipment) between the input and output streams of the reactor”. Here, one of the output streams of the reactor can be selected as rich stream in the component to be recovered, and an input to the reactor can be selected as a lean stream in this same component. When carrying out the mass exchange at this level, the entire scheme of separation and subsequent recycling can be smaller and even non-existent (if full integration between the input and the output is achieved). After defining the recycle structure and the separation system, but before defining the heat exchange network (because it generates new process streams and may change some process conditions) another ME allocation heuristic can be used. At this final design stage we can: “explore the implementation of a mass exchange (through a gaseous exchange membrane equipment) between the purge stream (a process output stream) and process input streams”. Here, the purge streams of the process can be selected as rich streams in the component to be recovered, and inputs to the process can be selected as lean streams in this same component. When carrying out the mass exchange at this level of the hierarchy, there are no significant changes in the design of the process. However, there can be changes in the flow rate of the process streams. Based in these mass exchange opportunities, the heuristic allocation of MEs was proposed. It is important to note that these heuristic allocations of MEs do not attempt to compete or replace the techniques of process optimization based on mathematical programming. The heuristic allocations of MEs should be seen as a previous step to mathematical programming. The heuristic generates alternatives processes (feasible) that can be evaluated later by resorting to process simulation (as done in this work), or by mathematical programming. For each process alternative generated, the variables usually studied can be optimized by mathematical programming, for example, stream flow rates, compositions, temperatures, pressures, reaction conditions, separation structure, etc.
Definition of the study problem Consider a simple process like the one shown in Fig. 1. In this process, a reaction takes place between two components (let's call them reactant A, and reactant B), both gaseous, which react to form another component (let's call it product C). This process has two inputs to the reactor, and one output (it is also assumed that they are completely in the gaseous state, although this is not necessary). The R-Input 1 is rich in reactant A and can contain other undesired or inert components. The R-Input 2 is rich in reagent B and can also contain other undesired or inert components. In general, one of the
13511
reactants entering into the reactor (usually the cheapest one, let us assume reactant B) is in excess to accelerate the reaction [2]. Therefore, reactant B will be found in abundance in the output stream of the reactor (R-Output 1), together with product C and small amounts of reactant An unreacted if the reaction progress is incomplete. At the output stream of the reactor (R-Output 1), we can also find some by-product, catalyst or inert component fed together with the reactants. If reactant B (it can be hydrogen or oxygen) added in excess is valuable, it will be separated from the output stream of the reactor (by means of the separation system) and recycled to the reactor inlet. However, prior to recycling, a purge should be carried out (to avoid the accumulation of inert components and/or byproducts in the process recycle) and then the stream can be compressed and recycled. In addition to gaseous recycling, liquid recycles (in dotted line) can be necessary, which may contain a reactant, a co-product, an inert or even part of the product (acting as heat carriers). Nevertheless, the liquids recycling is done by pumps (does not need expensive equipment as are compressors) and does not consume much energy, so it is not considered in this analysis. If this process were synthesized following the Douglas design procedures [1,2], the R-Output 1 stream is defined at an early stage of the design, when the recycle structure is decided (with the separation system not yet defined, but represented by a black box with a target separation duty). The R-Output 1 streams is highlighted in Fig. 1 drawing it as a bold line and labeling it. This is the stream for which a ME should be decided at an early stage of the design. The separation system following the reactor can consist of different operations such as condensation, distillation, phase separation, adsorption (e.g. pressure swing adsorption) or absorption (e.g. absorber, scrubber or stripper). Moreover, since there are components in the gaseous state, it is feasible that the separation system includes conventional permeation operations through membranes. These permeation operations, usually work with a considerable pressure difference (e.g. 30 bar or higher) through the membranes [24,25]. The most common approach is to take advantage of the available pressure of the rich stream (from which it is desired to recover a component), resulting in a low-pressure permeate stream. This permeate component to be recycled should be compressed up to the pressure of the reactor (basically equal to the pressure of the rich stream). In this case, the cost of compression equipment is lower than if the entire rich stream need be compressed since a smaller volume of gas is involved. In other cases, if the available pressure is not enough, all the rich stream should be compressed, giving higher compression costs. In these recovery and recycling systems, the main cost is usually given by the compressor (capital cost and energy consumed) and, to a lesser extent, by the required membrane area. In a later stage of the design following Douglas procedure [1,2], the synthesis of the separation system is done, which adds detail to the structure of the earlier design stage. At this stage the purge stream is completely defined. The Purge stream is highlighted in Fig. 1, drawing it as a bold line and labeling it. This is the stream for which a ME should be decided at a final stage of the design.
13512
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
Fig. 1 e Simple process.
Use of the heuristic allocation of a ME at an early design stage
be relatively low, giving a low partial pressure driving force for the exchange, requiring a large membrane exchange surface.
