- Email: [email protected]
Application of a new startup procedure using distributed heating along distillation column
Chemical Engineering and Processing 48 (2009) 1487–1494
Contents lists available at ScienceDirect
Chemical Engineering and Processing: Process Intensification journal homepage: www.elsevier.com/locate/cep
Application of a new startup procedure using distributed heating along distillation column Leandro O. Werle, Cintia Marangoni, Fernanda R. Steinmacher, Pedro H.H. de Araújo, Ricardo A.F. Machado ∗ , Claudia Sayer Federal University of Santa Catarina, Chemical Engineering Department, Mail Box 476, Florianópolis, SC 88040-970, Brazil
a r t i c l e
i n f o
Article history: Received 31 July 2009 Received in revised form 13 October 2009 Accepted 14 October 2009 Available online 23 October 2009 Keywords: Distillation column New startup procedure Distributed heating
a b s t r a c t This study proposes the use of distributed heating applied to the sieve trays along distillation columns, aiming to evaluate a new experimental configuration and procedure applied to the startup. Through this new configuration it is expected that the column heating time and energy will be reduced. Due to the characteristics of the process, such as inertia and non-linearity, almost all the process variables change rapidly, which increases the time required to reach the steady state. These factors make the objective of this study even more challenging. Most of the previously published studies on startup involve simulations, whereas the main contribution of this paper is that it provides experimental results. The experiments were carried out on a pilot distillation column with 13 sieve trays processing a binary mixture composed of water and ethanol, with a flow rate of 300 L h−1 in the fourth tray. It was verified that the use of a heating system which is able to act in a distributed way allows a reduction in the time (40%) and energy (33%) required for the startup procedure, compared to a conventional approach. © 2009 Elsevier B.V. All rights reserved.
1. Introduction The startup of a distillation column represents one of the most complicated dynamic operations in the chemical industry. It involves the control of complex heat and mass transfer operations, and a wide range of operating conditions has been described [1]. The startup procedure of a distillation column is time and energy consuming and the process is unproductive during this period. Thus, it is desirable to reduce the transient period, which can be very long, and consequently the startup time. The result is the optimization of the unit and its energy costs. It is well known that depending on the productive and operational features, distillation column startup may show a long transition period before the steady state is reached, that is, from hours to days. Also, control of a startup process is very challenging. Almost all process variables change quickly; control variables like heating power, reflux ratio or feed flow have to be adjusted at least once, but mostly several times, before reaching the correct steady state [2].
∗ Corresponding author at: Departamento de Engenharia Química e Engenharia de Alimentos, Universidade Federal de Santa Catarina, Campus Universitário, Trindade, CP 476, CEP 88040-970, Florianópolis, Santa Catarina, Brazil. Tel.: +55 48 3721 9554; fax: +55 48 3721 9554. E-mail address: [email protected] (R.A.F. Machado). 0255-2701/$ – see front matter © 2009 Elsevier B.V. All rights reserved. doi:10.1016/j.cep.2009.10.010
The startup of a continuous distillation column presents a series of operational problems, as described by Kister [3]. Other common problems, which have been observed by Ruiz et al. [4], are related to the instrumentation, difficulties with the process fluids and, more frequently, mechanical and hydraulic problems. These latter two are related to the complete filling of the liquid in the downcomer and the vapor in the perforations of the trays, the pressure increase with the accumulation of liquid in the tray (holdup), and others. There are no general rules regarding the startup strategies for the various kinds of distillation columns, but one widely accepted description is that of Ruiz et al. [4]. These authors divided the startup into three distinct phases: a discontinuous, a semicontinuous and a continuous phase. The first occurs when the heating is in progress, the plates are weeping and instruments are started. The second phase has a strong non-linear behavior since during this phase the hydraulic variables reach their steady state and the composition tends toward the desired value. Finally, the third phase occurs when the column can be considered to be in steady state and the quality control system is turned to automatic. Based on this description, some researchers have discussed the startup problem proposing different strategies aiming to reduce the transient period. Little previous work has been carried out on the startup of distillation processes, due to the difficulties involved in experimental investigations. Efforts are concentrated on theoretical and modeling proposals, particularly for reactive and heat-integrated processes. Furthermore, most studies are
1488
L.O. Werle et al. / Chemical Engineering and Processing 48 (2009) 1487–1494
performed manually or show the efficiency of the application of an advanced control technique. Reepmyer et al. [5] describes four strategies for conventional distillation column startup: (1) conventional: set all control variables to steady state values and wait; (2) total reflux: column is run in loop operation, no distillate removal; (3) total distillate removal: exact opposite to previous strategy, column is run without reflux; and (4) time optimized: developed for heat-integrated columns. Also, for conventional distillation Flender [6] reported that savings up to 80% in the startup time can be achieved if a column is started up utilizing the total distillate removal strategy. The results for startup times obtained by Reepmeyer et al. [5] testing total reflux and total distillate removal for reactive distillation, show that the savings are rather small. Eden [7] developed a procedure to generate startup sequences utilizing process knowledge for heat-integrated columns. This procedure was also used by Löwe and Wozny [8] in which a very similar optimized strategy for the startup of an energy integrated column system was applied and the results validated. Fabro [9] proposed the use of intelligent control techniques such as a neural network, fuzzy systems and genetic algorithms for distillation column startup, and achieved better control performance over the system during the simulations, in comparison to another advanced control and supervision architecture. Scenna and Benz [10] described multiple steady states for different initial conditions as starting points. In this publication, both cases are studied in reactive distillation columns. The column does not start from a cold and empty state, but from a defined state, with the trays filled, warm and with in-phase equilibrium. It is thus clear that different methods and strategies for distillation column startup have been studied in literature. They are focused on saving time by applying advanced control techniques, different operational conditions in reflux, or considering different distillation configurations such as heat integration. In fact, it is necessary to evaluate the operational characteristics of the distillation to achieve a better performance. In this regard, a previous study by Marangoni and Machado [11] describes the use of internal heat sources as a new operational approach to distillation columns. The authors reported that the use of an intermediate temperature control loop together with the dual-temperature classic control can lead to fast dynamics when the process is disturbed. This strategy is based on diabatic distillations. Most of the distillation columns are adiabatic with a reboiler at the bottom, a condenser at the top and adiabatic trays in between. Distillation with heat exchangers on all trays is called diabatic distillation, and is defined as a multistage separation process of a liquid or vapor mixture of two or more components by transferring a given amount of heat in such a way that the supply or the removal of heat is performed at two or more different levels [12]. Diabatic distillation has also been studied by computer simulation in order to determine the optimal distribution of heat to be transferred inside the column [13,14]. This configuration offers the benefits of an improved use of the heat of condensation and evaporation, an increase in the possibilities for energy integration between distillation processes and their surrounding processes, and finally a reduction in the operating costs of distillation installations. In this study, the startup problem of a distillation column is considered experimentally aiming to reduce the time and energy required. Thus, we use a new operational approach based on the concept of diabatic distillation through the heating in the distillation column trays in the stripping section. The condenser was maintained only as a source of heat removal at the top of the column. The procedure adopted for the column startup in the conventional operation was established by Steinmacher [15], based on the
classical procedure, with some modifications, such as starting with the accumulator partially filled with material in the desired composition as proposed by Sorensen and Skogestad [16] and also with bottom product withdrawal right from the beginning, as described by Fieg and Wozny [17]. Therefore, the objective of this study is to evaluate a new experimental configuration and procedure applied to the startup of distillation columns, through heat distribution along the unit, and compare them to a conventional procedure where heat is introduced only in the reboiler. As mentioned above, the startup of a distillation column can be considered in three distinct phases: the heating of all stages, the introducing of a reflux stream in manual mode, and the changing of the control loops to automatic mode. It is also important to emphasize that the main goal of this study is to analyze these three phases of the column startup, which require enormous amounts of energy. 2. Methodology The experiments were carried out in a pilot distillation unit, instrumented with a digital fieldbus communication protocol, processing an ethanol–water mixture (sub-cooled liquid) with a feed flow rate of 300 L h−1 , as described in the next section. 2.1. Description of the experimental unit In order to validate the proposed strategy, a column composed of 13 sieve trays was used. The pilot distillation unit operates continuously, with the feed stream inserted in the 4th tray, a total height of 2.70 m, and built in modules with 0.15 m height and 0.20 m diameter. A schematic diagram of the unit is shown in Fig. 1. The main feature of this distillation unit, which makes it different from others, is the use of heating points distributed along the column, which decentralize the heat supply from the bottom. For this, the unit is instrumented with electrical resistances with a power of 3.5 kW in per tray, according to Fig. 2b, which are started by power variators. Also, each module has a hole for the temperature measurement and for the sample collection. The unit is controlled and operated through a supervisory system, which allows the configuration of the control loops and monitoring of trends, as well as the choice of the selections to work with. The control configuration of the distillation column was formulated based on Nooraii [18], and is illustrated in Fig. 3. The following control loops were defined: (1) bottom level control through the manipulation of the bottom product flow rate; (2) reflux accumulator level control by manipulating the top product flow rate; (3) feed flow rate control as a function of the adjustment of the same stream flow rate; (4) feed temperature control through the manipulation of the fluid flow rate in the heat exchanger of this stage; (5) last tray (distillate) temperature control by means of the manipulation of the reflux flow rate; (6) reboiler temperature control through the vapor flow rate in the heat exchanger of this stage; and (7) temperature control of pre-defined stages of the column through the manipulation of the dissipated power in the tray electrical resistance. The first, second and third loops represent the column mass balance (inventory) control. The fifth and sixth loops comprise the quality control – in this case represented by the temperature. The use of these two loops in combination is referred to herein as conventional control. When these two loops are combined with the seventh loop mentioned above, it is considered herein as the distributed strategy. All control loops are instrumented with fieldbus protocols, along with the acquisition and indication of the bottom and distillate
L.O. Werle et al. / Chemical Engineering and Processing 48 (2009) 1487–1494
1489
Fig. 1. General scheme of the equipments and pilot distillation unit.
