Cost Analysis of

Heptene Production from

Propene and N-Butene

Contents:

1. Abstract
2. Introduction
3. Results
4. Discussion
5. Recommendations
6. Example Calculations
7. Sources
8. Additional Material
   1. One Stage
   2. Five Stages
   3. Nine Stages
   4. Cap Cost Pricing

Abstract

Analysis of the production process for heptenes from propylene and butene shows that the estimated equipment cost of the initially proposed design is $852K capital, with a profit of $52,000K/year. By replacing the external cooling system with internal reactor cooling coils and increasing the number of reactor stages, this profit can be increased to $55,000K/year with a production capital of $944K. Due to insufficient kinetic and conversion data, only an approximate assessment of process design could be made on the reactor system. Cost minimization analysis comprised the majority of our investigation.

Introduction

This process is designed to convert C3 and C4 unsaturated hydrocarbons (propylene and butenes) into 1-heptene and other unsaturated hydrocarbons. The reaction processes and reactor optimization are to be determined by laboratory trials over a range of feed concentrations and temperatures. This concentration and temperature dependence data is not available. Only the conversion data and inlet and outlet stream compositions are known for the initial process. No kinetic or equilibrium data is supplied for this process. The primary reactions include:

C₃H₆ + C₃H₆ = C₆H₁₂ formation of 1-Hexene

C₃H₆ + C₄H₈ = C₇H₁₄ formation of 1-Heptene

C₄H₈ + C₄H₈ = C₈H₁₆ formation of 1-Octene

C₃H₆ + 2C₄H₈ = C₁₁H₂₂ formation of 1-Undecene

This set of reactions is exothermic. Therefore, the initial design of the reactor is at 45° C to keep the reactions from running away. The major constraints on this problem are the lack of kinetics data for the reactions. This inhibits any attempt to design an alternate reactor system. The reactor itself must have a good heat removal rate in order to keep the system within reasonable bounds. All streams should be kept pressurized so air inflow to the system does not occur.

The feed stream is a process cut generated from the distillation of crude oil. The feed and product stream specifications for our process, with mass balance modifications, are as follows:

Note that only n-butene reacts in the above equations to produce the desired products, isobutene is a bystander. None of the saturated hydrocarbons participate in these reactions, instead they inhibit the process by consuming reactor volume and providing a heat sink as a buffer against runaway reaction. The products are separated for individual marketing, mostly as gasoline additive to increase octane ratings.

Our modeling extends over the plant from feed to just after the reactor process. This portion of the plant includes two storage vessels (one each for the C3 and C4 streams), a pump to feed the mixed cuts to the reactor, and the reactor system (explained below). All considered streams are pressurized into a liquid state to reduce reactor and piping volume. No portion of the process requires unusually high pressure or temperature conditions except for the C3 storage, which has a pressure requirement of 20 barg to maintain the light propane component in the liquid phase. The entire process is depicted below:

Figure 1

The reactor consists of a five-stage tower with a separate feed and heat exchange process for each stage. Each stage feeds into the stage below it by means of liquid "downcomers," similar to a distillation column with no vapor flow, only weeping to the lower stages. All stages have equal volume and can be considered "well-mixed." Heat exchange occurs by pumping the product through an external heat exchanger at each stage. The exchangers reduce the recycle stream to 40° C to maintain an overall reactor temperature of 45° C. This overall setup is depicted in figure 2:

The lack of kinetics data inhibits any attempt to design an alternate reactor system. Data in Chauvel’s article suggested a plug-flow reactor with multiple inject ports along the flow axis. However, there was no information on the relative composition of each injection stream and this idea was discarded as requiring more research than our efforts allowed.

The reactor must be kept within reasonable temperature bounds as the reactions are exothermic and could cause reaction runaway if heat removal systems fail. All streams should be kept pressurized so air inflow to the system does not occur, creating an explosion hazard. The elevated system pressure will also enable the use of smaller equipment as there is never an expansion into a vapor phase.

The original design suggested the use of carbon steel heat exchange tubes for use in cooling water duty. While there is not much temperature stress on that system (the cooling water varies between 35 and 45° C), it was suggested by Dr. C. A. Armeniades of Rice University, that we use stainless steel tubes instead. Corrosive effects were predicted due to the use of water and conditions indicated that scaling would likely occur if CS is used. Dr. Armeniades advised that cost savings would be created in long term analysis of the system, as the CS tubes would require more maintenance and more frequent replacement/overhaul work if used in cooling water service.

Results

Given the lack of specific kinetic data, the primary focus of this project was to solve for the existing system and research the kinetic process. Sources of data proved to be scarce and idealized in treatment. The idealized data focused on relative conversion percentages of products. The chart at the right shows the relative weight percent of product components, where the C₈₊ indicates the undecene compound, included as the "catch-all" for C₈₊ compounds.

Optimization attempts focused on modifications to the basic design of this operating system with an emphasis on equipment around the reactor. Reactor conditions (volume, pressure, mixing quality, and temperature) remained constant in all designs to minimize departure from known conversion data. One major modification that would reduce equipment cost is the installation of internal cooling coils to replace the external pump and heat exchanger system designed to remove heat from the reactor.