If we follow the design heuristic: explore the implementation of a mass exchange between the input and output streams of the reactor”; we note that the reactant B (added in excess) is copiously present in the reactor output stream (R-Output 1), and can be exchanged from this stream into the reactor input stream (R-Input 1), which is poor in this component (if it has any). In this case, the reactant B in R-Output 1 is at a high partial pressure, while in R-Input 1, it is at a small (if not null) partial pressure. Thus, there exists a considerable partial pressure gradient between these two streams that can be used to carry out a mass exchange of reactant B without the need for a compressor. This can be done resorting to a countercurrent gaseous mass exchange membrane equipment (as proposed previously [14]). This is shown in Fig. 2. Performing the exchange between the input and the output of the reactor has the advantage that a direct recycle of reactant B is carried out, and if the exchange is important, the entire separation system following the reactor is smaller. However, as the reactant B is found together with all the other effluents of the reactor, its partial pressure may in some cases
Use of the heuristic allocation of a ME at a final design stage The purge stream (Purge) of Fig. 1, is seen to be rich in reactant B. To recover reactant B present in the purge stream, several separation operations can be used. If conventional permeation operations are used, the permeate stream with reactant B is obtained at low pressure. To recycle this stream to the process, it should be compressed up to the inlet pressure of the reactor. Other configurations can be used, such as pressurizing the purge stream before the permeation unit. However, the compressor is still necessary. Conventional permeation operations can be used in other streams of the processes (not only in the purge stream) e.g. in an early stage of design over the gaseous stream exiting the separation system, or even on the stream exiting the reactor. However, using them at that stage is not usually found in the industry. If we follow the design heuristic: “explore the implementation of a mass exchange between purge stream and process inputs streams”; we can use the partial pressure gradient to exchange B between the purge stream (Purge) rich in this
Fig. 2 e Use of the heuristic allocation of a ME at an early design stage.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
13513
Fig. 3 e Use of the heuristic allocation of a ME at a final design stage.
component and the input stream to the reactor (R-Input 1) poor in this component. This configuration is shown in Fig. 3. Performing the exchange between the input stream (RInput 1) and the purge stream (Purge) has the benefit that reactant B already passed through the separation system and is already separated from product C. In this way, its partial pressure is relatively greater than in the R-Output 1 stream, giving a greater partial pressure gradient for the exchange. This exchange can be decided in a final stage of design, or in an existing process (without making considerable changes in its structure). A new trade-off can also be made by varying the purge fraction to send more purge stream to the exchange unit. As a disadvantage, it does not reduce the size of the separation system, so the costs of distillation columns and recycle compressors are not reduced. Again, component B can
only be exchanged up to the point when the partial pressure of it is equal at both sides of the membrane.
Joint use of the heuristic allocation of MEs at both design stages Depending on the process, the MEs can be used in both places. Fig. 4 shows an example of this. When applying the heuristic allocation at an early stage, the first exchange of reactant B is between the R-Output 1 stream (rich in B) and the input stream of the reactor (R-Input 1) (free of B). When applying the heuristic allocation at final stage, the second exchange generated is between the purge stream and the input stream to the reactor after performing the first exchange (R-Input 1.1). In this case, the input stream
Fig. 4 e Joint use of the heuristic allocation of MEs at both design stages.
13514
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
to the reactor (R-Input 1) first receives a part of the reactant B coming from the output stream of the reactor (R-Output 1) and then another part from the purge stream (Purge). Consequently, the partial pressure of reactant B in the R-Input 1.1 stream is no longer zero and will depend on the amount of reactant B received in the first exchange. In addition, since part of reactant B is already recovered in the first exchange, the purge stream (Purge) has less of this reactant, and this reactant is at lower partial pressure. The increase of the partial pressure of reactant B in the R-Input 1.1 stream and the decrease of it in the purge stream results in a lower partial pressure gradient to perform the second exchange. Fig. 5 shows an alternative configuration to that of Fig. 4. In this alternative, both heuristic allocations were also applied together, but the difference is that the lean stream (R-Input 1), first receives it from the exchange with the purge stream (Purge), and later from the output stream of the reactor (ROutput 1). With this alternative, the partial pressure of reactant B in the stream R-Input 1, is zero. Also, the amount of reactant B in the purge stream (Purge) remains smaller (as in the previous configuration), than it would be if the exchange between stream R-Output 1 and R-Input 1.1 stream was not performed. This alternative configuration increases the partial pressure gradient for the exchange at final design stage, but decreases it for the exchange at an early design stage. In this way, there are two possible alternatives for jointly using the heuristic allocation, just by changing the order of the first exchange for the lean stream in reactant B. Another alternative configuration for a simple process is shown in Fig. 6. In this alternative configuration, the lean stream in reactant B (R-Input 1) is split into two streams: R-Input 1-1 and RInput 1e2. Therefore, in both exchanges, the lean streams are free of the reactant B. Thus, the initial partial pressure gradient for the exchange is considerable in both cases. However, the new lean streams (R-Input 1-1 and R-Input 1-2) have smaller flow rates than R-Input 1 (from which they were split) and can more quickly be saturated of reactant B; that is, the logarithmic mean partial pressure gradient may be
smaller. In addition, it is necessary to decide in what proportion R-Input 1 is split towards each exchange. The previous description deals with a simple process with two reactants and one product, in which it is only wanted to recover one reactant using either individually or jointly the heuristic allocations. However, alternative designs are not limited to this. In a process in which it is desired to recover more than one component/reactant, both heuristics allocations could be used together (or not) for each component/ reactant, having the options described above for each of the components/reactants. In addition, if there is more than one input lean stream, then it could be chosen with which one to exchange or in what proportion to split each of them for the exchanges. In this way, the possible exchange options are multiplied, making the problem of design be a problem of superstructures of greater complexity, given the nature of the exchanges and how they affect a particular process.