stream flows and the pressures at the same stages. The temperatures of all the trays, reboiler, accumulator and feed are monitored by a programmable logic controller. The pressures were monitored in order to ensure the proper functioning of the equipment and the process. 2.1.1. Definition of local heating (stage selection) The proposed modification to distributed heating fulfills the industrial needs, in optimizing the distillation processes and reducing the energy costs incurred by disturbances. It is a simple alteration and it consists only of the local heating of some column trays, through resistances with the heat supply controlled by a power variator. As previously mentioned, the new configuration proposed herein is based on the concept of diabatic distillation, through the heating in the distillation column trays in the stripping section used. Therefore, the great difference of this proposal in relation to other diabatic distillation studies, besides the use of electrical resistances instead of heat exchangers, is the distribution of heat between the reboiler and the trays of the stripping section. In this case, heat exchangers are not used in the rectifying section. The
Fig. 2. Upper view of the modules with details of the tray: (a) conventional and (b) with electrical resistance.
condenser used was maintained as the only source of heat removal at the top of the column. To identify the most sensitive stage for the consequent application of the distributed heating, three different methods were applied [19]. In the first method, the differences between the temperatures of two successive trays were calculated throughout the column and the most sensitive tray was that which presented the greatest difference in relation to its adjacent trays. In the second method, a temperature profile for a given value of the manipu-
Fig. 3. Control configuration of the distillation unit.
1490
L.O. Werle et al. / Chemical Engineering and Processing 48 (2009) 1487–1494
lated variable (in this case, the reflux flow and the reboiler heat) is obtained. The most sensitive tray gives a symmetrical response to positive and negative equal variations. Finally, the third method analyzes the tray with the highest derivative of the temperature in relation to the stage when the process is disturbed. To analyze the first method, the temperature profile for three different conditions of ethanol feed composition (15, 25 and 35%) was observed. In the second and third methods it was necessary to disturb the process and evaluate its behavior. The feed flow used as the standard condition was 400 L h−1 , which was increased to 600 L h−1 and also decreased to 200 L h−1 . It is important to emphasize that the different methods can produce different responses. The definition was based on this analysis together with the characteristics of the plant. 2.2. Experimental procedures of startup The startup procedures for the distributed heating and conventional approaches were almost the same. The only difference was the use of an electrical resistance with a total power of 3.5 kW on an intermediate tray, from the beginning of the experiment. In the conventional approach, the column is operated with heat supply only at the reboiler, with the manual mode of the vapor valve providing 10 kW (8% the steam valve opening). When the proposed new configuration is applied, the heat is supplied at the bottom by the reboiler (with the same previous overall heat supply) and at an intermediate point, by way of the electrical resistance on the 2nd tray, started with 50% of its capacity (1.75 kW). As previously mentioned, the procedure adopted for the column startup in the conventional operation is based on a study by Steinmacher [15]. In this study the startup of the distillation column was considered to have three distinct phases: (1) heating of all stages; (2) introducing of the reflux stream in manual mode; and (3) changing of the control loops to automatic mode. These phases as well as the procedures are illustrated in Fig. 4 and detailed in the next section. 2.2.1. Discontinuous phase In the first phase, at the beginning of the procedure, volumetric compositions of ethanol in the feed and accumulator (partially filled with mixtures in the desired compositions) were 15 and 80%, respectively. During this phase, and throughout the startup, the feed stream is continuously introduced automatically and controlled at 300 L h−1 . To do this experimentally the process is closed, that is, the bottom and top product streams form the feed stream in the main tank (Fig. 1). This is the reason why the accumulator is already partially filled, thus minimizing load composition fluctuations (observed in previous experiments when this equipment was empty). The startup is initiated with the withdrawal of bottom product (automatic bottom level control). The bottom temperature control loop was maintained in manual mode. The liquid mixture from the feed descends throughout the column until it reaches the bottom, where the heat exchanger heats and vaporizes this stream. The vapor mixture ascends through the column, heating it tray-by-tray until it reaches the condenser. Here phase 1 is finalized. 2.2.2. Semi-continuous phase The condensed mixture is received by the accumulator, which is already partially filled (30%) with the material at the desired composition (80%, v/v ethanol) for the reflux stream. At this moment an increase in the level and temperature of the mixture in the accumulator is applied. When the accumulator temperature reaches the desired value (same as in the last stage), the total reflux is introduced until a stable situation is achieved (temperature derivative is zero over time and, for this system, the composition is also considered to be constant). This characterizes the end of phase 2. In
Fig. 4. Steps of the startup process for distillation column.
L.O. Werle et al. / Chemical Engineering and Processing 48 (2009) 1487–1494
1491
this phase, the temperature of the last stage control loop (through the manipulation of the reflux flow rate) is in manual mode, set at 30 L h−1 , and production of the top product has not started. The accumulator level control loop is not automatic. 2.2.3. Continuous phase At this moment, the top product withdrawal is started and the accumulator level bottom and top temperature control loops are changed to automatic mode, starting at phase 3. The steady state is determined when the accumulator temperature and top composition do not vary over time. These variables were used as the evaluation parameters. Volumetric composition measurements corroborate this analysis. The composition was measured with a densimeter for alcohol. 2.3. Energy balances Results are compared based on the energy costs. The energy balance, used to calculate the thermal energy of the reboiler, was obtained considering the steady state at the end of the startup, assuming the following simplifying hypotheses: (1) losses to the environment are negligible; (2) in each theoretical tray liquid and vapor are in equilibrium; (3) pressure is constant along the column. The contributions of thermal energy that enter the column are represented by feed stream, electrical resistance and the reboiler. The thermal energy leaving the column is represented by the distillate stream, condenser and bottom product stream. The heat is calculated based on the enthalpy, using classical relations. The values of the thermodynamic properties required for the calculation of mass and enthalpy balances were obtained through simulations with the software Hyprotech HYSYS® , version 7.0. The thermodynamic model for the liquid phase was UNIQUAC (UNIversal QUAsiChemical) and the vapor phase was considered ideal, mainly because of the low operation pressure of the column. 3. Results and discussions The first phase of this study was to determine the tray where the distributed heating could be applied. As cited before, this was achieved through a sensitivity analysis employing three different methods. The results obtained with the first method (successive trays) using three ethanol feed composition conditions demonstrated the possibility of using trays 1, 2, 3, 5 and 7. As the fifth and seventh trays are located in the rectifying section, they were discarded. It was assumed, following diabatic studies upon which this proposal was based, that in this section it is better to remove heat than to supply it. In addition, it was defined that only one tray would be used to test the new procedure for the startup of the column. To define this stage, since method 1 was not conclusive, the analysis of symmetrical response and maximum derivative (methods 2 and 3) was used. The derivative method again indicated trays 5 and 7, which had been previously discarded, but the symmetrical response method indicated tray 2 as the most appropriate for this study. Fig. 5 shows this analysis, where it can be observed that tray 2 is almost the same distance from steady state when the process is disturbed with positive and negative perturbations in the feed flow. Based on this sensitivity analysis, the distributed action proposed herein was used only in tray 2. Tray 1 was discarded because it is located near the reboiler and 3 was discarded to avoid more than one simultaneous effect. In order to evaluate the performance and verify the effect of the heat distribution along the column during the startup of the process, experiments were carried out with the procedures described in Section 2.2. For both experiments the same criterion was used to
Fig. 5. Results of sensitivity analysis using symmetrical response method: () negative disturbance, () steady state, and () positive disturbance.
