Optimization of an Ethylbenzene Process - SMBHC Thesis Repository

The objective function is a mathematical function that achieves a maximum or minimum for certain values of the decision variables selected. The best possible outcome of the objective function is a global optimum, where the objective function is the best for all decision variables. If the objective function is poorly chosen, all the calculations and work can be wasted.

The engineer seeks to achieve the global optimum of the objective function. In device design, it is essential that the right catalyst and operating conditions are selected to minimize the adverse effects that occur. In order to know which changes would have the greatest impact, a sensitivity analysis was applied to the base case to determine how gradual changes in certain parameters affect the plant's net present value.

Table 1: Data Required for Base Case (in Addition to PFD and Flow Tables)

Net$Present$Value$(in$ Millions)$

Percent$Change$(%)$

Now that we knew the changes that would be applied to our optimization case, we were able to apply the changes to the chemical process of the base case. As an example, I will focus on how we optimized the reactor part of the process and the effect on net present value. Using this parametric optimization, it was found that the higher the temperature of the reactor set, the more ethylbenzene is produced.

Although this was a decrease in reactor lengths from the base case, with the increased temperatures from the previously explained optimization, this was found to be the optimum length. After the optimization was completed, the present value of the plant was $22.5 million, a staggering increase from -$11 million in the base case. Through my use of the various optimization strategies, my team was able to change a net present value of -$11 million to.

Figure 3. Plot of conversion versus reactor temperature. The reactor temperature is representative of the entire reactor train, i.e

APPENDIX

PRODUCTION OF ETHYLBENZENE

It is proposed to build an ethylbenzene plant producing 80,000 tons of ethylbenzene per year. In the original scenario presented, or the base case, the process has a fixed feed ratio (8:1) of benzene to ethylene, a purity of 99.8% in the ethylbenzene product stream and a fixed annual production as previously mentioned. The diagram in figure is based on the main reaction in the process, where one mole of benzene (cost $1.04 per kilogram) reacts with one mole of ethylene ($0.72/kg) to form one mole of ethylbenzene ($1.348/kg).

On a kilogram-mole basis (ie, 1000 gram-mole), the economic potential of the process is calculated to be $41.74 per kilomole of ethylbenzene. With that in mind, the base case presents a problem: its net present value, or total return on investment with interest rate factored in, is a loss of $11 million. The goal of our optimization, or the profit maximization of the investment, is to make changes to the process that meet the feed and product parameters, but minimize the cost, and maybe.

Figure 1. Process concept diagram of the ethylbenzene process. Inlet streams are benzene and ethylene; product is ethylbenzene (3)

Results & Discussion

Base Case

This column, T-301, separates most of the benzene and lighter hydrocarbons from a stream of ethylbenzene and diethylbenzene. The size of the equipment was calculated from current information, temperature and pressure specifications, and other parameters, either known or determined from Pro/II. When we sized the heat exchangers, there were several variables to consider - the presence or absence of a phase change, required heat consumption and the classification of utility used.

By using a Qlatent:Qtotal ratio, the temperature difference between the latent and the sensible side of the heat exchanger was found. Once this internal temperature was reached, we found ∆TLM for the sensible and latent sides of the exchanger. The dimensions of the process vessels, from Table 2, are shown according to their length and diameter.

The sizing of the vessels on the two distillation columns was done differently, as two products left the vessel: the reflux stream and the. These flow rates, molecular weights and densities of the streams leaving the vessel were read from the PRO/II simulation and used to calculate the total flow rate leaving the vessel, from which the volume, length and diameter could then be determined. Sizing the towers required many different variables to calculate the number of trays, tower height, and tower diameter.

A range of diameters was determined for the top and bottom of the tower. Feedstock costs were the sum of ethylene feed costs, benzene feed costs, and catalyst costs. At the end of the project life, the salvage value was 10% of the recoverable FCI.

Based on this analysis of the base case, we decided that optimization of the factory needed to take place.