Example: HDA process Brief description of HDA process In this work, the use of the heuristics allocations to an alternative of the classic HDA process [14,15] is technically analyzed. The most extensively studied alternative of the HDA process is that which includes the separation of the produced diphenyl in a secondary reaction (as a co-product). This alternative has a train of three distillation towers to separate it. However, diphenyl does not have a market comparable to that of benzene, so it is an unwanted co-product. In this work, the last separation tower (toluene-diphenyl separation) is eliminated from the conventional process, respecting the same reaction conditions as well as the dimensions of the reactor. The recirculation of diphenyl has the consequence that it decomposes in benzene (in an equilibrium reaction), so that the production of benzene is slightly increased. For this reason, it is necessary to increase the amount of benzene separated from the benzene-toluene separation tower.
Fig. 5 e Alternative joint use of heuristics at both design stages.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
13515
Fig. 6 e Alternative joint use of heuristics at both design stages splitting the lean stream.
Otherwise, it would accumulate in excess in the process recycle, which is undesired. In the HDA processes alternatives, the raw materials (the fresh hydrogen and toluene) streams are mixed with the gaseous and liquid recycle streams, are heated up to the reaction temperature, and then fed to the reactor. The reactions that occur in it are: Toluene þ H2 / Benzene þ CH4 2Benzene 4 Diphenyl þ H2 For this reaction, the reactor is modeled as a plug flow reactor, taken the kinetic rate parameters from Bouton and Luyben [9]. Kinetic rate parameters are given in Eqs. (1) and (2), where R 1 and R 2 represents the reactions rates (expressed in units of kmol/s m3); PH, PT, PB and PD, are the partial pressures of hydrogen, toluene, benzene and diphenyl respectively (given in Pa, pascals); R is the universal gas constant expressed in units of cal/mol K, and T is the temperature (given in Kelvin), while the activation energies are given in cal/mol. 50976 2:4 exp R 1 ¼ PT P0:5 H RT
(1)
50976 50976 R 2 ¼ P2B 0:001 exp PD PH 0:0071 exp RT RT
and produced in the reactor), benzene, toluene, and diphenyl. The largest amount of hydrogen and methane are separated from the aromatics using a partial condenser, and flash separation at a temperature of 48.9 C and pressure of 33 atm. The separated hydrogen and methane are recycled, after purging a part of this stream (around 12%) to avoid the accumulation of methane in the process. This gaseous recycle stream also carries some benzene, toluene and small amounts of diphenyl. On the other hand, a part of the separated liquid is used to quench hot gases coming from the reactor, and thus avoiding hydrocracking. The remaining liquid goes to the distillation train. The first distillation column (a stabilizer) separates the hydrogen and methane that could not be separated in the flash separator. The second distillation column separates the produced benzene from toluene and diphenyl. Finally, toluene and diphenyl are recycled to the reactor. For this alternative (with these small changes), the benzene production goes from around 125 kmol/h to 128.74 kmol/h. Fig. 7 shows a flow sheet of this HDA process alternative (also very studied) [1,2], previous to the heat exchanger network synthesis. To implement the steady-state process alternatives, Aspen Plus V8.6 was used (with the Peng-Robinson equation of state as a physical property model). All physical properties data are taken automatically from Aspen property database.
