3 replies to this topic

[atifprocessengineer]

Hi,

I have a final year project regarding LPG production.

I have designed a de-ethanizer coloumn whose function it is to separate the Methane and Ethane from the feed mixture of C1 to C9 hydrocarbons.

I have calculated the feed pressure by the buble point method and this results as 80 atm. But when I calculated the bottom pressure, this comes out lower then the feed pressure.

Is this possible?

If yes, then what are its disadvantages?

[Art Montemayor]

atif:

I had to edit and correct your original script because it was hard to understand and it was full of "texting" script - which only makes it more difficult to read and interpret. Please stick to conventional, correct, spell-checked, and proper english script. Now to what I believe is your basic question(s):

If what you mean is that you have taken the thermal and compositon conditions of your feed stream and calculated a "feed" pressure, then I believe that by applying a bubble point calculation you have identified the pressure condition on the feed tray - NOT "the feed". You must be specific in engineering; otherwise, you will get all mixed-up and confused. The feed stream to a distillation column has to be at a pressure higher than that existing on the feed tray. This is fairly obvious and makes common sense because otherwise you would not be able to have feed stream flowing into the distillation column at the feed tray.

Your bubble point results are not possible in real life. You must have a driving force (a pressure difference) through the entire distillation column, and it is this driving force that enables the vapors generated in the reboiler to ascend up the column and out to the overheads condenser. This means, in other words, that the vapor pressure existing on the intermediate feed tray has to be an intermediate pressure - some value between the reboiler vapor pressure and the overheads condenser vapor pressure. This, of course, assumes that you have located the feed tray in an intermediate height.

Check your bubble point calculations. Your reboiler temperature has to be much higher than your feed tray temperature and this contributes to a higher vapor pressure within the reboiler. Go back and review the principles of distillation and focus on WHAT, exactly, is happening inside the column in a physical manner. This type of engineering analysis will guide you to make and apply the correct mathematical relationships.

Good luck.

[Zauberberg]

Just to add a few important issues on top of what Art has said:

- Basically it's impossible to design De-C2 towers to operate at pressures above 30-35 barg. The main reason is that the critical pressure of LPG components will be exceeded in the reboiler, which means they cannot be recovered as liquid product.

- Operating pressure of the tower will depend on 2 things: one of them is the available temperature of cooling medium in the overhead condensers, and the second one is requirement for C2- gas (De-C2 overhead product) compression. In practice, you normally want to run the tower at the highest possible pressure and thus avoid expensive refrigeration in the overhead condensers, and also the additional C2 compression cost (if there is any).

Best regards,

[Padmakar Katre]

Dear,
Addition to the above two comments, the operating pressure is always a function of the propane plus components(C3+) slip in de-ethanizer overhead and the available chilling medium. The industrial operating pressure for any de-ethanizer column is in the range of 25-32(32.5 max Reference is 'Steady state modeling of refining processes by Gerald Kaes') bar with marginal propane slip. Again it depends the ovhd condenser configuration, as the product is at its dew point or bubble point.You can go for lower pressure with lesser theoretical trays and higher column diameter as well the lower pressure steam/lower temperature heating medium but the condenser temperature is also very low so consider the propane or propylene refrigeration.