Figure 1. The process flow diagram for the base-case ethylbenzene production plant

Optimized Case

Due to the results of the changes and the resulting net present values, our group decided to further optimize the process by choosing the proposed change 1 and change 2. Thus, we investigated the effects of increasing the reactor inlet temperatures; from these changes we observed effects on reaction conversion and selectivity. By removing these two heat exchangers, we were able to save money on utilities and labor; but we had to change the material of two reactors.

We also investigated the elimination of the second distillation tower, T-302, and its associated recycle bottoms stream. When we realized that a very small amount of diethylbenzene was circulating in the recycle stream, compared to the rest of the system, we concluded that the recycle stream was a waste of resources. After eliminating the second tower, the recycle stream, and all equipment associated with that tower and stream, we had to make minor adjustments to keep the ethylbenzene product stream within the required specifications.

$100,000 bare module cost of the tower, several hundred thousand dollars due to the reduced load on the fired heating, the ~$100,000 bare module cost associated with that. Finally, after eliminating the second distillation column, we needed to ensure that the remaining distillation column could efficiently separate the tower feed (stream 16). However, there was another aspect of the factory that had a major impact on the impact of the investments.

Since one of the limitations of the project was a specific production rate of ethylbenzene, revenues could only be changed by increasing sales. Utilities costs for the optimized case, which ended up being $120,000/year, were another driver of the increased NPV, as they represented a significant decrease from the utility costs in the base case, $924,600/year. The majority of the utility costs came from the fired heater, which had to heat the stream to 370°C before the stream entered the reactor chain.

This reduced the required heat exchanger capacity from 25,500 MJ/hour to 19,400 MJ/hour, reducing the energy costs of the fired heating from $2.35 million/year to $1.8 million/year. Finally, the removal of four heat exchangers, two on T-302 and two eliminated before two of the reactors in the reactor chain, further reduced energy costs. The greatest impact on energy costs came from the reduced load on heated heating.

Table 5. Special conditions matrix, base case equipment.

Conclusion and Recommendations

It also gave us a conventional payback period, or the time required for the project to break even, to the value of 5.1 years. Given the option to consider Change 1 (new catalyst), Change 2 (cheaper feed costs, lower quality benzene feed), both changes, or no changes, our group chose both changes because of the increased net present value. As we proceeded with further optimization from both changes, we decided that a vital part of the optimization would include eliminating the bottoms recycle stream and improving the reactor chain section of the plant.

By analyzing the temperature effect on the single-pass conversion and selectivity of ethylbenzene in the reactors, we concluded that an increased temperature in the entire reactor chain section would allow the greatest conversion and reduce the amount of feed required. In addition, because the reactions taking place were exothermic, the first reactor would increase the temperature of the stream entering the second reactor, further increasing the conversion in each successive reactor. Although this action would require an increase in equipment costs, we deemed it a necessary call to achieve the highest conversion through the use of stainless steel materials for the second and third reactors.

As a direct effect of the higher conversion taking place in the reactors, less diethylbenzene was produced and thus there was no need for a second distillation column that separated the diethylbenzene and recycled it to another reactor. By eliminating the bottoms recycle stream we were able to eliminate the equipment costs for a distillation tower, reactor and pump, but what greatly benefited the NPV was the elimination of the use of the fired heater for the bottoms stream. Annual energy costs were reduced from $900,000 to $100,000, mainly due to the reduced heating load.

After completing our optimization, we would recommend further investigation of the cost of food to further increase the net present value. There is a significant amount of benzene escaping in the fuel gas stream; if this can be put back into the recycling stream, less fresh feed will be required, resulting in a lower cost of materials. Also, by adding more reactors or changing reactor parameters such as pressure drop, temperature, and length, we believe that feedstock feed can be reduced, if the single-pass conversion of each reactor can be further increased.

Due to time constraints, we were not able to analyze these changes in depth and how they affected the entire process.