Use of the heuristics allocations (2)
The homogeneous reaction takes place in the temperature range of 621 Ce667 C (below it the reaction rate is very low and above it triggers hydrocracking reactions) and at a pressure circa 36 atm. An excess of hydrogen of at least 5:1 is necessary to prevent the burning (or charring) in the reactor at this high temperature. The reactor gaseous effluent should be quickly quenched at a temperature of 621 C to prevent burning in the following heat exchanger. The output stream of the reactor contains hydrogen, methane (fed with hydrogen
The MEs are used with the goal of exchanging hydrogen between different streams of the process. Streams will be called rich if they have an appreciable hydrogen molar fraction, and lean streams, those in which the hydrogen molar fraction is low or negligible. Although the driving force it is the partial pressure gradient, as both rich and lean streams have a similar order of total pressure, the molar fraction provides a good reference. To model the hydrogen exchange, a custom mass exchange equipment with membranes in countercurrent configuration was
13516
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
Fig. 7 e Conventional HDA without diphenyl separation.
modeled with Aspen Custom Modeler V7.2 [14] and selected SAPO34 membranes [20]. For a more detailed description of this custom model, and for the membrane performance, refer to Refs. [14,19,20].
For the use of the heuristic allocation of a ME at an early design stage, the output stream of the reactor is taken as the rich stream, after being quenched and cooled only up to 325 C (and at a pressure of 34 atm), so that, in this way, all
Fig. 8 e HDA flow sheet with the use of the heuristic allocation of a ME at an early design stage.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
the components are in gaseous state. After the hydrogen exchange, the stream can be further cooled to the flash separation temperature (48.9 C). As a lean stream, the input stream of toluene to the reactor is taken, heated only up to 325 C. This stream is the summation of the fresh toluene stream, plus the recycled toluene with a small amount of benzene and diphenyl. This ME will be identified as “ME 1". Fig. 8 shows this alternative. For the use of the heuristic allocation of a ME at a final design stage, the purge stream of the process heated up to a temperature of 325 C is taken as the rich stream; and as a poor stream, the same one used for the first heuristic. This ME will be identified as “ME 2". Fig. 9 shows the flow sheet of this alternative. For the joint use of MEs, two configurations are studied. In a configuration (configuration A), the toluene input stream to the reactor (the lean stream), first exchanges hydrogen with the reactor output stream (ME 1), and later with the purge stream (ME 2). In the other configuration (configuration B), the order of the exchanges is inverted. That is to say, that the lean stream, first exchanges (receives hydrogen) with the purge stream, and later with the reactor output stream. These configurations are depicted in Figs. 10 and 11.
Technical analysis After used the heuristics allocations of MEs, finding a new global optimum can require modifying many process parameters, including temperatures, pressures, molar fractions, the extent of reactions, etc. However, as the goal of this work is to compare between the alternatives (with and without used of
13517
MEs) it is convenient not to affect the reaction conditions and other process conditions. For the reaction conditions not to change appreciably, it is necessary that the initial conditions of the reaction are similar. Therefore, the quantity of the components that are present in the reactor inlet will be kept as constant as possible. The alternatives considered for to the HDA process have hydrogen, methane, toluene, benzene and diphenyl at the reactor inlet. When recovering hydrogen by the use of the MEs in any of the design stages, the amount of it that enters from the hydrogen feed stream has to be reduced. But the hydrogen feed stream is not pure, it has 3% of methane. In this way, a reduction of this stream brings with it a reduction of methane entering to the reactor. The methane is also produced in the reactor by the main reaction. In turn, as the selectivity of the membranes is not ideal, they exchange also methane, in addition to hydrogen. In this way, the exchange of hydrogen produces variations of methane at the reactor inlet that should be controlled (if you want to maintain the reaction conditions without sensitive changes). Methane acts as a thermal heat carrier and it is traditionally controlled by varying the purge fraction (which is usually a variable subject to optimization). In this work, the purge fraction will be used as the control variable, to keep the amount of methane constant at the reactor input. The gaseous recycle stream, in addition to hydrogen and methane, also has toluene, a little benzene, and very little amount of diphenyl, so that a variation of the purge fraction, produces a variation of the amount recycled to the reactor thereof. In turn, the liquid recycle stream of toluene also has benzene and diphenyl. The variation in the amount of toluene at the reactor inlet can be controlled by varying the
Fig. 9 e HDA flow sheet with the use of the heuristic allocation of a ME at a final design stage.
13518
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
Fig. 10 e HDA flow sheet with the joint use of MEs at both design stages (configuration A).
amount that is supplied by the toluene feed stream. In contrast, benzene and diphenyl are not supplied by feed streams. These last components enter into the reactor only through the recycles (benzene due to incomplete separations
and diphenyl (mainly) because it is an unwanted co-product to be is reprocessed). The benzene can be controlled by varying very slightly the amount of it that is distilled vs. the amount recycled. The variations of diphenyl are due to the
Fig. 11 e HDA flow sheet with the joint use of MEs at both design stages (configuration B).
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
same causes that produce the variations of benzene. Thus, when controlling the amount of benzene recycled this control it is also acting in a similar way on diphenyl (since both are present in the same stream). The use of MEs, introduce changes in the heat exchanger network design. However, in previous works [14,15] the heat exchanger network was analyzed and no important changes were found, so it is not considered in this work.