Fig. 6. Feed temperature, conventional () and with distributed heating applied on tray 2 ().
evaluate the establishment of the steady state, based on the stabilization of the column and on the derivative of the temperature curves. The volumetric composition measurements corroborate this evaluation. The behavior of the feed temperature in both experiments can be verified in Fig. 6, which shows the feed temperature increase during the column startup. This increase occurs because part of the heated liquid from the bottom is sent to the main tank through the bottom product stream, thus the feed temperature increases gradually until the permanent regime is reached. In industry, the tanks are very large and the influence is smaller. However, there is also the heating time of the equipment itself. A detailed analysis of the results shown in Fig. 6 allows the verification of different slopes for the initial heating phase (linear step), with a higher rate for the configuration with distributed heating, as confirmed by the values of the linear coefficients of the curves, shown in Table 1. This fact indicates that the steady state is reached faster when the proposed distributed configuration is used. With the calculation of the derivative of the feed temperature the effect of the distributed heating can be better observed. These derivatives were obtained from sigmoidal functions, adjusted to Table 1 Coefficients of the linear stage of the heating in the temperatures of the feed stream and of tray 3. Model
Linear regression (conventional operation)
Linear regression (distribution heating operation)
Feed temperature R2 Tray 3 R2
Y = 34.09X + 31.23 0.99 Y = 31.53X + 29.76 0.99
Y = 40.51X + 30.52 0.99 Y = 35.15X + 30.15 0.99
1492
L.O. Werle et al. / Chemical Engineering and Processing 48 (2009) 1487–1494
Fig. 7. Derivatives of the feed temperatures in function of time, conventional () and with distributed heating applied on tray 2 ().
describe the feed temperature in each of the cases. The distributed heating allowed the steady state to be reached faster, as shown in Fig. 7. The temperature profiles of a stripping tray (tray 3) and a rectifying tray (tray 13) are represented in Fig. 8. In these experiments a difference was noted in the tray behavior along the sections of the column during the heating. The stripping trays show a linear heating region, and which is not observed in the rectifying trays. Tray 3 was chosen to illustrate the stripping section as it is the tray which best represents the heat addition effect. This stage is successive to tray 4 which receives the feed stream, as also indicated by the sensitivity analysis. The results shown in Fig. 8 indicate that before the beginning of the reflux, the temperature reached at the top of the column is slightly higher for the distributed configuration than for the conventional one. However, the temperatures after the introduction of reflux were almost the same in both situations. This confirms that the steady state and the reflux flow rate (30 L h−1 ) were the same in both cases. After the linear region (linear phase), shown in Fig. 8, an interval can be observed. This is the time required to heat the tray contents. In this phase, the heating is slow until the bubble point temperature of the mixture is reached, and the tray temperature increases rapidly. The same behavior was observed by Wang [20] and Steinmacher [15]. When using the distributed strategy, the time saving is the same for both the third and thirteenth trays (0.32 h), when compared with the conventional strategy. An analysis of the angular coefficients of the linear regression in the feed and tray 3 temperatures for conventional and distributed modes was carried out. The results are summarized in Table 1. In all of the curves the multiple regression coefficient obtained was
Fig. 8. Temperature profiles of the third tray: conventional (䊉) and distributed heating applied on tray 2 () and of the thirteenth tray, conventional () and distributed heating applied on tray 2 (), during the process startup.
Fig. 9. Column bottom temperature, conventional heating () and with distributed heating applied on tray 2 (), and manipulated variable conventional (䊉) and distributed heating applied on tray 2 ().
higher than 0.99. It is observed that, for both temperatures, angular coefficients are higher when the distributed strategy is used. Therefore, similarly to the results for the initial phase of the feed temperature increase, the distributed heating reduces the time required to reach the stationary tray temperatures. The results shown in Fig. 8 and Table 1 indicate that the distributed heating resulted in a better performance, a faster action and a reduction in the time required to reach the desired temperature at the top of the column. Fig. 9 shows the profile of the bottom temperature before and after the control loop is changed to automatic mode with the two studied configurations. The distributed heating again showed a faster action when compared with the conventional process. Using the intermediate heating, the desired temperature value (98 ◦ C) was reached 0.75 h faster than using the conventional configuration. This is of particular interest as it shows that the action of the resistance increased the heating rate of the bottom temperature. In addition, in Fig. 9 (manipulated variable) it can be observed that after changing the control loop to automatic mode the bottom valve operated with a smaller opening, thus compensating the energy added to tray 2. In order to better evaluate the effect of the distributed heating, the evolution of the derivatives of the bottom temperatures for the two experiments were compared. These derivatives were obtained from sigmoidal functions, adjusted to describe the bottom temperature in each of the cases, as shown in Fig. 10. It was verified that the proposed distributed configuration allowed the steady state to be reached faster. A period of 1.4 h was required for the bottom temperature to become stable with the conventional configuration and only 0.8 h for the distributed configuration. This represents only 57% of the time required to reach the desired bottom temperature. The reduction in the time required to reach the steady state when the distributed heating is applied is confirmed by the measurements of the volumetric ethanol fractions in the bottom column stream and their derivatives as a function of time (dX/dt) to characterize the steady state, as shown in Figs. 11 and 12, respectively.
L.O. Werle et al. / Chemical Engineering and Processing 48 (2009) 1487–1494
1493
Fig. 13. Startup phases: (1) discontinuous; (2) semi-continuous; (3) continues. Fig. 10. Derivatives of the bottom temperatures in function to time, conventional () and with distributed heating applied on tray 2 ().
Fig. 11. Volumetric ethanol fractions in the current of bottom of the column, conventional (), and with distributed heating applied on tray 2 ().
higher internal steam flow rates, since the contact of the feed stream with the heated resistance on plate 2 favors the steam formation inside the column, enhancing the advantage of the distribution of heat. In Fig. 13 the time required for each column startup phase can be verified in detail, as well as how the distributed heating influences the time reduction of each phase. In this figure the effect of the distributed in relation to the conventional heating, mainly in phases 1 and 3, can be clearly observed. In all three stages the reduction of the time, due to the effect of the addition of the intermediate heating point, was significant, when compared to the conventional procedure, being in the order of 43, 21 and 50%, respectively. Thus, a reduction of approximately 40% of the total startup time was reached for a distillation column operating with heat distribution, which suggests considerable gains in the process. We can observe, in Table 3, the advantage of using the new proposed configuration, when evaluated in terms of energy costs, for each of the three startup phases analyzed, during the transient period. In Table 3 Qr is the heat input to the bottom by the reboiler and Qr,e is the heat input to the bottom by the reboiler plus the heat contributed by the electrical resistance in second tray. The thermal energy values of each startup phase (Qphase ) are obtained by multiplying the heat supplied to the column (only the reboiler in the case of conventional operation, and reboiler and electrical resistance in the case of the distributed heating operation) by the duration of each phase. From this analysis, the advantage of the new configuration proposed herein is evident in relation to the conventional procedure, since the total energy required is 33% lower in the configuration with distributed heating. Finally, the thermal energies of the streams, shown in Fig. 14, were calculated after the startup process was concluded and the system reached steady state conditions (end of phase 3). It can be observed from the results shown in Fig. 14 that the amount of
Fig. 12. Derivatives of the volumetric fraction of ethanol of bottom in function time, conventional () and with distributed heating applied on tray 2 ().