 

Two alternate reactor configurations were selected based on conversion data from a graph provided by Chauvel. This graph represents conversion of n-butene to products based on number of CSTR stages in the reactor. Both one-stage reactor and a nine-stage reactor were modeled in attempt to determine the economic impacts of altering the process. Conversion was modeled for each species as following a feed-ratio according to the n-butene conversion, where only enough propylene was fed so as to be fully consumed in each reactor (i.e., there is always full conversion of propylene feed).

Our modeling covered different reactor designs using the conversion data given above. The base case was depicted in figure 1. The one-stage case and the nine-stage case are modifications of the base case in that they vary the number of reactor stages, feed composition, and the overall conversion of 1-butene. All of our modified designs included internal cooling coils. The following chart displays the differences between production streams in these models:

All cases give similar selectivity while increasing raw conversion from case to case. Each case also demands higher propylene feeds as the conversion increases. This process can be pushed far beyond conversion rates of 55%, but higher conversions encourage more trimerization and tetramerization, which does not yield the desired products. A manufacturing cost analysis of the investigated cases is given below:

The capital cost for building each of these four design cases is as shown below:

Discussion

The above sheets include heat exchangers as representative additional tower costs. The method for determining the costs of internally cooled towers involves calculating the cost for external heat exchangers and adding those costs to the price of a multi-stage tower. Results for specific stream data and process layouts are depicted in the appendices. Heat exchanger size was recalculated for each process as described under "Example Calculations" in the appendices.

Profits were calculated under the assumption that all chemicals could be sold for an average of $.60/lbm. Shell Chemical provided a rough price estimate of $.40-$.80/lbm. It should be emphasized that one of the largest contributing factors to price will be the purity of the product. This project does not specifically calculate purity, but the problem specifications range from 93-99% per product. These products are intended as gasoline additives. With these factors taken into account, a selling price of $.60/lbm/product seems reasonable. Additionally, our reactants are purchased at much lower grades than the grades we found purchasing information for in Chemical Marketing Weekly. Consequently, the reactant prices are adjusted to 1/3 of the advertised sales costs. Our plant is calculated as having a 10 year-lifespan in terms of depreciation, and our cost calculations follow the general methods as outlined in Turton, Bailie, Whiting and Shaewitz, pp.81.

In general, the process may be overpriced with regards to capital cost. No modifications to the plant were that radical, so the suggested improvements seem entirely reasonable. The yearly operational figures also seem reasonable within the perspective of this as a preliminary analysis.

Recommendations

The use of a 9-stage, internal coil reactor is recommended for the production of heptenes and associated chemicals via the Dimersol process. The extra 2% conversion of n-butenes generated by this method yields an additional $2,200K/year as compared to the base case, and an extra $13,000K/year compared to the 1 stage reactor. In contrast, the construction cost of this method is only $65K greater than the base case, and $500K greater than the cheapest scenario considered here. With these prices in mind, the larger reactor will pay for the extra cost in 3 months of its 10-year lifespan. Additional study is recommended with regards to reaction kinetics/equilibrium data, catalyst sources, and alternative methods of heptene generation such as alkylation, which is the current standard. Detailed market examination should be conducted to determine the marketability of the product chemicals.

Most calculations in this project were made through the use of programs such as HYSYS and CAPCOST. However, some external calculation was required for the heat exchangers and yearly operational costs. Most operational costs were calculated as performed in Turton, Bailie, Whiting and Shaeiwitz.

Heat Exchangers

Heat exchanger data is calculated based on the 5 stage external case. Each stage has it’s own exchanger which deals with an equal portion of the total heat generated in the reactor. Our examples are the one-stage and the nine-stage cases.

1. 

   We assume that the heat exchange rate per unit area remains constant at about 14.195e3 kJ/hm2 of area, and then calculate area for each internal exchanger.

2. 

   At this point, we assume our heat exchangers are ¾ inch stainless steel pipes. These pipes take up volume in each reactor stage as they are internal. This volume on a per stage basis is as follows.

3. 

Similarly, water flow rates are ratios of the 5-stage case.

D.

Chauvel, A., P. Leprince, Y. Barthel, C. Rimbault, and J-P Arlie. Manual of Economic Analysis of Chemical Processes. Translated by R. Miller and E.B. Miller, McGraw-Hill, NY, 1976. pp. 207-228.

Ngandjui, L. M. Tiako and Thyrion, F. C.. "Kinetic Study and Modelization of n-Butenes Oligomerization over H-Mordenite". Industrial & Engineering Chemistry Research. American Chemical Society, 1996, vol. 35, 1269-1274.

Ann M. Mitchell, Petroleum Engineer, Amoco Corporation, Whiting Refinery, Whiting, Indiana.

Dan J. Sajkowski, Optimization Superintendent, Amoco Corporation Whiting Refinery, Whiting, Indiana.