Comparison between the heuristics allocations at an early and a final design stage The variations of the main variables of the HDA process are explored in the range of 0e1000 m2 of membrane exchange area, for the cases where the heuristics are applied independently (see Figs. 8 and 9). This range was selected because at the maximum exchange area, the amount of hydrogen exchanged tends to reach a maximum value. The amount of hydrogen at the maximum area is 65.58 and 77.91 kmol/h for the ME 1 and ME 2 respectively. On the contrary, the quantity of methane exchanged increases almost linearly with the exchange area and does not tend to reach a maximum in the range analyzed. This characteristic indicates that beyond a certain exchange area, it stops being convenient to continue increasing it (increases the methane exchange, but not the hydrogen exchange). The amount of methane in the maximum area is 7.67 and 12.07 kmol/h for the ME 1 and ME 2 respectively. For the use of the ME 1, the amount of hydrogen fed to the process is reduced by increasing the exchange area by 3.37% (due to the hydrogen exchanged), while it is reduced 31.64% for the ME 2. The variations of toluene fed and benzene distilled are minimal (0.15% the larger variation), and can be explained in terms of the controls on the purge fraction and the amount of benzene distilled (with the goal of keeping constant the reaction conditions). The variation of the purge fraction necessary to avoid the methane accumulation is very slight for the use of the ME 1 (0.31%), while for the use of the ME 2 it is appreciable (7.54%). When the ME is used at an early design stage, the molar flow of the gaseous recycle stream and the hydrogen molar fraction of it are reduced considerably (3.96% and 4.77% respectively). When the ME is used at an early design stage, the molar flow of the recycle stream decreases only 1.02%, while the molar fractions of its components remain constant (since the exchange takes place in the purge stream). The recycle compressor power exhibit a reduction of 4.09% for the exchange at early design stage, while that for a final design stage the reduction is not appreciable. After these appreciations, a question arises: Why does the exchange with the purge stream (ME 2) reduces more the consumption of fresh hydrogen than the exchange with the outlet stream of the reactor (ME 1)? This can be explained mainly for two reasons: First, the amount of hydrogen exchanged with the purge stream is greater, since hydrogen is found at a higher concentration in the purge stream than in the reactor outlet stream, therefore, there is a greater partial pressure for the exchange (the driving force). Second, the
13519
hydrogen exchanged with the purge stream is a hydrogen that was purged from the process, thereof the amount of it that is recycled by the gaseous recycle stream is not directly affected by this exchange. While the hydrogen exchanged with the outlet stream of the reactor reduces the amount of hydrogen in the gaseous recycle stream. It is a change in the path of hydrogen recycling (it is as a direct recycle). It is for this reason that the recycle compressor power is considerably reduced. For the exchange with the purge stream, and when the purge is taken after the recycle compressor, there is no reduction in the recycle compressor power.
Joint use of heuristics allocations at both design stages In this subsection, the variations of the main variables of the HDA process are explored, also in the range of 0e1000 m2 of membrane exchange area, when using MEs at both design stages jointly, for the two configurations proposed, called above configuration A (see Fig. 10) and configuration B (see Fig. 11). The exploration of the variables is performed by varying the membrane area of ME 1, for different membrane areas of ME 2. For the configuration A, hydrogen exchange (in ME 1) increases as the area of ME 1 increases, but is not affected by variations in the area of ME 2. For a certain area of ME 2, the amount of hydrogen exchanged in ME 2 decreases when increasing the area of ME 1, and this decrease is larger at higher membrane areas of ME 2. This is mainly due to the fact that the lean stream receives a higher amount of hydrogen from ME 1 (when increasing its area), and therefore the partial pressure gradient for exchange in ME 2, is lower. Another effect, although of much less weight, is that when exchanging hydrogen in ME 1, there is less hydrogen left for the second exchange, whereby the rich stream for the second exchange (with the purge stream) has a slightly lower partial pressure of hydrogen (and consequently, a lower partial pressure gradient). The total flow rate of hydrogen exchanged (hydrogen exchange in ME 1 plus hydrogen exchange in ME 2) tends asymptotically to a value that is the maximum that the lean stream can receive (where the exchange area would be infinite and the partial pressure gradient would be zero). In the case in when the area of ME 