Fig. 14. Thermal energy of the currents and of the resistance calculated after the system reached the steady state for the conventional and distributed configurations.
It can be observed in Fig. 11, that there is a change in the bottom ethanol fraction which starts at 0.14% and ends at 0.08%. This is due to the introduction of reflux. Through this observation it is possible to analyze the period out of steady state – as illustrated by the derivative of this variable in relation to time. Is important to note that reflux is introduced in the 13th stage – more distant from the bottom. This means that the interaction between the bottom and top is responsible for the longer time required to reach steady state (or to maintain the column out this condition). As illustrated by other results, Fig. 12 shows that the distributed configuration allows that the transition period is shorter than the conventional one (around 0.75 h), corroborating the previous analysis. Table 2 shows the process variables obtained after the steady state was reached in both cases. Results shown in Table 2 indicate that the distributed configuration leads to a greater production of top product when compared with the conventional configuration. This occurs because the distributed heating operation maintains
1494
L.O. Werle et al. / Chemical Engineering and Processing 48 (2009) 1487–1494
Table 2 Parameters of the steady state reached for the two studied experiments. Variable/stream
Conventional operation
Flow rate (L h−1 ) Ethanol fraction (v/v) Temperature (◦ C)
Distributed heating operation
Feed
Top
Bottom
Feed
Top
Bottom
300.00 0.14 76.50
13.20 0.88 80.40
286.80 0.08 98.50
300.00 0.14 77.00
15.86 0.87 81.20
284.14 0.08 98.50
Table 3 Energy cost for each phase of the startup of the column. Variables
Phase 1 Phase 2 Phase 3 a b c d
Conventional operation
Distributed heating operation
Qr (kW)
Time (h)
Qphase (kW)
Qr,e (kW)
Time (h)
Qphase (kW)
10.00a 10.00a 15.00c
1.53 1.05 1.50
15.30 10.50 22.50
11.75b 11.75b 15.25d
0.88 0.84 0.78
10.34 9.87 11.90
Calculated with steam valve, manual mode (opening of 8%). Calculated with steam valve, manual mode (opening of 8%) + electrical resistance (50% of its capacity). Calculated with steam valve, automatic mode (±18% of opening). Calculated with steam valve, automatic mode (±14% of opening) + electrical resistance (50% of its capacity).
energy supplied to the stripping section column is slightly larger in the distributed configuration (15.25 kW, reboiler + electrical resistance), in relation to the 15.00 kW of the conventional configuration (practically the same amount of energy). Furthermore, the small difference among these values may also be due to the experimental measurement error, besides the approaches and hypotheses assumed in the calculation of the energy balance. The results shown in Fig. 9 help to clarify the explanation. After setting the bottom control loop to the automatic mode, the configuration with distributed action (heating at bottom and on tray 2) required a smaller opening of the bottom steam valve (to reach the same steady state). This occurred because energy was already being added inside the column through the electrical resistance. This finding indicates that there was in fact distribution of energy between the bottom and the resistor.
4. Conclusions Comparison between the two operation approaches, showed that the temperatures of the trays of the stripping section increase at a higher rate when the distributed heating with electrical resistance on tray 2 is used, reaching a permanent regime in a much shorter time than in the conventional process. In addition, the distributed heating allowed the steady state of the temperatures of the column feed, top and bottom, to be reached faster than in the conventional process. Composition measurements of bottom and top products, as well as temperature profile comparisons of these streams, allowed it to be verified that the steady state reached in both cases was the same. Thus, the introduction of distributed heating along the column is demonstrated to be a valid option for reducing the startup time. In this study, a reduction of approximately 40% in the startup time was observed in relation to the conventional process, allowing faster dynamics and lower operation costs, as well as a considerable decrease in the production of out-of-specification products. It was observed that the configuration with distributed heating required 33% lower energy for the 3 startup phases compared with the conventional configuration. The next step in this research is to test the control techniques and algorithms for the automatic startup of distillation columns, improving the steady state and allowing it to be reach in as short a time as possible, using this new procedure for the startup process with the heat distribution along the column.
Acknowledgements The authors are grateful for the financial support from the National Agency of the Petroleum (ANP), and the Studies and Projects Fund (FINEP), by means of the Human Resources Program of ANP for the Petroleum and Gas sector – PRH-34-ANP/MCT. References [1] G. Wozny, P. Li, Optimization and experimental verification of startup policies for distillation columns, Comp. Eng. 28 (2004) 253–265. [2] F. Reepmeyer, J.U. Rekle, G. Wozny, Time optimal start-up strategies for reactive distillation columns, Chem. Eng. Sci. 59 (49) (2004) 4339–4347. [3] H.Z. Kister, Distillation Operation, McGraw-Hill, New York, 1990, 729 pp. [4] C.A. Ruiz, I.T. Cameron, R. Gani, A generalized dynamic model for distillation column. III. Study of startup operations, Comp. Chem. Eng. 12 (1) (1998) 1–14. [5] F. Reepmeyer, J.U. Repke, G. Wozny, Analysis of the start-up process for reactive distillation, Chem. Eng. Technol. 26 (2003) 81–86. [6] M. Flender, Zeitoptimale Strategien für Anfahr- und Produktwechselvorgänge an Rektifizieranlagen, VDI Verlag GmbH, Düsseldorf, 1999, ISBN: 3-18rr361003-5. [7] M. Eden, A. Koggersbøl, L. Hallager, S.B. Jørgensen, Dynamics and control during startup of heat integrated distillation column, Comp. Chem. Eng. 24 (2000) 1091–1097. [8] K. Löwe, G. Wozny, A new strategy for product switchover and startup for a heat- and mass-integrated distillation system, Chem. Eng. Proc. 40 (2001) 295–302. [9] J.A. Fabro, L.V.R. Arruda, F. Neves Jr., Startup of a distillation column using intelligent control techniques, Comp. Chem. Eng. 30 (2005) 309–320. [10] N.J. Scenna, S.J. Benz, Start-up operation of reactive columns with multiple steady states: the ethylene glycol case, Ind. Eng. Chem. Res. 42 (2003) 873–882. [11] C. Marangoni, R.A.F. Machado, Distillation tower with distributed control strategy: feed temperature loads, Chem. Eng. Technol. 30 (2007) 1292–1297. [12] R. Rivero, Exergy simulation and optimization of adiabatic and diabatic binary distillation, Energy 26 (6) (2001) 561–593, ISSN: 0360-5442. [13] G. Koeijer, A. Rosjorde, S. Kjelstrup, Distribution of heat exchangers in optimum diabatic distillation columns, Energy 29 (2004) 2415–2440. [14] G.M. De Koeijer, R. Rivero, Entropy production and energy loss in experimental distillation columns, Chem. Eng. Sci. 58 (8) (2003) 1587–1597, ISSN: 0009-2509. [15] F.R. Steinmacher, C. Marangoni, R.F.A. Machado, P.H. Araújo, Startup evaluation and definition of operational conditions of a distillation column, Ann. Braz. Cong. Petrol. Gas (2004). [16] E. Sørensen, S. Skogestad, Optimal startup procedures for batch distillation, Comp. Eng. 20 (1996) S1262–S1275. [17] G. Fieg, G. Wozny, Experimental and theoretical studies of the dynamics of startup and product switchover operations of distillation columns, Chem. Eng. Proc. 32 (1993) 283–290. [18] A. Nooraii, J.A. Romagnoli, J. Figueroa, Process, identification, uncertainty characterization and robustness analysis of a pilot scale distillation column, J. Process Control 9 (3) (1999) 247–264. [19] W.L Luyben, Evaluation of criteria for selecting temperature control trays in distillation columns, J. Process Control 16 (2006) 115–134. [20] L. Wang, P. Li, G. Wozny, S. Wang, A startup model for simulation of batch distillation starting from a cold state, Comp. Chem. Eng. 27 (2003) 1485–1497.
Contents lists available at ScienceDirect
Chemical Engineering and Processing: Process Intensification journal homepage: www.elsevier.com/locate/cep
Application of a new startup procedure using distributed heating along distillation column Leandro O. Werle, Cintia Marangoni, Fernanda R. Steinmacher, Pedro H.H. de Araújo, Ricardo A.F. Machado ∗ , Claudia Sayer Federal University of Santa Catarina, Chemical Engineering Department, Mail Box 476, Florianópolis, SC 88040-970, Brazil
a r t i c l e
i n f o
Article history: Received 31 July 2009 Received in revised form 13 October 2009 Accepted 14 October 2009 Available online 23 October 2009 Keywords: Distillation column New startup procedure Distributed heating
a b s t r a c t This study proposes the use of distributed heating applied to the sieve trays along distillation columns, aiming to evaluate a new experimental configuration and procedure applied to the startup. Through this new configuration it is expected that the column heating time and energy will be reduced. Due to the characteristics of the process, such as inertia and non-linearity, almost all the process variables change rapidly, which increases the time required to reach the steady state. These factors make the objective of this study even more challenging. Most of the previously published studies on startup involve simulations, whereas the main contribution of this paper is that it provides experimental results. The experiments were carried out on a pilot distillation column with 13 sieve trays processing a binary mixture composed of water and ethanol, with a flow rate of 300 L h−1 in the fourth tray. It was verified that the use of a heating system which is able to act in a distributed way allows a reduction in the time (40%) and energy (33%) required for the startup procedure, compared to a conventional approach. © 2009 Elsevier B.V. All rights reserved.
1. Introduction The startup of a distillation column represents one of the most complicated dynamic operations in the chemical industry. It involves the control of complex heat and mass transfer operations, and a wide range of operating conditions has been described [1]. The startup procedure of a distillation column is time and energy consuming and the process is unproductive during this period. Thus, it is desirable to reduce the transient period, which can be very long, and consequently the startup time. The result is the optimization of the unit and its energy costs. It is well known that depending on the productive and operational features, distillation column startup may show a long transition period before the steady state is reached, that is, from hours to days. Also, control of a startup process is very challenging. Almost all process variables change quickly; control variables like heating power, reflux ratio or feed flow have to be adjusted at least once, but mostly several times, before reaching the correct steady state [2].
∗ Corresponding author at: Departamento de Engenharia Química e Engenharia de Alimentos, Universidade Federal de Santa Catarina, Campus Universitário, Trindade, CP 476, CEP 88040-970, Florianópolis, Santa Catarina, Brazil. Tel.: +55 48 3721 9554; fax: +55 48 3721 9554. E-mail address: [email protected] (R.A.F. Machado). 0255-2701/$ – see front matter © 2009 Elsevier B.V. All rights reserved. doi:10.1016/j.cep.2009.10.010
The startup of a continuous distillation column presents a series of operational problems, as described by Kister [3]. Other common problems, which have been observed by Ruiz et al. [4], are related to the instrumentation, difficulties with the process fluids and, more frequently, mechanical and hydraulic problems. These latter two are related to the complete filling of the liquid in the downcomer and the vapor in the perforations of the trays, the pressure increase with the accumulation of liquid in the tray (holdup), and others. There are no general rules regarding the startup strategies for the various kinds of distillation columns, but one widely accepted description is that of Ruiz et al. [4]. These authors divided the startup into three distinct phases: a discontinuous, a semicontinuous and a continuous phase. The first occurs when the heating is in progress, the plates are weeping and instruments are started. The second phase has a strong non-linear behavior since during this phase the hydraulic variables reach their steady state and the composition tends toward the desired value. Finally, the third phase occurs when the column can be considered to be in steady state and the quality control system is turned to automatic. Based on this description, some researchers have discussed the startup problem proposing different strategies aiming to reduce the transient period. Little previous work has been carried out on the startup of distillation processes, due to the difficulties involved in experimental investigations. Efforts are concentrated on theoretical and modeling proposals, particularly for reactive and heat-integrated processes. Furthermore, most studies are
1488
L.O. Werle et al. / Chemical Engineering and Processing 48 (2009) 1487–1494
performed manually or show the efficiency of the application of an advanced control technique. Reepmyer et al. [5] describes four strategies for conventional distillation column startup: (1) conventional: set all control variables to steady state values and wait; (2) total reflux: column is run in loop operation, no distillate removal; (3) total distillate removal: exact opposite to previous strategy, column is run without reflux; and (4) time optimized: developed for heat-integrated columns. Also, for conventional distillation Flender [6] reported that savings up to 80% in the startup time can be achieved if a column is started up utilizing the total distillate removal strategy. The results for startup times obtained by Reepmeyer et al. [5] testing total reflux and total distillate removal for reactive distillation, show that the savings are rather small. Eden [7] developed a procedure to generate startup sequences utilizing process knowledge for heat-integrated columns. This procedure was also used by Löwe and Wozny [8] in which a very similar optimized strategy for the startup of an energy integrated column system was applied and the results validated. Fabro [9] proposed the use of intelligent control techniques such as a neural network, fuzzy systems and genetic algorithms for distillation column startup, and achieved better control performance over the system during the simulations, in comparison to another advanced control and supervision architecture. Scenna and Benz [10] described multiple steady states for different initial conditions as starting points. In this publication, both cases are studied in reactive distillation columns. The column does not start from a cold and empty state, but from a defined state, with the trays filled, warm and with in-phase equilibrium. It is thus clear that different methods and strategies for distillation column startup have been studied in literature. They are focused on saving time by applying advanced control techniques, different operational conditions in reflux, or considering different distillation configurations such as heat integration. In fact, it is necessary to evaluate the operational characteristics of the distillation to achieve a better performance. In this regard, a previous study by Marangoni and Machado [11] describes the use of internal heat sources as a new operational approach to distillation columns. The authors reported that the use of an intermediate temperature control loop together with the dual-temperature classic control can lead to fast dynamics when the process is disturbed. This strategy is based on diabatic distillations. Most of the distillation columns are adiabatic with a reboiler at the bottom, a condenser at the top and adiabatic trays in between. Distillation with heat exchangers on all trays is called diabatic distillation, and is defined as a multistage separation process of a liquid or vapor mixture of two or more components by transferring a given amount of heat in such a way that the supply or the removal of heat is performed at two or more different levels [12]. Diabatic distillation has also been studied by computer simulation in order to determine the optimal distribution of heat to be transferred inside the column [13,14]. This configuration offers the benefits of an improved use of the heat of condensation and evaporation, an increase in the possibilities for energy integration between distillation processes and their surrounding processes, and finally a reduction in the operating costs of distillation installations. In this study, the startup problem of a distillation column is considered experimentally aiming to reduce the time and energy required. Thus, we use a new operational approach based on the concept of diabatic distillation through the heating in the distillation column trays in the stripping section. The condenser was maintained only as a source of heat removal at the top of the column. The procedure adopted for the column startup in the conventional operation was established by Steinmacher [15], based on the
classical procedure, with some modifications, such as starting with the accumulator partially filled with material in the desired composition as proposed by Sorensen and Skogestad [16] and also with bottom product withdrawal right from the beginning, as described by Fieg and Wozny [17]. Therefore, the objective of this study is to evaluate a new experimental configuration and procedure applied to the startup of distillation columns, through heat distribution along the unit, and compare them to a conventional procedure where heat is introduced only in the reboiler. As mentioned above, the startup of a distillation column can be considered in three distinct phases: the heating of all stages, the introducing of a reflux stream in manual mode, and the changing of the control loops to automatic mode. It is also important to emphasize that the main goal of this study is to analyze these three phases of the column startup, which require enormous amounts of energy. 2. Methodology The experiments were carried out in a pilot distillation unit, instrumented with a digital fieldbus communication protocol, processing an ethanol–water mixture (sub-cooled liquid) with a feed flow rate of 300 L h−1 , as described in the next section. 2.1. Description of the experimental unit In order to validate the proposed strategy, a column composed of 13 sieve trays was used. The pilot distillation unit operates continuously, with the feed stream inserted in the 4th tray, a total height of 2.70 m, and built in modules with 0.15 m height and 0.20 m diameter. A schematic diagram of the unit is shown in Fig. 1. The main feature of this distillation unit, which makes it different from others, is the use of heating points distributed along the column, which decentralize the heat supply from the bottom. For this, the unit is instrumented with electrical resistances with a power of 3.5 kW in per tray, according to Fig. 2b, which are started by power variators. Also, each module has a hole for the temperature measurement and for the sample collection. The unit is controlled and operated through a supervisory system, which allows the configuration of the control loops and monitoring of trends, as well as the choice of the selections to work with. The control configuration of the distillation column was formulated based on Nooraii [18], and is illustrated in Fig. 3. The following control loops were defined: (1) bottom level control through the manipulation of the bottom product flow rate; (2) reflux accumulator level control by manipulating the top product flow rate; (3) feed flow rate control as a function of the adjustment of the same stream flow rate; (4) feed temperature control through the manipulation of the fluid flow rate in the heat exchanger of this stage; (5) last tray (distillate) temperature control by means of the manipulation of the reflux flow rate; (6) reboiler temperature control through the vapor flow rate in the heat exchanger of this stage; and (7) temperature control of pre-defined stages of the column through the manipulation of the dissipated power in the tray electrical resistance. The first, second and third loops represent the column mass balance (inventory) control. The fifth and sixth loops comprise the quality control – in this case represented by the temperature. The use of these two loops in combination is referred to herein as conventional control. When these two loops are combined with the seventh loop mentioned above, it is considered herein as the distributed strategy. All control loops are instrumented with fieldbus protocols, along with the acquisition and indication of the bottom and distillate
L.O. Werle et al. / Chemical Engineering and Processing 48 (2009) 1487–1494
1489
Fig. 1. General scheme of the equipments and pilot distillation unit.