2 is zero, the amount of hydrogen feed can be reduced by increasing the area of ME 1. For larger areas of ME 2, the hydrogen feed increases when the area of ME 1 increases, although it is always smaller than in the case when the areas of both MEs are zero. This behavior is due to the fact that the ME 2 reduces more than ME 1 the quantity of the hydrogen feed, and that when increasing the area of ME 1, the hydrogen exchange in ME 2 is reduced. Finally, the recycle compressor power is only affected by the variation of the area of ME 1. This is consistent with the fact that the exchange with the purge stream does not reduce the recycle compressor power, and in this case, the increase of the area of ME 2 does not produce variations in the quantity of hydrogen exchanged in ME 1. On the other hand, for the configuration B, the amount of hydrogen exchanged in ME 1 decreases considerably when with increasing the area of ME 2. A slight reduction of the
13520
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
hydrogen exchanged (in ME 2) is noted when increasing the area of ME 1. This slight reduction is due to a slight reduction of the hydrogen partial pressure in the purge stream, since part of the hydrogen was exchanged in ME 1. The total flow rate of hydrogen exchanged tends asymptotically to a value that is the maximum amount that the lean stream can receive, but their upper bound is lower than for the configuration A. This establishes that the configuration B has less potential to exchange hydrogen than the configuration A. The amount of hydrogen feed decreases more when the area of ME 2 increases, than when that area of ME 1 increases. Finally, when increasing the area of ME 2, there is a smaller reduction of the recycle compressor power (because of the ME 1 exchanges less hydrogen). From the actual analysis, it can be noted that each ME and each configuration (A and B) has its own characteristics. The most economical alternative will depend on (not exclusively) the particular costs of energy, hydrogen (which is reducing its consumption), the cost of conventional equipment and to some extent on the cost of the membranes. For the HDA process, the cost of the compressor represents a large proportion of the investment cost, almost one-third of the total capital cost for some process alternatives [10]. Therefore, a small reduction of this equipment (and of the consumed energy by this equipment) may be more important (in the total annualized cost) than a reduction in the amount of hydrogen fed to (or consumed by) the process. Table 1 shows a comparison of the effects of used the heuristic allocations of MEs independently and jointly (configurations A and B). It assumes that when the
MEs are used independently, 1000 m2 of membrane area was selected for each ME; while that when the MEs are used jointly (configurations A and B), 500 m2 of membrane area was selected for each ME. It can be noted that the ME is used at an early design stage is which that most reduces the recycle compressor power (and its cost), while the ME at final design stage reduces more the amount of hydrogen fed. On the other hand, configuration A presents a significant reduction in hydrogen consumption and retains much of the reduction in the recycle compressor power. Configuration B has similar characteristics than configuration A, but less pronounced, from which it deduces that configuration A has more potential. Given the characteristics of the ME, it is possible that the economic optimum is in configuration A, with some area distribution between the MEs to be determined by optimization.
Other configurations and analyzes performed The alternatives were also analyzed taking the purge stream before the recycle compressor. The advantage is that the recycle compressor has lower power (446 kWe393.08 kW in the conventional design without the heuristics). These alternatives have as a disadvantage a smaller pressure in the purge stream (33.0 atm instead of 41.25 atm), resulting in a lower partial pressure gradient for hydrogen and a lower amount of hydrogen exchanged thereof, in ME 2. Thus, the amount of hydrogen exchanged when used the ME at a final design stage is 64.05 kmol/h for an area of 1000 m2 (previously it was 77.91 kmol/h). This alternative does not cause any
Table 1 e Comparison between alternatives (purge stream taken after recycle compressor). At an early design stage Exchange area (m2) Hydrogen exchanged (kmol/hr) Methane exchanged (kmol/hr) Hydrogen fed (kmol/hr) -Reduction (%) Recycle compressor power (kW) -Reduction (%) Purge fraction -Increase (%)
At a final design stage
Configuration A
Configuration B
ME 1
ME 2
ME 1-ME 2
ME 1-ME 2
1000 65.78 7.69 223.48 3.37 427.75 4.09 0.1191 0.31
1000 77.91 12.05 158.10 31.64 446.00 0.00 0.1276 7.54
500e500 78.18 9.47 197.70 14.52 433.04 2.91 0.1248 3.68
500e500 69.75 9.60 174.45 24.57 441.92 0.91 0.1246 3.52
Table 2 e Comparison between alternatives (purge stream taken before recycle compressor).