stream flows and the pressures at the same stages. The temperatures of all the trays, reboiler, accumulator and feed are monitored by a programmable logic controller. The pressures were monitored in order to ensure the proper functioning of the equipment and the process. 2.1.1. Definition of local heating (stage selection) The proposed modification to distributed heating fulfills the industrial needs, in optimizing the distillation processes and reducing the energy costs incurred by disturbances. It is a simple alteration and it consists only of the local heating of some column trays, through resistances with the heat supply controlled by a power variator. As previously mentioned, the new configuration proposed herein is based on the concept of diabatic distillation, through the heating in the distillation column trays in the stripping section used. Therefore, the great difference of this proposal in relation to other diabatic distillation studies, besides the use of electrical resistances instead of heat exchangers, is the distribution of heat between the reboiler and the trays of the stripping section. In this case, heat exchangers are not used in the rectifying section. The
Fig. 2. Upper view of the modules with details of the tray: (a) conventional and (b) with electrical resistance.
condenser used was maintained as the only source of heat removal at the top of the column. To identify the most sensitive stage for the consequent application of the distributed heating, three different methods were applied [19]. In the first method, the differences between the temperatures of two successive trays were calculated throughout the column and the most sensitive tray was that which presented the greatest difference in relation to its adjacent trays. In the second method, a temperature profile for a given value of the manipu-
Fig. 3. Control configuration of the distillation unit.
1490
L.O. Werle et al. / Chemical Engineering and Processing 48 (2009) 1487–1494
lated variable (in this case, the reflux flow and the reboiler heat) is obtained. The most sensitive tray gives a symmetrical response to positive and negative equal variations. Finally, the third method analyzes the tray with the highest derivative of the temperature in relation to the stage when the process is disturbed. To analyze the first method, the temperature profile for three different conditions of ethanol feed composition (15, 25 and 35%) was observed. In the second and third methods it was necessary to disturb the process and evaluate its behavior. The feed flow used as the standard condition was 400 L h−1 , which was increased to 600 L h−1 and also decreased to 200 L h−1 . It is important to emphasize that the different methods can produce different responses. The definition was based on this analysis together with the characteristics of the plant. 2.2. Experimental procedures of startup The startup procedures for the distributed heating and conventional approaches were almost the same. The only difference was the use of an electrical resistance with a total power of 3.5 kW on an intermediate tray, from the beginning of the experiment. In the conventional approach, the column is operated with heat supply only at the reboiler, with the manual mode of the vapor valve providing 10 kW (8% the steam valve opening). When the proposed new configuration is applied, the heat is supplied at the bottom by the reboiler (with the same previous overall heat supply) and at an intermediate point, by way of the electrical resistance on the 2nd tray, started with 50% of its capacity (1.75 kW). As previously mentioned, the procedure adopted for the column startup in the conventional operation is based on a study by Steinmacher [15]. In this study the startup of the distillation column was considered to have three distinct phases: (1) heating of all stages; (2) introducing of the reflux stream in manual mode; and (3) changing of the control loops to automatic mode. These phases as well as the procedures are illustrated in Fig. 4 and detailed in the next section. 2.2.1. Discontinuous phase In the first phase, at the beginning of the procedure, volumetric compositions of ethanol in the feed and accumulator (partially filled with mixtures in the desired compositions) were 15 and 80%, respectively. During this phase, and throughout the startup, the feed stream is continuously introduced automatically and controlled at 300 L h−1 . To do this experimentally the process is closed, that is, the bottom and top product streams form the feed stream in the main tank (Fig. 1). This is the reason why the accumulator is already partially filled, thus minimizing load composition fluctuations (observed in previous experiments when this equipment was empty). The startup is initiated with the withdrawal of bottom product (automatic bottom level control). The bottom temperature control loop was maintained in manual mode. The liquid mixture from the feed descends throughout the column until it reaches the bottom, where the heat exchanger heats and vaporizes this stream. The vapor mixture ascends through the column, heating it tray-by-tray until it reaches the condenser. Here phase 1 is finalized. 2.2.2. Semi-continuous phase The condensed mixture is received by the accumulator, which is already partially filled (30%) with the material at the desired composition (80%, v/v ethanol) for the reflux stream. At this moment an increase in the level and temperature of the mixture in the accumulator is applied. When the accumulator temperature reaches the desired value (same as in the last stage), the total reflux is introduced until a stable situation is achieved (temperature derivative is zero over time and, for this system, the composition is also considered to be constant). This characterizes the end of phase 2. In
Fig. 4. Steps of the startup process for distillation column.
L.O. Werle et al. / Chemical Engineering and Processing 48 (2009) 1487–1494
1491
this phase, the temperature of the last stage control loop (through the manipulation of the reflux flow rate) is in manual mode, set at 30 L h−1 , and production of the top product has not started. The accumulator level control loop is not automatic. 2.2.3. Continuous phase At this moment, the top product withdrawal is started and the accumulator level bottom and top temperature control loops are changed to automatic mode, starting at phase 3. The steady state is determined when the accumulator temperature and top composition do not vary over time. These variables were used as the evaluation parameters. Volumetric composition measurements corroborate this analysis. The composition was measured with a densimeter for alcohol. 2.3. Energy balances Results are compared based on the energy costs. The energy balance, used to calculate the thermal energy of the reboiler, was obtained considering the steady state at the end of the startup, assuming the following simplifying hypotheses: (1) losses to the environment are negligible; (2) in each theoretical tray liquid and vapor are in equilibrium; (3) pressure is constant along the column. The contributions of thermal energy that enter the column are represented by feed stream, electrical resistance and the reboiler. The thermal energy leaving the column is represented by the distillate stream, condenser and bottom product stream. The heat is calculated based on the enthalpy, using classical relations. The values of the thermodynamic properties required for the calculation of mass and enthalpy balances were obtained through simulations with the software Hyprotech HYSYS® , version 7.0. The thermodynamic model for the liquid phase was UNIQUAC (UNIversal QUAsiChemical) and the vapor phase was considered ideal, mainly because of the low operation pressure of the column. 3. Results and discussions The first phase of this study was to determine the tray where the distributed heating could be applied. As cited before, this was achieved through a sensitivity analysis employing three different methods. The results obtained with the first method (successive trays) using three ethanol feed composition conditions demonstrated the possibility of using trays 1, 2, 3, 5 and 7. As the fifth and seventh trays are located in the rectifying section, they were discarded. It was assumed, following diabatic studies upon which this proposal was based, that in this section it is better to remove heat than to supply it. In addition, it was defined that only one tray would be used to test the new procedure for the startup of the column. To define this stage, since method 1 was not conclusive, the analysis of symmetrical response and maximum derivative (methods 2 and 3) was used. The derivative method again indicated trays 5 and 7, which had been previously discarded, but the symmetrical response method indicated tray 2 as the most appropriate for this study. Fig. 5 shows this analysis, where it can be observed that tray 2 is almost the same distance from steady state when the process is disturbed with positive and negative perturbations in the feed flow. Based on this sensitivity analysis, the distributed action proposed herein was used only in tray 2. Tray 1 was discarded because it is located near the reboiler and 3 was discarded to avoid more than one simultaneous effect. In order to evaluate the performance and verify the effect of the heat distribution along the column during the startup of the process, experiments were carried out with the procedures described in Section 2.2. For both experiments the same criterion was used to
Fig. 5. Results of sensitivity analysis using symmetrical response method: () negative disturbance, () steady state, and () positive disturbance.
Fig. 6. Feed temperature, conventional () and with distributed heating applied on tray 2 ().
evaluate the establishment of the steady state, based on the stabilization of the column and on the derivative of the temperature curves. The volumetric composition measurements corroborate this evaluation. The behavior of the feed temperature in both experiments can be verified in Fig. 6, which shows the feed temperature increase during the column startup. This increase occurs because part of the heated liquid from the bottom is sent to the main tank through the bottom product stream, thus the feed temperature increases gradually until the permanent regime is reached. In industry, the tanks are very large and the influence is smaller. However, there is also the heating time of the equipment itself. A detailed analysis of the results shown in Fig. 6 allows the verification of different slopes for the initial heating phase (linear step), with a higher rate for the configuration with distributed heating, as confirmed by the values of the linear coefficients of the curves, shown in Table 1. This fact indicates that the steady state is reached faster when the proposed distributed configuration is used. With the calculation of the derivative of the feed temperature the effect of the distributed heating can be better observed. These derivatives were obtained from sigmoidal functions, adjusted to Table 1 Coefficients of the linear stage of the heating in the temperatures of the feed stream and of tray 3. Model
Linear regression (conventional operation)
Linear regression (distribution heating operation)
Feed temperature R2 Tray 3 R2
Y = 34.09X + 31.23 0.99 Y = 31.53X + 29.76 0.99
Y = 40.51X + 30.52 0.99 Y = 35.15X + 30.15 0.99
1492
L.O. Werle et al. / Chemical Engineering and Processing 48 (2009) 1487–1494
Fig. 7. Derivatives of the feed temperatures in function of time, conventional () and with distributed heating applied on tray 2 ().
describe the feed temperature in each of the cases. The distributed heating allowed the steady state to be reached faster, as shown in Fig. 7. The temperature profiles of a stripping tray (tray 3) and a rectifying tray (tray 13) are represented in Fig. 8. In these experiments a difference was noted in the tray behavior along the sections of the column during the heating. The stripping trays show a linear heating region, and which is not observed in the rectifying trays. Tray 3 was chosen to illustrate the stripping section as it is the tray which best represents the heat addition effect. This stage is successive to tray 4 which receives the feed stream, as also indicated by the sensitivity analysis. The results shown in Fig. 8 indicate that before the beginning of the reflux, the temperature reached at the top of the column is slightly higher for the distributed configuration than for the conventional one. However, the temperatures after the introduction of reflux were almost the same in both situations. This confirms that the steady state and the reflux flow rate (30 L h−1 ) were the same in both cases. After the linear region (linear phase), shown in Fig. 8, an interval can be observed. This is the time required to heat the tray contents. In this phase, the heating is slow until the bubble point temperature of the mixture is reached, and the tray temperature increases rapidly. The same behavior was observed by Wang [20] and Steinmacher [15]. When using the distributed strategy, the time saving is the same for both the third and thirteenth trays (0.32 h), when compared with the conventional strategy. An analysis of the angular coefficients of the linear regression in the feed and tray 3 temperatures for conventional and distributed modes was carried out. The results are summarized in Table 1. In all of the curves the multiple regression coefficient obtained was
Fig. 8. Temperature profiles of the third tray: conventional (䊉) and distributed heating applied on tray 2 () and of the thirteenth tray, conventional () and distributed heating applied on tray 2 (), during the process startup.
Fig. 9. Column bottom temperature, conventional heating () and with distributed heating applied on tray 2 (), and manipulated variable conventional (䊉) and distributed heating applied on tray 2 ().
higher than 0.99. It is observed that, for both temperatures, angular coefficients are higher when the distributed strategy is used. Therefore, similarly to the results for the initial phase of the feed temperature increase, the distributed heating reduces the time required to reach the stationary tray temperatures. The results shown in Fig. 8 and Table 1 indicate that the distributed heating resulted in a better performance, a faster action and a reduction in the time required to reach the desired temperature at the top of the column. Fig. 9 shows the profile of the bottom temperature before and after the control loop is changed to automatic mode with the two studied configurations. The distributed heating again showed a faster action when compared with the conventional process. Using the intermediate heating, the desired temperature value (98 ◦ C) was reached 0.75 h faster than using the conventional configuration. This is of particular interest as it shows that the action of the resistance increased the heating rate of the bottom temperature. In addition, in Fig. 9 (manipulated variable) it can be observed that after changing the control loop to automatic mode the bottom valve operated with a smaller opening, thus compensating the energy added to tray 2. In order to better evaluate the effect of the distributed heating, the evolution of the derivatives of the bottom temperatures for the two experiments were compared. These derivatives were obtained from sigmoidal functions, adjusted to describe the bottom temperature in each of the cases, as shown in Fig. 10. It was verified that the proposed distributed configuration allowed the steady state to be reached faster. A period of 1.4 h was required for the bottom temperature to become stable with the conventional configuration and only 0.8 h for the distributed configuration. This represents only 57% of the time required to reach the desired bottom temperature. The reduction in the time required to reach the steady state when the distributed heating is applied is confirmed by the measurements of the volumetric ethanol fractions in the bottom column stream and their derivatives as a function of time (dX/dt) to characterize the steady state, as shown in Figs. 11 and 12, respectively.
L.O. Werle et al. / Chemical Engineering and Processing 48 (2009) 1487–1494
1493
Fig. 13. Startup phases: (1) discontinuous; (2) semi-continuous; (3) continues. Fig. 10. Derivatives of the bottom temperatures in function to time, conventional () and with distributed heating applied on tray 2 ().
Fig. 11. Volumetric ethanol fractions in the current of bottom of the column, conventional (), and with distributed heating applied on tray 2 ().
higher internal steam flow rates, since the contact of the feed stream with the heated resistance on plate 2 favors the steam formation inside the column, enhancing the advantage of the distribution of heat. In Fig. 13 the time required for each column startup phase can be verified in detail, as well as how the distributed heating influences the time reduction of each phase. In this figure the effect of the distributed in relation to the conventional heating, mainly in phases 1 and 3, can be clearly observed. In all three stages the reduction of the time, due to the effect of the addition of the intermediate heating point, was significant, when compared to the conventional procedure, being in the order of 43, 21 and 50%, respectively. Thus, a reduction of approximately 40% of the total startup time was reached for a distillation column operating with heat distribution, which suggests considerable gains in the process. We can observe, in Table 3, the advantage of using the new proposed configuration, when evaluated in terms of energy costs, for each of the three startup phases analyzed, during the transient period. In Table 3 Qr is the heat input to the bottom by the reboiler and Qr,e is the heat input to the bottom by the reboiler plus the heat contributed by the electrical resistance in second tray. The thermal energy values of each startup phase (Qphase ) are obtained by multiplying the heat supplied to the column (only the reboiler in the case of conventional operation, and reboiler and electrical resistance in the case of the distributed heating operation) by the duration of each phase. From this analysis, the advantage of the new configuration proposed herein is evident in relation to the conventional procedure, since the total energy required is 33% lower in the configuration with distributed heating. Finally, the thermal energies of the streams, shown in Fig. 14, were calculated after the startup process was concluded and the system reached steady state conditions (end of phase 3). It can be observed from the results shown in Fig. 14 that the amount of
Fig. 12. Derivatives of the volumetric fraction of ethanol of bottom in function time, conventional () and with distributed heating applied on tray 2 ().