2
Exchange area (m ) Hydrogen exchanged (kmol/hr) Methane exchanged (kmol/hr) Hydrogen fed (kmol/hr) -Reduction Recycle compressor power (kW) -Reduction Purge fraction -Increase
At an early design stage
At a final design stage
Configuration A
Configuration B
ME 1
ME 2
ME 1-ME 2
ME 1-ME 2
1000 65.78 7.67 223.48 3.37 376.83 4.14 0.1191 2.98
1000 64.05 9.19 170.58 26.25 390.10 0.76 0.1254 5.63
500e500 66.20 8.18 209.42 9.45 380.11 3.30 0.1222 2.98
500e500 65.12 8.33 183.75 20.55 386.75 1.61 0.1219 2.72
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
13521
Table 3 e TAC comparison between alternatives (purge stream taken after recycle compressor). Cost of membrane area (U.S.$/m2)
3000 1500 750
Alternative
At an early design stage
At a final design stage
Configuration A
Configuration B
Optimum Exchange area (m2) Minimum TAC (U.S.$) Optimum Exchange area (m2) Minimum TAC (U.S.$) Optimum Exchange area (m2) Minimum TAC (U.S.$)
ME 1
ME 2
ME 1-ME 2
ME 1-ME 2
0 75,680,369 0 75,680,369 100 75,674,827
250 75,534,170 500 75,344,919 900 75,164,777
0e250 75,534,170 0e500 75,344,919 0-9}00 75,164,777
0e250 75,534,170 0e500 75,344,919 0-9}00 75,164,777
Table 4 e TAC comparison between alternatives (purge stream taken before recycle compressor). Cost of membrane area (U.S.$/m2)
3000 1500 750
Alternative
At an early design stage
At a final design stage
Configuration A
Configuration B
Optimum Exchange area (m2) Minimum TAC (U.S.$) Optimum Exchange area (m2) Minimum TAC (U.S.$) Optimum Exchange area (m2) Minimum TAC (U.S.$)
ME 1
ME 2
ME 1-ME 2
ME 1-ME 2
0 75,538,163 0 75,538,163 100 75,533,466
200 75,459,601 450 75,303,442 850 75,142,831
200 75,459,601 450 75,303,442 850 75,142,831
200 75,459,601 450 75,303,442 850 75,142,831
disadvantage for the mass exchange of ME 1 (its performance retained same). Table 2 shows a comparison of these alternatives.
Cost analysis To provide a brief assessment of Total Annual Cost (TAC) of the alternatives, we considered the raw material cost, utilities cost and installed cost of large equipment. We take special attention on the varying cost of raw material, MEs and recycle compressor. To annualize the installed costs, we considered a capital charge factor of 0.351. For raw material, we considered a price of 0.5058 U.S.$./kg for hydrogen and 0.6399 U.S.$./kg for toluene [26]. The cost of energy consumed by the compressors was determined considering an electrical-mechanical efficiency of 0.9 and a cost of 0.18 U.S. $/kW-h. For MEs, we refer to Babita et al. [27], which report a cost of 3000 U.S.$./m2 of membrane area. To consider the uncertainty of the membrane area cost after their industrialization, we consider a cost reduction of one half and one quarter. Tables 3 and 4 show the TAC for the alternatives, when the purge stream is taken after and before of the recycle compressor, respectively. From these tables, it can be seen that the economic benefit of the incorporation of the MEs with SAPO34 membranes are marginal. That is because, although the SAPO34 have good selectivity, also have a low permeance, and the reduction of consumption of raw materials and energy of the recycle compressor, is offset by the cost of the MEs units added. Under the condition analyzed (maintaining constant the process reaction, costs and prices, etc.), the alternative with MEs
allocated at the final design stage has a lower TAC. Moreover, when the purge stream is taken after the compressor, a lower TAC it is reached. Note than the alternatives in configuration A and B, have zero area for ME 1, even at the lower cost of membrane area. This is because, under the conditions analyzed, it is more efficient their use at the final design stage.
Conclusions This work presents a general discussion of the incorporation of the heuristic allocations of mass exchangers (ME) within the hierarchical design procedure of Douglas [1,2]. The additional contribution of this work, with respect to the papers cited [14e18] is a more detailed description of the general problem of the incorporation of the MEs, using them independently and jointly in the early and final stages of the design procedure, including its order of application. The heuristics allocations were applied to an alternative of the HDA process, keeping constant most of the process variables (especially the reaction conditions), to independence the impact (hydrogen recovery and energy saving) of their used. When analyzing results obtained by using the heuristic allocation at an early stage of the design procedure, it is noted that a significant reduction in the power of the recycle compressor (of the order of 4.09%). In addition, a reduction of the amount of hydrogen fed to the process (of the order of 3.37%) is obtained. When applying the heuristic allocation at a final stage of the design procedure (or to an already existing process), the benefit was a
13522
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 4 ( 2 0 1 9 ) 1 3 5 0 9 e1 3 5 2 2
significant reduction in the amount of hydrogen fed into the process (of the order of 31.64%), but there is no appreciable decrease in the power of the recycle compressor. Using the heuristics allocations jointly in both design stages, both benefits are obtained, although less pronounced (2.91% for the recycle compressor power and 14.52% for the hydrogen fed to the process). This suggests that, if the goal is to maximize the profits, the joint use of the MEs in configuration A would be the most appropriate. However, it is not always like that. The most profitable alternative will depend on the process conditions, the optimization criteria, and especially, the permeance (and selectivity) of the membranes, and of the costs of raw materials and energy. It is important to note that, the implementation of these MEs does not require any compressor to provide the driving force necessary for the exchange of hydrogen. These results suggest the convenience of studying new applications of ceramic membranes in a countercurrent configuration, both in the recycling structures and separation systems, and in the design of hydrogen recovery systems from the process purge stream.
Acknowledgment The authors thank the financial support of the National Council of Scientific and Technical Research of Argentina, through Project PIP 0352.
references
[1] Douglas JM. A hierarchical decision procedure for process synthesis. Am Inst Chem Eng J 1985;31:353e61. [2] Douglas JM. Conceptual design of chemical processes. New York: McGraw-Hill; 1988. [3] Biegler LT, Grossmann IE, Westerberg AW. Systematic methods for chemical process design. Prentice Hall PTR; 1997. [4] Re-engineering the chemical processing Plant: process intensification edited by Andrzej Stankiewicz and Jacob A. Moulijn. Marcel Dekker: New York. 2004. $165. 529 pp. ISBN 0-8247-4302-4. Org Process Res Dev 2004;vol. 8:299. [5] Tian Y, Demirel SE, Hasan MMF, Pistikopoulos EN. An overview of process systems engineering approaches for process intensification: state of the art. Chem Eng Process Process Intensification 2018;133:160e210. [6] Rajagopal S, Ng KM, Douglas JM. A hierarchical procedure for the conceptual design of solids processes. Comput Chem Eng 1992;16:675e89. [7] Douglas JM. Synthesis of separation system flowsheels. AIChE J 1995;41:2522e36. [8] Douglas JM. Process synthesis for waste minimization. Ind Eng Chem Res 1992;31:238e43. [9] Bouton GR, Luyben WL. Optimum economic design and control of a gas permeation membrane coupled with the
[10]
[11]
[12]
[13] [14]
[15]
[16]
[17]
[18]
[19]
[20]
[21]
[22]
[23]
[24] [25] [26] [27]
hydrodealkylation (HDA) process. Ind Eng Chem Res 2008;47:1221e37. Murthy Konda NVSN, Rangaiah GP, Lim DKH. Optimal process design and effective plantwide control of industrial processes by a simulation-based heuristic approach. Ind Eng Chem Res 2006;45:5955e70. Pantoja CE, Nariyoshi YN, Seckler MM. Membrane distillation crystallization applied to brine desalination: a hierarchical design procedure. Ind Eng Chem Res 2015;54:2776e93. El-Halwagi MM, Manousiouthakis V. Synthesis of mass exchange networks. Am Inst Chem Eng Journal 1989;35:1233e44. El-Halwagi MM. Process integration. New York: Elsevier Science Inc.; 2006. Fischer CD, Iribarren OA. Mass integration as a design heuristic: improvements in the HDA process. Ind Eng Chem Res 2011;50:12664e77. Fischer CD, Iribarren OA. Hydrogen recovery from the purge stream of an HDA process using the concept of mass exchange networks. Int J Hydrogen Energy 2013;38:6381e90. Fischer CD, Iribarren OA. Improvements in the design of the cyclohexane production process implementing mass exchange with a counter current gas permeation membrane. Ind Eng Chem Res 2013;52:1698e704. Fischer CD, Iribarren OA. Hydrogen recovery from the purge stream of a cyclohexane production process using a mass exchange heuristic. Int J Hydrogen Energy 2014;39:20094e104. Fischer CD, Iribarren OA. Improvements in the design of the ammonia synthesis process implementing counter current gas permeation modules. Ind Eng Chem Res 2012;51:16410e8. Welk ME, Nenoff TM, Bonhomme F. Defect-free zeolite thin film membranes for H2 purification and CO2 separation. In: van Steen IMC E, Callanan LH, editors. Studies in surface science and catalysis. Elsevier; 2004. p. 690e4. Li S, Falconer JL, Noble RD. SAPO-34 membranes for CO2/CH4 separations: effect of Si/Al ratio. Microporous Mesoporous Mater 2008;110:310e7. Hamad TA, Agll AA, Hamad YM, Bapat S, Thomas M, Martin KB, et al. Hydrogen recovery, cleaning, compression, storage, dispensing, distribution system and End-Uses on the university campus from combined heat, hydrogen and power system. Int J Hydrogen Energy 2014;39:647e53. Fischer CD, Iribarren OA. Oxygen integration between a gasification process and oxygen production using a mass exchange heuristic. Int J Hydrogen Energy 2016;41:2399e410. Fischer CD, Iribarren OA. Oxygen integration of autothermal reforming of ethanol with oxygen production, through ion transport membranes in countercurrent configuration. Comput Chem Eng 2017;99:245e54. Yampolskii Y, Freeman B. Membrane gas separation. The Atrium: John Wiley & Sons Ltd; 2010. Pabby AK, Rizvi SSH, Sastre AM. Handbook of membrane separations. Boca Raton, FL: CRC Press; 2009. ICIS. Chemical insdustry news & chemical market intelligence. Babita K, Sridhar S, Raghavan KV. Membrane reactors for fuel cell quality hydrogen through WGSR e review of their status, challenges and opportunities. Int J Hydrogen Energy 2011;36:6671e88.