Fig. 14. Thermal energy of the currents and of the resistance calculated after the system reached the steady state for the conventional and distributed configurations.
It can be observed in Fig. 11, that there is a change in the bottom ethanol fraction which starts at 0.14% and ends at 0.08%. This is due to the introduction of reflux. Through this observation it is possible to analyze the period out of steady state – as illustrated by the derivative of this variable in relation to time. Is important to note that reflux is introduced in the 13th stage – more distant from the bottom. This means that the interaction between the bottom and top is responsible for the longer time required to reach steady state (or to maintain the column out this condition). As illustrated by other results, Fig. 12 shows that the distributed configuration allows that the transition period is shorter than the conventional one (around 0.75 h), corroborating the previous analysis. Table 2 shows the process variables obtained after the steady state was reached in both cases. Results shown in Table 2 indicate that the distributed configuration leads to a greater production of top product when compared with the conventional configuration. This occurs because the distributed heating operation maintains
1494
L.O. Werle et al. / Chemical Engineering and Processing 48 (2009) 1487–1494
Table 2 Parameters of the steady state reached for the two studied experiments. Variable/stream
Conventional operation
Flow rate (L h−1 ) Ethanol fraction (v/v) Temperature (◦ C)
Distributed heating operation
Feed
Top
Bottom
Feed
Top
Bottom
300.00 0.14 76.50
13.20 0.88 80.40
286.80 0.08 98.50
300.00 0.14 77.00
15.86 0.87 81.20
284.14 0.08 98.50
Table 3 Energy cost for each phase of the startup of the column. Variables
Phase 1 Phase 2 Phase 3 a b c d
Conventional operation
Distributed heating operation
Qr (kW)
Time (h)
Qphase (kW)
Qr,e (kW)
Time (h)
Qphase (kW)
10.00a 10.00a 15.00c
1.53 1.05 1.50
15.30 10.50 22.50
11.75b 11.75b 15.25d
0.88 0.84 0.78
10.34 9.87 11.90
Calculated with steam valve, manual mode (opening of 8%). Calculated with steam valve, manual mode (opening of 8%) + electrical resistance (50% of its capacity). Calculated with steam valve, automatic mode (±18% of opening). Calculated with steam valve, automatic mode (±14% of opening) + electrical resistance (50% of its capacity).
energy supplied to the stripping section column is slightly larger in the distributed configuration (15.25 kW, reboiler + electrical resistance), in relation to the 15.00 kW of the conventional configuration (practically the same amount of energy). Furthermore, the small difference among these values may also be due to the experimental measurement error, besides the approaches and hypotheses assumed in the calculation of the energy balance. The results shown in Fig. 9 help to clarify the explanation. After setting the bottom control loop to the automatic mode, the configuration with distributed action (heating at bottom and on tray 2) required a smaller opening of the bottom steam valve (to reach the same steady state). This occurred because energy was already being added inside the column through the electrical resistance. This finding indicates that there was in fact distribution of energy between the bottom and the resistor.
4. Conclusions Comparison between the two operation approaches, showed that the temperatures of the trays of the stripping section increase at a higher rate when the distributed heating with electrical resistance on tray 2 is used, reaching a permanent regime in a much shorter time than in the conventional process. In addition, the distributed heating allowed the steady state of the temperatures of the column feed, top and bottom, to be reached faster than in the conventional process. Composition measurements of bottom and top products, as well as temperature profile comparisons of these streams, allowed it to be verified that the steady state reached in both cases was the same. Thus, the introduction of distributed heating along the column is demonstrated to be a valid option for reducing the startup time. In this study, a reduction of approximately 40% in the startup time was observed in relation to the conventional process, allowing faster dynamics and lower operation costs, as well as a considerable decrease in the production of out-of-specification products. It was observed that the configuration with distributed heating required 33% lower energy for the 3 startup phases compared with the conventional configuration. The next step in this research is to test the control techniques and algorithms for the automatic startup of distillation columns, improving the steady state and allowing it to be reach in as short a time as possible, using this new procedure for the startup process with the heat distribution along the column.
Acknowledgements The authors are grateful for the financial support from the National Agency of the Petroleum (ANP), and the Studies and Projects Fund (FINEP), by means of the Human Resources Program of ANP for the Petroleum and Gas sector – PRH-34-ANP/MCT. References [1] G. Wozny, P. Li, Optimization and experimental verification of startup policies for distillation columns, Comp. Eng. 28 (2004) 253–265. [2] F. Reepmeyer, J.U. Rekle, G. Wozny, Time optimal start-up strategies for reactive distillation columns, Chem. Eng. Sci. 59 (49) (2004) 4339–4347. [3] H.Z. Kister, Distillation Operation, McGraw-Hill, New York, 1990, 729 pp. [4] C.A. Ruiz, I.T. Cameron, R. Gani, A generalized dynamic model for distillation column. III. Study of startup operations, Comp. Chem. Eng. 12 (1) (1998) 1–14. [5] F. Reepmeyer, J.U. Repke, G. Wozny, Analysis of the start-up process for reactive distillation, Chem. Eng. Technol. 26 (2003) 81–86. [6] M. Flender, Zeitoptimale Strategien für Anfahr- und Produktwechselvorgänge an Rektifizieranlagen, VDI Verlag GmbH, Düsseldorf, 1999, ISBN: 3-18rr361003-5. [7] M. Eden, A. Koggersbøl, L. Hallager, S.B. Jørgensen, Dynamics and control during startup of heat integrated distillation column, Comp. Chem. Eng. 24 (2000) 1091–1097. [8] K. Löwe, G. Wozny, A new strategy for product switchover and startup for a heat- and mass-integrated distillation system, Chem. Eng. Proc. 40 (2001) 295–302. [9] J.A. Fabro, L.V.R. Arruda, F. Neves Jr., Startup of a distillation column using intelligent control techniques, Comp. Chem. Eng. 30 (2005) 309–320. [10] N.J. Scenna, S.J. Benz, Start-up operation of reactive columns with multiple steady states: the ethylene glycol case, Ind. Eng. Chem. Res. 42 (2003) 873–882. [11] C. Marangoni, R.A.F. Machado, Distillation tower with distributed control strategy: feed temperature loads, Chem. Eng. Technol. 30 (2007) 1292–1297. [12] R. Rivero, Exergy simulation and optimization of adiabatic and diabatic binary distillation, Energy 26 (6) (2001) 561–593, ISSN: 0360-5442. [13] G. Koeijer, A. Rosjorde, S. Kjelstrup, Distribution of heat exchangers in optimum diabatic distillation columns, Energy 29 (2004) 2415–2440. [14] G.M. De Koeijer, R. Rivero, Entropy production and energy loss in experimental distillation columns, Chem. Eng. Sci. 58 (8) (2003) 1587–1597, ISSN: 0009-2509. [15] F.R. Steinmacher, C. Marangoni, R.F.A. Machado, P.H. Araújo, Startup evaluation and definition of operational conditions of a distillation column, Ann. Braz. Cong. Petrol. Gas (2004). [16] E. Sørensen, S. Skogestad, Optimal startup procedures for batch distillation, Comp. Eng. 20 (1996) S1262–S1275. [17] G. Fieg, G. Wozny, Experimental and theoretical studies of the dynamics of startup and product switchover operations of distillation columns, Chem. Eng. Proc. 32 (1993) 283–290. [18] A. Nooraii, J.A. Romagnoli, J. Figueroa, Process, identification, uncertainty characterization and robustness analysis of a pilot scale distillation column, J. Process Control 9 (3) (1999) 247–264. [19] W.L Luyben, Evaluation of criteria for selecting temperature control trays in distillation columns, J. Process Control 16 (2006) 115–134. [20] L. Wang, P. Li, G. Wozny, S. Wang, A startup model for simulation of batch distillation starting from a cold state, Comp. Chem. Eng. 27 (2003) 1485–1497.